c© 2007 Wiley-VCH Verlag GmbH & Co. KGaA, Weinheim10.1002/14356007.a12 169.pub2

Gas Production 1

Gas Production

Heinz Hiller, Lurgi, Frankfurt am Main, Germany (Chap. 1)

Rainer Reimert, (formerly Lurgi) Engler-Bunte-Institut der Universitat Karlsruhe, Karlsruhe, Germany(Chap. 1, 4 and 7.3)

Friedemann Marschner, Lurgi, Frankfurt am Main, Germany (Chap. 2)

Hans-Joachim Renner, Lurgi, Frankfurt am Main, Germany (Chap. 2)

Walter Boll, Lurgi, Frankfurt am Main, Germany (Chap. 2 and 5.3)

Emil Supp, Lurgi, Frankfurt am Main, Germany (Chap. 3, 5.1 and 5.2)

Miron Brejc, Lurgi, Frankfurt am Main, Germany (Chap. 3)

Waldemar Liebner, Lurgi, Frankfurt am Main, Germany (Chap. 3)

Georg Schaub, (formerly Lurgi) Engler-Bunte-Institut der Universitat Karlsruhe, Karlsruhe, Germany(Chap. 4)

Gerhard Hochgesand, Lurgi, Frankfurt am Main, Germany (Chap. 5 and 6.3)

Christopher Higman, Lurgi, Frankfurt am Main, Germany (Chap. 5.1, 5.2 and 7.2)

Peter Kalteier, Lurgi, Frankfurt am Main, Germany (Chap. 5.1 and 5.2)

Wolf-Dieter Muller, Lurgi, Frankfurt am Main, Germany (Chap. 5.3)

Manfred Kriebel, Lurgi, Frankfurt am Main, Germany (Chap. 5.4)

Holger Schlichting, Lurgi, Frankfurt am Main, Germany (Chap. 5.4)

Heiner Tanz, (formerly Lurgi, Frankfurt), IVTP, Dreieich, Germany (Chap. 5.5, 5.6, 5.7 and 5.8)

Hans-Martin Stonner, Lurgi, Frankfurt am Main, Germany (Chap. 1.5, 6.1 and 6.2)

Helmut Klein, Lurgi, Frankfurt am Main, Germany (Chap. 6.2)

Wolfgang Hilsebein, Lurgi, Frankfurt am Main, Germany (Chap. 7.1)

Veronika Gronemann, Lurgi, Frankfurt am Main, Germany (Chap. 7.1)

Uwe Zwiefelhofer, Lurgi, Frankfurt am Main, Germany (Chap. 7.2)

Johannes Albrecht, Lurgi, Frankfurt am Main, Germany (Chap. 7.3)

Christopher J. Cowper, British Gas, London, United Kingdom (Chap. 8)

Hans Erhard Driesen, Engler-Bunte-Institut der Universitat Karlsruhe, Karlsruhe, Germany (Chap. 8)

Related Articles → Coal Liquefaction and → Coal Pyrolysis are separate keywords.

1. Introduction . . . . . . . . . . . . . . 31.1. Types of Gases; General Overview

of Production Methodsand Characteristics . . . . . . . . . 3

1.1.1. Water Gas and Producer Gas . . . . 31.1.2. Synthesis Gas and Reduction Gas . 41.1.3. Town Gas and Medium-Btu Gas . . 71.1.4. Biogas and Landfill Gas . . . . . . . 91.1.5. Rich Gas and Substitute Natural Gas

(SNG) . . . . . . . . . . . . . . . . . . 91.2. Raw Materials for Gasification . . 91.3. Physicochemical Basis for

Gas Production . . . . . . . . . . . . 12

1.4. Characteristics of the BasicProcesses . . . . . . . . . . . . . . . . 16

1.5. Product Gas Treatment . . . . . . . 181.5.1. Purification Processes . . . . . . . . . 181.5.2. Conditioning . . . . . . . . . . . . . . 191.5.3. Byproducts . . . . . . . . . . . . . . . 202. Steam Reforming of Natural Gas

and Other Hydrocarbons . . . . . 212.1. Feedstocks . . . . . . . . . . . . . . . 212.2. Natural Gas and Other Gaseous

Hydrocarbons . . . . . . . . . . . . . 222.2.1. Principles . . . . . . . . . . . . . . . . 24

2 Gas Production

2.2.2. Catalysts, Catalyst Poisons,Desulfurization . . . . . . . . . . . . . 26

2.2.3. Tubular Reformers . . . . . . . . . . . 272.2.4. Production of Fuel Gas and

Synthesis Gas . . . . . . . . . . . . . . 292.2.5. Special Reforming Processes . . . . 302.3. Tubular Steam Reforming of

Liquid Hydrocarbons . . . . . . . . 312.3.1. Commercial Processes . . . . . . . . 322.3.2. Fuel Gas and Synthesis Gas from

Liquid Hydrocarbons . . . . . . . . . 322.3.3. Special Processes . . . . . . . . . . . 322.4. Prereforming . . . . . . . . . . . . . 332.4.1. Principles . . . . . . . . . . . . . . . . 332.4.2. Catalysts . . . . . . . . . . . . . . . . . 342.4.3. Prereforming of Natural Gas . . . . 342.4.4. Prereforming of Naphtha;

Rich Gas Process . . . . . . . . . . . 352.5. Autothermal Catalytic Reforming 373. Noncatalytic Partial Oxidation

and Special Gasification Processesfor Higher-Boiling Hydrocarbons 40

3.1. Raw Materials . . . . . . . . . . . . . 403.2. Partial Oxidation of

Hydrocarbons . . . . . . . . . . . . . 423.2.1. Principle . . . . . . . . . . . . . . . . . 423.2.2. Types of Processes . . . . . . . . . . . 433.2.3. Influencing Raw Gas Composition . 473.2.4. Submerged Flame Process . . . . . . 513.3. Hydrogenating Gasification . . . . 514. Gas Production from Coal, Wood,

and Other Solid Feedstocks . . . . 524.1. Fundamentals . . . . . . . . . . . . . 534.1.1. Thermodynamics of Chemical

Reactions . . . . . . . . . . . . . . . . 534.1.2. Kinetics . . . . . . . . . . . . . . . . . 544.2. Classification and General

Characteristics of GasificationProcesses . . . . . . . . . . . . . . . . 56

4.2.1. Criteria for Classification . . . . . . 564.2.2. Criteria for Process Assessment . . 574.2.3. Mathematical Modeling of

Gasification Reactors . . . . . . . . . 574.3. Characterization of Solid

Feedstocks for Gasification . . . . 594.4. Moving- or Fixed-Bed Processes . 624.5. Fluidized-Bed Processes . . . . . . 664.6. Entrained-Flow Processes . . . . . 704.7. Molten-Bath Processes . . . . . . . 724.8. Underground Coal Gasification . 744.9. Environmental Aspects of

Gasification . . . . . . . . . . . . . . 755. Gas Treating . . . . . . . . . . . . . . 76

5.1. Carbon Monoxide ShiftConversion . . . . . . . . . . . . . . . 77

5.1.1. Fundamentals . . . . . . . . . . . . . . 775.1.2. Catalysts . . . . . . . . . . . . . . . . . 785.1.3. Clean Gas Shift Conversion . . . . . 795.1.4. Raw Gas Conversion . . . . . . . . . 815.1.5. General Comments on Reactor

Arrangements . . . . . . . . . . . . . . 825.1.6. Noncatalytic Quench Conversion . 825.2. Carbonyl Sulfide Conversion . . . 825.3. Methanation and Methane

Synthesis . . . . . . . . . . . . . . . . 835.3.1. Definitions and Applications . . . . 835.3.2. Principles of Methanation . . . . . . 845.3.3. Methanation as a Step in Hydrogen

Purification . . . . . . . . . . . . . . . 855.3.4. Methanation of Rich Gas . . . . . . . 865.3.5. Methane Synthesis from Gases with

High Carbon Monoxide Content . . 865.4. Absorption Processes . . . . . . . . 905.4.1. General . . . . . . . . . . . . . . . . . . 905.4.2. Processes for Carbon Dioxide and

Sulfur Compound Absorption . . . 935.4.2.1. Physical Absorption Processes . . . 945.4.2.2. Chemical Absorption Processes . . 1015.4.2.3. Physical – Chemical Absorption

Processes . . . . . . . . . . . . . . . . 1075.4.2.4. Comparison of Various Absorption

Processes for Hydrogen Productionand IGCC . . . . . . . . . . . . . . . . 108

5.4.3. Liquid-Phase Oxidation Processes . 1125.4.4. Removal of Gas Impurities of Low

Concentration . . . . . . . . . . . . . . 1165.5. Adsorption Processes . . . . . . . . 1185.5.1. Fundamentals . . . . . . . . . . . . . 1185.5.2. The “Classic” Method . . . . . . . . 1205.5.3. Pressure-Swing Adsorption . . . . . 1215.5.4. Adsorption on Activated Carbon . . 1225.6. Cryogenic Processes . . . . . . . . 1225.6.1. Partial Condensation . . . . . . . . . 1235.6.2. Liquid Methane Wash Process . . . 1245.6.3. Liquid Nitrogen Wash Process . . . 1255.7. Gas Separation by Membranes . 1265.8. Addition of Inerts or Other

Substances . . . . . . . . . . . . . . . 1276. Handling of Byproducts . . . . . . 1286.1. Aqueous Condensates . . . . . . . . 1286.1.1. Mechanical Treatment . . . . . . . . 1296.1.2. Extraction and Adsorption of

Organic Substances . . . . . . . . . . 1316.1.3. Removal and Recovery of Ammonia

and Sulfur . . . . . . . . . . . . . . . . 1326.1.4. Biological and Final Treatment . . . 1346.1.5. Removal of Heavy Metals . . . . . . 135

Gas Production 3

6.1.6. Example of an IndustrialApplication . . . . . . . . . . . . . . . 135

6.2. Hydrocarbon Condensates . . . . 1366.2.1. Distillation . . . . . . . . . . . . . . . 1376.2.2. Hydrorefining . . . . . . . . . . . . . . 1376.3. Gaseous Byproducts . . . . . . . . . 1397. Typical Examples of Complex Gas

Production Plants . . . . . . . . . . 1417.1. Methanol Production from

Natural Gas . . . . . . . . . . . . . . 1417.1.1. Methanol Production Based on

Catalytic Autothermal Reforming . 1427.1.2. Comparison of Conventional Steam

Reforming and CombinedReforming Processes forMethanol Production . . . . . . . . . 143

7.2. Hydrogen Production Based onHeavy Residues . . . . . . . . . . . . 143

7.3. Combined Cycle Power SystemBased on Coal . . . . . . . . . . . . . 147

7.3.1. Introduction . . . . . . . . . . . . . . . 1477.3.2. Fundamentals . . . . . . . . . . . . . . 148

7.3.3. Installations and Design Studies . . 1508. Analysis and Quality Control . . . 1528.1. Quality Specifications . . . . . . . . 1528.1.1. Combustion Characteristics . . . . . 1538.1.2. Minor Constituents . . . . . . . . . . 1538.2. Test Methods . . . . . . . . . . . . . 1538.2.1. Determination of Combustion

Characteristics . . . . . . . . . . . . . 1538.2.1.1. Calorimetry . . . . . . . . . . . . . . . 1538.2.1.2. Density . . . . . . . . . . . . . . . . . . 1548.2.1.3. Wobbe Index . . . . . . . . . . . . . . 1548.2.2. Analytical Methods . . . . . . . . . . 1558.2.2.1. General Methods for Determination

of Several Components . . . . . . . . 1558.2.2.2. Specific Methods for Determination

of Individual Components . . . . . . 1558.2.2.3. Methods forDetermination ofMinor

Components . . . . . . . . . . . . . . . 1568.2.2.4. Determination of Trace Constituents 1589. Acknowledgement . . . . . . . . . . 15810. References . . . . . . . . . . . . . . . 158

1. Introduction

1.1. Types of Gases; General Overviewof Production Methods andCharacteristics

Gases for industrial or heating uses can be classi-fied in various ways. The most appropriate clas-sification appears by calorific (heating) valueranges, in kJ/m3 (STP):

Water and producer gas 4600 – 12 500Synthesis and reduction gas ca. 12 500Town gas and medium-Btu gas 16 700 – 20 000Biogas and landfill gas 18 000 – 29 000High-Btu gas and substitute natural gas (SNG) 25 000 – 37 000

A distinction has deliberately been made bet-ween water gas and synthesis gas, although wa-ter gas was formerly used in large quantities forsynthesis as well. However, the so-called watergases are based on very specific raw materialsand belong to a different generation of produc-tion methods compared to ynthesis gases pro-duced in modern plants. Reduction gas refers togases used in iron production.

Synthesis gas is the basis for most of the hy-drogen (→ Hydrogen) produced, and it is also

used as a fuel gas, e.g., in gas turbines (see Sec-tion 7.3).

Biogas and landfill gas (→ Methane) are pro-duced during biological degradation of organicmatter. They may be collected and purified ei-ther to be used further (renewable energy) or toprevent air pollution.

1.1.1. Water Gas and Producer Gas

The terms water gas and producer gas were firstused in solid fuel gasification technology. Bothtypes of gas were very important from the 1920sto the 1940s. They were the basis for producingsynthesis gas, town gas, and fuel gas. The tech-nology is of only minor importance today be-cause the raw material basis has shifted largelyfrom solid fuels to liquid and gaseous fuels, andhigh-performance processes have been devel-oped for coal-based synthesis gas, town gas, andgrid gas.

Production of Water Gas.Water gas is usu-ally produced in an alternating operation, i.e.,by alternately hot blowing a fuel layer with airand gasifying it with steam. Among its vari-ous modifications, the production of so-calledblue water gas from coke was the most commonmethod, yielding gas with high hydrogen and

4 Gas Production

carbon monoxide contents which was suitablefor use as synthesis gas. By addition of higherhydrocarbons, the calorific value of water gascould be increased (carbureting) so it could beused as town gas.

Although most water gas plants were basedon coke, plants also operatedwith anthracite andnoncaking highly volatile types of coal (coal wa-ter gas or double gas). Coal with a high volatilematter content gives gases with higher calorificvalues because gasification gas is supplementedby a considerable amount of devolatilization gas.

In the water gas process with its cyclic modeof operation, coke has two functions: (1) it re-acts with blowing air during warm-up and withsteam during gasification, and (2) serves as aheat regenerator.

Because of the expensive rawmaterial (coke),complicated batch operation, and relatively lowoutput per gasifier (6000 – 9000 m3 (STP)/h),this method of producing synthesis and towngases has been replaced by more up-to-dateprocesses. The gas production of synthesis gasplants based on either natural gas (in tubular re-formers; see Chap. 2) or heavy oil (by partialoxidation; see Chap. 3), and of modern coalgasification plants (see Chap. 4) is ten timeshigher than that of water gas plants.

Production of Producer Gas. Producer gasis obtained by supplying a gasification agentconsisting of air and steam continuously, ratherthan cyclically, to a gas producer of the samegeneral design as a water gas producer. This gashas a lower calorific value than water gas be-cause it contains atmospheric nitrogen. Producergas was obtained from a variety of rotary grategas producers from the 1920s to the 1940s. Be-cause of its high nitrogen content, which cannotbe eliminated at economically justifiable cost,this gas is not suitable for use as synthesis gas;it was employed almost exclusively as fuel gasand as an additive to town gases in municipalworks.

In addition to coke, noncaking coals werealso used as feedstocks, and a few plants werealso built for the gasification of caking bitumi-nous coals under atmospheric conditions and un-der pressure (Lurgi pressure gasification).

Nature of Water Gas and Producer Gas.The following is a typical analysis of a watergas, in percent by volume:

Carbon dioxide 5 Hydrogen 50Carbon monoxide 40 Nitrogen 5

A typical analysis of a producer gas, in per-cent by volume, is

Carbon dioxide 5.5 Hydrogen 10.5Carbon monoxide 29.0 Nitrogen 55.0

The higher heating values (HHV) for wa-ter gas based on coke and for producer gas are10 500 – 12 500 kJ/m3 (STP) and 4600 – 6700kJ/m3 (STP), respectively.

The nitrogen content of water gas, amount-ing to ca. 5 vol%, does not originate from cokeor coal for the most part, but is due to the factthat air remains in the coke pores on transitionfrom the blowing to the gasification period. Thisnitrogen content can be lowered by discarding asmall fraction of the gas produced at the begin-ning of the gasification period.

1.1.2. Synthesis Gas and Reduction Gas

The gas production method used depends on therawmaterials. The rawmaterials are mainly nat-ural gas, naphtha, heavy vacuum residue, andcoal (see Chaps. 2, 3, 4). In addition, uncon-ventional fuels such as solid or liquid wastematerials or biomass gain interest. However, inthe recent decades most of the world produc-tion of synthesis gas was based on natural gasand sulfur-rich heavy vacuum residues. Becauseof more stringent environmental protection re-quirements, these residues may no longer beburned in boiler plants in many countries, sotheir market value is often low. Countries thatdo not have these two raw materials, particu-larly countries in which no natural gas is avail-able, often use naphtha. Depending on certainboundary conditions such as the relationship bet-ween oil and coal prices, as well as local fac-tors and strategic considerations, plants basedon coal will continue to be built.

Production of Synthesis Gas. Because syn-thesis gas is processed under elevated pressure,a trend toward operating gas production plantsat increasingly high pressure could be observed.Gas production at higher pressure reduces the

Gas Production 5

cost of compressing synthesis gas to the requiredsynthesis pressure.

Natural Gas as Raw Material (Chap. 2).Synthesis gases are produced from natural gasmainly in tubular reformers. If the atural gascontains virtually no nitrogen, a temperatureof 750 – 800 ◦C is sufficient for gas suitable inammonia synthesis.Residual unreactedmethaneis reformed in a secondary reformer (catalyticpartial oxidation with air) with simultaneousadjustment of its nitrogen content.

For normal methanol synthesis, reforming isperformed in one step in a tubular reactor at 850– 900 ◦C in order to leave as little methane aspossible in the synthesis gas.

For large methanol synthesis plants, Lurgihas introduced a two-step combination that givesbetter results. In the primary tubular reformer,lower temperature (ca. 800 ◦C) but higher pres-sure (2.5 – 4.0 MPa instead of 1.5 – 2.5 MPa)are applied. The secondary reformer operateswith a fixed catalyst bed as in the case of am-monia synthesis gas, but pure oxygen is used forpartial oxidation.

More recently, Lurgi developed another two-step gas production scheme. It is based on cat-alytic autothermal reforming with an adiabaticprereformer and has economical advantages forvery large methanol plants (Section 7.1).

The various tubular reformer designs differprimarily in the method of firing the catalyst-containing tubes (from the side or top, terraced,or from the bottom). Subsequent developmentsof combined reforming resulted in the integra-tion of the primary and the secondary reform-ing step in one apparatus using the thermal en-ergy of the effluent of the autothermal processfor heating the steam reformer part (M.W. Kel-logg, Uhde). Exxon tested the simultaneous re-forming and partial oxidation of natural gas in areactor with fluidized catalyst. Apart from tubu-lar reformer technology with or without a sec-ondary reformer, partial oxidation (in one step)is still used sporadically for reforming naturalgas. This can be performed either noncatalyt-ically (Shell, Texaco) or catalytically (Lurgi).The latter is used especially when low ratios ofhydrogen to carbon monoxide are required.

Naphtha as Raw Material (Chap. 2). Thepurity of naphtha allows the use of catalysts.Therefore, the production of synthesis and re-duction gases from naphtha is carried out almost

exclusively catalytically and at elevated pres-sure.

Externally heated tubular reformers con-taining special catalysts are used to producehydrogen-rich gases (e.g., ammonia synthesis).Two processes stand out from the others: theydiffer in the types of tubular reformers recom-mended and, primarily, in the type of nickel cat-alyst used. Topsøe uses a nickel catalyst withoutalkali; ICI employs one with a definite alkalicontent.

In future naphtha might replace methanol asa raw material for hydrogen produced on boardof vehicles driven by fuel cells. Due to the muchsmaller size of such production units as com-pared with plants in the petrochemical industryand due to restrictions by volume and by utilities(e.g., steam and water), the production schemeand above all the reactors and heat exchangersfor naphtha reforming in vehicles look rather dif-ferent.

Apart from single-stage tubular reforming,two-stage processes have also proved success-ful for the production of gases rich in carbonmonoxide, such as those required for methanoland oxo synthesis. An example of a two-stageprocess is the Recatro process, developed byBASF – Lurgi. Here, naphtha is converted withrelatively little steam in an initial low-temper-ature adiabatic stage over a special nickel cat-alyst to produce a rich gas containing ca. 60– 65 vol% methane plus hydrogen and carbondioxide. This gas is supplied together with theremaining steam to a conventional tubular re-former to be reformed to carbon monoxide-richsynthesis gas in the presence of a conventionalcatalyst. The two-stage operation makes carbondioxide recycling unnecessary in many cases orreduces it considerably.

A few plants using partial oxidation with(Lurgi) or without (Shell, Texaco) catalysts havebeen built.

Heavy Fuel or Residual Oil as Raw Material(Chap. 2). For the production of synthesis andreduction gases from heavy and residual oils,two processes are accepted worldwide: the Shelland the Texaco process, both operating underelevated pressure and noncatalytic. The Texacoprocess was originally designed for partial ox-idation of natural gas; the Shell process, devel-oped later, was tailored from the beginning togasification of heavy oil. The differences bet-

6 Gas Production

ween the two processes are virtually negligibleand involve only details of equipment design.

Coal as Raw Material (Chap. 4). Processesemploying coal that are used in practice oper-ate on a noncatalytic basis. Various processesoperating under elevated pressure are available:Lurgi, BGL (British Gas/Lurgi), HTW (HighTemperature Winkler), U-Gas, Texaco, Shelland Prenflo. The Koppers – Totzek process andthe original Winkler process operate at atmo-spheric pressure.Coal gasification plants for theproduction of synthesis gas built since the 1950sare based essentially on the Koppers or Lurgiprocesses.

In viewof the distribution of fossil fuels in theworld—almost ten times as much coal exists aspetroleum and natural gas, and it is distributedrelatively uniformly—and the sharp increase inoil prices in the 1970s, efforts have been madein many countries to develop new coal gasifi-cation processes tailored to the requirements ofparticular countries and their coal deposits.

Inmost processes, the heat required for gasifi-cation is provided by partial oxidation with pureoxygen. These processes are based essentiallyon the following three process arrangements:

1) moving bed,2) one or more fluidized beds, and3) entrained-flow systems in which coal dust

and the gasification agent are conveyed cocur-rently.

Combinations of these three process arrange-ments are being investigated, as are processesapplying high-temperature melts.

The overall arrangement of various coal gasi-fication plants, including downstream units, isdetermined essentially by the composition ofthe raw gases. The gases produced at increasedpressure by the Lurgi and by the BGL gasi-fication processes contain methane, which canbe converted to carbon monoxide and hydrogenfor downstream ammonia or methanol synthe-sis, or can be removed before entering the high-pressure compressor. The Koppers – Totzek, theTexaco, as well as the other entrained-flow pro-cesses yield a gas without methane, but witha higher specific oxygen consumption. As inall high-temperature gasification processes, gasfrom the latter two processes has a high car-bon monoxide content, requiring considerableconversion measures for certain syntheses. The

gasification processes applied to coal are alsosuitable for petrol coke if they can cope with itslow reactivity and they are also applied whenbiomass or waste materials are used as fuels.

Production of Reduction Gas. In principle,reduction gas production is very similar to thatof synthesis gas. However, iron ore is usuallyreduced under atmospheric conditions, so re-duction gas can be produced at atmosphericor slightly increased pressure. Reduction gasesused for the production of sponge iron by variousdirect reduction processes (→ Iron, Chap. 2.6)must meet certain requirements. Generally, theproportion of oxidizing components such assteam and carbon dioxide in the gasmixture pro-duced should be as low as possible, always <10%. This goal can be achieved by cooling thegases produced from coal, oil, naphtha, or natu-ral gas and removing carbon dioxide. However,because reduction gases in the shaft furnacesof steel works are needed at high temperature,more cost-effective solutions are being soughtand used. In the Midrex process natural gas isconverted at ca. 1000 ◦C in the tubular reformer,without the addition of steam, thereby producingonly hydrogen and carbon monoxide; the con-version is such that the reformed gases can beused directly in the shaft furnace.

As the recovery of byproducts from cokeoven raw gas (benzene, phenol) loses economi-cal attractiveness, reduction gas is considered tobe produced from coke oven raw gas by catalyticpartial oxidation.

Composition and Aftertreatment. Synthe-sis gas requirements depend on the intended ap-plication.

Ammonia synthesis units now generally op-erate between 13.0 and 22.0 MPa and requiregas with a hydrogen: nitrogen ratio of 3: 1. Im-purities are limited in modern processes to < 1ppm for sulfur and 2 – 10 ppm for oxygen-con-taining components such as oxygen itself, car-bon dioxide, and water. With a view to the econ-omy of synthesis (i.e., the purge gas quantitiesto be branched off), the methane content of syn-thesis gas must not exceed 0.5 – 0.8 vol%.

This means that the reformer and gasifica-tion gases produced from various raw materialsmust be conditioned appropriately. If the meth-ane content is originally considerably high it can

Gas Production 7

be lowered by secondary reformingwith oxygenor air. Because all reformedgases contain carbonmonoxide, they have to be converted, irrespec-tive of the rawmaterials fromwhich they are pro-duced; this is done by catalytically reacting car-bon monoxide with steam to form hydrogen andcarbon dioxide (shift conversion). Depending onits residual content after shift conversion, eitherthe carbonmonoxide concentrationmust be low-ered (to the ppm range) in a methanation stageor it must be removed at low temperatures by aliquid nitrogen wash. Hydrogen sulfide, organicsulfur compounds, and carbon dioxide have tobe removed by appropriate washing procedures.

Similar conditioning and purification pro-cesses have to be applied for the production ofhydrogen for fuel cells (→ Fuel Cells).

Methanol is generated from carbon monox-ide as well as from carbon dioxide:

2 H2+CO�CH3OH

3H2+CO2 �CH3OH+H2O

The preferred reaction path goes viacarbon monoxide, some carbon dioxideis required for kinetic reasons. The ratio(cH2 − cCO2)/(cCO + cCO2) should generallybe between 2.0 and 2.2. The theoretical valuederived from adding both reaction equationsafter pondering each of them arbitrarily is 2.0.A higher carbon monoxide content is requiredfor fuel methanol synthesis (Lurgi Octamix andothers), in which a mixture of methanol andhigher alcohols is produced as an additive tomo-tor fuel. For high-pressure methanol synthesis(35.0 MPa), which is still used in a few cases, thepresence of methane adds an economic burden(as in ammonia synthesis). It is less importantin low-pressure (5.0 – 10.0 MPa) methanol syn-thesis (ICI, Lurgi) because compression costsfor this gas component are practically negligiblebecause of the lower compression ratio required.The content of sulfur compounds in the gas usedfor low-pressure methanol synthesis, which isbased on a special copper-containing catalyst,must be extremely low. The amount of hydro-gen sulfide and organic sulfur compounds in thesynthesis gas produced from raw materials withrelatively high sulfur content can be lowered toca. 0.1 ppm by special processes (Rectisol forcoal and oil gasification gases; hydrogenationof the organic sulfur compounds in the presence

of Comox and Nimox catalysts, and subsequentadsorption prior to the reforming of naphtha andnatural gas in the tubular reformer).

Purities applicable in methanol synthesis arealso required for oxo synthesis. However, be-cause the latter often requires a hydrogen: car-bon monoxide ratio of ca. 1: 1, the carbon: hy-drogen ratio of the feedstock may require thatcarbon dioxide be recycled or fed to the gasifi-cation unit.

For iron ore reduction gas, the sulfur contentmust be reduced to a few parts per million. Table1 shows typical analysis data for raw gases fromvarious starting materials.

Table 1. Typical raw gas analyses

Feedstocks and processes Content, vol%

H2 CO CH4 C2+ N2+Ar

CO2 H2S +COS

CoalLurgi pressure

gasification43 12 11.5 1 0.3 32 0.2

BGL 28 57 7.1 0.4 4.2 3 0.3Koppers – Totzek 30 55 0.1 1.4 13.2 0.3Texaco 37 41.3 0.1 0.6 20.7 0.3

Oil gasificationShell – Texaco 46 47 0.5 0.6 5.5 0.4

NaphthaTubular reformer 67 19 3 11Recatro (BASF –

Lurgi)66 22 4 8

Natural gasTubular reformer 73 16 4 7

1.1.3. Town Gas and Medium-Btu Gas

Because natural gas (→ Natural Gas) has be-come the preferred fuel gas for households andindustry, the production of town and medium-Btu gas from coal and liquid fuels is of minorimportance today. High-temperature coking —formerly general practice in town gas works—has disappeared almost completely (→ Coal Py-rolysis). The following is a brief outline of gasifi-cation processes for the production of town gas.

Basis of Production.When charcoal was re-placed in the last century by coke from bitumi-nous coal to smelt iron in blast furnaces, largequantities of degasification gas were producedby the carbonization of caking coals. Initially,these gases were largely flared; later on, theywere distributed through piping systems supply-ing gas particularly to the municipalities in the

8 Gas Production

neighborhood of steelworks. The gas grids werelarge enough to provide sufficient gas from thecoke ovens even at times of peak demand; onSundays and holidays, the gas was flared in thecoking plants.

With the increasing demand for grid gas, uti-lization of the gas system improved, so thattowns had to build their own gas production fa-cilities to meet peak requirements. These facil-ities were based mainly on water gas, which insome cases was carbureted.

Both solid and liquid or gaseous fuels wereused as raw materials for town gas production.Liquid hydrocarbons such as fuel oil and, lateron, naphtha were preferred, but coal was alsoused.

Plants using fuel oil are based mainly on acyclic process arrangement and became verysuccessful since the 1950s. Some of them usestraight thermal processes without catalysts, butmost employ a catalyst bed. During a blowingphase, a bed of catalytic or noncatalytic materialis heated by the combustion of fuel oil with air.After switching to the gasification period, fueloil is passed over the hot bed along with steam.The ONIA GEGI and SEGAS processes and theKoppers Cowper process have become widelyknown [1].

The disadvantages of these cyclic processesare their relatively low output and difficultiescaused by excessive wear of the switchovermechanism. In addition, as environmental lawstightened toward the end of the 1950s and be-ginning of the 1960s, plant operators turned toprocesses using lighter hydrocarbons and con-tinuous operation. The development of specialcatalysts (ICI, Topsøe) permitted gas productionfrom readily available liquid feedstocks (such aslighter hydrocarbons and naphtha having a boil-ing end point up to 200 ◦C) to be carried out inexternally heated tubular reformers. Apart fromits higher efficiency, this mode of operation hadan additional advantage because the gases pro-duced were available at increased pressure, thusreducing the expenditure for their distributionin pressure gas systems. The methane content ofthe gases could be increased by keeping reac-tion temperatures in the tubular reformer low, sothat after a cold carbureting stage with propane– butane, standard town gas could be produced.

Cold carburetingwas often not applicable be-cause either propane – butane was not available

at a reasonable price, or certain standards withrespect to dew-point were not met, therefore,modified processes aiming at increased produc-tion of methane from naphtha (BASF/Lurgi,British Gas, Japan Gasoline Corp.) were devel-oped. Because methane content increases withdecreasing temperature according to thermody-namic equilibrium, appropriate catalysts had tobe developed. In these processes, a so-called richgas having a methane content of ca. 65 vol% isproduced from naphtha and steam by a slightlyexothermic reaction in a shaft reactor at 400– 490 ◦C. Because the calorific value and spe-cific gravity of this gas are too high, comparedwith pertinent standards for town gas, part ofit must be reformed in a downstream tubularreformer using commercial catalysts and thenremixed with the residual methane-rich gas. An-other method of obtaining a leaner gas is to addnitrogen or air.ManyEuropean cities built plantsof this type with town gas capacities up to 1 ×106 m3 (STP)/d.

Plants that converted propane – butane totown gas via externally heated tubular reformerswere built particularly for small capacities. Partof the propane – butane was added to increasethe calorific value of the reformed gas after car-bon monoxide shift conversion. A few plantsproduced town gas by reforming natural gas inexternally heated tubular reformers for transi-tional periods. This was done mainly when theconsumers’ existing piping systems and burnerswere to be used, even if a loss in efficiency hadto be accepted in converting natural gas to towngas.

Oil gasification by partial oxidation (Shell,Texaco, Lurgi; see Section 3.2.2) has not be-come important in town gas production becauseof the high investment cost of such plants. Thegases produced, consisting essentially of hydro-gen and carbon monoxide must be carbureted tocomply with standards.

Along with the development of gas produc-tion processes, downstream gas conditioningand gas cleaning units were adapted as well.Carbon monoxide shift conversion plants oftenreduced the carbon monoxide content to < 3vol%, sometimes even to < 1 vol%, to elimi-nate the danger of poisoning.

The composition of town and grid gas is givenin Tables 42 and 43 (Chap. 8). Group A in thesetables corresponds to town gas; Group B, to grid

Gas Production 9

gas; and the first gas family, to town and gridgases combined.

1.1.4. Biogas and Landfill Gas (→ Methane)

Biogas is the gaseous reaction product fromanaerobic digestion of wastewater having a highBOD value or of sludge from wastewater treat-ment plants. It is used solely as fuel for combus-tion engines to meet the power requirements ofthe treatment plant.

Landfill gas is produced by the degradationof organic matter in the waste disposed of inlandfills under anaerobic conditions. Becauseaccess to oxygen is not strictly controllable,aerobic degradation may occur simultaneously,producing water and carbon dioxide. Landfillgas is recovered via pipelines buried in the land-fill. For existing landfills, wells are drilled andlined with tubes to allow the gases to be ex-tracted. Landfill gas is used mostly as fuel gasfor cogeneration and in combustion engines af-ter suitable purification.

Composition. Biogas is composed of 60 –70 vol% methane and 30 – 40 vol% carbondioxide, depending on the feed and digester tem-perature (mesophilic – thermophilic).

Landfill gas is composed of 50 – 55 vol%methane, 41 – 45 vol% carbon dioxide, and 1– 4 vol% nitrogen.

Depending on the nature of landfill waste, thegas may also contain considerable amounts ofhalocarbons.

1.1.5. Rich Gas and Substitute Natural Gas(SNG)

Basis of Production. As a result of thedoubts concerning gas supplies and their avail-ability to the industrial nations in the 1970s ef-forts were made to produce synthetic natural gasfrom the plentiful coal supplies. These effortswere feasible where natural gas accounted fora high percentage of primary energy and cheapdomestic coal was available. Thus, a plant wasbuilt in North Dakota (United States) which hasgone into operation in 1984, producing ca. 3.9× 106 m3 (STP)/d of SNG from lignite. In this

plant, lignite is gasified in Lurgi pressure gasi-fiers, and the methane content required is ad-justed by reacting carbon monoxide and hydro-gen from the raw gas in a methanation stagedownstream of the gas cleaning unit. In addi-tion, the Lurgi gasification process has been de-veloped further to produce SNGmore economi-cally by increasing gasification pressure to 10.0MPa (Ruhr 100, see Section 4.4). Here, morethan 50 vol% of methane recoverable from coalis produced in the gasification stage proper.

Heavy vacuum residues can also be usedas feedstock for SNG production. However, asin coal-based plants, capital investment is rela-tively high. The discovery of huge new naturalgas resources in some areas of the world duringthe past few years has made the production ofSNG less attractive.

Composition. Rich gases have a highercalorific value than grid gas, with gross calorificvalues from about 25 000 kJ/m3 (STP) upward.A rich gas produced from naphtha (British Gasor BASF – Lurgi process) has the following typ-ical analysis, in percent by volume:

Carbon dioxide 21.6 Hydrogen 12.5Carbon monoxide 0.3 Methane 65.6

Substitute natural gas should differ as littleas possible from natural gas. It must be condi-tioned so that it can be substituted for naturalgas. This depends on the following parameters:higher heating value, density, and Wobbe index(see Section 8.2.1.3).

In addition the specified content of impuri-ties, particularly sulfur, must not be exceeded.Typical SNG produced in a coal gasificationplant in North Dakota has the following analy-sis, in percent by volume (unless otherwise stip-ulated):

Methane 96.3 Nitrogen + argon 1.1Carbon monoxide 0.02 Carbon dioxide 1.2Hydrogen 1.38 Hydrogen sulfide 4 ppm

1.2. Raw Materials for Gasification

Gaseous Fuels. Natural gas as raw materialis of paramount importance for gasification orreforming (the term reforming is used more

10 Gas Production

commonly for gases) in externally heated tubu-lar reformers to yield synthesis gas for variousapplications. Plants also reform natural gas bypartial oxidation. The composition of naturalgases is not important for the technology to beapplied, as long as their inerts content is negligi-ble. However, if the nitrogen content rises above5 – 10 vol%, this may have consequences forthe design of ammonia synthesis gas plants op-erated on the tubular reformer principle.

In certain new energy scenarios, natural gasstill plays an important role. Thus, the energyrequired for natural gas reforming may be sup-plied—at least partially—by solar collectors orhigh-temperature nuclear reactors. Liquidmotorfuels can be synthesized from synthesis gas. Inthis way, thermal energy, which is difficult tostore and utilize, is converted to energy carriersbetter adapted to the requirements.

Coke oven gas (→ Coal Pyrolysis) is of mi-nor importance for producing synthesis gas be-cause special expenditures are necessary to re-move various impurities from the coke oven gas,and compression costs are high. In older plants,coke oven gas was separated into its individualcomponents by fractionated condensation anddistillation at low temperature rather than beingreformed (the content of methane and higher hy-drocarbons is low anyway); these componentswere then used for synthesis gas production.

Liquid Fuels. In most cases, various frac-tions of crude oil are used for gasification. Onlyin very rare cases is unrefined crude oil useddirectly for gasification because neither its eco-nomic nor its technical considerations are opti-mum.

Liquefied petroleum gas (LPG, a propane –butane mixture; → Liquefied Petroleum Gas) isused for the production of town gas, particularlyin smaller capacity plants, and for peak shaving.It is virtually sulfur-free and can be processedwithout significant problems. Light naphtha isutilized for larger town gas plants and for theproduction of synthesis gas. Generally, fractionsboiling at 40 – 110 ◦C are used, but full-rangestraight-run naphthawith a boiling end point of150 – 200 ◦C may also be employed. Depend-ing on its origin, this naphtha has a sulfur con-tent between 30 and 300 ppm. Naphtha must becompletely desulfurized because almost all con-ventional processes for the production of townor

synthesis gas include a reaction stage involvingcatalysts that are very sensitive to sulfur poison-ing.

In some cases, LPG and naphtha have alsobeen converted to town and synthesis gas by par-tial oxidation (thermally or catalytically) withoxygenor oxygen-enriched air. For catalytic par-tial oxidation, desulfurization to 20 – 50 ppm issufficient.

Synthesis gas is produced from high-boilingresidues by partial oxidation with oxygen. Orig-inally, the high-boiling fraction (initial boilingpoint above 350 ◦C) from atmospheric distilla-tion units of refineries (→ Oil Refining)— theheavy fuel oil—was used for this purpose. Todate even higher viscosity fractions from vac-uum distillation and conversion plants are beingused increasingly. Because of the more stringentrequirements imposed by environmental legisla-tion, these heavy oils can no longer be burned inboiler houses and are, therefore, cheap rawmate-rials for gasification, which also eliminates var-ious pollutants in the residue such as sulfur andheavy metals. To be used in gasification plants,the heavy oil must be suitable for pumping andspraying. Therefore, the oils are preheated to at-tain the required viscosity.

The ash content of the heavy oils and the com-position of the ash affect the operating temper-ature of gasification plants and the service lifeof the refractory lining. Vanadium and nickel inash are particularly harmful.

Apart from petroleum fractions, other liquidfuels may be used for gasification, although theyare not very important. These include:

1) tars from bituminous coal and lignite for gasi-fication by partial oxidation;

2) fractions fromshale oil (→ Oil Shale) produc-tion, particularly heavy fractions, for gasifica-tion by partial oxidation;

3) methanol for producing peak shaving gas incatalytic town gas plants;

4) hydrocarbon condensates from natural gasproduction for gasification by partial oxida-tion; and

5) waste from chemical plants (e.g., isobutyr-aldehyde).

Purity Requirements for Liquid andGaseous Fuels. Raw materials for gasificationplants using catalysts must generally be very

Gas Production 11

pure. Therefore, in most cases, such plantsare equipped with special upstream purificationunits. A certain tolerance regarding the con-tent of impurities can be accepted in cyclic pro-cesses that regenerate catalysts in the respectivecycles. However, continuous processes impor-tant to modern technology require exceptionallypure fuels, especially with respect to sulfur com-pounds. Catalysts used for these processes thuspermit the use only of gaseous fuels (natural gas)and of liquid hydrocarbons up to a boiling endpoint of ca. 200 ◦C.

Generally, the lower the operating temper-ature of the catalytic process, the lower is theresidual sulfur content that can be tolerated inthe raw materials. How adverse an effect sulfurcompounds in the gas have on the active sites ofcatalysts can be derived from equilibrium con-ditions governing the reaction of the metallicphase of the catalyst with the sulfur compounds.Most gasification catalysts contain nickel as anactive substance. At constant partial pressure ofsulfur compounds in the gas, nickel sulfide for-mation increases with decreasing temperature; asulfided catalyst no longer provides the desiredresults regarding reaction direction (danger ofcarbon formation), speed, and selectivity.

Therefore, light hydrocarbons are desulfur-ized to < 0.1 ppm of sulfur upstream of plantsoperated at low catalyst temperature (400 –700 ◦C). The sulfur content may often be higherby one or two orders of magnitude in catalyticprocesses operating above 850 ◦C.

Solid Fuels. Various solid fuels are not allequally suitable for gasification. In addition, dif-ferent gasification processes have different fuelrequirements.

Because endothermic gasification reactionsshould be as complete as possible, the suitabilityof solid fuels for gasification improves as theirreactivity increases. This reactivity can be corre-lated with the oxygen content, which decreaseswith increasing carbonization (i.e., with increas-ing age of the fuel).

The classification and designation of varioustypes of coal are different in individual industrialcountries (→ Coal, Chap. 4).As a rule, however,they depend on volatile matter, which almost al-ways means classification by age of the coal.

Additionally, the following properties of fu-els are essential for assessing their suitability forgasification:

Ash content Coal grain sizeWater content Melting behavior of the ashSulfur content Behavior of the coals on heatingCaking properties

Because coal gasification adds considerablyto the costs of upgraded energy (i.e., the gas),coal used for this purpose should be as cheapas possible. Wherever possible, only coal that isnot suitable for other applications and, therefore,has a lowmarket price should be used. Coal pro-duced by open-pit mining is of special interestfor gasification.

Almost all solid fuel such as wood, peat, lig-nite, young and old bituminous coal, and coke,including petrol coke, has been used for gasifi-cation. In the future, more use should be madeof lignites and younger noncaking or caking bi-tuminous coal, as well as coal with a high ashcontent (20 – 50 wt %) or sulfur-rich coal whichis rarely suitable for other applications.

With regard to various types of petrol coke(delayed, fluid, Lurgi – Ruhr gas (LR) coke, →Oil Shale, ), their high sulfur and heavy-metalcontent (→ Petroleum Coke) renders combus-tion increasingly difficult in view of pollutioncontrol measures. Petrol coke is, therefore, turn-ing into a new potential raw material for gasifi-cation.

Gasification of wood has been revived be-cause of increasing interest in renewable rawmaterials. Emphasis is mainly on the productionof clean fuel gas (low sulfur content of wood),but synthesis gas production is also possiblewithmodern fluidized-bed processes.

Suitability of Fuels for Gasification. Con-siderable differences exist in the complexity ofgasification technology applied, depending onwhether solid, liquid, or gaseous fuels are used.A general rule is that capital cost and consump-tion figures are higher for solid than for liquidfuels, and for liquid higher than for gaseous fu-els. Therefore, in the production of particulargases,mere comparison of specific energy pricesof various fuels is insufficient; individual costsfor gasification must be considered as well (seeTable 2).

12 Gas Production

Table 2. Comparison of capital cost and energy requirements foran ammonia plant and a methanol plant with a capacity of 1000 t/dbased on gaseous, liquid, and solid fuel

Natural Heavy Coalgas oil

Ammonia plantCapital cost,% 100 170 225Energy requirements,% 100 115 135

Methanol plantCapital cost,% 90 150 200Energy requirements,% 95 105 125

To make end products competitive, the spe-cific energy price of coal must be lower thanthat of heavy oil, and the latter must be lowerthan that of natural gas. In principle, this alsoapplies to other solid and liquid fuels with re-gard to the complexity of the conversion plantbased on these particular feedstocks.

Even if fuels of the same physical state arecompared, great differences exist in the proper-ties of individual substances, which also affecttheir suitability for gasification.

1.3. Physicochemical Basis for GasProduction

Gasification can be described by overall chem-ical reactions, from which conclusions may bedrawn regarding the conversion of materials andheat of reaction.

Gasification of solid and liquid fuels involvesheterogeneous reactions at the phase boundaryof particles or droplets and homogeneous reac-tions in the gas phase. The latter include reac-tions of components that devolatilize at operat-ing temperature.

The following principles form the basis forcalculating gasification processes:

1) Material Balance. In establishing a materialbalance, the total number of atoms of the in-dividual elements in the reactants is equalizedto that in the products, irrespective of the dis-tribution of the elements between the individ-ual components (principle of conservation ofmass).

2) Energy Balance. Energy balance is based onthe principle of conservation of energy. Con-sequently, energy supplied to the system (en-thalpy of reactants and heat supplied) is equalto energy removed from it (enthalpy of prod-ucts and heat removed). In determining en-

thalpy, formation and phase change energiesmust be considered.

3) Chemical Equilibrium. The equilibrium com-position of the gasification product is deter-mined by the temperature-dependent equilib-rium constants of reactions going on in thesystem and by the operating pressure.

4) Chemical Kinetics. Because equilibrium isnot attained when residence time is too short(for example, because the temperature is toolow), the composition of the gasification prod-uct in this case is determined by the kinetics ofthe reactions. In addition, transport phenom-ena, such as boundary layer or pore diffusion,may have some influence (see page 15.

Practical calculation leads to nonlinear equa-tion systems in which the number of unknownquantities is determined essentially by the num-ber of components in the raw gas. The solu-tion may be found iteratively or— if appropri-ate routines are used—simultaneously. Calcu-lations are particularly complicated when chem-ical kinetics must be considered.

Reaction Enthalpy of Gasification. Gasifi-cation reactions can be described bymanydiffer-ent equations. To calculate the reaction enthalpyof gasification, however, a limitation to certaincharacteristic reaction groups is sufficient. Thereaction enthalpy of all other reactions can bederived by combining these basic equations, andindividual reactions have no influence on the to-tal heat balance of a sequence of reactions if thesame feedstocks are turned into the same prod-ucts (Hess’s law).

The reaction groups characteristic of gasifi-cation can be subdivided as follows (∆H refersto 0 ◦C and 0.1 MPa as reference state; ∆H isnegative for exothermic reactions and positivefor endothermic reactions):

1) Reactions with molecular oxygen (combus-tion)

C+1/2O2�CO∆H = −110.62 kJ/mol (1)

CO+1/2O2�CO2∆H = −283.15 kJ/mol (2)

H2+1/2O2�H2O∆H = −242.00 kJ/mol (3)

CnHm+(n+m/4) O2�nCO2+m/2 H2O (4)

Gas Production 13

For example,

CH4+2O2�CO2+2H2O∆H

= −802.86 kJ/mol (4a)

C2H6+7/2O2�2 CO2+3H2O∆H

= −1428.8 kJ/mol (4b)

C3H8+5O2�3 CO2+4H2O∆H

= −2045.4 kJ/mol (4c)

The reaction enthalpy of combustion of othergaseous or liquid fuels is described in the lit-erature (e.g., [2 – 4]).For partial combustion of gaseous or liquidfuels to carbon monoxide and steam,

CnHm+(n/2+m/4)O2�nCO+m/2 H2O (5)

the reaction enthalpy can be calculated fromthe enthalpies of Equations (4) and (2) as fol-lows:

∆H (Eq. 5) = ∆H (Eq. 4)−n∆H (Eq. 2)

2) Reactions with steam

C+H2O�CO+H2∆H = +131.38 kJ/mol (6)

CO+H2O�CO2+H2∆H = −41.16 kJ/mol (7)

(homogeneous water gas reaction, shift con-version)

CnHm+nH2O�nCO+(m/2+n)H2 (8)

For example,

CH4+H2O�CO+3H2∆H

= +206.28 kJ/mol (8a)

C2H6+2H2O�2 CO+5H2∆H

= +347.50 kJ/mol (8b)

C3H8+3H2O�3 CO+7H2∆H

= +498.06 kJ/mol (8c)

The enthalpy of the reaction of fuels withsteam to produce carbon dioxide and hydro-gen

CnHm+2 nH2O�nCO2+(m/2+2 n)H2 (9)

can be determined by

∆H (Eq. 9) = ∆H (Eq. 8)+n∆H (Eq. 7)

because Equation (9) may be considered acombination of Equations (8) and (7).

3) Reactions with carbon dioxide

C+CO2�2 CO∆H = +172.54 kJ/mol (10)

(Boudouard reaction)

CnHm+nCO2�2 nCO+m/2 H2 (11)

For example,

CH4+CO2�2 CO+2H2∆H

= +247.44 kJ/mol (11a)

C2H6+2CO2�4 CO+3H2∆H

= +429.82 kJ/mol (11b)

C3H8+3CO2�6 CO+4H2∆H

= +621.53 kJ/mol (11c)

Generally, the reaction enthalpies for hydro-carbon gasification with carbon dioxide (Eq.11) are obtained from the enthalpies of Equa-tions (8) and (7) as follows:

∆H (Eq. 11) = ∆H (Eq. 8)−n∆H (Eq. 7)

4) Hydrocarbon decomposition reactions (sootformation)Hydrocarbon decomposition is described by

CnHm�nC+m/2 H2 (12)

For example,

CH4�C+2H2∆H

= +74.91 kJ/mol (12a)

C2H6�2 C+3H2∆H

= +84.74 kJ/mol (12b)

C3H8�3 C+4H2∆H

= +103.92 kJ/mol (12c)

14 Gas Production

may be considered the reverse of hydrocarbonformation from its elements and is essentialfor the gasification of hydrocarbons.The reaction enthalpy can be calculated fromthe combustion enthalpies of the reactants:

∆H (Eq. 12) = ∆H (Eq. 4)−n∆H (Eq. 1)

−n∆H (Eq. 2)−m/2∆H (Eq. 3)

In calculating the main gas components, sidereactions of sulfur compounds (H2S, COS, etc.)and nitrogen compounds (NH3, HCN) are gener-ally neglected because they have no great effecton the numerical distribution of the main gascomponents. If detailed information on the dis-tribution of secondary components is required,further calculations can be made and the resultsused for the main gas components.

Chemical Equilibrium. The gasification re-actions represented by Equations (1) – (12) arenever complete in either of the two possible di-rections but tend to reach an equilibrium ex-pressed by the equilibrium constant Kp . ForEquation (8 a), for example, this constant isgiven by

Kp =pCO·p3H2

pCH4 ·pH2O=

xCO·x3H2

xCH4 ·xH2OP 2 = f (T )

where xi represents the mole fraction and pi thepartial pressure of each of the four components,P the total pressure, and T the temperature.

For Equation (7), the carbon monoxide shiftconversion (which in nearly any case has to beconsidered) the equilibrium constant is given by

Kp =pCO2 ·pH2

pCO·pH2O=

xCO2 ·xH2

xCO·xH2O= f (T )

In comparison to the steam reforming reaction(8 a) the equilibrium composition of the shiftconversion is not influenced by the total pres-sure, because simply spoken, the number ofmolecules on the right and on the left side ofEquation (7) is identical. The temperature de-pendence of the equilibrium constant (log Kpvs. 1/T ) is shown in Figure 1.

According to the laws of thermodynamics,the equilibrium constant Kp of a reaction can becalculated by means of the Gibbs free energy ofthe reaction or the free formation energies of thereactants.

Figure 1. Water gas equilibrium, Kp = f (1000/T )

Strictly speaking, the equilibrium constantKp in the above form applies only to ideal gases.At increased pressure (total of ca. 1.0 MPa ormore) and near the boiling points of the liquids,the appropriate partial fugacities should be usedinstead of partial pressures. These fugacities al-low for the real behavior of individual gas com-ponents. However, in practice, partial pressurescan be used for the calculation with sufficientaccuracy in most cases.

In heterogeneous reactions, partial pressuresof the solids are not taken into account becausethey are regarded as a function of temperaturealone, which is included in the equilibrium con-stant Kp . However, the free formation energyof the solids must be considered in a thermody-namic calculation of the equilibrium constants.

The three reactions described by Equations(7), (8 a) and (10) are almost always sufficientto calculate the theoretical product gas compo-sition of various gasification processes. Table 3shows the equilibrium constants of these threereactions at different temperatures. In additionequilibrium constants for heterogeneous meth-ane decomposition (Eq. 12 a) are given as anexample of hydrocarbon decomposition. Thisequilibriummay also be derived from Equations(7), (8 a), and (10). Additionally equilibriumconstants forCOShydrogenation and hydrolysisare included (see Section 5.2).

COS+H2�H2S+COCOS+H2O�H2S+CO2

The temperature dependence of the equilib-rium constants can be calculated with an equa-tion of the following type

Gas Production 15

Table 3. Temperature-dependent equilibrium constants of various reactions

Equation (7) (8 a) (10) (12 a) COShydrogenation

COS hydrolysis

Unit for Kp bar2 bar barTemperature, ◦C

100 12.4 4.39×104

200 19.1 4.26×103

300 38.5 6.13×10−8 3.60×10−7 4.36×10−3 24.4 940400 11.5 5.54×10−5 8.55×10−5 5.56×10−2 28.0 323500 4.82 9.16×10−3 4.99×10−3 3.76×10−1 30.3 146600 2.51 4.88×10−1 1.16×10−1 1.66 31.5 79.1700 1.52 11.8 1.42 5.41 32.1 48.7800 1.02 159 11.0 14.1 32.2 32.9900 7.41×10−1 1.38×103 59.8 31.2 32.0 23.7

1000 5.71×10−1 8.63×103 249 60.61100 4.60×10−1 4.12×104 8451200 3.84×10−1 1.59×105 2.43×103

1300 3.30×10−1 5.16×105 6.08×103

1400 2.91×10−1 1.48×106 1.36×104

1500 2.61×10−1 3.69×106 2.79×104

logKp = A+B/T+C logT+D·T+E·T 2

where T is the absolute temperature in Kelvin.For the various equilibria used in this article thevalues of the constants A – E are given in Table4. (The data of Tables 3 and 4 were taken fromthe ASPEN program [5] which uses in this casethe correlations cited in [6]. Minor differenceswith respect to other literature data may occur.)

Table 4a. Values of coefficients for calculating the equilibriumconstants of various reactions as a function of absolute temperature

Equation (7) (8a)A −8.000980 −18.98177B 2456.620 −9520.569C 1.984098 10.91012D −0.3329441×10−3 −0.319556×10−2

E 0.563315×10−7 0.3631894×10−6

Table 4b. Values of coefficients for calculating the equilibriumconstants of various reactions as a function of absolute temperature

Equation (10) (12a)A 2.220419 −21.61666B −8709.775 −2875.50C 2.495100 9.505546D −0.6399432×10−3 −0.3694027×10−2

E 0.5280003×10−7 0.5233708×10−6

Table 4c. Values of coefficients for calculating the equilibriumconstants of various reactions as a function of absolute temperature

Equation COS hydrogenation COS hydrolysisA 5.956145 4.485945B −566.2154 1568.825C −1.309130 −1.681947D 0.5766548×10−4 0.7452208×10−3

E −0.1222679×10−7 −0.1160356×10−6

A general assessment of tendencies in gasifi-cation reactions is facilitated by the Le Chatelier– Braun principle that a system in equilibriumtries to evade a change forced upon it. For ex-ample, in Equation (8 a), an increase in totalpressure results in increased methane formationbecause the number of gaseous molecules onthe right side of Equation (8 a) is higher thanthat on the left. A pressure increase in reactionsproceeding without a change in the number ofmolecules, such as the homogeneous water gasreaction (Eq. 7) has almost no effect.

Reaction Sequence for Various Gasifica-tion Processes. No statements can be madeon the basis of thermodynamics regarding thechronological sequence of gasification reac-tions. Calculations must include reaction ki-netics. Frequently, however, the required dataare not available. Approximation methods mustthen be used to calculate product gas compo-sition. A common method involves use of Kpvalues which, according to experience, reflectthe actual final state but pertain to a tempera-ture other than that prevailing at the end of thereactor.

Some processes employ catalysts to acceler-ate the reaction and enhance reaction kinetics inorder to approach equilibrium. Selective cata-lysts promote desirable reactions at the expenseof undesirable ones (e.g., carbon formation ac-cording to Boudouard). Generally, reaction rateincreases as temperature increases. Gasification,

16 Gas Production

therefore, approaches equilibriummore easily athigher than at lower temperature.

Tubular Reactor Reforming. When reform-ing hydrocarbons, especially methane, in exter-nally heated tubular reactors, conversion of thehydrocarbons with steam to yield carbon oxidesand hydrogen (according to Eqs. 8 and 9) mustbe considered basic reactions. The compositionof the product gas is determined by the simul-taneous equilibrium between Equations (7) and(8 a).

Because soot formation should be sup-pressed, the Boudouard reaction (Eq. 10) andhydrocarbon decomposition (Eq. 12 and, espe-cially, 12 a) are essential for an assessment ofthe reaction system.

Rich-Gas Process. When naphtha is re-formed catalytically with steam in a shaft reac-tor to yield gases with a high methane content,the endothermic gasification reactions (Eqs. 8and 9) must be considered as starting reac-tions, whose products are then transformed inan exothermic methanation reaction describedby Equation (8 ). The final condition of the re-forming process is determined by Equations (7)and (8 a).

Oil Gasification. In oil gasificationwith oxy-gen and steam, partial combustion of hydrocar-bons (Eq. 4), as well as steam and carbon diox-ide gasification of carbon (Eqs. 8 and 11), mustbe considered to proceed more or less simulta-neously. Product gas composition is determinedby the final state of the simultaneous equilibriumamong the reactions described in Equations (7),(8 a), and (10).

Coal Gasification. The first step in a coalgasification reaction (e.g., with oxygen andsteam in countercurrent mode of operation) isexothermic combustion of carbon to carbondioxide, according to Equations (1) and (2). Thecarbon dioxide formed then reacts with hot car-bon to yield carbonmonoxide in an endothermicreaction (Eq. 6). This in turn reacts with steam toproduce either hydrogen and carbon dioxide (ac-cording to Eq. 7) or methane and carbon dioxide(by the reverse of Eq. 11 a). The latter occursparticularly at increased gasification pressure.Carbon monoxide and hydrogen are also pro-duced by direct endothermic carbon gasificationwith steam (Eq. 6). In entrained-flow gasifica-tion processes, the reaction sequence describedis mostly overlapping. The mode of reaction es-

sentially determines the temperature profile in agas producer.

The composition of gasification gas is deter-mined by a more or less accurate adjustment ofthe simultaneous equilibrium among the shiftconversion reaction (Eq. 7), themethane reform-ing reaction (Eq. 8 a), and the Boudouard reac-tion (Eq. 10).

1.4. Characteristics of the BasicProcesses

In most cases, the aim is to produce gases withhigh hydrogen and carbon monoxide content.This applies to all synthesis gases and also totown gas. Depending on the particular require-ments, the carbon monoxide content may be re-duced with steam to yield hydrogen and carbondioxide. To produce such gases from varioussolid, liquid, and gaseous fuels, heat must besupplied to support the endothermic reactions.Thus, different processes can be classified ac-cording to the type of heat supply.

On the other hand, processes for producinggases with high methane content, such as richgas and SNG, in principle require no exter-nal heat supply because methane formation isexothermic. Therefore, the process characteris-tics of this group are different.

Trends in the past decades have shifted fromprocesses operated under atmospheric condi-tions to those operated at elevated pressure. Thishas resulted in economic advantages becausethe subsequent chemical syntheses [ammonia,methanol, oxo (→ Oxo Synthesis ) and Fischer–Tropsch synthesis (→ Coal Liquefaction)] op-erate at more or less elevated pressure, which re-moves the cost of compressing the high-volumesynthesis gas. Moreover, plants operating underpressure can be designed for higher capacity dueto their smaller gas volume, which again meanssavings in capital costs.

An essential criterion is whether catalystsare used. With catalysts, a better approximationto reaction equilibria is attained and undesiredbyproducts are mostly avoided. However, cata-lysts can thus far be used successfully only forgaseous feeds and for distillate liquid feeds.

Cyclic Processes. Cyclic processes use heatcarriers consisting of refractory brick lining (in

Gas Production 17

regenerators, air heaters, etc.) or beds of ceramicmaterial, which may be appropriately impreg-nated to impart catalytic properties to them. Re-generative heating is provided by hot flue gasesfrom the combustion of liquid or gaseous fuels.During gasification, the heat carrier transfers itsheat to the raw material to be converted. Theprocess is suitable for liquid and gaseous feed-stocks. Any carbon black deposited during thereforming period can be burned off after the sys-tem is switched back to heating. The process isnot suitable for operation at elevated pressure.

Processes using circulating heat carriersare similar to cyclic gasification except for beingcontinuous processes circulating the heat carrierin a closed loop. This carrier is heated by hot fluegases and then fed to a reaction section in whichits heat is used for gasification. In addition to sen-sible heat, someheat carriers used in the past alsoproduced heat through reaction with gas compo-nents (producing CaCO3 from burned lime andCO2). Any carbon formed in the process can beburned off in the heating zone. The process isgenerally suitable for gaseous, liquid, and solidfuels. The circulating heat carrier may consistof granular catalytic or noncatalytic solids, or ofliquids (such as slag baths) and circulating gas.Although system design does not altogether ex-clude elevated operating pressure, the necessarydesign and operating conditions would be diffi-cult to meet.

Autothermal Processes. Autothermal pro-cesses produce the heat required for gasifica-tion through partial combustion of the fuel tobe converted in the reactor. Oxygen or air— ifnitrogen in the air is not a ballast in the prod-uct gas— is added for this purpose. The processmay operate with or without catalysts. If cata-lysts are used, the temperature can be kept lowerbecause a closer approach to equilibrium is en-sured. However, the catalytic process cannot beused to gasify heavy and residual oils becausemost of these oils contain ash particles that con-taminate the catalyst. Also, because of the ashcontent, catalysts for coal gasification are not ap-plied, so far. Either the catalyst would have to bevery cheap so that it could be disposed of alongwith the ash, or it would have to be recovered bya very complex procedure.

Processes applying baths ofmolten slag, iron,or salts as heat carriers and reaction promotershave been used and are still being developed.A characteristic of autothermal processes is thatthe oxygen added to generate heat is chemicallybound in the product gas, so the ratio of hydro-gen to carbon monoxide in the product is lowerthan in other processes. This type of process isapplied to solid, liquid, and gaseous fuels.

Processes Involving Externally HeatedWalls. Processes using externally heated wallsinclude tubular reforming, which is used widelytoday.Tokeep the tube stresses resulting fromel-evated pressure and temperature to a minimum,all tubular reformers employ a catalyst filling inthe tubes, which allows lower reaction tempera-ture.

Modernmaterials permit tubes to be designedfor wall temperatures of ca. 900 ◦C (in specialcases even 1000 ◦Cat atmospheric pressure) andpressures up to 4.0 MPa. All gaseous hydrocar-bons (with low unsaturated hydrocarbon con-tent) and distillate liquid fuels are suitable feed-stocks. When liquid hydrocarbons are used, de-mands on the catalysts are extremely high. Vir-tually ash-free fuels are used almost exclusivelyto heat the tubes because ash constituents (salts,metals) would damage tube walls.

Reactorswith externally heatedwalls are alsoused in gas production from solid feedstocks.This process technology offers the opportunityto employ nonfossil energy (heat from nuclearreactors, solar energy) for gasification.

Rich-Gas Processes. Because no externalheat supply is required for rich-gas processes,the equipment and process loops can be cho-sen freely. In most systems, a catalyst is appliedbecause methane formation is favored by lowertemperature. Apart from the unfavorable equi-librium position, the risk of carbon deposits re-sulting from the Boudouard reaction is higher atincreased temperature, and other byproducts areformed as well. All the processes operate at el-evated pressure because this promotes methaneformation.

Because catalysts are highly sensitive withrespect to attrition, reactors are normally of thefixed- rather than the fluidized- or entrained-bedtype. Raw materials must be virtually free of all

18 Gas Production

substances that may poison catalysts, particu-larly sulfur compounds.

1.5. Product Gas Treatment

Each gasification plant requires gas treatment fa-cilities to purify product gases and also, in manycases, to condition them. Whereas the gas pu-rification system eliminates the components thatwould affect downstream processing or utiliza-tion of these gases, the purpose of a condition-ing system is to adjust gas components to theappropriate ratio. The ratio of hydrogen to car-bon monoxide, especially, must be matched tothe requirements of downstream synthesis or re-duction units and other gas consumers.

The task of gas purification and conditioningis most complex downstream of plants gasify-ing solid fuels. Capital investment in these pu-rification units can be much higher than that re-quired for the gasification plant itself. Treatmentof gases produced through gasification of highmolecularmass liquid hydrocarbons (heavy oils,vacuum residues) is also rather expensive. Gen-erally, the smaller the molecules from which agas is produced, the lower is the capital invest-ment required in treatment facilities. The costsof gas purification are also strongly dependenton the amount and type of impurities in the feed-stock, particularly its content of sulfur and traceelements.

Whereas most catalytic gasification pro-cesses have some purification facilities (remov-ing sulfur, lead, chlorine) upstream of the gasifi-cation stage, the impurities in noncatalytic pro-cesses should be removed only from the productgas.

Impurities in the Gasification Gas. De-pending on the type and composition of feed-stock and the type of gasification process used,a gas purification plant must handle the follow-ing impurities:

– Dust (coal gasification)– Carbonization products, tar, gas liquor (coalgasification)

– Carbon (Heavy-oil gasification)– Oxygen (coal gasification, oil gasification)– Cyanic compounds (coal gasification, oilgasification)

– Nitrogen oxides (coal gasification, oil gasifi-cation)

– Carbon dioxide (all gasification processes)– Hydrogen sulfide, carbonyl sulfide, organicsulfur

– compounds (all noncatalytic gasification pro-cesses)

– Ammonia (all gasification processes)

Other gas components, even if they occuronly in trace amounts, require specific measuresto prevent long-range damage or malfunction ofthe plant. This is particularly true for the gasifi-cation of coal and heavy oils. Such componentsmay include volatile metal compounds and highmolecular mass organic compounds.

1.5.1. Purification Processes

Because of the large number of impurities, alongwith their widely varying content and combi-nations, a variety of gas purification processesare available to meet the requirements of down-stream processing (synthesis, in particular) andof industrial and municipal consumers (towngas). Ammonia synthesis, for instance, toleratesonly ca. 5 ppm (mL/m3) of oxygen compounds(e.g., O2, CO, CO2, and H2O), and the sulfurcontent must be < 1 ppm (mL/m3).

Methanol synthesis, on the other hand,requires a sulfur content of < 0.1 ppm(mL/ m3) and the stoichiometric ratio(cH2 − cCO2)/(cCO + cCO2) must be adjustedto ca. 2 (see Section 7.1.2). To achieve this ra-tio, the gas must contain a certain amount ofcarbon monoxide—a factor that influences theselection of gasification and purification pro-cesses and also raises the question of whetheran additional process stage is required to con-vert carbon monoxide into hydrogen and carbondioxide. Carbon dioxide need not be removedcompletely so the general plant design can besimplified considerably.

Reduction gases for direct reduction of ironore may contain sulfur on the order of 7 ppm(mL/m3), but the content of water and carbondioxide should not exceed 5 vol%.

If several gas purification plants are com-bined for optimum effect, perhaps including gasconditioning facilities as well, possible interac-tion should be checked carefully. Interactionscould occur because of solvents contained in the

Gas Production 19

gas, depending on their vapor pressure, differentoperating behavior, etc. If fuel gas is producedfor a combinedpower cycle (seeSection 7.3), thegas purification plant must meet the load changerequirements of such plants.

Gas purification processes may be classifiedas follows:

Adsorption. Adsorption describes the reten-tion of impurities by solids, either through phys-ical deposition or through chemical reaction(chemisorption) (→ Adsorption , see also Sec-tion 5.5). Whenever solids have to be reused(e.g., for economic reasons), batchwise opera-tion is necessary because the laden solids mustbe regenerated.

Sulfur compounds are eliminated by usingbeds of iron hydroxide and zinc oxide (espe-cially for desulfurizing gaseous and liquid feed-stocks for catalytic gasification). Activated car-bon is favored in special cases for cleaning gasescontaining high molecular mass organic sulfurcompounds.

Molecular sieves (→ Adsorption; → Zeo-lites) have been developed for selective removalof particular gases. They can be used not only forcarbon dioxide, sulfur compounds, and steam,but also for diatomic gases.

Gas purification systems based on the ad-sorption (and chemisorption) principle are pre-ferred when small quantities of impurities mustbe eliminated or small gas throughputs must bepurified.

Absorption. A general distinction is madebetween physical and chemical absorptionprocesses. Physical processes (→ Absorption,Chap. 1.4 , see also Section 5.4) are particu-larly suitable for high pressure, high impuritycontent, and large gas throughput. The processesavailable differ in solvents used, mode of plantoperation, operating temperature, method of en-ergy recovery, and potential for recovering sev-eral components more or less separately. Theselection of appropriate solvents is determinedto a large extent by their specific solubility co-efficients for the desired gases and by their se-lectivity.

Chemical absorption systems (→ Absorp-tion, Chap. 1.5) can be used at elevated or atmo-spheric pressure. They are particularly suitablewhenever the absolute quantity of impurities to

be removed is small. Conventional processesdiffer mainly in the degree of purity that canbe achieved and the amount of energy (normallysteam) required for regeneration.Numerous pro-cesses also use some combination of physicaland chemical absorption principles.

Special Wash Processes. Special applica-tions include the removal of solids entrainedwith gases in the form of dust or carbon black,dust-laden tar, and high molecular mass or un-saturated hydrocarbons that tend to form gums.

1.5.2. Conditioning

For nearly all applications, gasification gases notonly must be purified but also conditioned to ob-tain the optimum proportions of desirable gascomponents (Sections 5.1, 5.2, 5.3).

Carbon Monoxide Shift Conversion.Prominent among gas conditioning processesis catalytic carbon monoxide shift conversion,which is based on water gas equilibrium (Eq. 7).By this process, the hydrogen content of a gascan be increased as desired at the expense of itscarbon monoxide content. This is important notonly for ammonia synthesis but also for townand grid gas production. The process permitsthe gas to be adjusted to the standard densityand Wobbe index, and also enables nontoxic,low-carbon monoxide town gas to be produced,in conformity with current regulations.

Carbon monoxide shift conversion requiresthe addition of steam, which can be done inone of several ways. Steam at the appropriatepressure may be added directly to the gas tobe converted. If steam is not available in suf-ficient quantities and under economically jus-tifiable conditions, a so-called saturator coolerloop is often used. With this arrangement, gasentering the conversion section is humidified bycirculating hot water, and unreacted steam in thegas leaving this stage is recondensed and recy-cled to the saturator. A distinction is also madebetween high- and low-temperature conversion,the latter being provided downstream of the for-mer to reduce the residual carbonmonoxide con-tent of the gas to a minimum (e.g., for ammoniasynthesis). Shift conversion catalysts for sulfur-free gases are based on iron (high temperature

20 Gas Production

shift) or copper (low temperature shift). They aresupplied as oxides and reducedduring operation.For gases containing sulfur compounds the shiftcatalysts are based on cobalt and molybdenum.They are active only as sulfides.

Methanation. Another gas conditioningmethod is methanation, the conversion of car-bon monoxide, carbon dioxide, and hydrogen tomethane according to Equations (8 a) and (9). Amethanation stage is included in many ammoniasynthesis plants to ensure a high degree of purityof the synthesis gas. Similarly, a methanationstep is indispensable if a high-methane gas forSNG production is to be obtained.

Other Processes. The following processesmay also be regarded as gas conditioning.

If the calorific value of town gas, grid gas, orSNG is insufficient, the gasification gas may becarbureted by adding C3 and C4 hydrocarbonsor, less frequently, small quantities of high mo-lecular mass hydrocarbons. On the other hand,too high a calorific value can be reduced byadding nitrogen, air, or oxygen.

For certain syntheses, gases must be dried.Drying is also necessary in pipelines, to protectthe piping from corrosion where a risk of watercondensation in the presence of carbon dioxideexists. In addition, the formation of gas hydratesis avoided.

Unlike town gas produced in the past fromcoke oven batteries, the town gas obtained to-day through gasification no longer contains anychemical compounds that produce a manifestodor. Therefore, such gases are now odorizedartificially so that leaks occurring in the pipingor at the consumer’s facilities can be detectedeasily.

To recover particular components from a gasstream (e.g., H2 frompurge gas), pressure-swingadsorption (PSA) (→ Adsorption) or membraneprocesses (→ Membranes andMembrane Sepa-rationProcesses) can be used. ThePSAprincipleis based on the fact that the adsorption capac-ity of molecular sieves (→ Zeolites) differs fordifferent gases. Generally, heavier or polar gascomponents are adsorbed preferentially. Whena gas mixture is passed over a molecular sievebed, the component least amenable to adsorp-tion (e.g., H2) remains in the gas whereas othercomponents are adsorbed.

Gas separation by membranes is based onthe differences in permeation coefficients ofdifferent gas components. Permeation denotesthe sequence of absorption (solution) of thecomponent in the membrane material, diffusionthrough themembrane, and desorption. The per-meation rate of a species increases with its sol-ubility and its diffusion coefficient in the mem-brane material. Membranes may consist of ce-ramic or metallic materials, organic liquids, orpolymers.

Both processes, PSA and membrane separa-tion, can be used to advantage for gases at ele-vated pressure.

To remove traces of higher hydrocarbons(e.g., compressor lubricants) from gases, beds ofactivated carbon are used. For recovery of low-boiling compounds, as well as purification fromtraces of higher boiling compounds, cryogenicprocesses (→ CryogenicTechnology) canbe ap-plied. These employ primarily liquidmethane ornitrogen to absorb higher boiling gases such ascarbon monoxide.

1.5.3. Byproducts

Some gasification processes yield byproductsalong with the desired gas (see also Chap. 6).A comparison of different gasification processesshould, therefore, always consider such byprod-ucts and assess their importance. Environmentalprotection, too, plays an important role becauseit imposes stringent conditions on pollutant dis-charge.

Many gasification plant feedstocks containsulfur, and sulfur often constitutes the main by-product. Whether this is best recovered in theform of elemental sulfur or sulfuric acid is worthconsidering.

Other coal gasification processes have cokeas a byproduct.

Gasification involving coal carbonizationprocesses can produce a large number of byprod-ucts such as tar, oil, gasoline, phenols, cresols,ammonia, and organic acids. Certain liquid by-products may even contain solids whose elimi-nation would result in additional costs. Becauseof this, processes have been developed to recyclecertain liquid fractions to the gasification plantif their market value is low.

Gas Production 21

Heavy-oil gasification plants yield carbon asa byproduct. Attempts to use it for special in-dustrial applications have been successful, butdemand for the specific qualities of carbon islimited. Therefore, processes are being devel-oped to recycle most of this carbon to the gasifi-cation stage. These and processes where carbonis burnt in conventional power stations at littleextra cost are of major interest.

Gas liquor is a product of almost all gasifi-cation processes. Depending on its compositionand on plant requirements, this gas liquor caneither be treated for use as boiler feed water orcooling water, or be discharged to receiving wa-ter via chemical or biological effluent treatmentplants.

2. Steam Reforming of Natural Gasand Other Hydrocarbons

In the steam reforming process hydrocarbons arecatalytically converted by reaction with steaminto hydrogen and carbon oxides. This is themost common method for producing hydrogenor hydrogen: carbon oxidemixtures in themanu-facture of important basic chemicals (e.g., am-monia and methanol), oil refining, and in manyother industrial applications (e.g., iron ore re-duction, hydrogenation of fats, production ofoxo alcohols).

Nowadyas natural gas predominates by far asa feedstock over other hydrocarbons (e.g., naph-tha, LPG, refinery gases) or coke oven gaswhichare used only under special circumstances.

The catalytic steam reforming process intubular furnaces was invented in 1926 – 1928by BASF [7]. This process was applied in theUnited States for the first time in the early 1930sat two commercial plants: (1) to produce hydro-gen fromnatural gas for hydrogenation purposesand (2) to synthesize ammonia. The process wasconducted under lowpressure (0.4 – 1 MPa) andat temperatures close to 800 ◦C until the early1950s; pressures up to 4 MPa and temperaturesup to 950 ◦C are used today.

A special type of steam reforming which hasbeen in use even earlier than tubular steam re-forming is autothermal reforming, also calledcatalytic partial oxidation. This process differsfrom catalytic steam reforming in that the re-quired reaction heat is not supplied from outside

(furnace) but by internal partial combustion ofthe feedstock with oxygen or air, admixed to theprocess feed. As autothermal reforming is car-ried out in refractory-lined vessels (cold shell),higher pressures can be applied than in tubu-lar steam reforming. Outlet temperatures up to1000 ◦C are usual. Due to the additional expen-diture for oxygen supply and safe process con-trol this technology was of minor general im-portance in the past. Of common use, however,is the process modification with air as oxidantin ammonia synthesis. More recently, the oxy-gen based gasification has been introduced intomethanol synthesis. “Combined reforming”, acombination of tubular and autothermal reform-ing, is recognized as the most efficient syngastechnology for large scale methanol plants at themoment.

A comprehensive overview of recent devel-opments in steam reforming technology is givenin [8].

Whereas in former years steam reformershave been built nearly exclusively at locationswhere there was requirement for hydrogen orsynthesis gas and products derived therefrom,a big change has occurred since the 1970s, inthat particularly world-scale production plantsfor ammonia and methanol have been built atlocations, where natural gas is available in abun-dance (e.g., Trinidad, Borneo, Cape Horn).

Total hydrogen production by catalytic re-forming currently (ca. 2000) amounts to over400 × 109 m3/a worldwide. Its uses can be bro-ken down approximately as follows: 60% forammonia production, 10% for methanol pro-duction, 25% for petrochemical processes (hy-drotreating, hydrodesulfurization, hydrocrack-ing), and 5% for other industrial purposes. Thecost of commercial production is about $ 0.10per cubicmeter of hydrogen. The catalytic steamreforming process is especially low in emis-sions. Only the off-gas from the steam reformercontains, apart from carbon dioxide, a certainamount ofNOx and, dependingon the feedstock,sulfur dioxide.

2.1. Feedstocks

Natural Gas. Natural gases (→ Natural Gas) are used as gasification feedstocks, either di-rectly or after appropriate pretreatment. Table 5

22 Gas Production

gives examples of pretreated natural gases andother hydrocarbon feedstocks. High-boiling hy-drocarbons are largely eliminated before trans-portation to prevent condensation in the distri-bution system.

Particular attention must be paid to sulfurcompounds in natural gas. A gas suitable for cat-alytic reforming has to be virtually free of thesecompounds as sulfur is a strong poison to theemployed catalysts. High percentages of nitro-gen are undesirable as well, except in ammoniasynthesis gas. For hydrogen – carbon monoxidegases, nitrogen must be removed in a complexoperation. A certain percentage of carbon diox-ide may be tolerated and even be beneficial inmethanol production.

The condensation fractions recovered fromcasinghead gases (see Table 5) extend into thenaphtha range and contain higher hydrocarbonsthat may create problems in catalytic reforming.These fractions are handled similar to naphtha.

Reevaporated liquefied natural gas (LNG) isfree of all accompanying components such ascarbon dioxide, steam, sulfur compounds, andC5+ hydrocarbons, which must be removed forprocess reasons prior to liquefaction. It does notcontain any gas with a lower boiling point thanmethane, such as nitrogen or noble gases (he-lium). However, when obtained from the publicgrid, LNGmay contain sulfur compounds (usedas odorants; e.g., tetrahydrothiophene) and ni-trogen (from calorific value – density condition-ing).

Other Gases (see Table 6). Gases of differ-ent qualities are derived from refineries, cokingplants, the chemical industry, steelworks, etc.They are suitable not only as fuel gases but, in-creasingly, as gasification feedstocks or as rawmaterials for the chemical industry. As a typi-cal example, hydrogen used in a refinery can beproduced by the conversion of refinery gases orbutane.

purge gases) from ammonia and methanolsynthesis or olefin production are also used forgas production.Unsaturated hydrocarbons in theresidual gases are usually saturated by hydro-genation before they are catalytically gasifiedwith steam. The top gas from some processesfor direct reduction of iron ore is upgraded byadding fresh natural gas and used for gas produc-tion. Coke oven gas, which was formerly used

mainly as town gas, is utilized today for ammo-nia production or the reduction of iron ore.

Naphtha. At the beginning of the 1960s,light naphtha was available in large quantitiesfor the production of town gas, ammonia, andmethanol. Since the 1973 oil crisis, it has beenused almost exclusively as a rawmaterial for theproduction of chemicals (especially ethylene).Exceptions are town gas and synthesis gas pro-duction in isolated areas where no natural gas isavailable.

A distinctionmust bemade between straight-run virgin naphtha as obtained directly fromcrude oil distillation and naphtha cuts from theproduction facilities of a refinery, for example,from a reforming unit (high-aromatic naphtha).

Naphthas differ according to geographicalorigin of the crude. The main criteria for assess-ing the suitability of a naphtha as a raw mate-rial for gas production are its boiling range (lowboilers can be converted more easily); carbon:hydrogen ratio; aromatic, olefin, and naphthenecontents; sulfur content; and content ofmetals orother substances (e.g., lead, chlorine, which willpoison the catalyst). Table 7 shows the principalcharacteristics of various naphthas.

Generally, low-boiling naphthas rich in alka-nes are preferred for catalytic gasification be-cause they are easier to convert and the cata-lyst has a longer service life. In individual cases,naphtha with a boiling end point up to 220 ◦C,mixtures of naphtha and kerosene, or keroseneitself may be considered as raw materials.

Ammonia and Methanol. Ammonia andmethanol are occasionally used as feedstocksfor hydrogen production (in small units and forpeak shaving gas).

2.2. Natural Gas and Other GaseousHydrocarbons

In the 1920sBASF developed a catalytic processfor endothermic conversion of gaseous hydro-carbons with steam in externally heated tubes,using nickel as a catalyst. Exchanges of experi-ence with Standard Oil (New Jersey) and ICI inthe 1930s led—particularly in the United Statesand England— to the first commercial applica-tion of this process in the production of hydro-

Gas Production 23

Table 5. Properties and composition of (treated) natural gases used as feedstocks for catalytic gas production

Table 6. Residual gases (Properties and composition of various residual gases and of a butane fraction

24 Gas Production

Table 7. Properties and composition of various naphtha feedstocks for gasification

Light naphtha Boiling range of full-range naphtha Heavy Treated naphtha a Condensate

naphtha HydrotreatedRaffinate fromcasingheadgas

Density (15 ◦C),kg/L

0.66 0.68 0.680.72 b 0.70 – 0.73 0.71 – 0.73 0.71 0.76 0.73 0.74

Initial boiling point,◦C

31 41 39 35 – 46 60 39 – 45 43 – 65 44 95 45 99 31

Intermediate boilingpoint, ◦C c

b 70 69 96 – 105 b 98 127 91 134 110 138

Final boiling point,◦C

89 108 –125

123150 – 165 66 180 – 185 165 – 187 200 190 197 206 344

Molar mass 98 – 102C/H mass ratio 5.17 5.24 5.375.70 – 5.77 b 5.4 5.8 5.54 6.0 6.0 5.61Alkanes, vol% b ca.

75.5

b b 56.3 95 d b 81.3 48 72 b 71

Olefins, vol% 0 0 0.10 – 0.2 0.1 0.3 – 0.6 1 1.5 b b

Naphthenes, vol% b ca. 2220 b 37.4 b 7.8 41 8 40 b

Aromatics, vol% 1.5 2.5 3.13.7 – 6.2 6.3 3.2 – 4.8 11 – 15 9.4 11 20 11 b

Chlorides, mg/kg 1.2 3.0 –4.0

1.21 – 3 4 16 – 30 b 0.1 b b b

Lead, mg/kg 0.02 0.02 0.90.01 0.1 0 b <0.001

b b b

Sulfur, mg/kg 3.0 15 –20

14 11 – 54 120 8 – 10 560 – 635 100 9 b 68

Higher heatingvalue, kJ/kg

48 235 47960

47855

47 270 – 47400

b 47 835 – 48275

46 685 – 46895

b 46 520 b b

Lower heating value,kJ/kg

b 44465

b 44 090 – 44170

b 44 630 b b 43 335 b b

a From reformate.b Not measured.c Boiling point after 50% reduction in volume.d Paraffins and naphthenes.

gen for hydrogenation and ammonia synthesis.These plants operated at atmospheric pressure.New reformer tube materials, especially the de-velopment of centrifugally cast tubes, permittedhigher pressure in the gas production section,which in turn improved plant efficiency consid-erably.

2.2.1. Principles

Catalytic conversion of hydrocarbons withsteam to produce gases with high hydrogen andcarbon monoxide content, according to the re-actions (8) and (8 a)—and always accompa-nied by the shift conversion (Eq. 7) (see Sec-tion 1.3)— is known as steam reforming. Thisconversion is highly endothermic. It occurs in anumber of parallel, externally heated tubes (re-former tubes) filled with a nickel catalyst.

After being preheated in heat exchangers orfired heaters to ca. 500 ◦C, the steam – hydro-carbon mixture is heated further over nickel

catalyst in the tubes to the desired end tem-perature of normally 800 – 900 ◦C. Higher hy-drocarbons (Eq. 8) are completely converted tocarbon monoxide and hydrogen, methane onlypartially—corresponding to the reaction con-ditions. The composition of the reformed gasat the reformer outlet reflects the equilibria ofEquations (8 a) and (7). For specified condi-tions (pressure, reformedgas temperature, steam– hydrocarbon ratio), the composition of the re-formed gas can, therefore, be calculated fromcomponent balances and equilibriumconversionrates. The steam: hydrocarbon ratio is an essen-tial parameter in the steam reforming reactionandmust not be lowered arbitrarily, because lowratios lead to carbon deposits on or in the cat-alyst. Figure 2 shows the influence of pressureand temperature on the composition of the re-formed gas for a defined steam:methane ratio ina methane reforming unit.

Gas Production 25

Figure 2. Equilibrium composition for methane reformingat 500 – 1000 ◦C and pressures (in the direction of thearrows) of 0.1, 0.5, 1, 1.5, 2, 2.5, 3, 3.5, 4 MPa. Molarsteam: methane ratio 3.333 corresponding to 5 kg H2O perkilogram of carbon

Kinetics. The kinetics of catalytic steam re-forming of methane are summarized in [9, 10]and [11]. Many authors observe that the shiftconversion reaction described by Equation (7)proceeds very fast so that the gas compositionmay be assumed to be in equilibrium at everypoint in the reformer tube—a proposition thatis not true for methane conversion which ap-proaches to equilibrium only towards the outletof the tubes. Sulfur in the gas has no decisive in-fluence on the shift conversion, but can decreasethe reaction rate of the hydrocarbon conversiondramatically, especially at the inlet side of thetube.

Methane reforming can be described as afirst-order reaction, irrespective of operatingpressure. At low temperature, the molecular dif-fusion rate is much higher than the reaction ve-locity so that, theoretically, the catalyst activitycan be fully utilized. If the temperature is high,the conversion rate is determined by pore diffu-sion, and catalyst efficiency is reduced. Hence,large-surface catalysts offer an advantage withregard to catalytic activity. For further details of

reaction mechanism and kinetics, see [12] and[13].

Carbon Deposition. Carbon deposition is afrequently experienced phenomenon with steamreforming. Though generally various reactionsmay cause solid carbon formation from car-bon compounds containing gaseous phases [e.g.,Boudouard reaction (2 CO � C + CO2), hy-drocarbon decomposition (CnHm � n C + m/2H2), the heterogeneouswater gas reaction (CO +H2 � C + H2O)], hydrocarbon decompositionis practically the exclusive cause for carbon for-mation in the course of reactions in the steam re-forming process. Therefore carbon deposition isobserved mainly in the first half of the reformertubeswhere hydrocarbon concentrations are stillhigh and the approach to equilibrium low.

Investigations have shown [14] that thethermodynamic limits of carbon formation(graphite) do not apply below 700 ◦C becausethere the relevant reactions are dominated by ki-netics. So the carbon-free area may increase, de-pending on catalyst properties. Figures for a ura-nium – nickel – α-alumina catalyst are given in[15] and for a number of other catalysts in [13].

Figures 3 and 4 show experimental “equi-librium constants” for the Boudouard reactionand methane decomposition measured by vari-ous authors, aswell as calculatedfigures for ther-modynamic equilibrium according to [16]. Fig-ure 4 also depicts the “working line” formethanesteam reforming according to [17], by applyingtwo catalysts with different activities. Points onthe curves indicate the location of the measur-ing point as a proportion of tube length (viewedfrom the gas inlet).

In steam reforming of natural gas, the tem-perature at the catalyst inlet is such that car-bon might theoretically be deposited. This doesnot occur, however, below ca. 650 ◦C, but onlyat higher temperature when an inactive catalystfails to decompose sufficient methane. Gener-ally, the formation of carbon deposits is not in-evitable even at operating temperatures where itmight be expected.

Figures 3 and 4 also illustrate the more fa-vorable initial situation and possible further re-forming of a naphtha-based rich gas with addedcarbon dioxide (to obtain high-COgases) if cata-lyst activity remains unimpaired (line (c) in Fig.

26 Gas Production

3 or line (a) in Fig. 4) or if no further methaneconversion occurs (line (b) in Figs. 3 and 4).

Figure 3. Equilibrium constant of the Boudouard reaction(2 CO � C + CO2)a) Condition of rich gas from naphtha at inlet to tubu-lar reformer; b) Equilibrium reached for reaction (7)only; c) Equilibria reached for reactions (7) and (8 a);d) According to [14]; e) According to [13]; f) For nickel –uranium catalyst according to [15]; g) For sulfided catalystaccording to [13]

Figure 4. Methane decomposition (CH4→ C + 2 H2)a) Equilibria reached for reactions (7) and (8 a);b) Equilibrium reached for reaction (7) only; c) Equilibriumconstant for carbon in graphite form according to [16];d) and e) Working lines for methane reforming in a reac-tor tube; d) High-activity catalyst; e) Low-activity catalystwith red hot tube section from ca. 650 ◦C to equilibrium;figures at the data points are proportions of tube length asviewed from inlet; f) Equilibrium constant determined byexperiment according to [13]; g) Connecting line betweenequilibrium constants determined by experiment at 410 and525 ◦C according to [14]

2.2.2. Catalysts, Catalyst Poisons,Desulfurization

Steam reforming is generally carried out in thepresence of nickel catalysts. These catalysts areusually in the form of thick-walled Raschigrings, with 16-mm diameter and height and a 6 –8-mmhole in themiddle. However, especially inmodern steam reformerswith high heat flux, cat-alysts with high geometric surface (HGS-type)are preferred. As already mentioned in Section2.2.1, the access to the inner surface of catalystscontinuously decreases with increasing temper-ature due to diffusion limitation. This is com-pensated by introducing catalysts with a highersurface to volume ratio. Typical shapes of thistype of catalysts are spokedwheels, gearwheels,or rings with several holes. These catalysts areadditionally advantageous with regard to theirlow pressure drop.

A general distinction is made between im-pregnated and precipitated catalysts. The latterare the more traditional types which are pro-duced by coprecipitation of all constituents.Pre-cipitated catalysts have a Ni-surface which is byone order of magnitude higher than that of im-pregnated catalysts and for this reason have anexcellent activity already at moderate temper-atures. But they are more prone to sintering athigher temperatures anddonot have themechan-ical strength of supported catalysts.

Impregnated catalysts are manufactured byseparately impregnating a support with nickelas the active constituent. The main advantage ofthis type of catalyst is that the support can be cal-cined at high temperatures before impregnation,resulting in predetermined excellent mechani-cal properties which are important especiallyat high temperature application. As the condi-tions of steam reforming have generally becomemore severe, impregnated catalysts predominateby far at present. Frequently used support ma-terials are alumina (the α-modification is pre-ferred today), magnesium oxide, or spinel typesubstances (e.g., magnesium aluminum oxide).In some cases, catalysts are stabilized by addi-tion of a hydraulic binder (calcium aluminumoxide). Crushing strength is decreased by thehydration of magnesium oxide at low tempera-ture (< 500 ◦C) and by changes from the γ to

Gas Production 27

the α-modifications of alumina above 500 ◦C.A silicon dioxide content of < 0.3 wt % in thecatalyst is essential, because silicon dioxide be-comes volatile in the presence of steam and mayfoul the downstream heat exchangers throughwhich the reformed gas flows.

Nickel is at first present in the form of nickeloxide; after the system has been heated, it is re-duced in the reformer tube by hydrogen formedwhen natural gas is added to steam. Some cat-alysts require a special method of reduction forwhich a definite steam: hydrogen ratio must bemaintained (see Fig. 5), according to [12] and[13].

Figure 5. Reduction conditions for nickel oxide

Catalyst Poisons andDesulfurization. Cat-alytic activity is affected seriously even by verylow concentrations of catalyst poisons in gasesto be reformed. Such catalyst poisons are sulfur,arsenic, copper, vanadium, lead, and chlorine orhalogens in general. Sulfur, found in practicallyall hydrocarbon feedstocks, in particular lowerscatalyst activity [18].

Earlier desulfurization systems used impreg-nated activated carbon as adsorbent and oper-ated at ambient temperature. However, becausetheir efficiency differs according to the particu-lar sulfur compounds in the gas and because therequirement with regard to the residual sulfurcontent has increased, zinc oxide desulfurizationsystems operating at 350 – 400 ◦C are generallypreferred today. Iron-based adsorbents are usedin only a fewcaseswhere special conditions (i.e.,low operating temperature, possibility of regen-eration, sulfur composition) prevail.

Zinc oxide reactors are very reliable in ab-sorbing hydrogen sulfide and, with limitations,sulfur compounds such as carbonyl sulfide andmercaptans. To achieve optimum desulfuriza-

tion conditions, certain space velocities and ef-fective flow ratesmust be ensured. Sulfur pickupranges from 15 to > 30%, that is almost com-plete conversion of zinc oxide to zinc sulfide ispossible.

Organic sulfur compounds such as mercap-tans and thiophenes require hydrogenation overcobalt – molybdenum or nickel – molybdenumcatalysts. These catalysts are often arranged ina separate vessel. Hydrogen or hydrogen con-taining gas are added to the process feedstockat temperatures of ca. 350 – 380 ◦C. The or-ganic sulfur compounds are converted to hydro-gen sulfide and the corresponding saturated hy-drocarbons. Hydrogen sulfide produced in thehydrogenation stage is then absorbed by zincoxide. The hydrogenation reactor can be usedsimultaneously to hydrogenate unsaturated hy-drocarbons in the raw gas. This reaction isstrongly exothermic. Because the temperaturerange for the hydrogenation stage is limited to250 – 400 ◦C, unsaturated hydrocarbon contentis limited as well. Ammonia, carbon monox-ide, or carbon dioxide impurities in hydrogenaffect desulfurization and can lead to undesir-able side reactions such as methanation [12]. Aneconomic solution particularly for natural gaseswith low sulfur content is the combination ofhydrogenation and zinc oxide absorption. Thismeans that both catalysts can be combined in onebed in the same reactor. Residual sulfur contentis normally < 0.2 mg/m3 [19].

2.2.3. Tubular Reformers

Tubular reformers are mostly top- and wall-firedbox-type units, apart from older types or specialdesigns (Fig. 6). Whereas in wall-fired reform-ers, rows of reformer tubes are heated mainly bythe radiant sidewall, top-orbottom-fired reform-ers provide most of the heat through radiation ofthe burner flame and the hot flue gases. Top-firedreformers may, therefore, have several parallelrows of tubes, whereas wall-fired reformers canhave only one row. Wall-fired reformers with alarge number of tubes consist of several units ad-jacent to one another; waste heat from the fluegases is recovered in a common system. Mod-ern reformers in large plants that produce am-monia and methanol synthesis gas from naturalgas contain hundreds of tubes arranged in rows.

28 Gas Production

Figure 6. Reformer typesA) Top-fired reformer; a) Inlet header(s); b) Pigtail; c) Burner; d) Catalyst-filled reformer tube; e) Flue-gas collector;f) Refractory lined outlet manifold; g) Outlet pigtail; h) Outlet manifoldB) Wall-fired reformers; a) Inlet header; b) Flue-gas duct; c) Wall burners; d) Outlet manifold; e) Terraced wall; f) Burner

Firebox. The reformer box consists of sev-eral layers of refractory insulation and insulatingmats with an exterior metal skin. The hot facerefractory must be resistant to temperatures >1200 ◦C; the temperature is reduced to ca. 75 ◦Con the outside of the metal skin.

Burner geometry, flame length and diameter,tube-to-tube and row-to-row spacing, fired tubelength, and distance from the flame to the re-former wall determine the homogeneity of heattransfer to the tubes. Modern computers arecapable of linking up this reformer geometrywith heat-transfer processes inside the reformertubes (radial heat conductivity) to derive com-plex equation systems from which heat trans-fer at all points of the firebox can be calculated(see Fig. 7). These equations yield local tem-perature profiles for flame – flue gas (g – f), re-former wall (d), tube wall (b), (c), reformed gas(a) or catalyst, and local heat flux (e). These fig-ures form the design basis for the reformer. Ifthe tube wall temperature and local heat flux areknown, and heating and cooling rates are esti-mated, the wall thickness of the reformer tubescan be calculated.

One important factor is spacing between thereformer tubes, referred to as tube pitch. In cal-culating tube pitch, irregularities in heat fluxaround the tube must be considered [20].

Burners and Firing Systems. Top-fired re-formers use turbulent open-jet burners withforced air supply. Appropriate selection of pa-

rameters such as jet diameter, fuel and air supplyrates, and burner geometry produces a relativelyshort, stable flame ca. 2 – 2.5 m in length; thereformer tubes, therefore, receive their heat toa considerable degree from flue gas radiation.Accurate calculation of heat transfer requires aknowledge of flue gas backflow and flame emis-sion behavior.

Figure 7. Approximate temperature profiles across a re-formera) Reformed gas (average over tube cross section); b) Tubewall inside; c) Tube wall outside; d) Heater wall; e) Localheat flux; f) Flue gas; g) Flame

Flue gases are discharged via ducts inside thefirebox and cooled to ca. 200 ◦Cinheat exchang-ers, which are used to preheat the gaseous feed-stock, generate steam, etc.

Wall-fired reformers employ other burner de-signs or flame shapes, such as resist burner cups

Gas Production 29

into which the fuel is burned or flames burningparallel to the side wall.

Reformer Tubes. Normally, the tubes haveinside diameters of 75 – 125 mm, wall thick-nesses of 10 – 20 mm, and fired lengths of 9 –15 m, depending on reformer type (top- or wall-fired). The fixed tube support is usually locatedoutside the firebox at the top or bottom; whenheated, the tube expands longitudinally from thissupport. Counterweights or spring hangers oftenrelieve the tubes of ca. 50%of their deadweight.

Tube wall thickness is calculated on the basisof internal pressure and of 100 000-h creep-to-rupture strength; see [20]. More sophisticatedcalculations allowing for mechanical and ther-mal stresses, and in some cases even establishinga link between them, are being used increasingly[21, 22].

Table 8 summarizes tube materials currentlypreferred, along with their rupture strengths.Tubes are welded to the required overall lengthfrom centrifugally cast tube sections > 3 m inlength. The calculated “sound wall thickness”must be increased by an allowance for fabrica-tion tolerances.

The pressure drop inside a reformer tube de-pends on the gas flow rate and geometry of thecatalyst; it is normally between 0.15 and 0.5MPa.

Pressurized Convection Reformers. Ama-jor factor determining the size of reformer tubesis the pressure difference between the reformedgas and the nearly pressureless flue gas. Obvi-ously, the wall thickness of the reformer tubescan be reduced if the pressure in the firebox isincreased. This concept is now mature enoughfor application on an industrial scale [23]. Re-former tubes with a much larger diameter thanusual can be employed; the catalyst is arrangedin an annular space, and reformed gas is recycledinternally by straight or helically coiled tubes (toincrease internal heat exchange). Such reform-ers can be designed for hydrogen output up to5000 m3 (STP)/h; beyond this, the wall thick-ness of the pressure vessel makes the concepteconomically unattractive.

Higher output can be achieved by connectingseveral units in parallel. The pressurized com-bustion reformer requires little space and can,therefore, be prefabricated in modules. How-

ever, combustion air and off-gas from the PSAunit required for the firing system must be com-pressed to high pressure, whereas the flue gaspressure energy can be recovered only partly inexpanders.

A promising route for the future might be toincrease process pressures to a level at whichcompression of synthesis gas is no longer re-quired [24].

Inlet Header and Outlet Manifold Sys-tems. The preheated hydrocarbon: steam mix-ture (400 – 530 ◦C) is fed to the reformer tuberows via inlet headers with branches to each in-dividual reformer tube. These inlet pigtails aredesigned to take up the expansion of the inletheader and of the tubes (if the fixed support is un-derneath the reformer), as well as their own ex-pansion. Header and pigtails are made of carbonsteel (or alloyed carbon steel). Flexible metalhoses are used instead of inlet pigtails only inspecial cases.

Hot reformed gases are discharged from be-low the catalyst supporting grid in the tubes tooutlet manifolds. The outlet manifolds are madeof centrifugally cast (seeTable 8) or drawnmate-rial (Incoloy 800 H). Brick-lined manifolds areused as well; in this case, reformer tubes aremounted directly on the manifolds or connectedby small-diameter straight tubes, or “flexitubes”[25]. Another design has manifolds installed inthe bottom of the firebox; product gas is ex-tracted through the reformer top by means ofone or more risers.

On leaving the outlet manifold, the hot re-formed gases are cooled in heat exchangersto generate steam or to preheat boiler feedwa-ter, raw gas, etc. They are then fed to carbonmonoxide shift conversion units or carbon diox-ide scrubbers for further conditioning.

2.2.4. Production of Fuel Gas and SynthesisGas

Operation of the reformer under defined reac-tion conditions, in conjunction with additionalprocess steps, produces the desired gas compo-sition for fuel gas or synthesis gas applications.The variety of possibilities is indicated in Ta-ble 9, which lists typical reaction parameters fortubular reforming.

30 Gas Production

Table 8. Creep-to-rupture strength (average for minimum values) σRmin 100 000 for centrifugal castings, N/mm2

Material designation Trade names Temperature, ◦C

(DIN 17 006)800 850 900 950 1000 1050 1100

G-X 40 CrNiSi 2520 HK 40 25.0 18.0 12.0 8.0 5.0 3.0G-X 35 CrNiSiNb 2424 IN 519 35.0 25.0 17.0 11.0 6.6G-X 50 CrNiSiNb 3030 Marker RE30N 40.0 30.0 21.0 13.5 8.2 4.7 2.4G-X 45 NiCrNb 3525 Pyrotherm G 35/25 NbTZ, KHR 35C 38.0 29.0 21.0 13.5 5.0 4.7 2.3G-X 40 NiCrNbTi 3425 Manaurite XM 47.0 35.7 25.7 17.5 11.2 6.8 3.9G-NiCr 48 W NA 22 H, Paralloy H48T 34.0 25.0 17.5 12.0 7.8 4.8 2.8G-NiCr 50 Nb * IN 657 28.5 20.5 13.0 7.5 3.8G-X 10 NiCrNb 3220 ** Pyrotherm G20 32Nb 27.0 18.0 11.0 6.4

* This material (which has good resistance to oil ash corrosion) is normally used as the hot face layer of centrifugally cast compound tubes;the cold face is made of high-strength material.** Material for hot but unheated outlet manifolds.

Table 9. Production of various types of gas by tubular reforming of gaseous hydrocarbons

Product gas forproducing

Typical process stages * Typical H2O/Cratio, mol/mol

Typical pressureat reformer tubeoutlet, MPa

Typicalreformed gastemperature,(primaryreformer), ◦C

CO desulfurization, tubular reforming with recycled CO2, CO2scrubbing, CO separation (e.g., low-temperature separation, CoSorb,PSA, or membranes)

2 – 3 1 – 3 850 – 900

H2 desulfurization, tubular reforming, HT and LT conversion, CO2scrubbing, methanation (standard arrangement)

4 – 5 1.5 – 3 800 – 900

desulfurization, tubular reforming, HT conversion, PSA, ormembranes

3 2 – 2.5 800 – 900

Oxosynthesisgas

desulfurization, tubular reforming, CO2 scrubbing, partial separationof hydrogen by PSA, addition of imported CO2 upstream ofreforming if necessary

2.5 1 – 2 850 – 950

Methanolsynthesis gas

desulfurization, tubular reforming, or additional secondary reformer 2.5 – 2.8 2 – 3 850 – 900 700 –800

NH3 synthesisgas

desulfurization, tubular reforming (primary reformer), secondaryreforming, HT and LT conversion, CO2 scrubbing, methanation

3.5 3 – 4 780 – 830

Town gas desulfurization, tubular reforming, HT conversion, CO2 scrubbing 3.0 1 – 2.5 650 – 750Reduction gas desulfurization, tubular reforming, or top gas recycling upstream of

tubular reforming1.25 – 1.5 0.65 0.2 – 0.3 0.15 –

0.2850 – 1000 900– 1000

* PSA = pressure swing adsorption; HT = high temperature; LT = low temperature.

2.2.5. Special Reforming Processes

Reforming with Additional Carbon Diox-ide. Reforming with additional carbon dioxideis used to produce gases with an increased car-bon monoxide content. By recycling carbondioxide from the reformer product gas, partiallyor totally, the hydrogen: carbon monoxide ratiocan be adjusted in a range defined by the com-position of the natural gas. By total carbon diox-ide recycle the complete carbon content of theprocess feed is converted to carbon monoxide.In this case the shift conversion reaction (Eq.7) is eliminated and the hydrocarbon conver-sion follows exclusively Equation (8). Productgases with particularly high content of carbonmonoxide are obtainedby adding carbondioxidefrom an external source in addition to the carbon

dioxide recycle to the reformer feed. This com-bined steam – carbon dioxide reforming entailsa greater risk of carbon deposits, which must becountered by selecting appropriate reaction pa-rameters and catalysts (see Section 2.3.3).

Whereas reduction gases having high hydro-gen and carbon monoxide content are normallyproduced by methane reforming with only aslight amount of steam being added (< 1.5 molof steam per mole of methane), the Midrex pro-cess in particular (direct reduction of iron oxide,→ Iron, Chap. 2.6.1) relies heavily on combinedsteam: carbon dixoide reforming.

Midrex Reforming. In the Midrex process,iron ore reduction and the production of reduc-tion gas are integrated. Natural gas ismixedwithoff-gas from iron reduction (cleaned and humid-

Gas Production 31

ified gas from the direct reduction furnace). Thelatter contains unreacted hydrogen and carbonmonoxide in addition to steam and carbon diox-ide produced by oxidation of carbon monoxidein the furnace. Typical for the process is a steam:carbon ratio (regarding the molar carbon con-tent of only methane and higher hydrocarbons)of 0.65 and a carbon dioxide: carbon molar ratio(see above) of ca. 0.8. Pressure at the reformeroutlet is ca. 0.2 MPa. Special catalysts (Midrexand others) ensure that gases can be reformedwithout carbon formation.

To this end, either catalysts of different ac-tivities must be arranged successively in the re-former tubes, or the activity of the catalyst mustbe adjusted by means of a controlled sulfur con-tent so that it will be lower at the tube inlet than atthe outlet. Reformed gas is conditioned prefer-ably to a hydrogen: carbonmonoxidemolar ratioof ca. 1.6. Processes for the production of directreduction gases are described in [26].

Steam Reforming with Addition of Air. Inthe past, town gas was produced with additionof process air to adjust gas density; examples arethe Koppers natural gas reforming process andtheDr. Otto butane reforming process. However,with the conversion of fuel gas systems to naturalgas, this technology has become less important.It is still used to make synthesis gas for ammo-nia production from low-boiling hydrocarbonsup to the naphtha boiling range in small com-pact plants containing a combination of primaryand secondary reformers (see below).

Lately, natural gas reforming with additionof air has been used temporarily to adjust gasdensity during the changeover period from towngas to natural gas until the conversion of gas ap-pliances to natural gas is completed.

The standard technology for producing am-monia synthesis gas from low-boiling hydrocar-bons is the two-stage reforming process. Thefirst stage is catalytic steam reforming of hy-drocarbons in a tubular reformer (primary re-former). Product gas from the primary reformer,which has a temperature of ca. 800 ◦C, is thenreformed with air in an autothermal secondaryreformer at an outlet temperature of ca. 1000 ◦C(Section 2.5).

MethaneReformingwithNuclear or SolarEnergy. In high-temperature nuclear reactors,

recycled helium is heated from 350 to 950 ◦C.For catalytic reforming, hot helium is circulatedin indirectly heated heat exchangers in counter-current to methane and steam flowing throughthe reformer tube, releasing its sensible heat andbeing cooled from 950 to 600 ◦C. The preferredreformer tube design has an inner helical tubethroughwhich reformedgas is discharged toheatthe catalyst-filled tube; see [27, 28]. Such reac-tors, termed EVA reactors, have been tested inpilot plants since 1971 and are considered suit-able for commercial reformed gas production.Energy is supplied to the heating helium in thecore of a high-temperature nuclear reactor. InCRS (i.e.,CentralReceiverSystem) solar plants,this occurs in the receiver.

Reformed gas can be used to produce basicchemicals (H2, NH3, CH3OH) or, in conjunc-tion with methanation, as a heat-transfer system(ADAM – EVA Systems, see Section 5.3.4). Noknown commercial applications have yet beenfound.

2.3. Tubular Steam Reforming of LiquidHydrocarbons

In the mid-1950s, ICI initiated widespread useof naphtha as feedstock for the steam reformingprocess. This was enabled by the developmentof special nickel catalysts in conjunction with avery efficient naphtha desulfurization. Initially,the ICI process was frequently used for towngas production. Due to the abundance of naturalgas, naphtha reforming is meanwhile of minorimportance. Only at some locations with no ac-cess to natural gas there is still some hydrogenand syngas production based on naphtha.

Principles. Direct reforming of naphtha isnot generally different from natural gas reform-ing. The conversion follows Equation (8) in thesameway as already described for the higher hy-drocarbons contained in natural gas (see Section2.2.1). Like natural gas reforming naphtha re-forming is carried out in externally heated tubesover a nickel catalyst and yields, via a complexreaction mechanism, a gas mixture consistingof hydrogen, carbon monoxide, carbon dioxide,methane, and steam. The gas composition re-flects the equilibria of Equations (8 a) and (7) at

32 Gas Production

outlet conditions. If less active or partially deac-tivated catalysts are used, the equilibrium con-centration of methane (as per Eq. 8 a) is not at-tained, and the methane content of the reformedgas is higher than at equilibrium. Due to thehigher carbon content in the process feed naph-tha reforming results in a product gas with in-creased carbon monoxide (and dioxide) contentcompared to natural gas feed.

2.3.1. Commercial Processes

Amixture of desulfurized and evaporated naph-tha with superheated steam (molar steam: car-bon ratio 3: 1 or higher) is supplied to reformertubes at 450 – 550 ◦C.

Catalyst. Due to the high carbon: hydrogenratio in the feedstock naphtha direct reformingis characterized by an increased tendency of car-bon formation. This is mainly caused by naph-tha cracking and dehydrogenation reactions, es-pecially on acidic centers of catalysts based onalumina which is the most frequently used sup-port material. The breakthrough of naphtha re-forming has been enabled by the developmentof alkalized catalysts. The acidic centers of thesupport are neutralized by adding potassium tosuppress naphtha cracking. Moreover, the steamconversion of carbon once formed by naphthacracking is strongly enhanced by the presenceof potassium. However, alkali also delays com-plete reforming to carbon monoxide and hydro-gen. Therefore, the portion of the reformer tubethat is filled with potassium-laden catalyst is re-stricted to the minimum required for reliablenaphtha conversion; the rest of the tube is filledwith more active alkali-free catalyst.

Today, a controlled continual release ofpotassium is achieved by incorporating a specialpotassium compound (KAlSiO4) in the catalyst,thereby reducing the potassium vapor pressure[12]. Alternatively, alkali-free catalysts mostlybased on magnesium oxide are used.

Desulfurization. Naphtha generally con-tains organic sulfur compounds, and the sulfurcontent may be as high as several hundred partsper million. It must, therefore, undergo hydro-genating desulfurization with hydrogen at 350– 400 ◦C. Cobalt – molybdenum catalysts or, if

gases with a certain carbon oxide content areintended for hydrogenation, nickel – molybde-num catalysts are used. The sulfur converted tohydrogen sulfide is then adsorbed on zinc ox-ide, which has been used commercially as a hotdesulfurization agent since the 1940s.

2.3.2. Fuel Gas and Synthesis Gas fromLiquid Hydrocarbons

The description given in Section 2.2.4 and Ta-ble 9 applies analogously to the production ofgas fromnaphtha. These plants are distinguishedfrom those based on natural gas by the followingfeatures:

1) More complex desulfurization systems;2) Use of a special catalyst in the tubular re-

former and, in conjunction with this, a specialstart-up system;

3) Fewer reformer tubes per quantity of hydro-gen and carbon monoxide produced at equalheat loads per unit area (natural gas requiresconsiderably more reforming energy than hy-drocarbons in the naphtha boiling range); and

4) Larger carbon dioxide washing systems(naphtha has a higher carbon content thanmethane).

Details of modern integrated ammonia plants,particularly the steam system, are describedelsewhere (→ Ammonia) [9].

In towngas production by the ICI process, thereformed gas is produced with a heating valuebelow that of town gas. The calorific value of thereformed gas must, therefore, be increased

1) By adding low-boiling hydrocarbons (e.g.,LPG) and

2) By producing gases with a higher calorificvalue in special process stages for additionto the reformed gas. This has been achievedby means of a downstream gas recycle hy-drogenator (GRH) or a parallel rich-gas stagebased on naphtha (see Section 2.4), or by re-acting reformed gas from the tubular reformerwith additional naphtha and steam in a sec-ondary reformer [29].

2.3.3. Special Processes

Steamless Carbon Dioxide Reforming.Feedstocks with high carbon: hydrogen ratios

Gas Production 33

(such as naphtha) and low steam: carbon ratiofavor the production of synthesis gas with highcarbon monoxide and low hydrogen content (asrequired, for instance, for the oxo synthesis). Inpractical applications, however, this systemsoonreaches the point at which carbon formation oc-curs, particularly at high pressure. This difficultyhas until now been overcome by reforming thegases under established process conditions (i.e.,CO2 recycling with less steam than in regularsteam reforming), washing out carbon dioxide,separating hydrogen and carbon monoxide ina cryogenic or PSA unit, and finally remixingthem in the desired ratio. In this concept unde-sirable hydrogen surplus can often be used onlyto fire the reactor.

The hydrogen production can be minimizedby converting hydrocarbons only with carbondioxide according to the reactions

CH4+CO2 � 2 CO+2H2

or as example for higher hydrocarbons

C3H8+3CO2 → 6 CO+4H2

This leads to hydrogen: carbon monoxide ra-tios of 1: 1 or even less.

Steamless carbon dioxide reforming requiresstill more sophisticated conditions and catalyststhan naphtha direct reforming [30]. It has beenrealized in the Calcor C process that employsspecial catalysts and ambient or only slightlyelevated pressure. This system recovers the nec-essary carbon dioxide from the reformed gas ina monoethanolamine (MEA) or diethanolamine(DEA) wash unit, recycles it to the reformingstage, and adds imported carbon dioxide as re-quired [31].

Downstreamof the reformer and carbondiox-ide wash unit, raw gas from a C3– C4 feed-stock, for example, will contain hydrogen andcarbon monoxide at a ratio of ca. 0.65, with amethane content below 1 vol%. For higher car-bonmonoxide concentrations, ranging to> 99.5vol%, carbon monoxide and hydrogen are sepa-rated in a PSA unit or cold box. Plants designedfor 160 and 250 m3 (STP) of carbon monoxideper hour have been constructed to date.

Higher Feedstocks. Catalytic reforming ofhydrocarbonswith a higher boiling range has of-ten been investigated. In most cases, only short-term tests were made, which did not reach the

stage of commercial application. The use of veryhigh-boiling hydrocarbons is described in [32].

In other tests, naphtha, crude oil, and heavyresidues have reportedly been converted to re-formed gases with a hydrogen content of ca.60 – 65%without desulfurization or removal ofheavy metals [33]. The alternative of noncat-alytic conversion of high boiling hydrocarbonsor even residues and the high availability of nat-ural gas have slowed down the interest in thesefeedstocks at present.

2.4. Prereforming

Prereforming is a term which has been intro-duced into common nomenclature of gasifica-tion processes in the 1980s. In the most usedsense it means the application of an adiabaticlow temperature steam reforming step upstreamof a conventional steam reformer using naturalgas as feedstock. Regarding effect and condi-tions this process step corresponds to the steamconversion of naphtha in the so-called rich gasprocess which has been applied for nearly 40years. Even the same or very similar catalystsare used [34].

Prereforming can be defined as a steam re-forming process at limited temperatures (<700 ◦C) resulting in an intermediate product gas,themain compound of which is methane besidessteam. The process is in thermodynamic equilib-rium with respect to the methane (Eq. 8 a) andshift conversion reaction (Eq. 7). The productdoes not contain higher hydrocarbons. Normallythis intermediate gas is further processed in asteam reformer, but in special cases it can alsobe used for other applications, particularly forthe production of fuel gas (e.g., town gas, SNG).

2.4.1. Principles

Prereforming follows the same sequence of re-actions as tubular steam reforming characterizedby Equations (8), (8 a), and (7). But in contrastto the latter, the equilibrium is established at farlower temperatures. The main feature of prere-forming is the irreversible, complete conversionof the higher hydrocarbons fed to the process.Due to the comparably low temperature themaincompound of the product gas is—besides un-converted steam—methane. For this reason the

34 Gas Production

steam conversion of naphtha is called “rich gasprocess”, which reflects the fact that the result-ing product gas is rich in methane. The remain-ing compounds are hydrogen, carbon dioxide,traces of carbon monoxide, and the inerts con-tained in the feed.

Depending on the feedstock prereformingcan be endo- or exothermic. Generally the steamconversion of hydrocarbons to carbon oxidesand hydrogen is endothermic. But as prereform-ing is carried out at only moderate temperatures,the carbon oxides are further converted to meth-ane, a reaction that is highly exothermic. So, ifnaphtha is used as process feed, prereforming isexothermic whereas in the case of natural gas,the main compound of which is already meth-ane, the overall reaction is endothermic.

2.4.2. Catalysts

As prereforming is a steam reforming processat low temperatures, special catalysts are re-quired to achieve sufficiently high reaction rates.Common properties of these catalysts are a highnickel content (30 – 70 wt %) and a high spe-cific surface (e.g., 200 m2/g); the support mate-rial may be made, for instance, of alumina withminor additions of potassium or magnesium sil-icate. The geometric form is also adjusted tothe special requirements among which the con-version of higher hydrocarbons is the most im-portant. To restrict diffusion ways small pellets(e.g., 5 × 5 mm tablets) are generally used.

Because of the lower temperature, prereform-ing catalysts are still more sensitive to poison-ing than conventional steam reforming cata-lysts. This is especially true for sulfur poison-ing, which is at least partly reversible at thehigher temperature level of tubular steam re-forming, but irreversible at prereforming con-ditions. For this reason special care has to betaken in the desulfurization step, the residual sul-fur content of which should not exceed 0.2 mgper kilogram of gas. Naphtha is desulfurized byusing the refinery-proven technology of hydro-genating desulfurization (see Section 2.2.2).

2.4.3. Prereforming of Natural Gas

Of the different variations of this process, pre-reforming of natural gas is most frequently used

at present. Driving force for the application ofthis technology is the general constraint for in-creased economy.

The prereformer is installed upstream of thetubular steam reformer to ease its operation. Asall higher hydrocarbons contained in the naturalgas are converted, the risk of carbon formation,which is the most critical point in steam reform-ing, is considerably lowered. This allows the de-crease of the steam: carbon ratio and an increasein the heat load on the reformer tubes, whichmeans less energy consumption and smallerequipment. In addition hydrogen is producedby partial conversion of natural gas and further-more residual traces of catalyst poisons are ab-or adsorbed in the prereformer. This gives, ef-fectively, a catalyst of optimum activity at theinlet of the tubular reformer, where high activityis required most.

Prereforming is generally carried out in adia-batic reactors with a typical inlet temperatureclose to 500 ◦C. Due to endothermic hydrocar-bon conversion the outlet temperature is lowerby 25 to 40 ◦C, depending of the content ofhigher hydrocarbons in natural gas. Figure 8shows a characteristic temperature profile of anadiabatic prereformer.

Figure 8. Characteristic temperature profile in a natural gasprereformera) Fresh catalyst; b) Spent catalyst

An option that is frequently made use of inprereforming is the reheating of prereformed gas

Gas Production 35

Figure 9. Tubular steam reforming with prereformer and reheating of prereformed gasa) Prereformer; b) Reformer tube; c) Furnace; d) heat recovery

before feeding it to the steam reformer. As steamreforming technology generally has a surplus ofthermal energy which can only be utilized forsteamgeneration, this is an additional possibilityto improve the overall energy efficiency of a gasproduction plant. In conventional steam reform-ing the reintegration of available process heat ei-ther fromflue or fromproduct gas is restricted bythe risk of cracking, that may occur when heat-ing natural gas – steammixtures to temperaturesabove ca. 550 ◦C. This risk depends essentiallyon the content of higher hydrocarbons. Due tothe absence of hydrocarbons exceptmethane andthe increasedhydrogen content, prereformedgascan be reheated to 650 ◦C without any problem(Fig. 9).

An alternative to adiabatic operation is carry-ing out prereforming in heat exchanger type re-actors by supplying heat from an outside source.This source can be the product gas of the down-stream main reformer (steam reformer or sec-ondary reformer) or the flue gas of the fireboxof the steam reformer. In this case the outlet tem-perature of the prereformer can be up to 700 ◦C.By this arrangement even more available sensi-ble heat can be reintegrated in the process thanby reheating.

2.4.4. Prereforming of Naphtha; Rich GasProcess

The rich gas process is the traditional implemen-tation of prereforming. In addition to prereform-ing in its actual sense this process has been used

particularly for the production of fuel gases dur-ing the decades of its application.

Development of the rich-gas (high-methanegas) process started in England in themid-1950s[35]. Parallel developments took place in Ger-many and Japan from about 1960. Today, thecommercially available processes are those ofBritish Gas, BASF – Lurgi, and Japan Gasoline.

Initially, the process was used mainly forthe production of town gas. Soon its advan-tages were also applied to the production ofhigh-carbon monoxide synthesis gases. In theearly 1970s, the rich-gas process was used inthe United States to produce substitute (or syn-thetic) natural gas (SNG).

The rich gas process is carried out commer-cially at 350 – 550 ◦C and 1 – 5 MPa. The out-let temperature of a rich-gas reactor is ca. 30– 100 ◦C higher than the inlet temperature ofthe hydrocarbon/steam mixture. Like with pre-reforming of natural gas the composition of therich gas can be calculated from the equilibriumconstants of the reversible reactions (Eqs. 8 a, 7)according to [36] (see Fig. 10).

The sequence of reactions occurring typicallyin an adiabatic shaft reactor can be observed byfollowing the temperature profile. At the inletof the catalyst bed the temperature profile is atfirst determined by endothermic cracking andgasification reactions, which cause the temper-ature to drop by ca. 20 ◦C. Further in the courseof the reaction, exothermic hydrogenation reac-tions prevail, leading to a temperature increaseof ca. 60 ◦C above inlet temperature (Fig. 11).

36 Gas Production

Figure 10. Equilibrium composition of rich gas C:H = 6;2.0 kg H2O per kilogram naphthaa) 2.5 MPa; b) 4.0 MPa; c) 5.5 MPa

Figure 11. Characteristic temperature profile in a rich gasreactor

During operation, the temperature profilemoves through the catalyst bed to the gas out-let. When the upper temperature point (i.e.,the end point of the temperature increase) ap-proaches the end of the catalyst bed, naphthabreaks through. Depending on the downstreamreaction stage (e.g., tubular reformer or metha-nation), naphtha breakthrough can be toleratedonly up to a certain concentration.

Migration of the temperature profile is causedby exhaustion or blocking of the active nickelcenters because of recrystallization, polymer de-posits, or catalyst poisoning (sulfur, chlorine, ar-senic, etc.). If caused by polymer deposits, thisprofile migration can be reversed to some extentby regenerating with steam or hydrogen.

Three processes have predominated commer-cially in the past:

BASF – Lurgi Process. This process is usedcommercially for the production of town gas,high-carbon monoxide synthesis gas, carbonmonoxide, hydrogen, and SNG. The first com-mercial plants have been in operation since 1965[37].

This process uses a relatively stable rich-gascatalyst which is fully active as low as 400 ◦C.Another characteristic of this process is the pre-ferred use of indirectly heated naphtha evapora-tors.

CRG Process. The catalytic rich-gas (CRG)process is based on developmental work by F.J. Dent (Midland Research Station of the GasCouncil). The process was used mainly for pro-ducing town gas and SNG over a nickel cata-lyst on an alumina carrier material. The temper-ature profile moves almost unchanged throughthe entire CRG catalyst bed. The first commer-cial plants have been in operation since 1965[36].

MRG Process. A special feature of themethane rich-gas (MRG) process is the use ofa selective hydrogenation catalyst in the hydro-genating naphtha desulfurization stage, whichis capable of operating with hydrogenation gashaving a low hydrogen content [38].

Substitute Natural Gas (SNG). The firstcommercial plants for production of SNG werebuilt in the United States. They use a rich-gas

Gas Production 37

reactor, generally followed by two methanationstages, carbon dioxide washing, and gas drying.Steam: naphtha ratios of 2.0 kg/kg or slightlyhigher are applied. An improvement was the in-stallation of a second rich gas stage after the firstone. To the second rich gas stage only naphthais added (in about the same proportion as to thefirst rich gas stage) or naphtha and only a smallamount of steam. Thus, the overall steam: naph-tha ratio decreases to ca. 1.1 – 1.4 kg/kg. Thetemperature range of the second stage variesaccording to the process. The usual inlet tem-perature is 350 – 380 ◦C; the outlet temperatureis about 100 ◦C higher. The two-stage modelowers steam consumption (i.e., increases effi-ciency) and the amount of steam to be condenseddownstream. It improves the rich gas composi-tion by decreasing the hydrogen content afterthe second rich gas stage.

Usually, rich gas from the rich-gas stage(s) iscooled to ca. 300 ◦C and methanated in a shaftreactor under adiabatic reaction conditions byusing nickel catalysts. Because undecomposedresidual steam in the rich gas is present dur-ing this reaction, the process stage is termed wetmethanation. After cooling and steam conden-sation, themethane content can be increased fur-ther in a second methanation stage (dry metha-nation). The dry methanation stage can be ar-ranged upstream or downstream of the carbondioxide scrubber (see Section 5.3.3).

Thepossibility ofmultistage rich-gas produc-tion and multistage methanation offers a greatvariety of combinations and process alternativesfor SNG production. Figure 12 shows the over-all diagram of an SNG plant with combinedtwo-stage rich-gas production and single-stagemethanation.

Town Gas. The rich-gas process was origi-nally used for town gas production. The stan-dard production concept of town gas from naph-tha, with total conversion of naphtha in a rich-gas stage followed by partial reforming of therich gas in an externally heated tubular reformerto yield a hydrogen-rich gas (see Fig. 13), wasdeveloped quite early. Hydrogen-rich reformedgas from the tubular reformer is treated furtherin the usual manner (CO shift conversion, CO2scrubber, etc.).

The ICImethodof injecting evaporatedmeth-anol into town gas upstream of the high-tempe-

rature conversion stage has been applied in sometown gas plants as an elegant way of increasingcapacity to cover peak requirements in winter.This has enabled the capacity of existing plantsto be increased by 20 – 30%.

Methanol can be converted to town gas ina rich-gas stage. However, in view of the com-paratively high price of methanol, this is eco-nomically justifiable only for covering peak re-quirements in winter. Although capital invest-ment costs for the plant are low, the process hasfound only limited application.

2.5. Autothermal Catalytic Reforming

Catalytic reforming with air or oxygen has beenin use longer than either tubular reforming orpartial oxidation. Representatives of this firstgeneration technology are the ONIA-GEGI andthe SEGAS processes, both used for the pro-duction of fuel (town) gas. Feedstock in thesecases was gas oil or even crude oil, resulting ina high deactivation rate of the catalysts in use bycoking. Correspondingly these processes wereoperated discontinuously with periodic regener-ation. Plant designs today, however, differ sub-stantially from this type of plant, as target prod-uct (syngas) as well as feedstock (natural gas)have changed completely and continuous oper-ation is a “must”.

Characteristic for autothermal reforming isthe high flexibility of the process. This generallyenables a wide range of reaction conditions withregard to feedstock, steam: carbon ratio, tem-perature and pressure. Typical operating condi-tions are outlet temperatures close to 1000 ◦Candpressures between 3 – 4 MPa. Suitable feed-stocks are natural gas, LPG, and naphtha up toa boiling end point of ca. 200 ◦C. Furthermoreautothermal reformers are often fed with gases,which have already passed a steam reformer(“primary reformer”). In this case, the autother-mal reformer is called “secondary reformer”.To-day, the combination of tubular steam reformingand autothermal reforming is the standard tech-nology for the production of ammonia synthesisgas.Autothermal reforming ismoreflexible thantubular reforming because the higher allowableoperating temperature can compensate for anyincrease in methane slip which higher pressurewould otherwise cause. Furthermore, the higher

38 Gas Production

Figure 12. SNG planta) Desulfurization; b) Rich gas reactor; c) Hydrogenating gasification stage; d) Methanation; e) CO2 wash

Figure 13. Diagram of a town gas plant operation on the rich gas process based on naphthaa) Desulfurization; b) Rich gas reactor; c) Tubular reformer; d) HT conversion; e) Naphtha evaporator; f) CO2 wash

operating temperatures can also compensate fora wider equilibrium approach (difference bet-ween actual and equilibrium temperature at thereactor outlet) whichmay be caused bymild sul-fur poisoning of the catalyst. The maximum op-erating temperature is not limited by the tubematerial, as in tubular reformers, but only by thestability of the catalysts and the refractory liningof the reactor.

Oxidizing agents may be oxygen or air. Theselection depends on whether the presence ofnitrogen in the reformed gas is desirable, e.g.,for the production of ammonia. By proper par-tition of primary reforming and secondary re-forming with air, a final syngas with stoichio-metric hydrogen: nitrogen ratio is obtained. Theapplication as secondary reforming in ammoniasynthesis is by far the most frequent commercialapplication of autothermal reforming.

There are two versions of the process differ-ing in reactor design and arrangement of the cat-alyst bed. In one type of reformer the reactants(feedstock, steam, and oxygen or air) are dis-charged directly from a mixer to the catalyst ina fixed bed reactor. In the second type the mixeracts as a burner discharging into an empty, usu-ally conical space above the catalyst bed, whichonly fills the lower portion of the reactor. Thelatter modification has generally proven as themore versatile and is used predominantly.

Typical applications of the twoversions of theprocess are compared inTable 10.Amethane tailgaswas selected as feedstock for the all-catalyticversion, whereas reformed gas from a tubular re-former for the production of ammonia synthesisgas (typical use of a secondary reformer) wasemployed for the version with precombustion.

Gas Production 39

Table 10. Data for two reformed gases obtained by autothermalreforming

Methane tail gasreforming withoxygen addition

Reforming withair addition andprecombustion(secondaryreformer)

Gas for reformingComposition, vol%

CO2 7.11CO 3.90 4.25H2 1.80 34.32CH4 80.00 3.90C2H4 0.25C2H6 0.05N2 1.70Ar 12.30H2O 50.42

Flow rate, m3/h 55 000 160 000Pressure, MPa 3.0 3.04Temperature, ◦C 500 780

Oxygen or airFlow rate, m3/h 27 725 30 400Temperature, ◦C 240 550

SteamFlow rate, kg/h 113 200Temperature, ◦C 485

Reformed gas (dry)Composition, vol%

CO2 13.23 10.59CO 16.99 11.67H2 63.23 55.52CH4 1.08 0.19N2 0.65 21.77Ar 4.82 0.26

Flow rate, m3/h 148 400 109 000Reformed gas (moist)

Flow rate, m3/h 251 450 196 000Pressure, MPa 2.65 2.97Temperature, ◦C 925 965

The catalysts used consist of thick-walledRaschig rings 16 mm in diameter and height,having a hole of 6 – 8 mm in the middle (stan-dard size). The nickel content is about 3 –12 wt %. Aluminum oxide (preferably the α-modification) is used as carrier material, and thecontent of silicon dioxide should be< 0.2% be-cause of its volatility with steam.

Figure 14 shows the design of a moderncommercial autothermal reformer. The refrac-tory lining is almost always multi-layered; high-strength bricks or other lining materials (casta-bles) are used on the hot face, whereas mate-rials of lower bulk weight and high insulatingcapacity are used on the cold face (inner wall ofpressure vessel). Special attention must be paidto ensuring a close fit between lining and pres-sure vessel (vapor stops), because the consider-able pressure drop in the catalyst bed involves

the danger of bypass streams between lining andpressure vessel, which can lead to overheating(see Fig. 14). In one frequently used modifica-tion the reformer is provided with a water jacketthat generally cools at atmospheric pressure butoccasionally is slightly pressurized.

Figure 14. Autothermal reformer (secondary reactor)a) Burner; b) Refractory lining; c) Water jacket; d) Catalyst;e) Grid

The catalyst grid is usually made of heat-resistant refractory material. Catalyst can be re-moved by suction from the top.

The burner (mixer) is the most critical part ofthe autothermal reformer. This is true for bothtypes of reactor design. Inmost cases, the burneris designed as a single-tube mixer. Various typesof burner nozzles are available, all intended to

40 Gas Production

produce a homogeneous mixture of the oxidant(oxygen or air) and the gasification agent beforethey enter the catalyst bed. The burner internalsare made of highly heat-resistant steel or caststeel; some burner types are cooled.

To achieve sufficient mixing of the differ-ent inlet gas streams is a special problem in theflameless version of the process, because a cer-tain free space below the mixer is required. Aslinear gas velocities have to be high in this areato avoid ignition, the mechanical stress on thecatalyst is high, leading to abrasion and disin-tegration. This effect causes an increase in theempty space below the mixer that has to be com-pensated for by periodical refilling of catalystduring operation which is difficult to perform.The advantage of this operation mode is that themaximum temperature in the reactor does notessentially exceed the adiabatic end temperature(reactor outlet temperature).

In the open flame version of autothermal re-forming generally a good mixture of the feedgases is achieved. The problem to be solved inthis case is to keep the flame, which has temper-atures even in excess of 3000 ◦C, away fromthe reactor wall and the catalyst surface. Anenormous progress has been made in the lastyears by applying model calculations (Compu-tational Fluid Dynamics) for burner and flame[39]. Therefore, this modification of autother-mal reforming has become the clearly superiortechnology.

Besides ammonia synthesis—where au-tothermal reforming (secondary reforming) isstandard technology—and a specific applica-tion in the gasification of Fischer – Tropsch tailgas, autothermal reforming is also applied insome large-scale methanol plants. The syngasproduction in these cases follows the mode of“combined reforming”, a combination of tubu-lar steam reforming and autothermal reform-ing comparable to ammonia syngas production.Combined reforming enables the production ofa syngas with an optimum ratio of hydrogen tocarbon oxides and minimizes the feed gas con-sumption in this way. However, pure oxygen isrequired as oxidant. This requires additional in-vestment cost, if the oxygen has to be producedon site. Therefore combined reforming is eco-nomical particularly for large-size plants with acapacity above 1000 t/d (see Section 7.1).

As already stated, a main advantage of au-tothermal reforming is the “cold” shell of thereactor, which allows the application of high re-action temperatures and pressures. Whereas atemperature close to 1000 ◦C is already a cer-tain limit in view of the heat resistance of theused catalysts, the potential of applying higherpressures is not yet finally utilized. Pressures upto 8 MPa, usual in noncatalytic partial oxida-tion, seem to be possible also for the catalyticversion and will make it increasingly attractivein the near future. The design of a plant withone pressure level for syngas production (onlyautothermal reforming) and methanol synthesishas been patented.

3. Noncatalytic Partial Oxidationand Special Gasification Processesfor Higher-Boiling Hydrocarbons

3.1. Raw Materials

Partial oxidation is the most commonly usedprocess for gasification of heavy oil althoughvirtually all hydrocarbon mixtures regardless oforigin, are suitable feedstocks.

CrudeOil,CrudeOilDistillates andResid-ual Oils. Crude oil (→ Oil Refining) is a mix-ture of many different types of hydrocarbon.These can be classified in one of three ways: asalkanes, (and isoalkanes), naphthenes, and aro-matics; by the number of carbon atoms in themolecule; or by their boiling point, which is ap-proximately proportional to this number. How-ever, crude oils also contain varying quantitiesof other substances, which also play a role in theusability of the oil and, especially, of the residue.The most important impurities are sulfur com-pounds; others are oxygen- andnitrogen-bearingcomponents, as well as minerals (sand and dis-solved salts).

In addition, some impurities not naturallypresent in the crude stem from its transporta-tion; the main example is water, usually seawa-ter with a high salt content. Residual oils mayalso contain impurities that arise from wear ofcatalyst, refractory material, and equipment inconversion units.

As a general rule, low-boiling componentsare distilled from the crude oil and marketed as

Gas Production 41

light naphtha, heavy naphtha, kerosene, and gasoil. However, some crudes contain only smallamounts of light ends; the extremes are plasticasphalts (→ Asphalt and Bitumen). These re-sources, from which marketable products canbe obtained only by applying expensive refiningtechniques, are suitable as feedstocks for par-tial oxidation with practically no pretreatment,except removal of water, salts, or solids.

Crude oil distillates in the boiling range ofca. 35 – 360 ◦C are valuable raw materials forthe chemical industry and are also processed intomotor fuels or used as light fuel oil. Like crudeoil, these distillates are used for gas productiononly in special cases, for example, if a surplusof various distillates cannot be balanced out inthe refineries.

The gas recycle hydrogenator (see page 51),in which only hydrocarbons with a final boilingpoint ≤ 325 ◦C can be gasified, is an exampleof the way in which process considerations canrestrict feedstock suitability.

Residual oils are the preferred raw materialfor partial oxidation. First, they are generallycheaper than distillates; second, the percentageof residues that cannot be used for combustionpurposes is increasing because of pollution con-trol. For example, in some places, power stationsare only permitted to burn heavy oils with a sul-fur content < 0.8%.

Refinery technology, too, contributes greatlyto the increasing quantity of high-boiling, high-sulfur residues. To recover an ever-increasingproportion of light and middle distillates fromcrude oil, processes have been developed thatalso enable recovery of naphtha and gas oilfrom distillation residues. Among these pro-cesses are liquid-phase cracking (→ Oil Refin-ing, Chap. 3.2), catalytic reforming (→ Oil Re-fining, Chap. 3.4), solvent deasphalting (→ OilRefining, Chap. 3.6), and hydrocracking (→ OilRefining, Chap. 3.5). In some cases, the residuefrom these processes may even be solid coke;it contains practically all the impurities in crudeoil, apart from those sulfur compounds that havealready been driven off during distillation andsalt fromseawater,which is removedbeforehandin the desalter.

On assessing the suitability of residual oilsfor gasification, the quantity and type of ash,feedstock viscosity, and above all, the ratio ofalkalis to vanadium which is important for the

lifetime of the refractory lining must be consid-ered. Properties such as boiling range, carbon –hydrogen ratio, and sulfur content do not influ-ence suitability but are important for assessingdesign parameters. Some typical residual oilsand other feedstocks currently used for partialoxidation are listed in Table 11.

Other Feedstocks. In addition to crude oiland its manifold products, a number of otherhydrocarbon mixtures, as well as carbon- andhydrogen-rich compounds, are used for gasifi-cation of heavy oils. Chief among these are cer-tain liquid byproducts from coal gasification andupgrading.

In gasification of bituminous coal by theLurgi pressure gasification process (see Section4.4) aqueous condensates aswell asmediumoilsand coal tar are obtained as byproducts.Mediumoils have a variety of uses, but the generally fluidtar can only be recycled to the gasification reac-tor or used as a feedstock for heavy-oil gasifica-tion. Precleaning, at least by settling, is neces-sary because of the generally high ash contentof the tar, which may be as much as 15%.

Lignite tar oil is a byproduct of the car-bonization of lignite or lignite briquettes. Likecoal tar, it may contain a larger or smaller quan-tity of dust, depending on the nature of the rawmaterial and the characteristics of the carboniza-tion process. Carbonization of lignite is uneco-nomic today. However, certain plants are stilloperated; in some cases, the char and lightercondensation products are sold and the lignitetar oil is used as a feedstock for gas productionby partial oxidation. In gasifying coal and lig-nite tars, attention must be paid to the contentof low-melting ash, which may require slaggingoperation.

In a few cases, undesirable or unsalablebyproducts of synthesis processes are used forpartial oxidation, preferably when the gas pro-duced by oil gasification is suited to such synthe-sis processes. An example of this is oxo synthe-sis, in which oxo alcohols and various byprod-ucts are produced from propylene and carbonmonoxide: hydrogen mixtures. One of these by-products is isobutyraldehyde, which is formedin considerable quantities. It has already beenused for partial oxidation with great success.

The heavy products of tar sand extraction (→Tar Sands) and oil shale retorting (→ Oil Shale)

42 Gas Production

Table 11. Specific data for typical feedstocks (basis: Lurgi oil gasification, gasification pressure: 6 MPa; no soot recycle)

Feedstock type

Feedstock properties Natural gas Vacuumresidue

Propaneasphalt

Petroleumcoke (35%water)

Orimulsion a

(29% water)

Specific gravity (15 ◦C) 1.05 1.14 1.03 1.002Kinematic viscosity, m Pa · s 180 620 150 300

at ◦C 200 200 80 20C/H ratio, kg/kg 3.01 8.72 9.02 26.65 8.88S content, wt % 4.1 6.1 3.4 4.00V content, mg/kg 210 280 1000 300Ni content, mg/kg 70 120 220 50Na content, mg/kg 30 40 20 – 120 80Ash content, wt % 0.07 0.12 0.05 0.03Higher heating value HHV (dry), MJ/kg 53.32 39.82 39.65 42.63 42.84Feed flow rates, kg per 1000 m3 (STP) b of CO + H2

Hydrocarbon 286 360 372 426 368Oxygen c 338 378 387 416 403Steam (380 ◦C) 126 173 20 16Gasification mode boiler boiler boiler quench quenchProduct gas composition, mol% (dry)H2 62.0 45.4 44.0 32.9 43.9CO 33.5 49.1 48.5 58.0 48.7CO2 3.1 3.9 4.5 5.9 4.8CH4 0.8 0.3 0.2 0.2 0.2N2 + Ar 0.6 0.2 1.3 1.4 1.4H2S 1.0 1.4 0.9 0.9COS 0.1 0.1 0.1 0.1Product steam d (saturated)Gross, ex boiler, kg 755 929 990 755 793Pressure, MPa 9 13 13 1.1 1.1a Aqueous emulsion of a highly viscous crude from the Orinoco basin in Venezuela.b 1 m3 (STP) = 1 m3 at 0 ◦C and 0.101325 MPa.c Expressed as 100% oxygen; actual O2-purity: 99.5 mol%.d Total steam from heat recovery boilers (basis: boiler feed water at 105 ◦C); medium pressure boilers in quench mode are optional, productgas may be fed directly to CO shift unit.

may also be suitable as feedstocks. Their use,however, lies in the future.

Natural gas is still used as a feedstock forthe production of synthesis gases with a higherratio of carbon monoxide to hydrogen than isachievable by steam reforming. For economicalreasons, partial oxidation of natural gas is usedin small plants and in regions where natural gasis cheap. Lately it is considered as a means ofconverting “stranded gas” into marketable liq-uid products like methanol or synthetic fuel.

3.2. Partial Oxidation of Hydrocarbons

3.2.1. Principle

Partial oxidation is, in principle, the reaction ofhydrocarbons with an amount of oxygen insuf-ficient for complete combustion at temperaturesbetween 1600 and 1350 ◦C and pressures up to

15 MPa. It operates as a continuous process. Thebasic reactions are as follows:

CnHm+n/2O2�nCO+m/2 H2 (13)

CnHm+nH2O�nCO+(m/2+n)H2 (14)

CnHm+nO2�nCO2+m/2 H2 (15)

Theminimum amount of oxygen required forcomplete conversion of the hydrocarbons is in-dicated by Equation (13); 0.5 kmol of oxygen isrequired for every kilomole of carbon. If the re-action proceeds according to this equation,whenhigh-boiling distillates or residual oils (i.e., feed-stocks with a large number of carbon atomsper molecule) are being gasified sufficient heatis generally produced to warm the reactants toabout 1500 ◦C. Carbon monoxide and hydro-gen are the main products until the hydrocar-bons have been completely converted; only thencan carbon dioxide and water be formed from

Gas Production 43

surplus oxygen. Opinions differ on the reactionsequence; some test results indicate that carbondioxide and water are primary reaction products[40].

To prevent excessive temperature increase,steam is usually added, which reacts endother-mically with the hydrocarbons according toEquation (14). This leads to the formation ofmore hydrogen than would be expected fromconversion according to Equation (13).

The main elements of the hydrocarbon mix-tures—carbon, hydrogen, oxygen, and sulfur—are converted independent of the nature ofthe feedstock, to the following thermodynami-cally stable compounds: carbon monoxide, car-bon dioxide, hydrogen, water, methane, hydro-gen sulfide, and carbonyl sulfide. The propor-tions of various components in the gas mixtureare determined by a range of equilibria, whichinclude the shift conversion:

CO+H2O�CO2+H2 (7)

the methane equilibrium:

CH4+H2O�CO+3H2 (8a)

the hydrogen sulfide – carbonyl sulfide equi-librium:

H2S+CO2 �H2O+COS (16)

and

CO+1/2O2 �CO2 (2)

CH4+CO2 � 2 CO+2H2 (11a)

These equilibria are largely established in thegasification reactor between 1500 and 1350 ◦C.They remain practically unchanged during veryrapid cooling by quenching with water or by in-direct heat transfer to water boiling at a compar-atively low temperature. Below 900 ◦C, even anapproximation to equilibrium could be achievedonly by accepting very long residence times orby using catalysts. The latter cannot yet be ap-plied commercially because of soot formation.The principal advantages of partial oxidation are(1) the possibility of converting hydrocarbonsthat cannot be vaporized and (2) the complete de-struction of all carbon – carbon bonds; the onlyhydrocarbon remaining in the raw gas is meth-ane.

Methane formation is about ten times as highas indicated by Equation (8 a); this is probablydue to the slowness of the reaction according toEquation (11 a). As expected, increasing pres-sure promotes methane formation; at 1400 ◦Cand 3.0 MPa, the methane content of the prod-uct gas is ca. 0.3 vol%, whereas it is 0.55 vol%at 6.0 MPa. Methane content is also influencedby changes in reaction temperature and additionof larger or smaller amounts of steam.

Under conditions prevailing in the reactionzone, no free carbon should be present, accord-ing to either the Boudouard equilibrium

C+CO2 � 2 CO (10)

or the reaction

C+H2O�H2+CO (6)

Soot can be present in the raw gas only at be-low 1200 ◦C and around 5.0 MPa. Whereas theamount of soot produced during partial oxida-tion of methane is practically zero, in heavy-oilgasification about 0.5 – 2.5 mass% of feedstockis contained in the raw gas as free carbon. Someauthors assume that this soot formation is duesolely to thermal cracking of a proportion of thehydrocarbons; others attribute it partially to theBoudouard limit being reached during cooling.Ash particles in the heavy oils also seem to actas nuclei of condensation or catalysts for sootformation [41, 42].

In heavy-oil gasification, soot formation canbe nearly eliminated by increasing the reactiontemperature, at the cost of a considerable in-crease in oxygen consumption and a correspond-ing increase in carbon dioxide content at the ex-pense of carbon monoxide.

Sulfur components of the gas mixture arelargely hydrogenated according to Equation(16).Under normal reaction conditions, ca. 95%of the sulfur is converted to hydrogen sulfide andthe remaining 5% to carbonyl sulfide. No sulfurdi- or trioxide is detectable in the raw gas frompartial oxidation.

3.2.2. Types of Processes

Three processes are commercially establishedfor the production of gasmixtures (primary com-ponents: H2 and CO) from gaseous or liquid hy-drocarbons by partial oxidation: the Texaco, the

44 Gas Production

Shell, and the Lurgi process. The main steps ofall three processes are now practically identical.

The original differences between the pro-cesses were that Texaco used natural gas as thesole feedstock and changed to liquid hydrocar-bons only after the process of steam reformingof natural gas became available, whereas Shellused heavy oils from the beginning. Lurgi’s pro-cess in contrast was developed for lignite tar oilsand slurries and covers now the broadest rangeof feedstocks.

What still distinguishes the processes todayare the methods of heat recovery and soot man-agement. While all three use boiler configura-tions for high-grade heat recovery, only Texacoand Lurgi have available quench configurationsfor heavy-duty services. On the other hand, allthree can offer some kind of soot recirculationsystem while only Lurgi and Shell have builtmore economic once-through systems, whichare also environmentally favorable.

Over 320 plants operating on these three pro-cesses have been built to date.

Texaco Process. Originally, the Texaco pro-cess did not have a waste-heat recovery boiler(Fig. 15). Feedstocks are preheated as in theother two processes. Fired heaters are often pre-ferred to steam-heated ones for heavy oil aswell,probably by analogy with natural gas crackingpractice. Besides partial oxidation of natural gasand residue oil, this process has been developedfurther for partial oxidation of coal slurries (fordetails, see Chap. 4).

Process steam is used to atomize the heavyoil in the reactor, so the oil must be pressurizedto only a few bars above reactor pressure.

Oxygen is supplied to the reactor througha central tube; the oil – steam mixture is fedcoaxially through an annular gap. The gas mix-ture leaving the reaction zone is cooled to 250 –300 ◦Cbyquenchingwith hotwater in a quench-ing section of the reactor. Most of the sootand ash is removed from the gas at the sametime, whereas the water vapor content increasesconsiderably due to evaporation of part of thequenchwater. Thewater vapor content of the gasis sufficient for conversion of the carbonmonox-ide to a residual content of ca. 1.8 – 2.5 vol% ina downstreamhigh-temperature conversion unit.

After leaving the quenching tank, the gas isscrubbed again in a special venturi tube and a

soot scrubber (a packed column), which removethe soot to a residual content of 1 – 2 mg/m3. Ifthe gas is to be fed to a high-temperature con-version unit, the soot is scrubbed out with hotwater.

Shell Process. The basic flow sheet of thisprocess resembles the one presented in Figure16. The processes differ in the burner technol-ogy, the soot removal and the soot/ash manage-ment. The Shell process was right from the be-ginning equipped with a waste-heat boiler andis still available in this configuration only.

By means of reciprocating or centrifugalpumps the heavy oil feedstock (as single stream)is fed through a preheater to the reactor. Theso-called co-annular burner is of the steam-atomizing type: oil, steam – oxygenmixture andsteam are routed through concentric, alternatingchannels (slits) whose number determines thecapacity.Oil viscosities at the burner of up to 300mm2/s are manageable. The burner is cooled bya pressurized circulating cooling water system.

The refractory-lined reactor and the waste-heat recovery boiler are similar to those of theother processes. Soot is removed from the rawgas in two stages: most of it is washed out withwater in a quench tube [instead of the venturiscrubber d) of Figure 15] and collected in a sootseparator. In the scrubber, a packed column, theremaining soot is removed to a residual contentof < 1 mg/m3. Preferably soot-free water fromthe soot recovery unit is supplied to the top of thescrubber, and soot-laden water from the scrub-ber outlet is used for the quench tube. By coolingthe water fed to the scrubber, the gas is cooledto slightly above ambient temperature. The sootslurry is treated in a once-through soot/ash re-covery andmanagement system as described be-low.

Lurgi Process. This process was originallydeveloped for the gasification of tars, oils, andslurries formed as byproducts of fixed bed gasi-fication. This lends to an inherent robustness (asin slagging service) and flexibility (digesting un-mixable feeds) which the refinery-derived pro-cesses are lacking. TheLurgi process is availablein both, quench and boiler configurations.

Figure 16 shows the boiler configuration ofthe Lurgi process with metals ash recovery sys-tem. The feedstock is fed through preheaters

Gas Production 45

Figure 15. Texaco gasificationa) Preheater; b) Reactor; c) Venturi scrubbing section; d) Soot scrubber; e) Decanter; f) Flash distillation; g) Naphtha column;h) Water clarification

Figure 16. Lurgi oil gasification (boiler configuration) with metals ash recovery system (MARS)a) Preheater; b) Reactor; c) Waste heat boiler; d) Venturi scrubber; e) Fine wash scrubber; f) Slurry tank; g) Filter; h) Multiplehearth furnace; i) Return water vessel; j) Wastewater stripper

(a) to the reactor (b) by centrifugal pumps (orreciprocating pumps for smaller capacities). Apressure of only about 0.5 MPa above reactorpressure is needed for operation of the steam-atomizing burner. Preheating of feedstock andoxygen is mainly done for optimization of theprocess heat balance, i.e., to save on the costsof heavy oil and oxygen otherwise consumedfor internal heating. The lowering of feedstockviscosities is an additional—but welcome—

effect. The burner itself works well even withviscosities up to 500 mm2/s, meaning that ev-ery pumpable feedstock can be gasified. Thistop mounted multi-gun burner has the capabil-ity of handling different, unmixable fluids, e.g.,emulsions and slurries in addition to heavy oils.The broken feed line represents such an optionalsecondary feed.

The oxidant is preheated as well and mixedwith steam as moderator prior to being fed to the

46 Gas Production

burner. The burner and reactor are tuned byfluid-dynamic simulation to achieve mixing of thismixture with the feedstock within the smallestpossible volume. Thus, the reactor space is op-timally used for gasification to completion. Theburner is water-cooled using a dedicated circu-lating cooling water system at a pressure abovethat prevailing in the reactor. The entire reactor isrefractory-lined. In the nonslagging service as-sociated with boiler operation a refractory mate-rial with high alumina content has achieved thelongest lifetimes.

The raw gas leaves the reactor at about1350 ◦C and passes directly into the heat re-covery boiler (c) where 10 – 14 MPa steam isformed while cooling the gas to a few degreesabove saturated steam temperature. This boileris specifically designed for the high gas inlettemperatures and the particulates-charged gasat high velocities. A small part of the steamthus generated is used for feedstock and oxi-dant preheating, the bulk is superheated for usein steam turbine drives or in a combined cyclepower plant. The rest of the heat is recoveredin a boiler feed water economizer. Commonly,most of the heat content of the gas below thedew point is used to heat up make-up water orturbine condensate in a water preheater after theventuri scrubber.

The raw gas from the reactor contains a cer-tain amount of free carbon (soot). The plant isnormally designed for a soot content in the gasequivalent to 0.5 wt % of hydrocarbon reactorfeedstock. The soot in the gas is removed in atwo-stage water wash, which consists of a ven-turi scrubber (d) and the fine wash scrubber (e).The gas leaves the scrubber at 40 ◦C at a resid-ual soot content < 1 mg/m3. The soot slurry,which has an ash-soot content of 1 wt % car-bon, is routed to the metals ash recovery system(MARS, see below).

The Lurgi oil gasification is also available inthe quench configuration where the hot reactiongas is brought down rapidly to saturation tem-perature bydirectwater quench.This quench canbe selected for slagging service as in the originalapplication (coal – oil slurries, wastes) or for theless severe high-ash, high-salt residue service. Inslagging service the reactor is lined with a spe-cial high-chromium refractorywhich has provenits resistance in many years of operation.

Like in the Shell process, the gasmixture pro-duced contains approximately equal proportionsof carbon monoxide and hydrogen, ca. 5% car-bon dioxide, and small amounts of methane, ni-trogen, argon, and hydrogen sulfide. Althoughnot equally well suited to all syntheses (mean-ing those that require CO and H2, as well asthose that require one of the two, either alone,or either mixed with a third gas such as N2),its composition is nevertheless neutral for manyuses. Methanol synthesis gas or hydrogen canbe produced by converting part or all of the car-bon monoxide. If the necessary amount of nitro-gen is added, ammonia synthesis gas can be pro-duced. If the hydrogen – carbon monoxide ratioremains unchanged, the gas is suitable, after pu-rification, for oxo synthesis and for productionof maximum quantities of carbon monoxide bylow-temperature separation or absorption.

Soot/Ash Recovery and Management. Inthe Texaco process, soot is extracted from thesoot water with naphtha in a closed system andrecovered for recirculation to the reactor. Shellused an open system based on pelletizing in itsoriginal process but was subsequently obliged tochange to a closed system also based on extrac-tion with naphtha.

Due to difficulties with increased viscositiesof very heavy feedstocks, which makes carbonrecycle impossible, and due to poorer economicsof the recycle processes, once-through systemswere developed by Lurgi and Shell. These sys-tems basically incinerate the soot after filtrationand recover the heavy metals as oxides.

Texaco Method (Fig. 15). Soot water ismixed with naphtha, the soot and some of theash passing into the naphtha phase. After the twophases are separated, practically soot-free waterfrom the decanter is depressurized in a flashtank, to drive off dissolved naphtha. Naphtha isrecovered by condensation and recirculated tothe process; the same is true for naphtha-freewater.

Thenaphtha – sootmixture from thedecanteris mixed with heavy oil, and naphtha is distilledoff overhead in a distillation column.Oil collectsat the bottom of the distillation column togetherwith some ash and a residual amount of naphtha.The oil is mixed with more heavy oil and fed tothe reactor via the oil heater.

Gas Production 47

Lurgi Metals Ash Recovery System (MARS,Fig. 16). The problem of uneconomic carbon re-cyclewas solved byLurgi through application ofits vanadium roasting technology. A proprietarymultiple hearth furnace is used for controlledcombustion of the soot under conditions wherethe vanadium oxides neither melt nor corrode.

The soot slurry from the gasification (in both,quench and boiler configurations) is flashed toatmospheric pressure in a slurry tank (f). Theslurry is filtered (g) leaving a filter cake withabout 80 wt % residual moisture and a clear wa-ter filtrate, which is partly recycled as quench orscrubbing water. The filter cake is subjected toa controlled incineration process in a multiplehearth furnace (h). The product contains typi-cally about 75 wt % vanadium pentoxide (plusthe nickel and iron from the crude), which canbe sold to metal reclaimers. The process is au-tothermal, the heat of combustion being suffi-cient to evaporate the moisture content of thefilter cake. Excess water is stripped of dissolvedgases in a wastewater stripper (j) and pumped tofinal treatment.

Shell Method. Shell uses a similar system offiltration and incineration. Themain distinctionsto the Lurgi process as described above are in theselection of key equipment like furnace and fil-ters.

Selection Criteria for Quench Configura-tion Versus Heat Recovery Configuration. Inaddition to process-specificmarketing strategiesthere are technical and economic criteria for theselection between quench and boiler configu-rations (see Table 12). These must be studiedcarefully for every application of the Texaco andLurgi processes that alone offer alternatives.

Generally speaking, the heat recovery config-uration can be selected when the feedstock con-tains low amounts of precipitation-prone con-taminants, which could block the boiler tubes.Most refinery heavy residues fall in this cate-gory. In this case it is possible to run the gasifica-tion at highest efficiency by generating consid-erable amounts of high-pressure steam, whichalso yields highest efficiency in a possible down-stream integrated gasification combined cycle(IGCC) power plant. The additional investmentin the boiler is recovered quickly.

Table 12. Comparison of quench and heat recovery configuration

Quench configuration Heat recoveryconfiguration

FeedstocksGas, residues, wastes,slurries (sludges, coal,coke), extreme ashand/or salt contents

highest flexibility limited by possiblesalt precipitation

Product range:H2 + CO, H2, CO fastest (cheapest)

route to H2 and NH3

highest flexibility

Energy utilization:Steam production MP * steam available

trade off efficiency vs.cost

HP ** steam, heatrecovery at highesttemperature, highestefficiency for IGCCpossible

Investment cost: lowest cost forgasification unit

HP ** waste-heatboiler (highefficiency) at extracost

* MP = medium pressure.** HP = high pressure.

On the other hand, should the feedstock con-tamination be above the waste heat boiler’s lim-its, the quench configuration has to be chosen.It also may be selected when such contaminatedfeeds are expected in later operation of the re-finery or, when the customer wants to keep openall options for feedstocks like coke/coal slurriesand wastes. The quench configuration thus of-fers highest (potential) flexibility at lowest in-vestment cost traded off against a loss in energyefficiency.

The product side also influences the selec-tion, e.g., for hydrogen or ammonia as productit is more economical to route the quenched, i.e.,hot saturated gas directly into a rawgas shift con-version unit (see Section 5.1) instead of cool-ing it down and desulfurizing it before the shift.This is because the water quench raises all thesteam necessary for the CO shift reaction di-rectly within the reactor effluent. If a syngas richin CO is required, the heat recovery configura-tion is the appropriate choice.

3.2.3. Influencing Raw Gas Composition

Different types of gas are produced not so muchby changing operating conditions (e.g., pres-sure and temperature) as by changing the pro-portions of reactants, adding carbon dioxide,and employing further process steps immedi-ately downstream of actual gasification. Com-mercially produced gases can be classified in

48 Gas Production

three large groups: synthesis, fuel, and reduc-tion gases.

An overview of gas yields and compositionsin relation to the feedstock and gasifying media,together with other important data, is presentedin Table 13.

Synthesis Gas. Synthesis gasmust nearly al-ways be low in undesirable gas components suchas methane, nitrogen, and argon, and high ineither carbon monoxide or hydrogen, or both.Medium to high pressure is desirable as well.Whereas the nitrogen and argon contents of un-purified synthesis gas depend almost exclusivelyon the quality of air separation (nitrogen com-ponents of heavy oils generally cause a rise ofless than 0.1% in the nitrogen content of thegas), the methane content is largely a functionof gasification pressure (see Section 3.2.1).

The proportions of gasification steam, oxy-gen, and heavy oil are the same for almost allsynthesis gases, regardless of whether hydro-gen- or carbon monoxide-rich gases or gaseswith a defined ratio of hydrogen to carbonmonoxide are produced. For the last category,which includes primarily synthesis gases for theproduction of oxo alcohols, carbon dioxidemustbe recycled to the reaction zone to increase thecarbon monoxide content of the synthesis gasand obtain the correct ratio of hydrogen to car-bon monoxide (which can range from 1: 1 to 1:1.5 for certain oxo processes).

To manufacture synthesis gases in which no(or only partial) conversion of carbon monox-ide to hydrogen and carbon dioxide with steamoccurs, use of the waste-heat boiler to cool thehot reformed gas is preferred to quenching withwater or steam. Examples are gases for carbonmonoxide production, some oxo processes, andmethanol synthesis.

A special case is partial conversion of car-bon monoxide immediately following the gasi-fication reaction. This can be brought about byquenching with defined quantities of water orsteam, either in the bottom section of the gasi-fication reactor or in a quench tube located bet-ween the reactor and the waste-heat boiler.

Here, best results are achieved by using steamand cooling the gas to ca. 1000 ◦C.At lower tem-perature, the water gas equilibrium, (Eq. 7) can-not be reached without a catalyst. A very flexi-ble and economic alternative is the use of a direct

quench reactor in parallel to a reactor/waste-heatboiler train.

If, on the other hand, production of hydrogen-rich synthesis gas is desired, direct quenchingof the hot synthesis gas, followed by catalyst-based shift conversion is the preferred route.An example is the production of ammonia fromheavy oil in countries with cheap coal reservesand little oil of their own, (e.g., India), whichare able to supply coal-based energy at low cost(whether directly from steam or indirectly viaelectric power).

Oxygen with a purity of > 95% is used forgasification. In fact, the oxygen content is usu-ally > 99%, because the cost of achieving thehigher purity is always less than the extra costof gas purification and synthesis (e.g., by ad-ditional consumption of synthesis gas) to com-pensate for hydrogen or carbonmonoxide lossesfrom increased discharge of purge gas.

The possibility of manufacturing ammoniasynthesis gas from heavy oil by gasificationwithenriched air (i.e., with a N2 – O2 mixture con-taining > 21% O2) has also been proposed butnot exploited commercially. The higher costs ofpurifying and conditioning the gas that resultfrom the increased volume of synthesis gas be-cause of its nitrogen content outweigh the sav-ings in gasifying medium. (Enriched air can beproduced from air by mixing with pure oxygenor by adsorptive depletion of nitrogen.)

Fuel Gas. In contrast to synthesis gas pro-duction, a low content of inerts is not particularlyimportant for the manufacture of fuel gas. Onlythe required heating value determines the gasi-fying medium, pure oxygen or air. Combustionof larger volumes of gas instead of gas with anincreased heating value is often justifiable eco-nomically. However, not only the heating valuebut also the density and thus the Wobbe index(see Section 8.2.1.3) of the fuel gas are almostalways stipulated for reasons of interchangeabil-ity with fuel gases from different sources. Whenfuel gases containing carbon monoxide are usedin the private sector, it is mandatory to minimizetheir toxicity by limiting the carbon monoxidecontent, generally to between 1 and 3 vol%.

Hot Carburetion. Fuel gas specificationsfor the private sector require heating values rang-ing from 17 500 (town gas) to 34 000 kJ/m3

Gas Production 49

Table 13. Typical results of partial oxidation (basis: Shell gasification)

Type of gas: Untreated raw gas Synthesis gas, Fuel gas, Fuel gas, Reduction(H2-rich) carbureted low-calorific gas

Feedstock: Bunker-C fuel(without sootrecirculation)

Vacuumresidue (withsootrecirculation)

Bunker-C fuel (with soot recirculation)

Operating pressure, MPa 6 6 6 6 1.5 35Preheating temperature, ◦C

Heavy oil – natural gas 245 280 245 245 245 245O2– air 245 260 245 245 400 245Steam 310 360 310 310 310 310CO2 245Feedwater 160 105 160 160 160 160

Feed quantitiesFeedstock, pure kg 100 100 100 100 100 100O2 (98 – 99.5%), m3 77.6 70.7 74.1 73.7 81.2Air, m3 450Steam, kg 45 50 45 45 30Quenching steam – water, kg 100 * 60 **

CO2, m3 50

Light naphtha, kg 65Product gas volume, m3 303 290 327 354 602 305Gas composition, vol%

CO2 4.32 4.63 13.34 13.32 2.36 7.68CO 46.55 48.92 33.29 31.02 22.34 52.29H2 47.15 44.94 51.49 26.68 16.71 38.03CH4 0.60 0.3 0.60 27.79 0.03 0.60N2+ Ar 0.56 0.2 0.54 0.49 57.70 0.60H2S + COS 0.81 1.01 0.73 0.70 0.16 0.80

Lower heating value (LHV), kJ/m3 11 365 10 105 16 897 4670 11 077Steam output, kg 275 258 301 252 580 320Steam pressure, MPa 10.5 8 10.5 10.5 10.5 10.5Steam exported, kg 230 213 118 207 540 290

* Steam.** Liquid.

(natural gas). However, gasification of heavy oilproduces heating values of only 4000 – 11 000kJ/m3, depending on whether air or oxygen isused as the gasifying medium. Thus, the heatingvalue must be improved. This can be achieved—after appropriate gas conditioning—eitherby methanation without addition of other fu-els or by hot carburetion with addition of liquidhydrocarbons having a low final boiling point.Methanation is necessary if heating values> 20000 kJ/m3 are required. For the production oftown gas with heating values between 16 500and 19 000 kJ/m3, however, hot carburetion ismore beneficial from the heat economy aspect,because it operates with practically no addi-tion or subtraction of heat. This is in contrastto the strongly exothermic methanation process,in which about 10 500 kJ (i.e., ca. one-fifth theheating value of the feedstock) must be dissi-pated per cubic meter of methane formed.

Due to the widespread availability of natu-ral gas— in some places as LNG—carburetionhas lost most of its significance as a commercialprocess.

As can be calculated with the aid of Table13, the thermal efficiency (i.e., ratio of the heat-ing value of the product gas to the heating valueof the fuel) is 81% for normal gasification andincreases to 89%with hot carburetion, at the ex-pense of reduced generation of steam fromwasteheat.

FuelGas for IGCCPowerPlants. Aspecialarea of fuel gas production is for gas turbines inpower plants. Due to increasingly stringent airpollution control regulations, operating coal- oroil-fired steam boilers without flue gas desulfu-rization systems is impossible. Because clean-ing large gas volumes at slightly above atmo-spheric pressure is very expensive, attempts havebeen made to find a more economic solution by

50 Gas Production

applying fuel gasification with subsequent gascleaning under pressure. As shown in Figure 17,partial oxidation of heavy oil can be combinedbeneficially with a gas turbine-driven generator,not only from the point of view of gas cleaningbut also with regard to generation of high-pres-sure steam. For the arrangement illustrated, anefficiency well above 46% is achieved, 2 – 4%better than for a conventional steamboiler powerplant [43].

IGCC Power Plants (Fig. 17). A heavy refin-ery residue is gasified with oxygen from an airseparation unit (i) in a Lurgi oil gasificationplant (a) at 6.5 MPa, producing a lean gas witha heating value of 11 000 kJ/m3. Oxygen andresidue are preheated to 280 ◦C. Cooling in thewaste-heat recovery boiler produces high-pres-sure steam at 13.0 MPa. After the soot is re-moved as described above and the gas is cooled,carbonyl and hydrogen sulfides and all othertrace contaminants are removed to 20 ppm inone step in a Lurgi Rectisol unit (b) (see Section5.4.2.1). The desulfurized gas passes through agas expander (f) generating electric power andrefrigeration to cool the methanol within Recti-sol. This dual use expansion adds significantlyto the overall efficiency of power generation.

The low-calorific gas is burned in a gas tur-bine (d). For NOx -control the gas is further di-luted by saturation with water and—optionally—with nitrogen from the air separation unit.Waste heat from the gas turbine off-gas and thesaturated steam from the gasification are usedin the waste-heat steam generator (h) to produceadditional steam and to superheat steam. The en-ergy of the superheated steam is used in a con-densing steam turbine (g) that drives a generator(e). For maximum efficiency the steam systemwill be designed for three levels with one or tworeheat stages.

The gas turbine is also coupled to an air com-pressor in which combustion air for the gas tur-bine is compressed. A part of this compressedhot air can be routed through a cooler – saturatorsystem producing steam for fuel gas saturationbefore entering amediumpressure air separationsaving the separate air compressor. The nitrogenfrom the air separation in turn can be recycledback to the gas turbine for NOx -control and tomaintain themass flowof the turbine. Air and ni-trogen integration as described yields higher ef-ficiency than a nonintegrated system with inde-

pendent air separation unit (ASU) compressor,but involves increased complexity of the plant.Therefore different designswith varying degreesof integration exist to accommodate for all re-quirements.Many versions of this IGCCprocessexist, some of which export steam or synthesisgas togetherwith electrical power.Coproductionof this type results in better overall economicsfor a plant.

A good example for this concept is a refer-ence plant for the Lurgi oil gasification (SVZ,Schwarze Pumpe, Germany) where the oil gasi-fier together with coal-waste gasifiers producesa synthesis gas which is fed to a combined cy-cle and a methanol plant [44]. Another exampleis a large refinery in Pernis (The Netherlands)where a Shell gasification feeds two thirds of itscapacity to a hydrogen production plant and onethird to an IGCC power plant [45].

Reducing Gas. Gasproduced fromheavyoilis also used in ore reduction. The carbon diox-ide and water vapor content or reducing gasesmust be as low as possible. Water vapor canbe removed easily from the gas by condensa-tion, which presents no problem if the gas hasbeen freed of soot, sulfur, and carbon dioxide.However, for reduction of less sensitive ores,uncooled or only partly cooled, soot-laden gasmanufactured from low-sulfur oils is usedwhen-ever possible. Thus, water vapor content may bean important element of the gas composition.The quality of reducing gases is measured bythe degree of oxidation η, which is defined asfollows:

η0 =

(cH2O+cCO2

) ·100cH2O+cCO2+cH2+cCO

(%)

Degrees of oxidation ≤ 5% are the goal (seeTable 13).

Reducing gases, which are always produceddirectly alongside the ore reduction furnace, dif-fer from synthesis and fuel gases in pressure andtemperature. Synthesis and fuel gases must bemaintained at high pressure and below ambienttemperature: for synthesis gases, with regard togas purification and compression; for fuel gases,with regard to gas purification, distribution, andstorage. For admitting reducing gas to the shaftfurnace, a slight overpressure is sufficient; onthe other hand, a temperature up to 1000 ◦C isdesirable. The pressure range for manufacture

Gas Production 51

Figure 17. Lurgi oil gasification for combined cycle power plantsa) Oil gasification; b) Gas purification; c) Sulfur recovery; d) Gas turbine; e) Generator; f) Gas expander; g) Steam turbine;h) Heat recovery steam generator; i) Air separation; j) Air compressor (compressed air for air separation may be extractedfrom air compressor for gas turbine); k) Nitrogen compressor (optional); l) Motor; m) Combustion chamber

of reducing gas is currently 0.3 – 0.5 MPa, butefforts are in progress to achieve ore reductionat ca. 2 MPa.

3.2.4. Submerged Flame Process

The BASF submerged flame process is also apartial oxidation process; it was developed toproduce acetylene (→ Acetylene, Chap. 4.2.3)and ethylene from natural gas and crude oil.Its distinctive feature is maintenance of a flamebeneath the surface of the oil sump, to ensurehigh temperature and short residence time. Hotgases produced by partial combustion and crack-ing, which have a temperature of ca. 1500 ◦C,are quenched almost instantly to 250 ◦C in thehot oil sump. Thus, the thermodynamic equilib-rium for the high temperature is “frozen”. As inother partial oxidation processes, soot is formed,which collects in the oil sump and is recirculatedto the burner. After condensable cracking prod-ucts are removed, the cracked gas, which stillcontains a small amount of soot, is scrubbedwithcondensate. The condensate is then recirculatedto the submerged burner.

The process operates between 0.2 and 1.0MPa and yields a cracked gas with the followingapproximate composition, in percent by volume,which can be processed into synthesis gas:

Carbon dioxide 7 – 9 Ethylene 6 – 7Carbon monoxide 40 – 45 Propylene 1 – 2Hydrogen 23 – 30 Higher olefins and

acetylenes0 – 5

Methane 4 – 5

From the residual gas, 3.5 t of ammonia, for in-stance, can be produced per tonne of acetylene.

3.3. Hydrogenating Gasification

Hydrogenating gasification, like partial oxida-tion, is a continuous process.

Principles. In contrast to partial oxidation,hydrogenating gasification is not based on com-bustion of part of the feedstock; instead, a hy-drogenation reactionwhich proceeds exothermi-cally overall takes place with addition of hydro-gen. Liquid hydrocarbons are cracked primar-ily to form methane and ethane. This process,therefore, is beneficial for the production of fuelgases with a high heating value: on the otherhand, direct manufacture of carbon monoxide-and hydrogen-rich synthesis gas is not possible[46].

Gas Recycle Hydrogenator (GRH) Pro-cess. The British Gas Corp. began to developthe GRH process in the early 1960s [47]. Thisenables the conversion of liquid hydrocarbonswith a final boiling point up to ca. 325 ◦C toa gas mixture consisting mainly of hydrogen,methane, and ethane under controlled tempera-ture and pressure in a shaft reactor, termed the

52 Gas Production

hydrogenator. Hydrogenating gasification pro-duces no soot. The reaction takes place at 1.8– 7.0 MPa and 350 – 500 ◦C at the inlet or ca.750 ◦C at the outlet, depending on the boilingrange of the feedstock. The latter must contain acertain amount of sulfur (about 10 ppm), whichacts as a poison on reactor components made ofnickel alloy, to prevent them from acting as cat-alysts that might lead to soot formation from thehydrocarbons. After desulfurization, aromaticsremoval and methanation, an SNG may be pro-duced with a higher heating value (HHV) of 39100 kJ/m3, and a pressure of 6.0 MPa. The ther-mal efficiency of the process is 85%.

Fluidized-Bed Hydrogenation (FBH) Pro-cess. The FBH process, also developed by theBritish Gas Corp., enables gasification of crudeand extremely heavy oils. Feed oil and hydrogenare superheated and fed into the FBH reactor, inwhich hydrogenation of the hydrocarbons takesplace in a fluidized bed formed from coke bet-ween 700 and 820 ◦C. The most important pro-cess variables are temperature, residence time,and oil: hydrogen ratio. Part of the hydrogen issupplied to the reactor through a grate from be-low to fluidize the coke.

Most of the hydrogen is mixed with the oiland sprayed from below into a central tube. Bythis means, rapid circulation is achieved in thereactor. The oil evaporates at the outlet of thespray nozzle. Any oil that does not evaporate ad-heres immediately to the flowing coke particlesand is entrained with them. If the raw materialcontains residue-forming components, coke par-ticles are also formed. Therefore, a small amountof coke must continually be withdrawn from thereactor.

Gas leaving the reactor is quenchedwith lightaromatics, and the condensate is removed fromthe rich gas. Most of the sulfur in the oil is con-verted to hydrogen sulfide. A gas produced fromAlgerian crude with the addition of ca. 6.5 kg ofhydrogen per kilogram of oil has a HHV of 39400 kJ/m3 at a pressure of 5.0 MPa. For SNGproduction, gas from the FBH plant must un-dergo a second hydrogenation step to reduce itshydrogen content [48].More details on both pro-cesses, GRH and FBH, may be found in [49].

4. Gas Production from Coal, Wood,and Other Solid Feedstocks

This chapter deals with the production of gasesfrom solid carbonaceous materials such as coal,peat, wood, solid wastes and solid residuesfrom petroleum upgrading (delayed coke, etc.).The gasification process occurs via pyrolysisreactions (thermal decomposition) and subse-quent heterogeneous (gas – solid) reactions ofthe solid residue from pyrolysis with reactivegases (O2, steam, CO2, H2).

Besides carbon, solid feedstocks containother elements such as hydrogen, oxygen, nitro-gen, sulfur, heavy metals, and trace elements invarying amounts. The feedstocks are convertedto a gaseous product suitable for use either as asource of energy or as a rawmaterial for the syn-thesis of chemicals and liquid or gaseous fuels.The gaseous product can be handled with maxi-mum convenience and purified easily from com-pounds that would cause pollution problems.Thus, gasification is an upgrading process thatgreatly extends the uses of solid fuels.

During their long history, the science andtechnology of solid feedstock gasification haveadvanced in various stages induced by externalfactors. At first directed toward the productionof fuel or town gas for lighting and heating, theinterest in gasification shifted in the first half ofthis century to synthesis gas for the productionof chemicals and liquid motor fuels.

After a period of increasingly intense com-petition from oil and natural gas, gas produc-tion from solid feedstocks has become less im-portant since the 1940s. In the late 1960s andearly 1970s, some industrialized countries beganto expect natural gas shortages which, togetherwith oil price increases, changed the attitude to-ward utilization of solid fuels, particularly coalavailable in many parts of the world. More andmore, the finite nature of fossil fuel reserves, theneed for conservation, and good resource man-agement became accepted and the significanceof the world’s coal reserves was recognized. In-terest in coal utilization increased, primarily inthe fields of gasification and liquefaction (→Coal Liquefaction), and large sums of moneybecame available from governments for the re-search, development, and demonstration of newprocesses and the improvement of existing ones.More recently, since large resources of natural

Gas Production 53

gas have become available, the interest in coalgasification has decreased considerably.

Petroleum-derived residues have gainedsome attention as feedstocks for gasification,with the increasing yield of residues in petro-leum refining because of decreasing oil qualityand the need for their economic and environ-mentally acceptable utilization, for example, forH2 production. Use of wood as a feedstock hasplayed a continuous role in small-scale gasifi-cation units (e.g., for the production of gaseousfuels) and more recently gained interest in thecontext of biomass energy applications.

Advances in gasification have been madepossible by progress in other fields. Fundamen-tal research on the mechanism and kinetics ofchemical reactions, mass and heat transfer, fluiddynamics, and mathematical modeling has con-tributed to a better understanding of the phenom-ena occurring in gasification processes long aftertheir successful operation started.

This chapter introduces fundamental aspectsand process technologies of gas production fromsolid feedstocks. Given the variety of new de-velopments in the past decades, only a selectioncan be presented here to help identify and dis-cuss basic principles, state of the art, and futuretrends. For more details, the reader is referredto review articles on gasification fundamentals[50 – 52], gasification processes [53 – 59], orboth [60, 61], and to numerous publications onthis subject.

4.1. Fundamentals

Gasification of solid feedstocks involves thecomplex interaction of chemical reactions andphysical processes such as heat and mass trans-fer.

Figure 18 shows a scheme of the main chem-ical reactions that occur when carbonaceousma-terials are exposed to reactive gases at high tem-perature. The yield and composition of the prod-uct are determined by the combined effects offeedstock properties and reaction conditions onthe thermodynamics and kinetics of individualreactions. For example, the yields of liquid by-products (e.g., tar, phenols) depend on the re-action conditions that prevail during pyrolysisand secondary reactions of the volatile products(e.g., temperature, gas-phase residence time).

4.1.1. Thermodynamics of ChemicalReactions

Some reactions of the scheme in Figure 18 areendothermic, in particular the gasification re-actions of char with steam or carbon dioxide,which are significant in most gasification pro-cesses. (For heat effects of individual reactions,see Chap. 1)

The resulting heat requirement is met eitherby reactions with oxygen (in autothermic pro-cesses) or by an external heat source (allothermicprocess). Minimization of overall heat require-ments has been one incentive among others forimproving process efficiencies and economics.

For example, the idea of applying reactionconditions that favor the overall reaction

2 C+2H2O→CH4+CO2

which is nearly thermoneutral, has affectedgasification process development efforts duringthe times when production of CH4 from coalappeared attractive.

In some of the chemical reactions of Fig-ure 18, conversion is limited by equilibriumconstraints under typical gasification conditions(temperature, pressure). Most significant in thisrespect are the reactions of char with steam, car-bon dioxide, and hydrogen. Equilibrium con-stants for these and other reactions, aswell as theeffect of temperature, are presented in Chapter1.

Thermodynamic analysis may aid in exam-ining the limits for conversion and product gascomposition as affected by such process vari-ables as temperature, pressure, and stoichiome-try. For example, the effect of temperature andpressure on equilibrium gas composition for in-dividual char – gas reactions with steam, carbondioxide and hydrogen is shown inFigure 19. Thecurves reflect two general equilibrium princi-ples: (1) higher temperature favors products ofendothermic reactions and (2) higher pressureshifts the equilibrium in the direction of loweramounts of gaseous components. Some past in-centives of process development were based onthermodynamic analyses of gasification. For ex-ample, to produce methane-rich gases (substi-tute natural gas, SNG), high pressure and lowtemperature are advantageous under equilibrium

54 Gas Production

Figure 18. Main chemical reactions during gasification of solid carbonaceous materials

aspects. As a consequence, research and devel-opment efforts were directed toward applyinghigher pressure and enhancing feedstock reac-tivity (e.g., via catalysis), thereby lowering gasi-fication temperature.

Figure 19. Effect of temperature and pressure on equilib-rium gas composition in the presence of solid carbon

However, only the limits of conversion andproduct gas composition can be identified bythermodynamic analysis. In real gasificationsystems, equilibrium is generally not reachedbecause of kinetic constraints.

4.1.2. Kinetics

Pyrolysis Reactions. Among the chemicalreactions that occur during gasification, pyro-lysis is generally fast. This is illustrated in Fig-ure 20, which presents first-order rate constantsof some reactions shown in Figure 18. Underly-ing values for activation energies and preexpo-nential factors are taken from various literaturesources (Table 14). Formation of tar, methane,and hydrogen as volatile products from Mon-tana lignite is presented as part of pyrolysis.

The rates of primary pyrolysis product for-mation can be described by

dVi

d t= k0iexp (−EAi/RT ) (V ∗

i −Vi) (14)

where V i is the total amount of product evolvedfrom reaction i until time t, and V i * is the valueof V i at t → ∞.

Kinetic parameters for pyrolysis product for-mation have been determined for various kindsof coal and experimental conditions. Extensivecompilations of pyrolysis data and kinetic infor-mation drawn from them can be found in reviewpapers [51, 62].

Char – Gas Reactions. Of the char – gas re-actions, combustion is generally fastest, andgasification is generally faster with steam thanwith carbon dioxide and hydrogen.

Gas Production 55

Table 14. Examples of first-order rate constants for reactions occurring during gasification (for nomenclature, see text)

k0, EA, V *, References−1 J/mol wt %, (maf) b

Pyrolysis product formation tar 2.0×1017 315 281 3.5 (+ 2.9) [63] (lignite)CH4 4.7×1014 290 578 1.1 (+ 0.4) [63] (lignite)H2 1.6×1018 371 806 0.6 [63] (lignite)

Char – gas reactions C – O2 1.8×106a 113 050 [64] (bituminous coal)C – H2O 4.1×102a 146 545 [65] (subbituminous coal), [66]C – CO2 2.5×102a 146 545 [65] (subbituminous coal), [66]C – H2 8.4×10−4a 67 201 [65] (subbituminous coal), [66]

Tar cracking coal tar 1.2×109 164 968 [67] (bituminous coal)8.0×1012 272 992 [67] (bituminous coal)

a For (pi – peiq) = 1.013×105 Pa (see Eq. 15).

b maf = moisture and ash free.

Figure 20. Examples of first order rate constants for re-actions during gasification (pyrolysis product formation:Montana lignite [63], tar cracking: bituminous coal [67],char – gas reactions: bituminous and subbituminous coal[66]Pyrolysis product formation: a) Tar; b) H2; c) CH4;Char –gas reaction: e) C−O2; f) C−H2O; g) C−CO2; h) C−H2

Although detailed kinetics include noninte-ger reaction orders and inhibiting effects of re-action products, simple first-order rate laws canbe used for estimating and modeling purposes:dmcd t

= −k0exp (−EA/RT )mc(pi−peqi

)(15)

where pi is the partial pressure and peqi the equi-

librium partial pressure of reactant i, and mc isthe mass of solid carbon in the char. Values ofk0 and EA for char – gas reactions vary with thekind of solid feedstock, according to its reac-tivity. Values of gasification rate constants forchars derived from various feedstocks vary withup to eight orders of magnitude. High transientreactivities can be observed under rapid heatingconditions, (i.e., when pyrolysis and gasificationof the resulting char occur simultaneously). Ki-netic parameter values for char – gas reactionsin Table 14 are taken from a modeling study ofcoal gasification in moving-bed reactors.

A characteristic of gas – solid reactions is theexistence of a regime of temperature and pres-sure in which mass transfer affects overall ki-netics. Extra- and intraparticle mass-transfer re-sistance can be incorporated in Equation (15) asfollows:

dmcd t

= −keffmc(pi−peqi

)(16)

keff =1

1kg+ 1

η kr

(17)

where kg is an extraparticle (film) mass-trans-fer coefficient; kr, a first-order chemical reac-tion rate constant; and η, an effectiveness factorthat includes resistance to pore diffusion. Thetheoretical background can be found in reactionengineering textbooks (e.g., [68]; → Fluidized-Bed Reactors).

The temperature dependence of Equation(17) in its general form is shown in Figure21. Accordingly, the chemical reaction is rate-controlling at lower temperature because it hasthe highest activation energy and extraparti-cle diffusion is limiting at high temperature.

56 Gas Production

In between ranks the intraparticle mass-trans-fer regime. Because these regimes may all playa significant role in technical gasification theirlimits must be defined by means of kinetic reac-tor analysis. The effects of feedstock reactivity,particle size, gas flow rate, and total pressure arediscussed in [52].

Figure 21. Temperature dependence of effective rate con-stant for gas – solid reactions with porous particles, sche-matic representationA) Film-diffusion controlling (cG gas concentration, δ filmthickness)B) Pore diffusion controllingC) Chemical reaction rate controlling

Gas-Phase Reactions of Volatiles. Gas-phase reactions of volatile pyrolysis products in-clude a variety of individual reactions. Amongthem are cracking, reforming, combustion oftarry species and hydrocarbons, dealkylation,dehydroxylation, and shift conversion of carbonmonoxide.

The rates of these reactions depend, in a com-plexway, on the reactivity of the reacting speciesand on reaction conditions (partial pressures,temperature, catalytic effects).

Again, first-order laws are sufficient for esti-mating and modeling purposes:d cid t

= −k0exp (−EA/RT ) ci (18)

where ci is the molar (or mass) concentration ofspecies i in the gas phase.

Values of kinetic parameters for tar crackinggiven in Table 14 were determined with freshlyformed tar from bituminous coal [67].

Equations (14) – (16) and (18) can be usedfor describing source or sink terms in the bal-ance equations in reactor models (see Section4.2.3).

4.2. Classification and GeneralCharacteristics of Gasification Processes

4.2.1. Criteria for Classification

Processes for gasification of solid carbonaceousmaterialsmaybe classified in the followingways(the abbreviations given are used in subsequenttables that present information on individualgasification processes):

1) Method of contacting solid and gaseous reac-tants (type of gasification reactor)moving fixed bedfluidized bedentrained flowmolten bath

2) Kind of solid feedstockcoalwood – biomasssolid waste materialspetroleum-derived residues

3) Particle size of solid feedstockfinemediumcoarse

4) Gasifying mediumair (air)oxygen (ox)steam (st)steam – air (st/a)steam – oxygen (st/ox)hydrogen (hydr)

5) Method of supplying heat required by gasifi-cation reactionsinternal heating

autothermic (I/A)cyclic (I/C)heat-carrying fluids or solids (I/HC)

external heating (E)

Gas Production 57

heat transferred through walls of reactionvessel or heat exchanger (allothermic)

6) Operating pressureatmosphericelevated

7) Condition of gasification residue to be re-moveddry ash in nonslagging operationslag in slagging operation

8) Application of product gasfuel gas (medium or low Btu) (FG)fuel gas for combined cycle power plant(FG/ IGCC)synthesis gas (SG)substitute natural gas (SNG)source of hydrogen (H2)gas for direct reduction of iron (RG)

9) Developmental statuscommercially available processprocess in developmental stage

bench scalepilot-plant scaledemonstration scale

Obviously, these criteria are not independent,and others could be used instead. An exam-ple is a characteristic gasification temperature,which determines product gas yield and compo-sition, utility consumption, and the approach forpredicting product gas composition and processperformance (see Section 4.2.3).

For the purpose of this chapter, processes arepresented according to the method of bringingthe reactants into contact. This appears to bethe most appropriate way of identifying processprinciples and characteristics, and is the mostwidely applied and accepted practice. Figure 22shows examples of reactor types that representdifferent methods of contacting reactants.

4.2.2. Criteria for Process Assessment

Besides economics and availability, efficiencyand other process performance criteria charac-terize individual gasification processes and aidin their comparison and assessment. Some com-monly used criteria for practical purposes aredefined in the following material.

Efficiencies

(Lower heating values are often used insteadof higher heating values.)

Carbonconversion =1− carboningasificationresidue

carboninsolidfeedstock

Specific Gasification Rates

– Gas production rate per unit volume of gasi-fier

– Gas production rate per unit cross section ofgasifier

Specific Production and Consumption Fig-ures

– Gas (or byproduct) production per unit massof solid feedstock

– Utility consumption per unit mass of solidfeedstock

– Utility consumption per unit volume of prod-uct gas

4.2.3. Mathematical Modeling ofGasification Reactors

Mathematicalmodels have gained importance aspractical tools for designing gasification reactorsand improving their operation and control. Thegeneral aim is to predict and interpret gasifierperformance features such as

Product gas yield and compositionUtility consumption figuresEfficiencyResponse time to input changes

as affected by

1) Solid feedstock characteristics (e.g., proxi-mate and ultimate analysis, reactivity), and

2) Design and operational parameters (e.g., gasi-fier size, pressure, composition of gasifyingmedium).

58 Gas Production

Figure 22. Types of gasification reactor representing different methods of contacting reactants and operating temperaturesA) Moving bed (nonslagging, countercurrent); B) Fluidized bed; C) Entrained flow; D) Molten bath

Gas Production 59

The degree of detail and the complexity ofmathematical models depend on the particulartype of gasifier and the task to be addressed.Progress in computing and in efficient algo-rithms, aswell as the interest in coal gasification,led to numerous modeling activities and highersophistication of models during the 1970s andlater.

For gasification reactors, the following typesof models are most widely applied:

1) Blackbox models (empirical correlations),2) Equilibrium models, and3) Kinetic models.

Blackbox Models (Empirical Correla-tions). Experimental data form the basis forcorrelation of commercial-size gasifier perfor-mance with feedstock characteristics and opera-tional parameters. Such correlations are of greatpractical importance for all types of gasificationprocesses. They usually require little fundamen-tal knowledge and mathematical effort. How-ever, because they are not based on fundamentalprinciples, their usefulness for inter- or extrapo-lation is limited.

Equilibrium Models. Equilibrium modelsare valuable for identifying limits of conversionandproduct yields and for predicting reactor out-let conditions. This description is approximatebecause chemical equilibrium is usually reachedonly after long reaction times. Reaction ratesbecome slow when equilibrium is approached,leading to long residence times or large reactorvolumes to attain equilibrium conversion. Howwell an equilibrium approach can describe re-actor outlet conditions in a real system dependson the rates of chemical reactions, mass transfer,and convective flow of reactants.

Equilibrium models are well establishedfor high-temperature gasification systems(entrained-flow and molten-bath reactors),where chemical reactions are sufficiently rapidto approach thermodynamic equilibrium (e.g.,[69]).

In gasifiers where product gas compositionresults from reactions occurring at lower tem-perature (e.g., fluidized-bed or moving-bed re-actors), a defined deviation from chemical equi-librium can be incorporated by means of cor-rective factors for equilibrium constants or tem-

perature differences between actual and equilib-rium values [70, 71]. However, these restrictedequilibrium parameters generally cannot be cor-related with feedstock properties and operatingvariables on a fundamental basis. Therefore, ca-pabilities for extrapolation are limited.

Kinetic Models. Kinetic models offer thepossibility of describing local conditions in achemical reactor, thus representing the processthat eventually leads to reactor outlet conditions.Mathematical representation consists of differ-ential equations based on balance equations forindividual compounds and energy, including rateexpressions for chemical reactions (e.g., Eqs. 14– 18) and transfer of mass and heat. Because ki-netic models provide the most detailed descrip-tion of gasification reactors, they require a largenumber of parameters (activation energies, Ar-rhenius factors, mass- and heat-transfer parame-ters, etc.). Someparameters dependon feedstockcharacteristics and can be determined experi-mentally only with difficulty. Typical results ofkinetic models are temperature and componentconcentration profiles in axial and radial direc-tions. Examples of kinetic models can be foundfor moving-bed [66, 72], fluidized-bed [73, 74],and entrained-flow gasifiers [75, 76].

4.3. Characterization of SolidFeedstocks for Gasification

An overview of the molar ratios of carbon: hy-drogen: oxygen for solid feedstocks is presentedin Figure 23. Table 15 lists typical analyticaldata and Table 16 gives an overview of char-acteristic properties of different solid biomassandwastematerials. Some of themost importantproperties that characterize solid carbonaceousmaterials are discussed briefly in the following.The significance of these properties depends onthe type of gasification process.

Moisture Content. Solid carbonaceous ma-terials have moisture contents ranging from <5% (petroleum-derived residues and anthraciticcoals) up to about 50% (lignitic coals, peat,and wood).Moving-bed gasifiers can accommo-date the highest moisture content (ca. 35%) be-cause the feedstock is dried in the gasifier by

60 Gas Production

Table 15. Data from typical solid feedstocks for gasification

Delayed coke Anthracitic coal(Ruhr)

Bituminous coal(Pittsburgh)

Lignitic coal(North Dakota)

Wood(redwood)

Municipalsolidwaste

Analysis as-received, wt %C 82.6 88.6 75.4 48.7 35.8 47.6H 3.1 3.2 2.2 3.0 4.1 6.0O 0.2 3.3 8.1 14.6 29.7 32.9N 1.6 0.1 1.2 0.6 0.1 1.2S 3.4 0.9 2.2 0.2 0.1 0.3Ash 0.4 2.9 6.5 6.6 0.2 12Moisture 8.6 1.0 1.4 26.1 30.0 20V a 1000 30 b 30 b 30 b 15 n.a. c

Ni a 220 20 b 20 b 20 b 15 n.a. c

Molar ratioH/C 0.46 0.43 0.83 0.74 1.37 1.51O/C 0.002 0.03 0.08 0.22 0.62 0.58

Volatile matter, wt % 10.2 5.6 38.0 27.8 58.1 59.5Fixed carbon, wt % 80.8 90.5 54.1 39.5 11.7 8.5Reactivity e relative tobituminous coal 0.29 0.48 1.00 5.90 3.93 n.a. c

Ash softening tempera-ture e, ◦C 1155 1240 1260 1240 1150 n.a. c

Grindability, HGI f 50 – 100 50 60 60General characteristics wide distribution of

particle size, varieswith cuttingprocedure

particle-size distribution depends on mining – cutting andpreparation/beneficiation

a Parts per million (by weight).b Typical value for coal.c n.a. = not available.d Gasification rate in CO2 (standard conditions).e Oxidizing conditions.f HGI = hard growth index.

Table 16. Characteristic properties of nonconventional solid feedstocks for gasification

Reactivity Moisturecontent

Inorganicscontent

Inorganicscomposition

Inorganicsmelting temp.

Fixed carboncontent

Volatilescontent

Criticalcompounds

BiomassWood high high low Fe, Al, Si, Ca low low highWaste wood high low varying varying low low high insecticides

pesticidesEnergy crops high high moderate Si low low highAgriculturewaste

high low low to high Si, Na, K low low high

Sewage sludge high high high Si, Al, Fe, Ca,P

normal low high Hg, Cd, Pb,Cu, Cl, (N)

Waste materialsMunicipal waste high varying medium to

highSi, Ca, Al, Fe normal low high Cl, F, Zn, Pb,

Cu, Cr, S, HgHazardous wastevarying varying varying Si, Fe varying low varying Zn, Cu, Ba, SnUsed tires low low ZnPetroleum residuesCoke low low moderate low high low Ni, VHydrogenationresidue *

low low medium medium high low Ni, V

* Example: Veba Combi Cracking (using lignite coke as additive).

Gas Production 61

exiting product gas. Entrained-flow andmolten-bath gasifiers generally require that the mois-ture content be adjusted bypredrying tomaintainfree-flowing behavior during crushing, convey-ing, and charging to the gasifier. Fluidized-bedprocesses can cope with higher moisture con-tent which, however, in any case leads to lowerprocess efficiencies.

Figure 23. Molar ratios C−H−O of feedstocks forgasification

Ash Content. Ash is present in petroleum-derived residues and wood in very low amounts(< 1%) and in coals from 1 to> 50%. For gasi-fication, ash should be at a minimum becauseof the following reasons. High ash content isdetrimental to carbon conversion. It furthermoremust be heated to the gasifier outlet temperature,which leads to heat losses. In addition, provisionmust be made for withdrawing ash from the sys-tem,which increases in complexity and costwithincreasing operating pressure. An exception isthe dry ash moving-bed gasifier which can copewith high ash content (up to ca. 50%) and uti-lizes sensible heat in the exiting ash to preheatthe gasifying medium.

Ash fusion behavior is significant because insome gasification processes the ash must remainsolid for removal (moving-bed nonslagging andfluidized-bed processes), whereas in others it isremoved as a liquid slag (moving-bed slagging,entrained-flow, and molten-bath processes). Fu-sion temperature, therefore, must be thoroughly

taken into account in the design and operationof gasifiers. Ash composition must be consid-ered when selecting materials of construction,particularly for slagging gasifiers. Ash can varyin character from acidic to basic and can containelements such as iron, which may separate orrequire special handling. Also, in slagging oper-ation addition of reagents such as fluxes may bedesirable to improve flow characteristics of theslag or to adjust its chemical behavior.

Ash can be beneficial when particular con-stituents like potassium enhance the kinetics ofgasification, tar cracking, or other reactions.

Volatile Matter. Volatile matter refers togas, tar, oil, and phenolic compounds formed bypyrolysis reactions of the feedstock. The yield ofthese volatile products determines the remain-ing amount of char that must be gasified in (rel-atively slower) gas – solid reactions. Pyrolysisproduct distribution depends strongly on feed-stock characteristics and, to some extent, on re-action conditions during pyrolysis (temperature,heating rate, pressure, ambient gas composition,particle size [51]).

Fixed Carbon. The residual char remainingafter pyrolysis contains the fixed carbon thatmust be gasified in gas – solid reactions withsteam, hydrogen, carbon dioxide, air, or oxy-gen. Because these gas – solid reactions are en-dothermic for most gasification processes, thespecific heat requirement per unit mass of feed-stock (which in autothermic processes meansspecific oxygen or air consumption) depends onthe amount of fixed carbon in the feedstock. Ofthe different feedstocks presented in Table 15,wood and municipal solid waste have the low-est value of fixed carbon, whereas delayed coke,petroleum-derived residue from a high-severitycoking process, and anthracitic coal have thehighest values.

Caking Behavior. Caking and swelling oc-cur with certain coals, a phenomenon relatedto pyrolysis reactions. Certain bituminous coalswith a high content of volatile material becomefluid, and the particles may adhere together andswell. Other coals with similar volatile mat-ter content may retain the original shape, withindividual pieces caking together weakly. Stillother solid carbonaceous materials (e.g., lig-

62 Gas Production

nitic coals) can disintegrate during pyrolysis intosmall pieces. Accordingly, feedstocks are re-ferred to as strongly or weakly caking or non-caking. Fluidized-bedprocesses tend to be sensi-tive to strongly caking coals, whereas entrained-flow, molten-bath and moving-bed processes(with sophisticated mechanical stirrer systems)can generally accept feedstocks with cakingproperties.

Reactivity. Solid carbonaceous materialsand their residual char from pyrolysis exhibitmajor variations in their reactivity to steam, hy-drogen, carbon dioxide, and oxygen. This canbe attributed basically to three factors [77]:

1) Concentration of active sites (i.e., carbonatoms located at the edges of crystallites orbuilding blocks);

2) Diffusional limitations on how rapidly reac-tive gas molecules can reach active sites; and

3) Presence of catalytic inorganic impurities.

In technical gasification processes, reactivityaffects the temperature level at which the gasi-fier must be operated to ensure acceptable car-bon conversion. As a consequence, specific heatrequirements (i.e., oxygen or air consumptionin autothermic processes) are related to feed-stock reactivity, to an extent dependent on theparticular type of process. For example, low re-activity in combination with low ash fusion tem-perature may be prohibitive for dry ash removalgasification processes (e.g., fluidized-bed pro-cesses). Highly reactive feedstocks are particu-larly suited for methane production because lowoperating temperatures favor the formation ofmethane according to chemical equilibrium.

Of the feedstocks in Tables 15 and 16, wood,biomass waste, and lignitic coal are the most re-active, whereas petroleum-derived residues andanthracitic coal have the lowest reactivity.

Particle-Size Distribution. The particle-size distribution of a given carbonaceous ma-terial depends on its origin and characteristics:(1) system ofmining (for coals), (2) type of cok-ing process (for petroleum-derived residues),(3) kind of cutting (for wood), (4) kind of pre-treatment (for waste materials), (5) mechanicalstrength, and (6) method of preparation andtransport. For example, strip-mined bituminouscoal may contain up to 30% fines (< 3 mm)

whereas heavily mechanized deep-mined coalmay have 40% or more, and strip-mined lignitecan reach 60% [55].

In moving-bed gasifiers, relatively coarsefeedstock particles (3 – 50 mm) are preferred.The acceptance of fines is limited. Develop-ments have moved, however, toward acceptanceof a greater proportion of fines, either as agglom-erates or by direct injection of fines into the gasi-fier. Fluidized-bed, entrained-flow, and molten-bath processes generally require pretreatmentsteps such as drying, crushing, and grindingto achieve an optimum feedstock size distri-bution. Coal mining techniques that result inhigh fines production offer an advantage in thiscase, because they help reduce costs of the coalpreparation plant necessary for fine coal pro-cesses. Where a grinding step is needed, feed-stock grindability is an important property thataffects the kind and cost of equipment.

4.4. Moving- or Fixed-Bed Processes

In moving-bed gasifiers (also called fixed-bedgasifiers), the gasifying medium passes througha bed of granular or lump fuel. Figure 22 Ashows a nonslagging, countercurrent version,which exhibits excellent thermal efficiencies be-cause outgoing ash heats the incoming gases andoutgoing product gas heats the incoming solidfeedstock. The long residence times of solid par-ticles moving through the bed (typically 1 – 2h), together with the temperature profile of thecountercurrent system, allow high carbon con-version efficiencies.

Tables 17 and 18 contain information on se-lected moving-bed gasification processes andtypical performance data with coal and petro-leum-derived residues as feedstocks. Dry ashand slagging pressure gasifiers are shown as ex-amples in Figures 24 and 25.

In countercurrent moving-bed gasifiers, thesolid fuel descends through four not distinctlyseparate zones with varying temperatures (Fig.22 A) and gas compositions, where the chemicalreactions outlined in Figure 18 can occur.

Drying and Pyrolysis Zone. Raw feedstockcomes in contact with hot product gases, andmoisture is driven off. Subsequently, pyrolysisof the carbonaceous material occurs to yieldgaseous products and char.

Gas Production 63

Table 17. Selection of moving-bed gasification processes a

Process name British Gas – Lurgi(BGL) slagginggasifier

Lurgi pressuregasification

Wellman WellmanGalusha

Imbert Thermoselect

Licensor or contributor b British Gas Corp.,Lurgi

Lurgi WellmanThermalSystems

Dravo ImbertEnergietechnik

ThermoselectS.A.

Feedstock coal, petroleum coke,waste RDF c

coal, waste RDF ccoal coal wood municipal solidwaste

Gasifying medium st/ox st/ox, st/air st/air st/air, st/ox air (downdraft) oxMethod of heat supply I/A I/A I/A I/A I/A I/AOperating pressure, MPa elevated (2.0 – 3.0) elevated (2.0 –

10.0)atmospheric atmospheric atmospheric atmospheric

Condition of residue slag dry ash dry ash dry ash dry ash slagApplication of product gas FG, FG/IGCC, SG FG, FG/IGCC,

SGFG FG FG FG

Developmental status commercial commercial commercial commer-cial commercial commercialNumber of commercial-sizegas generators

2 > 165 > 40 > 150 very high 3

Reference [55, 78 – 80] [55, 83] [84] [78] [54, 87] [86]a For abbreviations, see Section 4.2.1.b in some cases historical.c RDF = Refuse-derived fuel.

Figure 24. Schematic of Lurgi dry ash pressure gasifiera) Coal lock; b) Drive; c) Scrubbing cooler; d) Stirrer;e) Water jacket; f) Ash lock

Gasification Zone. Char from the pyrolysiszone comes in contactwith hot combustionprod-ucts and steam from the zone directly below.Char – gas reactions occur mainly with steam,carbon dioxide, and to a lesser extent, hydrogen,so that the overall reaction is endothermic.

Combustion Zone. The combustion zonesupplies heat for the gasification zone directlyabove. The key reaction is that of carbon (inthe remaining char from gasification) with oxy-gen to produce heat and carbon oxides. Here,the temperature rises to a maximum, whichin nonslagging operation must be kept belowthe fusion point of the ash by supplying steamin excess of that required to gasify carbon. Inslagging operation, the steam supplied togetherwith oxygen only slightly exceeds that requiredto gasify carbon. This results in temperaturesof 1500 – 2000 ◦C, at which the ash melts anddrains off as a liquid slag.

Ash Zone. In dry ash removal gasifiers, theash bed, located at the bottom, heats the incom-ing gasifying medium by direct heat exchangeand acts as a distributor.

Lurgi Pressure Gasification Process.Awidely used version of the moving-bed gasi-fier is the Lurgi dry ash pressure gasifier (Fig.24), which has been used commercially sincethe 1930s. The main gasifier shell is surroundedby a water jacket in which process steam is pro-

64 Gas Production

Table 18. Typical performance data of moving-bed gasification processes, feedstocks: coal and petroleum-derived residues a

Process name British Gas – Lurgi Lurgi WellmanFeedstock bituminous coal

Pittsburghdelayed coke bituminous coal

South Africalignite NorthDakota

bituminous coalGermany

bituminous coalGermany

bituminous coal

Size of gasifier demonstration demonstration commercial commercial commercial pilot commercialVolatile matter,wt %

34.1 8.9 20.6 26.9 31.7 23.8 32.0

Fixed carbon,wt %

50.2 85.7 50.9 30.6 52.6 38.6 49.0

Moisture, wt % 5.0 3.6 10.0 36.3 3.0 5.0 7.0Ash, wt % 10.7 1.8 18.5 6.3 12.7 32.6 12.0C, wt % (daf b) 83.6 92.6 80.6 73.9 82.3 83.6 75.1H, wt % (daf) 5.7 3.8 4.0 5.2 5.1 6.0 5.9O, wt % (daf) 6.8 0.2 13.0 18.9 9.4 6.8 16.5N, wt % (daf) 1.6 1.4 2.1 1.0 1.3 1.1 1.3S, wt % (daf) 2.0 2.0 0.2 1.0 1.8 2.3 1.1Molar ratioH/C 0.818 0.492 0.596 0.844 0.744 0.861 0.943O/C 0.061 0.002 0.121 0.192 0.086 0.061 0.165

Fischer assay tar,wt % (daf) b 17.8 1.2 5.5 8.9 14.8

Feed throughput,t/h (daf) b 10.5 8 23.8 25.1 13.0 6.0 2.5

Reactordiameter, m

1.8 1.8 3.8 3.8 3.4 1.5 3.0

Operatingpressure, MPa

2.5 2.5 2.8 3.1 2.0 9.0 atmospheric

Gasifyingmedium

st/ox st/ox st/ox st/ox st/air st/ox st/air

Ratio of steam toO2

1.0 1.1 5.8 6.9 7.0

Ratio of steam toair

0.4 0.2

Crude gascomposi-tion, vol%CO2 3.0 1.0 27.8 32.0 10.0 30.2 4.0CO 56.4 60.0 23.0 15.8 21.5 17.5 29.0H2 28.0 30.0 28.6 39.8 20.0 35.0 16.0CH4 7.1 3.0 9.1 11.8 3.5 16.5 2.0CnHm 0.4 0.2 0.4 0.8 0.8 0.8N2 4.2 5.4 0.4 44.2 49.0

Crude gas yield,m3 (STP)/t

(daf b)2190 2640 2381 1980 3800 2075 4300

Tar – oil yield,kg/t (daf b) 42.0 63.7 33

Oxygenconsumption,m3/m3 (STP) 0.20 0.17 0.14 0.13 0.18

Airconsumption,m3/m3 (STP) 0.56

Cold gasefficiency, c %

85.1 88.1 80.9 86.8 76.3 81.9 ca. 80

Reference [88] [89] [90], [91] [92] [93] [84]a For abbreviations, see Section 4.2.1.b daf = dry and ash free.c For definition, see Section 4.2.2.

duced. A coal lock hopper is mounted on top ofthe gasifier to feed the coal (typically 3 – 50-mmparticles), and a motor-driven distributor is usedto spread incoming coal evenly over the coal bed.

Amechanical stirrer is included in some designsto allow the gasifier to handle caking coals. Amotor-driven grate at the bottom of the gasifieris used to withdraw the ash formed, which drops

Gas Production 65

Figure 25. Schematic of British Gas – Lurgi slagging gasifier (syngas production from solid waste materials, based on [80])a) Feed preparation (metal removal, drying, pelletizing); b) Gas purification and conditioning; c) Clean gas utilization (meth-anol synthesis, combined cycle power generation); d) Air separation

into the ash lock hopper. Steam and oxygen (orair) are introduced at the bottom of the gasifierand distributed through the coal bed via the ro-tating grate. The grate supports the coal bed andis continuously rotated to ensure constant, evenwithdrawal of ash. Raw product gas leaves fromthe top of the gasifier (typically at 400 – 600 ◦C)and flows through a scrubber – cooler where it iswashed by circulating gas liquor. Further cool-ing, cleaning, and conditioning of the gas mustbe done according to the desired application.

In the British Gas – Lurgi Slagging Gasifier,a slagging version of the Lurgi gasifier, the ashmelts as a eutectic with potentially added flux toform slag. The high temperature in the combus-tion zone is a result of low steam oxygen ratiosin the gasifying medium. Molten slag is tappeddiscontinuously through a taphole into a quenchvessel, where the slag granulates as it drips intothe quenchwater.One application of this processis the production of synthesis gas from variouswaste materials, as shown in Figure 25.

Compared to the dry ash process, the slag-ging version has higher capacities per unit vol-ume (see Fig. 32), has lower hydrogen carbonmonoxide ratios, andproduces less aqueous con-densate, tar, phenols, and hydrocarbons. Thesecan be recycled and destroyed by gasification.

For gasification of wood, most moving-bedprocesses apply the cocurrent (or downdraft)system in which product gas flows through ahot zone before leaving the gasifier (Fig. 26).The advantage of this system is the low tar con-

tent of the product gas, because most of the con-densable pyrolysis products are cracked whilepassing through the combustion zone. A highertar content in the product gas does not causea problem when a gas turbine is situated closeto the gasifier since tars are kept in the vaporphase. An example of a wood gasification pro-cess is given in Table 17. Moving-bed gasifiersfor wood gasification are suited to small-scaleapplication (generally less than 200 kg of woodper hour). Performance data for a moving-beddowndraft wood gasifier are given in Table 21.

Figure 26. Types of moving bed gasifier for wood gasifica-tionA) Countercurrent updraft; B) Cocurrent downdraft

For gasification of municipal waste togetherwith other waste materials, practical experienceis available with pressurized moving-bed gasi-fiers andwith the newThermoselect process (see

66 Gas Production

Table 21). In the latter case, rawmunicipalwasteis pyrolyzed before being gasified under slag-ging conditions. In the case of moving bed gasi-fication, pretreatment of the waste feedstock isrequired (extrusion/pelletizing, see Figure 25).Slagging operation ensures vitrification of theresidue containing the inorganic material in thefeedstock.

Some advantages of moving-bed gasifiersare: (1) the technology is mature with manycommercial designs available, including pres-sure operation; (2) the large fuel inventory pro-vides safety, reliability, and stability; (3) the dryash gasifier can cope with high ash content feed-stocks; (4) the gasifier may be banked for longperiods; (5) the carbon conversion efficiency ishigh; (6) the thermal efficiency is high becauseof countercurrent flow and thus the specific oxy-gen consumption is low (see Fig. 33); and (7) thecapability for turndown is high.

Some limitations of moving-bed gasi-fiers include (1) internal moving parts withsome mechanical complexity are employed;(2) gasifier capacity is limited by gas flow rates(see Fig. 32); (3) feedstock fines must be han-dled separately (e.g., via agglomeration); (4) theproduct gas contains tar, oil, phenols, and ammo-nia, which require separation and cleanup sys-tems; and (5) excess steam for temperature con-trol leads to thermal losses and requires specialcondensate treatment.

Besides the processesmentioned inTables 17and 18, other moving-bed processes exist, someof which are of historical interest only: BCURA,Chevron, Demag, Foster Wheeler, Gegas, GasIntegrale, Grand Forks, Kellogg, KGN [81], Kil-ngas [82], Leuna, Morgas, Otto, Pintsch, Riley– Morgan, Wilputte, Woodall Duckham for coalgasification [55, 58, 78], and various processesfor wood gasification [85, 94].

4.5. Fluidized-Bed Processes

Fluidized beds for gasification of solids arecharacterized by linear velocities of gasifyingmedium sufficient to lift the solid particles. Thisrequires smaller particle sizes than moving-bedprocesses, typically in the 0.1 – 5-mm range(Fig. 22 B). As a result of fluidization, the solidparticles move about randomly so that mixingof solids within the bed approaches that of a

perfectly mixed reactor. Therefore, the solidscomposition and the temperature of the bed arenearly uniform, bed temperature being chosenaccording to feedstock reactivity. Themaximumbed temperature is limited by the ash softeningand sticking behavior. The temperature of reac-tor exit gases is about the same as that of thebed; a heat recovery device is required for en-ergy economy.

Tables 19, 20 and 21 contain informationon selected fluidized-bed gasification processesand typical performance data with various feed-stocks. The number of processes in these tablesreflects the fact that fluidized beds have beenfavored in many recent process developmentprojects, mainly because they combine some ad-vantages of moving-bed (e.g., large fuel inven-tory) and entrained-flow (e.g., control of hydro-carbon formation) gasifiers. In addition to theprocesses listed, others combine fluidized bedswith various contacting methods, such as dilutephase entrainment ormoving bed.An example istheHygas gasification process. In Tables 19 and20 the DMT process is listed as an example foran allothermic process, with heat from an exter-nal source (e.g., from a nuclear reactor or fromhot flue gas) introduced via heat exchange bun-dles in the fluidized bed. These data are includedalthough presently there is little commercial in-terest in the process.

Fluidized beds with gas velocities far abovethe minimum fluidizing velocity (3 – 10 m/s),called circulating or fast fluidized beds [95], arealso used for gasification. They operate at lowerbeddensities and can thus copewith caking coalswithout particle agglomeration or sticking.

Winkler Gasification Process. The Winklergasifier, the first commercially applied fluidized-bed process, has been in service since the 1920sfor producing fuel and synthesis gases from coal(Fig. 27). The coal feed inlet is located in thelower part of the gasifier. Variable-speed screwsfeed crushed coal (< 9 mm) into the gasifier.Steam and oxygen (or air) are fed through noz-zles located at several levels in the fluidized bedwhich occupies only the lower third of the gasi-fier; the remainder serves as a disengaging zone.Secondary steam and oxygen injection occurabove the bed level to gasify unconverted carbonleaving the bed, tars, and hydrocarbons. Heavierparticles drop through the fluidized bed and passinto the ash discharge unit at the gasifier bottom,

Gas Production 67

Table 19. Selection of fluidized-bed gasification processes; feedstocks: coal and petroleum-derived residues, wood, and biomass waste *

Process name DMT(catalytic)

Foster Wheeler(Ahlstrom)

DMT HightemperatureWinkler

Lurgi CFB U-gas Winkler

Licensor or contributor ** DMT DMT Rheinbraun Lurgi IGT Davy McKeeFeedstock coal wood chips agriculture

forestry wastelignite, browncoal

coal, wood,waste

coal coal

Gasifying medium st air st st/air, st/ox st/air, st/ox st/air, st/ox st/air, st/oxMethod of heat supply E I/A E I/A I/A I/A I/AOperating pressure, MPa 4.0 2.0 0.3 – 2.0 1.0 – 2.5 < 0.2 < 3.5 < 0.4Condition of residue dry ash dry ash dry ash dry ash dry ash dry ash (ag-

glomerated)dry ash

Application of product gas SG FG (GT) FG (GT, GM,FC)

FG, FG/IGCC,SG

FG, FG/IGCC,SG

FG, FG/IGCCFG,SG,H2

Developmental status development commercial development commercial commercial commercial commercialNumber of commercial-size gasgenerators

≥ 1 2 3 ≥ 8 ca. 70

References [96, 97] [98] [99] [78, 102, 108] [89, 103, 104,109]

[55, 78, 85,105]

[55, 78, 110]

* For abbreviations, see Section 4.2.1.** In some cases historical.

whereas lighter particles are carried up throughthe bed with product gas. Coal residence timeis 20 – 30 min, based on coal feed rate. Carbonconversion is limited by feedstock reactivity andcontinuous loss of carbon with the ash.

Figure 27. Schematic of Winkler gasifiera) Coal screw; B) Ash screw

A modification of the atmospheric Winklergasification process has been developed (high-temperature Winkler process, HTW): It oper-

ates at elevated pressure (up to 1.0 MPa) andslightly increased temperature, i.e., conditionsthat considerably improveprocess efficiency andgas quality with respect to the requirements of asynthesis gas.

Figure 28. Schematic of U-Gas gasifier

U-Gas Gasification Process. A pressurizedversion of a fluidized-bed gasifier with con-trolled ash agglomeration is shown in Figure 28.The U-gas gasifier is a vertical cylinder with twocyclones for returning elutriated fines to the bed.A sloping grid at the bottom serves as the air andsteam distributor and the agglomerated ash out-let. Noncaking coals can be fed directly to thegasifier from the crusher (particle size< 6 mm).Caking coals may require pretreatment by con-

68 Gas Production

Table 20. Typical performance data of fluidized-bed gasification processes; feedstocks: coal and peat a

Process name DMT High-temperature Winkler U gas WinklerFeedstock bituminous brown coal peat bituminous lignite lignite

coal coalSize of gasifier pilot commercial commercial pilot, commercial commercial commercialVolatile matter, wt % 12.2 44 – 46 57 – 61Fixed carbon, wt % 68.2 37 – 40 20 – 24Moisture, wt % 1.2 12 (10 – 25) 15 (12 – 18) 8.0 8.0Ash, wt % 18.4 4 (2 – 18) 4 (3.5 – 4.5)C, wt %, (daf e) 86.9 69 60 71.1 71.2H, wt % 4.0 5.0 5.0 5.5 4.8O, wt % 4.9 25 31 18.9 20.2N, wt % 2.0 0.9 2.8 0.9 2.2S, wt % 2.2 0.1 0.2 3.8 1.6Molar ratioH/C 0.559 0.870 1.200 0.920 0.818O/C 0.042 0.272 0.387 0.199 0.213

Feed throughput,t/h (daf e) 22 21 1 – 5 29 41

Reactor diameter, m 2.71 2.4 6 6Operating pressure,MPa 4.0 1.0 1.0 1.6 – 2.4 0.105 0.105

Gasificationtemperature, ◦C 830 950 950 ca. 1000 815 – 1200 815 – 1200

Gasifying medium st st/ox st/ox c st/air st/air st/oxRatio of steam to O2 1.43Ratio of steam to air ca. 0.3 0.06Crude gas composition, vol%

CO2 22.3 22.8 21.3 9.9 7.7 19.5CO 13.7 35.7 31 19.6 22.5 36.0H2 56.2 35.9 29.7 17.5 12.6 40.0CH4 7.8 4.9 5.6 3.4 0.7 2.5CnHm 0.1 0.1N2 0.6 12.3 48.9 55.7 1.7

Crude gas yield,m3 (STP)/t (daf e) 2000 2100 3596 3058 1760

Tar – oil yield, kg/t (daf e)Oxygen consumption,m3/m3 (STP) 0.19 0.19 0.14 – 0.24

Air consumption,m3/m3 (STP) 0.43 – 0.76

Cold gas efficiency, d % 80 73 71 – 79 61.9 74.4Carbon conversion, d % 95 94 92 96 – 98 83.0 81.0Reference [96, 97] [102] [102] [105] [55, 56]a For abbreviation, see Section 4.2.1.b Predried.c Gasification with oxygen-enriched air.d For definition, see Section 4.2.2.e daf = dry and ash free.

tact with air in a second fluidized bed operatingat gasifier pressure and 370 – 430 ◦C. Here, anoxidized outer layer forms on the coal particlesand prevents agglomeration and possible block-age in the gasifier.

Coal is gasified with a mixture of air (or oxy-gen) and steam in a single-stage fluidized bed.Residence time of the particles is ca. 45 – 60min. The fluidizing velocity is about 1 m/s. Thekey to gasifier operation is the agglomerationand separation of low-carbon ash from the bed.

This is accomplished by keeping the tempera-ture near the gasifying medium inlet close to thesoftening point of the ash particles. Thus, ashparticles preferentially stick together, and the ag-glomerates grow until they are heavy enough tomove downward and fall out of the fluidized bed.

The fluidized-bed system is particularly at-tractive for gasification of wood and otherbiomass because it allows greater feedstock flex-ibility (e.g., with respect to particle size) andeasier scaleup thanmoving-bed gasifiers. There-

Gas Production 69

Table 21. Typical performance data of wood, biomass, and waste gasification processes a

Process name Imbert Lurgi/SVZ FosterWheelerAhlstrom

Lurgi CFB Thermoselect

System moving beddowndraft

grate gasifier fluidized bed fluidized bed moving bed

Feedstock wood RDF/variouswaste/coalmix

wood chips wood bark wood waste lignitecombustionresidue

municipalsolid waste b

Size of gasifier commercial commercial commercial commercial commercial

Volatile matter, wt % (70)Fixed carbon, wt % (15)Moisture, wt % 10 – 20 8 – 10 19 10 20 – 30Ash, wt % 2 – 5 2 78 20 – 30Feed throughput, t/h (daf e) ca. 10 3.5 – 4.5 mix mix 1.5 – 5

(70 – 100 MWth)Reactor diameter, m 3.6Operating pressure, MPa atmospheric 2.4 1.8 atmospheric atmospheric atmospheric

Operating temperature, ◦C f f 950 – 1000 ca. 800 900 – 950 600 – 1200

Gasifying medium air st/ox air air air ox

Ratio of steam to O2 – – – 0.0Ratio of steam to air 0.0 0.0 0.0 0.0 –

Crude gas composition, vol%CO2 10.0 31 – 35 12 15 22 – 27CO 20.0 15 – 22 16 15.4 35 – 39H2 19.0 29 – 35 13 14.8 32 – 36CH4 1.5 9 – 13 3 4.2 < 0.1CnHm 2.0 < 0.1N2 50.0 44 39.6 3 – 4

Crude gas yield, m3 (STP)/kg (daf e) 2000 – 3000 1588 – 1926Oxygen consumption, m3/m3 (STP) – 0.45 – 0.52 c

Air consumption, m3/m3 (STP) (0.48) g –Cold gas efficiency, d % > 70Carbon conversion, d % 93 > 96Reference [115] [80] [98] [116] [103] [86]a For abbreviations, see Section 4.2.1.b Without pretreatment.c With natural gas as additional fuel (ca. 135 m3 (STP)/t daf).d For definition, see Section 4.2.2.e daf = dry and ash free.f Temperature distribution.g Estimated value.

fore, several fluidized-bed concepts have beenproposed and developed; a selection is presentedin Table 21. As an example, Figure 29 showsa flow diagram for fuel gas production frombiomass and waste feedstocks in a circulatingfluidized-bed reactor integrated in a cement pro-duction plant. The feed material is fed to thelower part of the gasifier via a screw feeder. Airas gasifying agent is injected at the bottom. Be-cause of excellent heat-transfer conditions, thetemperature is uniform throughout the reactorand is high enough to convert the tarry and oxy-genated pyrolysis products. Typical plant capac-ities for fluidized-bed wood and waste gasifiersare in the range of several tons per hour. Theproduct gas is particularly suitable as feedstock

for boilers, lime kilns, gas engines, and gas tur-bines.

Advantages of fluidized-bed gasifiers are: (1)commercial designs are available; (2) the tech-nology is less complex than that of moving bedsand does not involve moving parts; (3) the largefuel inventory provides safety, reliability, andstability; (4) a variety of particle sizes can behandled; (5) the amount of tar and phenols islow; (6) gas composition is steady due to uni-form conditions in the bed; and (7) moderategasification temperatures can be used.

Limitations of fluidized bed-gasifiers includethe following: (1) capacity flexibility is lim-ited by entrainment at high gas velocities;(2) the inventory of solid carbon is lowered

70 Gas Production

Figure 29. Flow diagram of a circulating fluid bed process for the production of fuel gas from biomass and waste feedstocks,integrated in a cement production plant, based on [103].

by the high content of inert material in thebed; (3) caking coal may require pretreatment;(4) conflicting temperature requirements existfor low-reactivity feedstocks with low-softeningpoint ash; (5) inevitable loss of carbon in ashoccurs due to uniform solids composition of thebed; (6) the turndown range is limited by the gasvelocity required to maintain fluidization.

Besides the processesmentioned inTables 1920 21, other fluidized-bed processes are BASF– Flesch – Demag, Battelle – Union Carbide,Hydrogasification, BCR, CGT, CO2-Acceptor,Electrofluidic process, Exxon catalytic pro-cess, Hydrane, Hydrogasification (HKV), Hy-gas, KRW, Synthane, Tosco, University ofWyoming, West Virginia University, Winkler –Flesch for coal or petroleum coke [55, 58, 117,119], IGT’s RENUGAS and various other pro-cesses for wood and waste gasification [54, 94,118].

4.6. Entrained-Flow Processes

In entrained-flow gasifiers, solid particles arecarried or entrained by the reacting gases. Thus,solids and gases move in the same direction withapproximately the same velocity. To achievethis, the particles must be smaller than in othersystems (typically 70% < 0.1 mm). This, to-gether with high temperature (see Fig. 22 C),

allows gasification rates high enough to ensureacceptable carbon conversion during the shortsolids residence time in the gasifier (a few sec-onds).

The reactants (e.g., coal, oxygen, and steam)are typically introduced into the gasifier at highvelocity through one or more burners or noz-zles. The burners can be oriented in many dif-ferent ways, for instance, tangentially, radiallyopposed, and axially. Gasifier performance is af-fected to a large degree by the flow characteris-tics and mixing efficiency of the burners. Thesolid feedstock is either injected in the form ofa water slurry or as dry particles.

Flame temperatures at the burner outlets canbe as high as 2000 ◦C. This results in extremelyrapid conversion of feedstock particles and de-struction of virtually all higher hydrocarbons.Outside the immediate flame region, heat losses,further mixing with steam, and endothermic re-actions combine to lower the gas-phase temper-ature. Because of the high reaction temperature,oxygen consumption is usually greater than forother gasification systems (see Fig. 33). A sig-nificant portion of the feedstock ash melts be-cause of the high reaction temperature and isremoved from the reaction zone as liquid slag.

Tables 22 and 23 contain information on se-lected entrained-flow gasification processes andtypical performancedatawith coal or petroleum-derived residues as feedstock. Figures 30 and 31

Gas Production 71

show, as examples, a Koppers – Totzek gasifierand a schemeof theTexaco coal gasification pro-cess with gas cooler mode.

Figure 30. Schematic of Koppers – Totzek gasifiera) Gasifier; b) Waste heat boiler; c) Slag quench tank

Koppers – Totzek Gasification Process. Thetwo-headed Koppers – Totzek gasifier (Fig. 30)is horizontal and ellipsoidal in shape, with twoheads mounted on the ends. A waste-heat boileris mounted on top to recover heat from the hoteffluent gases. The usual waste-heat boiler sys-tem, which produces high-pressure steam as abyproduct of the process, is a radiant surfaceboiler followed by a fire tube boiler. The gasifiershell is protected from the high gasification tem-peratures by a castable refractory lining. Coalis dried to 2 – 8% moisture content (dependingon coal rank) and simultaneously pulverized to70 – 90%below 80 µm.A variable-speed screwfeeder feeds the pulverized coal into a mixingnozzle, where a premixed stream of oxygen andsteam is introduced. Steam and oxygen entrainthe pulverized coal, and the coal – steam – oxy-gen mixture is injected into the gasifier burn-ers at injection speeds above the speed of flamepropagation to prevent flashback. Two adjacentburners on each head improve turbulence be-havior and provide a safety feature to ensure

continuous ignition if one burner is temporarilyblocked. Coal particles are gasified in approx-imately 1 s or less, producing a temperature inthe flame zone of ca. 1900 ◦C. Heat losses andendothermic reactions cool the gasmixture to ca.1300 ◦C. Ash in the coal is liquefied because thegasifier temperature is maintained above the ashfusion temperature. Most of the ash flows downthe gasifier walls asmolten slag and drains into aslag quench tank. Part of the ash leaves the gasi-fier as fine slag particles entrained in the exitgas. Pressurized modifications of the Koppers– Totzek process are the Prenflo and the Shellgasification processes (the latter jointly devel-oped by Shell andKoppers), both using dry solidparticles as feed.

TTexaco Gasification Process. The exacocoal gasification process (based on experiencewith oil gasification; see Section 3.2.2) uses apressurized version of an entrained-flow reactorand a slurry system (Fig. 31). The gasifier isa vertical, cylindrical vessel with a refractory-lined upper portion where solid feedstock reactswith steam and oxygen. Hot product gas iscooled upon leaving the gasifier, either by di-rect quenching with water, resulting in internalsteam production, or by indirect cooling in gascoolers with external high-pressure steam gen-eration. The mode of operation is determinedby the method of downstream processing andby the end use of the product gas. The quenchmode is generally selected for the productionof hydrogen or ammonia; carbon monoxide inthe raw product gas reacts with steam in a car-bon monoxide shift step to produce additionalhydrogen. The gas cooler mode of operation isselected for applications that do not require car-bon monoxide conversion (e.g., syngas for oxosynthesis or clean fuel for power generation).The solid gasifier feed is processed through twostages of dry grinding (the second under in-ert atmosphere) and pneumatically conveyed toslurry mix tanks where it is mixed with waterto form a pumpable slurry (slurry solids loading50 – 70%). The slurry is fed continuously to thetop of the gasifier by a positive displacementpump.

Advantages of Entrained-Flow Gasifiers.Among the advantages of entrained-flow gasifi-cation processes are: (1) commercial designs areavailable; (2) the gasifier has no moving partsand a simpler geometry than a fluidized bed;

72 Gas Production

Table 22. Selection of entrained-flow gasification processes *

Process name Koppers – Totzek Prenflo Shell Texaco DesTec (DW)Licensor or contributor Krupp – Koppers

UhdeKrupp – Koppers UhdeShell Texaco Dynegy

Feedstock coal coal coal coal coalpetroleum coke(dry)

petroleum coke (dry) petroleum coke (dry) petroleum coke(slurry)

petroleum coke,(slurry)

Gasifying medium ox (st/ox) ox (st/ox) st/ox st/ox oxMethod of heat supply I/A I/A I/A I/A I/AOperating pressure, MPa atmospheric 3.0 2.0 – 4.0 < 8.5 < 1.5Condition of residue slag slag slag slag slagApplication of productgas FG,SG,H2 FG,SG,H2,FG/IGCC FG,SG,RG,FG/IGCC,H2 FG,SG,H2,FG/IGCC FG,SG,H2, FG/IGCC

Developmental status commercial developmentcommercial

development commercial commercial commercial

Number of commercial-size gas generators ca. 80 1 ≥ 1 ≥ 5 ≥ 2

Reference [55, 78, 120] [78, 120, 124] [78, 121, 122, 125] [55, 78, 126 – 129] [123]

* For abbreviations, see Section 4.2.1.

Figure 31. Schematic of Texaco coal gasification process with gas-cooler modea)Mill; b) Slurry tank; c) Slurry pump; d) Gasifier; e) Radiation cooler; f) Lock hopper; g) Convection cooler; h) Gas scrubber;i) Clarifier; j) Slag screen

(3) the gasifier has the highest capacity per unitvolume (see Fig. 32); (4) any type of coal maybe used without pretreatment; (5) no fines arerejected; (6) the product gas is free of tar andphenols (see Fig. 33); and (6) the slagged ashproduced is inert and has a low carbon content.

Limitations of entrained-flow gasifiers in-clude: (1) combustor nozzles and heat recoveryin the presence of molten slag are critical de-sign areas; (2) a small solids inventory requiresadvanced control techniques to ensure safe, re-liable operation; (3) pulverizing and drying ofsurface moisture are required; and (4) the highgasification temperature causes thermal losses,

necessitates higher quality construction materi-als, and leads to increased oxygen consumption(see Fig. 33).

Besides the processesmentioned inTables 22and 23, other entrained-flow processes are Bab-cock Wilcox, BiGas, Combustion Engineering,Foster Wheeler, GSP, and ORC for coal or pe-troleum coke gasification [55, 131].

4.7. Molten-Bath Processes

In molten-bath gasification processes, solidfeedstock and gasifying agents are injected into

Gas Production 73

Table 23. Typical performance data of entrained-flow gasification processes; feedstocks: coal and petroleum-derived residues a

Process name Koppers – Prenflo Shell TexacoTotzek

Feedstock coal mixture ofcoal/petroleumcoke

coal bituminouscoal (Illinoisno. 6)

Texas lignite bituminouscoal (Illinoisno. 6)

delayed coke

Size of gasifier commercial commercial pilot pilot g pilot g pilot h pilotVolatile matter, wt % 18.5 33 34 31 33.5Fixed carbon, wt % 56 52 41 27Moisture, wt % 1.0 b 9.4 6.0 16 c 33 c 13.0Ash, wt % 20.2 24.5 8.8 9 9 11.1 0.5C, wt % (daf f) 79.7 84.7 81.5 77.0 72.0 78.1 90.7H2, wt % 4.6 4.0 5.4 6.0 6.0 5.1 3.9O2, wt % 13.5 4.7 10.1 11.0 19.0 11.6 0.9N2, wt % 1.8 1.9 1.9 1.4 1.2 1.3 2.6S, wt % 0.5 4.7 1.1 4.3 1.0 3.8 1.9Molar ratioH/C 0.688 0.935 1.000 0.784 0.516O/C 0.127 0.107 0.198 0.111 0.007

Feed throughput,t/h (daf f) 71 1.7 5.7

Operating pressure, MPa i 2.5 2.4 2.1 – 2.8Gasification tempera-ture, ◦C 1870 > 2000 > 2000 > 1250

Gasifying medium st/ox st/ox st/ox st/ox ox ox oxRatio of steam to O2 0.196 0.14 0.04 0.080 0.0 1.0 d 0.75 d

Crude gas composition(dry), vol%CO2 6.9 3.8 1.9 2.0 4.6 20.7 13.3CO 64.3 61.9 58.4 65.1 61.8 41.1 53.4H2 27.1 22.5 26.2 30.5 32.4 36.6 31.8CH4 0.1 0.1 0.1 0.1 0.02CnHm

N2 0.9 11.7 13.4 0.9 0.9 0.4 0.95Crude gas yield,m3 (STP)/t (daf f) 2578 2230 2109 2029 2192 2667

Tar – oil yield,kg/t (daf f)

Oxygen consumption,m3/m3 (STP) 0.30 0.30 0.297 0.294 0.321 0.304

Cold gas efficiency, e % 69.3 78.4 81.8 81 81 70.3Carbon conversion, e % 90.0 99.3 99.3 99 94Reference [55] [124, 130] [130] [122] [122] [128] [126]a For abbreviations, see Section 4.2.1.b Predried.c Before drying.d H2O as water in slurry.e For definition, see Section 4.2.2.f daf = dry and ash free.g Commercial-scale data in [125].h Commercial-scale data in [129].i Exact pressure not always available, see also Table 22.

a hot liquid bath of molten slag, metal, or salt(Fig. 22 D). The high temperature normally re-quired to keep the bath molten promotes highreaction rates, which are often enhanced by cat-alytic properties of the melt. The material em-ployed for the melt largely determines the char-acteristics of each particular process. Slag actsmainly as a heat-transfer medium but can some-

times have a catalytic effect. Iron is used as abath material because of its strong affinity forsulfur, its high solubility of carbon in molteniron, and because any iron oxide present in themelt will donate oxygen.Molten salts, which aregenerally less corrosive and have lower meltingpoints thanmoltenmetals,may strongly catalyzethe basic char – steam reaction, permitting prac-

74 Gas Production

tically complete gasification of some feedstocksat relatively low temperature.

As an example, the Molten Iron Puregas(MIP) Gasification Process uses an iron inven-tory of 20 t for a coal throughput of 10 t/h. Thereactor is designed for gasification pressure upto 0.3 MPa and is brick lined. It contains liquidiron into which the feed coal is injected via aninert carrier gas, as well as facilities for feedingoxygen and other media that may be requiredsuch as lime for adjusting the basicity of theslag and steam for cooling the iron. Slag float-ing on the surface of the iron bath is dischargedcontinuously and subjected to secondary treat-ment. Product gas leaves the reactor at ca. 1400 –1500 ◦C and passes through a waste-heat recov-ery system followed by a dust removal system.Pilot-plant design enables optimum location offeeding points, either at the top (top blowing) orat the bottom (bottom blowing) of the reactor,to be investigated. This process draws on expe-rience available from steelmaking technology.

Advantages of molten-bath gasifiers Amongthe advantages of molten-bath gasifiers are(1) the large fuel or heat inventory providessafety, reliability, and stability; (2) the high-tem-perature bath ensures safe ignition; (3) any gradeof coal in a wide size range can be utilized (i.e.,crushed run-of-mine coal); (4) no pretreatmentor pulverization is required for gasifier opera-tion; and (5) product gas contains no sulfur (re-tained in molten bath) and no tar or phenols athigh temperature.

Limitations of molten-bath gasifiers includethe following: (1) cleanup of molten media (ifrequired) is complicated; (2) capacity flexibilityis limited bymelt entrainment in the product gasat high flow rates; (3) ash is removed as liquidslag, leading to loss of sensible heat; (4) the cata-lyst (if required) must be separated from ash forrecycle, and make-up catalyst must be added;(5) the melt may be highly corrosive to refrac-tory at the gasifier temperature; and (6) bottomblowing gasifiers require sophisticated coolingsystems for the tuyeres.

Process variations include besides theMoltenIron Puregas Process (MIP) other processes forcoal or petroleum gasification: Rummel [61],Rockgas Molten Salt [132, 133], Saarberg/Otto[78], Atgas, Kellogg, Klockner, Sun [55, 134].

Figure 32. Effect of pressure and type of gasification reac-tor on capacity per unit volume, based on published data,feedstock: various coalsa) Moving bed dry ash; b) Moving bed slagging;c) Fluidized bed; d) Entrained flow; e) Molten bath iron

4.8. Underground Coal Gasification[135 – 138]

The purpose of coal gasification in situ is to re-cover the energy or chemical values of the coalwithout mining. Gaseous products may be uti-lized for electric power generation, manufactureof organic chemicals, and synthesis of liquid orgaseous fuels. Advantages of in situ coal gasifi-cation are the elimination of underground min-ing operations and, possibly, the utilization ofcoal in beds that cannot be mined profitably. In-cluded among past attempts are projects in Eu-rope, the United States, and the Soviet Union.Although details vary from project to project,mixtures of oxygen and steam (or air and steam)must always be admitted into a hot zone in thecoal seam for gasification to occur, and rawprod-uct gas must be withdrawn from the coal seamfor surface processing to make the gas environ-mentally and technically acceptable. The maindifficulties in underground coal gasification arecreating andmaintaining adequate permeability,maintaining and controlling the reaction zone,minimizing the detrimental effects of fissuringand channeling, and maintaining steady product

Gas Production 75

Figure 33. Characteristic temperatures of different gasification systems and correlation with selected performance data, basedon published dataa) Phenols; b) NH3; c) CH4; d) O2-consumption

gas composition and yield.Despite past projects,underground gasification of coal is not usedwidely at present and is not expected to playa significant role in the near future.

4.9. Environmental Aspects ofGasification

Gasification offers a practical means of utilizingcoal, petroleum-derived residues, wood, or solidwaste material while at the same time meetingstringent environmental control requirements.Although the type and severity of environmentalcontrol problem vary with gasification process,routine commercial operations have demon-strated successful cleanup of the most difficultstreams.

Air Emissions. Among the sulfur com-pounds, reduced species (H2S, COS, CS2) pre-dominate in gasifier raw gases and can be re-moved by a variety of processes (Chap. 5). Sul-fur compound concentrations are higher in gasi-fier gases than in combustion gases, which fa-cilitates their removal. For example, synthesisgas streams can be cleaned sufficiently for use

in processes that require almost total removal ofsulfur compounds because of the use of sulfur-sensitive catalysts. In addition, studies indicatethat virtually complete control of sulfur in coalgasification increases product cost only moder-ately [53].

Particulates are routinely removed from syn-thesis gas produced by coal gasification to be fedto catalyst beds for chemical synthesis. In thiscase, particulates are removed to a far greaterextent for catalyst protection than would be nec-essary to comply with environmental require-ments. Of greater concern are emissions fromfeed lock hoppers or feed handling and storagesystems; these, however, can be controlled if theequipment is designed properly.

Although nitrogen oxides are not formed toany appreciable extent in the reducing atmo-sphere of gasification, use of the product gas asfuel for combustion could pose anNOx emissionproblem due to nitrogen compounds in the gas.However, well-cleaned product gas from gasi-fication should yield an improvement over, forexample, the direct firing of coal.

Water Effluents. Gasification processes inwhich raw gas leaves the gasifier at relatively

76 Gas Production

low temperature (moving bed) result in greaterrequirements for wastewater treating because ofthe tar and phenols produced, yet such waste-water can be purified adequately by conven-tional technology. For example, water effluentstreams fromLurgi gasifiers have beenwell con-trolled on a large commercial scale at the Sasol IIcomplex (→ CoalLiquefaction) inSouthAfrica.Some gasification plants are designed for zerodischarge in arid areas where discharges mightseriously affect the quality of the receiving wa-ter.

Solid Wastes. All coal-consuming processesrequire the disposal of ash or slag resulting fromthe mineral content of coal. Some concern hasbeen expressed that solid waste from gasifica-tion might be declared hazardous because of itsleaching characteristics. However, leaching ofgasificationwaste generally results in lower con-centration ofmetals than leaching of power plantash, which has not been classified as hazardous.

5. Gas Treating

Gas treating described in this chapter is oftenpreceded by a separation stage for condensatesobtained from cooling stages immediately fol-lowing gas production (see Chap. 6). An over-all gas treating sequence is described in Section7.2.

Objectives. The method of gas treating isdetermined essentially by the following condi-tions:

1) origin (i.e., state and composition of the rawgas) and

2) utilization (i.e., specifications for the treatedgas).

As described in previous chapters, the com-position of raw gases differs considerably de-pending on the mode of production; this appliesalso to raw gas pressure and temperature (seeTable 24).

Requirements for the treated gas are mani-fold. They are more or less stringent dependingon whether the gas is to be used as a feedstockfor synthesis or as fuel gas. Generally, no spe-cial requirements exist for fuel gas composition

as long as the limits for calorific value and den-sity are met (see Chap. 8). Purity is specifiedwith a view to avoiding corrosion and depositsin the piping system or toxicity.

When syntheses followdownstreamof gasifi-cation, the stoichiometry of the reaction requiresa very specific composition, and the catalysts,which are generally sensitive, set very exactingstandards for purity of the reactants. In addition,economic aspects must be considered concern-ing the inerts contained in the gas, which unnec-essarily reduce the partial pressure of reactantsand tend to concentrate in recycle processes sothat a considerable amount of purge gas must bebranched off.

Stoichiometric Conditions. Most synthesesof basic chemicals in very large quantities startfrom hydrogen and/or carbon monoxide, in con-junction with further reactants if necessary. Theratio of hydrogen to carbon monoxide offers asimple characterization (see Table 25).

Purity. Sulfur compounds, which are severerespiratory and catalyst poisons and may causedangerous corrosion, are the subject of almost allpurity specifications. The content of hydrogensulfide (measured as sulfur) is generally limitedto 2 mg/m3 in long-distance pipeline systems.For syntheses, particularly those using sensitivecopper catalysts, the permissible concentrationis much lower: e.g., for the low-pressure meth-anol process, 0.1 ppm of total sulfur, measuredas hydrogen sulfide.

In some syntheses, oxygen compounds arealso undesirable: thus a content of

O2+CO2+2CO<20 ppm

is usually required for ammonia synthesis (→Ammonia). For reasons of corrosion, the oxygencontent in gas grids is limited to 0.5 – 3%. Ni-trogen oxides are also undesirable because theymay form unwanted or dangerous compoundswith unsaturated hydrocarbons. This applies notonly to syntheses but also to gas piping systemsand cryogenic gas treatment processes (→ Cryo-genic Technology; → Natural Gas). In view ofthe high process pressures, the nitrogen contentof hydrogen used for hydrocracking is often lim-ited to avoid formation of ammonia.

A number of other impurities, such as hydro-gen cyanide, organic acids (e.g., formic acid),

Gas Production 77

Table 24. Typical raw gas data

Raw material: Coal Vacuum residue Natural gas

Example: 1 2 3 4 5 6Production process: Koppers –

TotzekLurgi PressureGasification

Shell GasificationProcess

TexacoGasificationProcess

Tubular reformer

ex gasifier downstream ofscrubber

downstream ofCO shift

upstream of COshift

downstreamof CO shift

Pressure, MPa 0.105 2.9 5.4 8.5 2.2 2.1Temperature, ◦C 45 45 35 35 350 230Water dew point, ◦C 45 45 35 35 170 162Analysis, vol%

H2 28.7 43.2 45.9 63.0 76.1 78.3CO 57.0 11.6 48.6 1.8 10.6 0.4CH4 0.1 10.7 0.5 0.3 3.3 3.0C2+ 1.0N2+ Ar 1.4 0.3 0.2 0.1CO2 12.6 32.9 4.0 34.2 10.0 18.3H2S 0.2 0.3 0.8 0.6COS, ppm 300 150 500 10

Other components O2, SO2,HCN, NH3

O2, CmHn ,HCN, NH3

O2, HCN, NH3 NH3

Table 25. Typical hydrogen – carbon monoxide stoichiometric numbers for a few basic chemicals

Product Acetic acid a Oxo alcohols b Direct Fischer – Methanol e Ammoniareduction c Tropsch d

CH2/CCO 0 – 0.02 1.0 – 1.2 1.3 – 1.5 1.6 – 2.3 2.0 – 2.3 >105

Additional reactants methanol propylene iron oxides CO2 CO2 N2(=1/3 H2)a Monsanto process (→ Acetic Acid).b BASF and Ruhrchemie process; Shell process: 1.6 – 1.8.c Midrex process.d Lower range for fixed-bed synthesis (ARGE); upper range for Synthol (SASOL).e The reaction is reflected more accurately by

(CH2−CCO2

)/

(CCO+CCO2

)(→ Methanol).

or unsaturated and higher boiling hydrocarbons,are normally not included in purity specifica-tions for synthesis processes because they arealready undesirable in upstream process stagesof synthesis gas production.

The total content of inerts (N2, Ar, CH4),which inmost cases are not dangerous to synthe-sis but undesirable for economic reasons, mustoften be < 2%.

To meet all these requirements, several pro-cess stages are necessary for gas purificationor separation. They include catalytic processes(CO shift conversion, COS conversion, metha-nation; see Sections 5.1, 5.2, 5.3), as well asabsorption, adsorption, and other fractionationprocesses (Sections 5.4, 5.5, 5.6, 5.7, 5.8).

5.1. Carbon Monoxide Shift Conversion

Processes for the manufacture of synthesis gasfrom hydrocarbon feedstocks produce mixturesof hydrogen, carbon dioxide,methane, and vary-ing amounts of carbonmonoxide. For themajor-

ity of industrial processes, this carbonmonoxidecontent is higher than that required for synthesis.Exceptions include, for instance, the synthesis ofoxo alcohols and of acetic acid from methanoland pure carbon monoxide.

The following are typical examples of com-mercial processes in which carbon monoxide inraw synthesis gas is partly or largely converted:

1) Adjustment to the desired hydrogen: car-bon monoxide ratio in carbon monoxide-richgases, which are formed, for instance, by par-tial oxidation (e.g., methanol synthesis);

2) Conversion of toxic carbon monoxide to car-bon dioxide (e.g., town gas); and

3) Production of hydrogen from steam used forthe carbon monoxide shift reaction (e.g., hy-drogen production, ammonia synthesis).

5.1.1. Fundamentals

The water gas shift equilibrium among carbonmonoxide, carbon dioxide, hydrogen, and steam

78 Gas Production

depends on temperature but is almost indepen-dent of pressure in the industrial range (elevatedpressures up to 7.0 MPa).

Carbon monoxide shift conversion is anexothermic reaction:

CO+H2O�H2+CO2∆H0R = −41.2 kJ/mol

The equilibrium constant is described by

Kp =pH2 ·pCO2

pCO·pH2O

A number of models for the temperature depen-dency of the equilibrium constant and the heatof reaction appear in the literature. The temper-ature dependence of the equilibrium constant isshown in Chapter 1 (Fig. 1 and Table 3). Pres-sure influences activity and selectivity of the COshift catalysts.

The various formulas, which are generallyused for calculation ofKp(T) and∆HR(T), giveonly slight differences for the temperature rangeof commercial processes. For the low-tempera-ture range around 250 ◦C, the heat of reactionof shift conversion∆HR is ca.− 39.6 kJ/mol, itis about− 37.3 kJ/mol for the high-temperaturerange around 500 ◦C [139].

Above 950 – 1000 ◦C, the water gas equilib-rium is established rapidly enoughwithout a cat-alyst. In reforming processes using nickel-basedcatalysts, equilibrium is still attained at muchlower temperatures. However, because nickel-based catalysts promote methanation of carbonoxides, they cannot be used for process steps inwhich only carbon monoxide shift conversion isdesired.

The temperature ranges used in commercialcarbon monoxide shift conversions are:

a) High-temperature shift (HTS): ca. 300 ◦C(start of run, SOR, with new catalyst) to510 ◦C

b) Low-temperature shift (LTS): ca. 180 ◦C(SOR) to 270 ◦C

c) Raw gas shift (RGS): ca. 200 ◦C (SOR) to500 ◦C.

Somecatalystmanufacturers offer a so-calledmedium-temperature shift catalyst,which is nor-mally a further stabilized LTS catalyst to operateat temperatures up to about 320 ◦C. Maximumtemperature for high-temperature shift (HTS) isdetermined by the temperature resistance of thecatalyst (sintering); minimum temperature, by

the decrease in activity of this catalyst (large cat-alyst volumes are required). In low-temperatureshift (LTS), the upper temperature limit is set byaccelerated recrystallization of the copper cata-lyst (deactivation) above 270 ◦C; the lower limitby both the decline in catalyst activity and theneed to keep the gas inlet temperature enoughabove the dew point of the gas – steam mixtureto prevent condensation ofwater and subsequentdamage to the catalyst. Raw gas shift (RGS)has a wide temperature range of application andthe above explanations of temperature limits forHTS and LTS are similarly applicable for RGS.

In practice, at start of run equilibrium con-version can be reached whereas at end of runconditions typically an approach to equilibriumof about 40 ◦C is realized.

5.1.2. Catalysts

For carbon monoxide shift conversion in so-called clean gases (i.e., gases containing smallquantities of impurities such as sulfur or con-densable hydrocarbons) in the high-temperaturerange, iron – chromium oxide-based catalystswith trace additions of compounds containingsodium and potassium are used. At presentHTS catalysts are often copper promoted. Inthe low-temperature range (LTS) copper – zinc– aluminum oxide-based catalysts with tracesof potassium, etc. are commercially available[140, 141]. Some of these added elements actas promoters (enhancing activity, selectivity andproviding self-guarding, i.e., chemically bindingtrace elements in the feed gas that would lead tocatalyst poisoning) and some as stabilizers (in-creasing mechanical strength and active surfacecompositions at elevated temperatures and pres-sures) [142].

When new, catalysts often contain smallamounts of sulfur. At present, the catalysts aresold mainly in tablet form and are stronger thanthe broken catalysts (iron oxide) formerly in use.In addition, their smooth surface and regularshape result in a smaller pressure drop throughthe catalyst bed.

High-temperature shift catalysts are sup-plied in the oxidized condition (i.e., iron ispresent as Fe2O3 and chromium as Cr2O3).Generally, the reduction necessary to activatethe catalyst is not carried out before it is put into

Gas Production 79

the reactor, because the reduction process de-creases the mechanical strength by up to 30%.Therefore, the catalyst is reduced in situ withcarbon monoxide- and hydrogen-laden gases,usually the clean gas itself. To control the re-duction rate, the reducing gas is diluted withsuperheated steam or nitrogen (or both) to amixture containing below 5 vol% (CO + H2).During reduction the temperature profile of thecatalyst bed has to be monitored carefully.

Iron oxide catalysts are resistant to apprecia-ble amounts of impurities; even sulfur levels inthe range of 10 to > 100 ppm (mL/m3) can betolerated at the cost of activity decrease. Rapidtemperature and pressure changes caused, forexample, by water coming into contact with thehotter catalyst (water evaporates in the pores)must be avoided. These maloperations lead todisintegration of the structure and, generally, toa rapid increase in pressure drop because of theaccumulation of dust and pellet parts. All heat-ing and cooling steps aswell as pressure changesshould be carried out gradually.

Low-Temperature Shift Catalysts. Stringentgas purity requirements must be met when us-ing low-temperature shift catalysts because theyare very susceptible to poisoning by sulfur orhalogen compounds. Therefore, the gas must becleaned to sulfur and chlorine contents of < 0.1ppm (dry) upstream of LTS conversion. If thispurity is not already met by the feed (e.g., fromsteam reforming of natural gas) or achieved bywashing the raw gas (partial oxidation followedby Rectisol wash), a “guard bed” (usually a ves-sel filled with zinc oxide or LTS guard catalyst)must be installed upstream of the low-tempera-ture conversion unit.

Raw Gas Shift Catalysts. Carbon monoxideshift conversion of gases containing apprecia-ble amounts of sulfur or heavy hydrocarbonssuch as tar is known as raw gas shift conver-sion. For such gases catalysts consisting mainlyof cobalt and molybdenum are used instead ofthe iron – chromium type. The activity of cobalt– molybdenum-based catalysts increases signif-icantly between 4.0 and 8.0 MPa, whereas withHTS iron oxide catalysts, increasing the operat-ing pressure above 4.0 MPa has practically noeffect on the volume of catalyst required. RGScatalysts attain their full activity only in the sul-fided condition. Therefore, they must be presul-fided before installation or sulfided during start-

up. A minimum content of sulfur compounds inthe dry raw gas of ca. 100 to 1500 ppm (for gasesfrom partial oxidation; the value depends on thesteam: sulfur ratio) is required in normal opera-tion to maintain catalyst activity. The minimumsulfur content in the feed gas also depends onother operating conditions (e.g., temperature).The sulfur content has no upper limit. The RGSoffers the advantage of allowing sulfur and car-bon dioxide removal to be accomplished in onestep downstream the RGS unit [140].

The volume of catalyst required dependson the partial pressures of reacting compo-nents, the desired carbon monoxide content ofthe converted gas, and the reaction tempera-ture. Normally, operating parameters are chosento achieve space velocities of 1000 – 3000 m3

(STP) of gas per cubic meter of catalyst per hourin the HTS range and 2000 – 5000 m3 (STP) percubic meter of catalyst per hour in the RGS andLTS range. If necessary, the process is dividedinto several steps with intermediate cooling orwater quenching (see Section 5.1.5).

5.1.3. Clean Gas Shift Conversion

If the carbon monoxide-containing gas is sub-stantially free of sulfur and heavy hydrocarbons,clean gas shift conversion is applied. In case thegas had to be desulfurized in an upstream purifi-cation stage, a cooler – saturator system is fa-vorable. For naturally clean gases, (e.g., steamreformer outlet gas) only a simple HTS reactoris necessary.

Desulfurized feed gases derived from up-stream partial oxidation and a physical or chem-ical washing unit usually enter the CO shift unitat a temperature between 20 and 120 ◦C. Thus,even if they are saturated with water vapor, theaddition of a considerable quantity of high-pres-sure steam is required to effect satisfactory car-bon monoxide shift conversion. For example,to lower the carbon monoxide content of a gasavailable at ca. 5.4 MPa and ambient tempera-ture from 45 vol% (dry) to a residual ca. 3 vol%(dry) in two stages at a final reaction tempera-ture of 380 ◦C, the addition of 1.18 kg of exter-nal HP-steam per cubic meter (STP) of gas isnecessary. The cooler – saturator system, whichgenerates steam in the saturator by using the re-action heat of the carbon monoxide shift con-

80 Gas Production

version, enables the quantity of external processsteam to be lowered to 0.33 kg/m3 (STP) of gas.

Cooler – Saturator System. Shift feed gasenters the saturator (a) at ambient temperatureand rises through the tower packing while be-ing sprayed with hot water (see Fig. 34). Thegas becomes saturated with steam at ca. 225 ◦C.Heating of the gas – steammixture is continued,more steam (0.33 kg/m3 (STP) dry gas) is added,and the resulting mixture is fed to the first reac-tion zone (c). The temperature of the mixturerises due to the heat of reaction to about 500 ◦C.The first converter outlet CO concentration isabout 6 to 8 vol% (dry). The hot gas is fed first toa gas – gas heat exchanger (b) to preheat the in-coming mixture and then into a gas – water heatexchanger (e) to heat circulating water; by thismeans, it is cooled to ca. 350 – 360 ◦C. At thistemperature, it enters the second reaction zone(d), where carbonmonoxide is converted furtherto a residual content of ca. 3 vol% (dry), the gasbeing simultaneously heated to ca. 380 ◦C.

Figure 34. Flowsheet of a two-stage CO conversion plantwith cooler – saturator systema) Saturator; b) Gas – gas heat exchanger; c) Reactor 1;d) Reactor 2; e) Gas – water heat exchanger; f) Cooler;g) Waste Heat Recovery

Converted gas leaving the second reactionzone is cooled in a second gas – water heat ex-changer (e). Then it enters the cooler (f) at ca.230 ◦C. Here it is irrigated with colder waterwhile it flows upward through a packed bed. Bythis means, most of the steam condenses, andwater leaving the cooler is heated approximatelyto the dew point of the converted gas. The wa-ter leaves the cooler at ca. 210 ◦C. The cooleris often divided into two sections: the bottomsection is fed directly with most of the water

leaving the saturator; the top section is fed witha side stream that has been cooled to ca. 50 ◦C.The heat thus removed is used e.g., for boilerfeedwater preheating, demineralized water pre-heating, etc. (g). By using such an arrangement,the gas can be cooled to about 60 ◦C.

Water from the cooler is heated to ca. 250 ◦Cin the two gas – water heat exchangers (e) be-fore entering the saturator (a), which it leaves atca. 160 ◦C. Alternatively to the shown arrange-ment, in larger plants, the first gas – gas heat ex-changer (b) can be installed inside reactor 1, toavoid the hot transfer piping.

Countercurrent Cooling Reactor. In smallerplants, a countercurrent cooling reactor (Fig. 35)is often used instead of two reaction steps inseries. Whereas the reaction in a simple fixed-bed reactor proceeds adiabatically, which neces-sitates the two-stage mode of operation, a muchmore favorable temperature profile is achievedin the countercurrent cooling reactor.

Figure 35. Countercurrent cooling reactora) Gas entry into the cooling tubes; b) Surface of catalystbed; c) Catalyst support grid; d) Manhole

The gas – steam mixture, which is not pre-heated to catalyst inlet temperature, enters theannular space between reactor shell and cata-lyst vessel, flows downward through this space,and enters the tubes of the catalyst vessel.While flowing through these tubes which aresurrounded with catalyst, the mixture is heated

Gas Production 81

to ca. 380 ◦C. It then enters the catalyst bedfrom above. Due to the interaction between thetemperature rise caused by the heat of reactionand the cooling effect of the gas mixture flow-ing through the cooling tubes, the temperaturecurve illustrated in Figure 36 develops. This typeof reactor combines a high reaction rate at hightemperature in the top section with a favorableequilibrium in the bottom section.

Figure 36. Temperature profile in countercurrent coolingreactora) Gas entry into cooling tubes; b) Gas entry into catalyst;c) Gas outlet (denotations a to c correspond with Fig. 35)

Carbon Monoxide Shift Conversion ofSteam Reformer Outlet Gas. Shift conversioncan be carried out more easily on product gasesof a steam reformer for the following reasons:(1) the content of carbon monoxide is far lower(ca. 12 – 15 vol% in dry gas) compared to syn-gas produced by partial oxidation or coal gasi-fication; that means generation of less reactionheat in CO conversion; (2) the required steam isalready present in the steam reformer outlet gas,i.e., the gas has only to be cooled in the down-streamwaste-heat boiler to the inlet temperatureof the shift converter (ca. 300 – 350 ◦C); (3) thegas is essentially free of catalyst poisons, as thereformer feed has to be purified before, becauseof the sensitivity of the reformer catalyst.

An example is the production of synthesisgas for ammonia manufacture from natural gas.After reforming in the primary and secondary re-

former, and initial cooling in awaste-heat boiler,the reformed gas is, e.g., available at 2.9 MPaand 350 ◦C, and a steam content of 0.64 kg/m3

(STP). It contains 10.5 vol% (dry) carbon diox-ide, 11.7 vol% (dry) carbonmonoxide, and 55.5vol% (dry) hydrogen, the remainder being nitro-gen andmethane.After passing through theHTSconversion step, a carbon monoxide content of2.2 vol% (dry) is obtained at 405 ◦C. The gas,which still contains 0.52 kg of steam per cubicmeter (STP), is cooled to 220 ◦C and then fed tothe LTS step. There the carbon monoxide con-tent is decreased to ca. 0.3 vol% (dry), while thetemperature increases to 232 ◦C.

Shift conversion of these gases is carried outeither in one step in a HTS converter or in twosteps, with intermediate cooling, in HTS andLTS converters. More recently quasi isothermalLTS reactors have been installed, reaching lowCO contents at their outlet in a single step [143].As a byproduct a significant amount ofmethanolis normally produced [141, 144].

5.1.4. Raw Gas Conversion

Raw gas conversion (sometimes also termed“sulfided CO shift”) denotes the conversion ofgases containing significant quantities of sulfur,soot, or condensable hydrocarbons. There aretwo principal industrial applications: (1) shiftconversion of gases from heavy-oil gasificationthat have not been cooled and desulfurized buthaveonlyundergonehot quenching to add the re-quired steam and remove soot; and (2) shift con-version of the carbon monoxide in gases fromcoal pressure gasification that contain not onlysulfur but also saturated and unsaturated hydro-carbons, including tars. Cobalt – molybdenumbased catalysts have proved particularly effec-tive for converting both types of gas.

The carbon monoxide content (ca. 45 vol%,dry) of a gas from partial oxidation of heavyoil enters the carbon monoxide shift reactor atabout 250 ◦C with a steam: dry gas volume ra-tio of ca. 2.4. The carbon monoxide content canbe reduced to ca. 1.6 vol% (dry) over a cobalt– molybdenum based catalyst. Residual carbonmonoxide contents of 0.5 – 1.0 vol% (dry) areattainable in a two-stage system even at lowersteam: dry gas ratios. The sensible heat of theproduct gas is used to generate steam and to

82 Gas Production

preheat feedwater. Impurities in the gases forraw gas conversion, especially formic acid ne-cessitate the use of mostly austenitic materialsfor streams with temperatures below the wa-ter dew point. Carrying out shift conversion onraw gases produced by coal pressure gasifica-tion is a good way of “detoxifying” town gas.The amount of steam present in the gas at 3.0MPa and 180 ◦C, ca. 0.6 kg/m3 (STP), is suf-ficient to reduce the carbon monoxide contentfrom, e.g., 15.5 to ca. 5.0 vol% (dry). The cobalt– molybdenum based catalyst has a very usefulside effect: unsaturated hydrocarbons in the gasare hydrogenated almost completely. Parallel tothe CO shift conversion reaction COS is con-verted to H2S; i.e., COS contents in the shiftedgas reach low values close to equilibrium at cat-alyst outlet conditions.

5.1.5. General Comments on ReactorArrangements

Apart fromone-reactor applications several two-reactor systems are of commercial interest: twoHTS (or RGS) reactors in series or one HTSfollowed by one LTS reactor, in each case withintermediate cooling. In the latter case a “guardbed” filled with zinc oxide is often installed bet-ween the HTS and LTS reactors to ensure thatgas entering the LTS step is largely sulfur andhalogen-free. Alternatively, the “guard bed” isfilled with a smaller amount of LTS guard cata-lyst. SomeHTSprocesses includewater quench-ing between two reactor beds to cool the gas onthe one hand and to add steam as a reactant onthe other hand. In these concepts special cautionhas to be taken to avoid wetting of the catalyst.

5.1.6. Noncatalytic Quench Conversion

A less frequently usedmethod of carbonmonox-ide shift conversion on raw gases is noncatalyticquench conversion. Here, advantage is taken ofthe fact that the water gas equilibrium is attainedfairly rapidly above ca. 950 ◦C, even withoutcatalysts. If, for example, oxo synthesis gas witha hydrogen: carbon monoxide volume ratio of1.25 is to be produced from raw gas obtained bypartial oxidation of asphalt and containing 48.9vol% (dry) carbon monoxide and 43.9 vol%

(dry) hydrogen, then 0.4 kg of cold water perkilogram of asphalt is sprayed into the hot (ca.1570 ◦C) gas. As a result of cooling caused bythe heating and evaporation of water and the as-sociated carbonmonoxide shift conversion (heatof reaction), thewater gas equilibrium is attainedat 1250 ◦C.The carbonmonoxide content is then38.3 vol% (dry) and the hydrogen content 47.8vol% (dry), with a H2: CO ratio as desired.

5.2. Carbonyl Sulfide Conversion

In all gasification processes based on sulfur-con-taining feedstocks, carbonyl sulfide is producedaswell as hydrogen sulfide. For example, by par-tial oxidation of vacuum residues with a sulfurcontent of 3 wt %, a raw gas containing 0.71vol% (dry) of hydrogen sulfide and 0.03 vol%(dry) of carbonyl sulfide is formed. Many mod-ern gas washing processes remove carbonyl sul-fide together with hydrogen sulfide. A selectivecarbonyl sulfide conversion (i.e., without par-allel carbon monoxide conversion in a HTS orRGS converter, see Section 5.1) is limited to ap-plications where there is no need for deep desul-furization (e.g., fuel gas for gas turbines).

Conversion of organic sulfur compounds infeedstocks for tubular reforming by hydrogen-ation over cobalt – molybdenum and nickel –molybdenum catalysts is described in Section2.2.2.

Hydrogenation of carbonyl sulfide over ironoxide and cobalt – molybdenum based (see Sec-tion 5.1) catalysts occurs simultaneously withcarbon monoxide shift conversion, and equilib-rium in the form

H2+COS�H2S+CO∆H0R = +10.9 kJ/mol

being established. In the CO shift conversionrange at high temperatures of 350 – 500 ◦C, theequilibrium and the kinetics of this reaction leadto high COS conversion rates. When the gasmentioned above undergoes carbon monoxideshift conversion, the concentration of carbonylsulfide in the product gas typically reaches val-ues below 10 ppm (dry). (Table 3 lists a numberof values of the equilibrium constant for this re-action in relation to temperature.)

A selective transformation of carbonyl sul-fide to hydrogen sulfide is required in gases inwhich carbon monoxide shift conversion is not

Gas Production 83

desired. Catalysts have been developed for theselective conversion of carbonyl sulfide to hy-drogen sulfide which are inactive for the wa-ter gas shift reaction. These catalysts promotehydrolysis of carbonyl sulfide according to

COS+H2O�H2S+CO2∆H0R = −30.2 kJ/mol

This reaction is catalyzed by chromium– aluminum oxide-based promoted catalysts,which work in a sulfided state. The commer-cially available catalysts are mainly not subjectto poisoning by heavy metals and arsenic; buthalogen compounds will lead to reduction of ac-tivity, selectivity, and lifetime. These catalystsare used at 150 – 320 ◦C; satisfactory equilib-rium is attained at space velocities between 500and 2000 m3 (STP) per cubic meter of cata-lyst per hour [145]. As at lower temperaturesthe equilibrium conditions favor COS hydroly-sis, less steam in the feed gas is needed to reachdesired hydrolysis rates. (Values for equilibriumconstants as a function of temperature are shownin Table 3.)

A typical reduction of the carbonyl sulfidecontent in gases from partial oxidation [assumedgas compounds (dry): 45 vol% H2, 45 vol%CO, 5 vol% CO2, 0.71 vol% H2S, 300 ppm(mL/m3) COS] to about 10 ppm is achievableby COS hydrolysis at 170 ◦C. A steam contentof 0.086 kg per cubic meter (STP) of gas in thisexample corresponds to a saturation tempera-ture of 150 ◦C at 5.0 MPa. In parallel to COShydrolysis the contents of CS2 and HCN arealso significantly reduced. For some applica-tions a separate bed for hydrogenation/hydro-lysis of these compounds has to be installed ad-ditionally. For gases with low total sulfur con-tent there are zinc oxide-based “absorption cat-alysts” commercially available, which can hy-drolyze COS in the presence of steam (low con-centrations) before H2S is absorbed in the samebed.

5.3. Methanation and MethaneSynthesis

5.3.1. Definitions and Applications

Methanation is generally defined as the cat-alytic reaction of hydrogen with carbon monox-ide and/or carbon dioxide, forming methane and

water, according to the description given in Sec-tion 5.3.2. Two main applications of methana-tion exist, which require very different technicalsolutions:

1) Methanation for the elimination of carbonoxides in hydrogen production: upstream ofthis final purification step the concentrationsof harmful carbon oxides are rather low andthe exothermic methanation reaction there-fore causes only a moderate temperature in-crease, i.e., a single methanation step is suffi-cient (see Section 5.3.3).

2) Methanation for the production of SNG (Sub-stitute Natural Gas): in this application, alsocalled methane synthesis, hydrogen shall beconverted to methane as completely as possi-ble. Due to the strongly exothermic reactionand the high concentrations of all reactantsmeasures have to be taken to limit temper-ature rises. As also the thermodynamics, i.e.,the relevant equilibrium, ask for relatively lowtemperatures to yield high methane and cor-respondingly low hydrogen contents in theproduct, usually multiple-stage reaction sys-tems will be required (see Sections 5.3.4 and5.3.5).

Application (1) must be regarded as a wellestablished technique with little incentive to im-provements. It is widely used in hydrogen pro-duction, although nowadays pressure-swing ad-sorption (PSA) systems for hydrogen purifica-tion dominate themarket. Catalytic systemsmaybe damaged by too high temperatures due tobreakthrough of carbon oxides as a result of sud-den malfunctions of the upstream purificationsystems.

Application (2), the production of SNG, haslost commercial interest at present due to theincreased availability of natural gas. The priceof SNG cannot compete with that of natural gas.The development of new techniques for this kindof plants has therefore been terminated.

SNG production plants have demonstratedthe reliability of processes for conversion ofnaphtha to SNG (Baltimore, USA) or coal toSNG (North Dakota, USA). Special local con-ditions— such as isolated distribution systemsor missing storage capabilities—may be thereason for temporary implementation of SNGplants (for instance in West-Berlin before theGerman reunification, where only naphtha could

84 Gas Production

be stored at the time). A methanation systemat low temperatures (with ruthenium-based cat-alysts instead of normally applied nickel cata-lysts) may find application in space technology[146]. In this projected application carbon diox-ide from the breathing air is methanated in com-bination with a water electrolysis, thereby re-gaining oxygen for the biological cycle.

5.3.2. Principles of Methanation

The principle of catalytic synthetic productionof methane from carbon monoxide and hydro-gen was discovered in 1902 by Sabatier andSenderens. It is described by

CO+3H2�CH4+H2O

Carbon dioxide can also be converted tomethane according to

CO2+4H2�CH4+2H2O

Both reactions are linked by shift conversion,which is always observed simultaneously when-ever active catalysts are used:

CO+H2O�CO2+H2

Many observations indicate that the transfor-mation of carbon dioxide to methane is initiatedby a reverse shift conversion reaction with hy-drogen to yield carbonmonoxide and steam. Thecarbon monoxide formed then reacts to yieldmethane:

Both reactions for the formation of methaneare strongly exothermic, as shown by Table 26.This table also indicates that low temperatureand high pressure are required or favorable forhigh methane yield.

Kinetics. The mechanism and kinetics ofmethanation have been investigated, but thesestudies do not yield uniform results. The rateof conversion of carbon monoxide is often de-scribed formally by

rCO = k1·pmCO·pn

H2

with m = 0.7 – 1.0 and n = 0.3 – 0.5.

Table 26. Equilibrium constants and heats of reaction (accordingto [147] for

In the purification of synthesis gas with alarge excess of hydrogen, the reaction rate lawis simplified to

rCO = k2·pmCO

Summaries and details can be found in the liter-ature; see [148 – 150].

Kineticsmust also be consideredwhen inves-tigating the possibility of undesirable side reac-tions. Kinetic effects have to be used to avoidthe formation of carbon, because methanationusually begins in the temperature region wheresuitable thermodynamic conditions prevail forthe Boudouard reaction to occur:

2 CO�C+CO2

(see Section 1.3). If thermodynamic equilib-rium calculations indicate a possible buildup ofcarbon in the outlet region of the reactor, this un-desirable side reaction cannot be excluded andmay call for changes in process conditions. Atthe outlet zone ofmethanation reactors, attentionmust also be paid to the methane decompositionreaction

CH4�C+2H2

(see also Section 1.3.). However, a smallresidual hydrogen content andmoderate temper-ature normally prevent this reverse reaction.

Gas Production 85

Catalysts. For the removal of carbon oxidesfrom synthesis gases, catalysts with usually <15 wt % nickel are used predominantly. Acti-vated alumina, magnesium oxide and mixturesthereof are widely used as carrier materials. Forthe synthesis of methane from gases rich in car-bon monoxide, catalysts with a high nickel con-tent, similar to those used in reforming naph-tha to a methane-rich gas, are usually employed.Catalyst poisons,must be eliminated carefully tothe same level as in naphtha reforming (see Sec-tion 2.4.4) or even lower, because deactivationwith time depends primarily on such impuritiesin the synthesis gas. Catalysts of the naphthareforming type usually reach their limit of op-eration at about 500 ◦C. High-nickel catalystswith special aluminum oxide carriers and stabi-lization by zirconium oxide, which are capableof operating at 300 – 700 ◦C, are also available.The catalysts are used preferably in the formof cylindrical tablets measuring 5 × 5 or 8 × 8mm.

Catalysts based on ruthenium have been triedrepeatedly for methanation but have not foundthe broad application of nickel-based catalysts.Raney nickel coated reaction tube systems arealso of little interest, because replacement of thespent catalyst appears complicated.

Metal Carbonyl Formation. Precaution mustbe taken with nickel-containing catalysts to pre-vent formation of highly poisonous nickel car-bonyl [Ni(CO)4] (TLV 0.1 ppm (mL/m3) or 0.7mg/m3; see [151]).

In particular, low temperatures must beavoided because they favor the formation of car-bonyls, as shownbyTable 27.Generation of car-bonyls is favored by nickel in reduced form, hightotal pressure, or high partial pressure of carbonmonoxide [152]. The reaction is also promotedby sulfur, selenium, and tellurium. In practicaloperation ofmethanation plants, temperatures≤200 ◦C at the nickel catalyst must be avoided,and procedures for start-up and shutdown mustbe adjusted accordingly. Formation of iron car-bonyl also has to be prevented because of itstoxicity and general corrosion problems, and be-cause iron carbonyl decomposes and forms ironlayers on the active surface of the catalyst as thetemperature increases. Thus, carbon monoxide-laden synthesis gases must be heated in stain-less steel heat exchangers. Many years of plantoperation have shown that with adequate pre-

cautions, the formation of carbonyls can be sup-pressed successfully.

Table 27. Calculated nickel carbonyl content at equilibrium forvarious carbon monoxide concentrations as a function oftemperature for a total pressure of 1.48 MPa (according to [152]),ppm (mL/m3)

Temperature, CO content, mol%◦C 0.2 0.5 1.0 2.0 3.266 0.3 12 190 3 000 20 00093 0 0.2 3 49 320

121 0 0 0.1 1.6 11149 0 0 0 0.1 0.5177 0 0 0 0 0.03204 0 0 0 0 0

5.3.3. Methanation as a Step in HydrogenPurification

Carbon oxides are often poisonous to catalytichydrogenation processes. Their simultaneouselimination is possible by means of a metha-nation process step, which has been used widelysince the 1950s and is nowadays competingwithpressure-swing adsorption processes. Methana-tion remains the normal purification method forammonia synthesis gas production, especiallyby secondary air reforming. Here, the nitrogencontent prevents the use of other methods.

As shown in Figure 37, the inlet gas, con-taining ca. 0.5 – 0.8 mol% of carbon monoxideand the sameamount of carbondioxide, is heatedby methanated gas in a heat exchanger (b). Thefinal inlet temperature to the adiabatic reactor—typically 300 – 330 ◦C—may be achieved by afired trimheater (c). In the adiabatic shaft reactor(a) filled with nickel catalyst, carbon monoxideand carbon dioxide are converted with hydrogento methane and water vapor. Because the tem-perature increase in the reactor is ca. 70 K foreachpercent of carbonmonoxide and about 55 Kfor each percent of carbon dioxide, the result-ing outlet temperature can reach ca. 400 ◦C andthe temperature gradient is then large enough toshut off the trim heater in continuous operation.Instead of the fired trim heater, other arrange-ments of heat exchangers in combination withthe reforming plant are possible, as well as elec-trically operated start heaters. Purified hydrogenis then cooled to ambient temperature in the heatexchanger (b) and a final cooler (d), which alsocondenses most of the vapor formed.

86 Gas Production

Figure 37.Methanation as a purification step for hydrogenor synthesis gasa) Reactor; b) Heat exchanger; c) Fired trim heater;d) Cooler

In synthesis gas purification of this type, thetotal content of carbon oxides is limited to ca. 2vol%; otherwise, the temperature limits of thecatalyst would be exceeded and the final contentof methane in the hydrogen would be too high.Because of the loss of hydrogen bymethanation,carbon oxides contents no greater than 0.8% to-tal are aimed at. Additionally, the residual con-tent of carbon oxides is increased by higher out-let temperature and thus will not remain belowthe accepted tolerance of ca. 10 ppm (mL/m3)for the sum of carbon monoxide and carbondioxide.

5.3.4. Methanation of Rich Gas

Although methanation as a method of purifyingsynthesis gas yields only small absolute quan-tities of methane, methanation of rich gas andmethane synthesis from gases high in carbonmonoxide (e.g., produced by coal gasification)intentionally generate methane as product. Pro-duction of SNG by gasification of coal is be-ing considered as a long-term replacement forgaseous hydrocarbons of natural originwhen thelatter are no longer available in sufficient quan-tity. In contrast, production ofmethane from richgas (see Section 2.4.4) is restricted to specificareas where naphtha is abundant but no naturalgas is available.

When SNG is produced from naphtha, themethane content of the rich gas is typically ca.70 vol% (see Fig. 10, Section 2.4.4), whereashydrogen amounts to ca. 7% and carbon diox-ide 25%. Less than 0.5% carbon monoxide ispresent, and unreacted steam from naphtha gasi-fication amounts to ca. 1.30 mol of H2O permole of dry gas. Such a gas mixture can be en-riched further by admitting it at an inlet temper-ature around 300 ◦C (i.e., after cooling down-stream of gasification) to an adiabatic shaft reac-tor containing a nickel catalyst identical or sim-ilar to that used in naphtha reforming. Becausethe portion of unreacted steam is quite high andfurther steam is formed during the reaction, thedegree of methane formation is limited in thisso-called wet methanation step.

After cooling of the gas and condensation ofsteam, themethane content can be increased fur-ther by another methanation step (dry methana-tion), which may be up- or downstream of therequired carbon dioxide removal. The upstreamarrangement yields better methane quality, butthe danger of carbon formation is higher due tothe thermodynamic possibility of a reverse shiftreaction and consequent formation of carbon ac-cording to the Boudouard reaction. Installationof the dry methanation step downstream of thecarbon dioxide scrubbing system allows milderreaction conditions, but the carbon dioxide andhydrogen content of the product is increasedslightly because of the need to provide excesscarbon dioxide for the reaction of hydrogen.

A typical analysis of the final SNG is: meth-ane 98.5 vol%, carbon dioxide 0.5 vol%, hy-drogen 1.0 vol% and carbon monoxide < 0.1vol%.

Figure 12, Section 2.4.4, shows a one-stagemethanation system for rich gas installed up-stream of the CO2 removal system as part of thetotal plant concept for the production of SNGfrom naphtha.

5.3.5. Methane Synthesis from Gases withHigh Carbon Monoxide Content

Synthesis gases high in carbon monoxide aregenerated by various coal and oil gasificationprocesses. These processes naturally result indifferent qualities of synthesis gas according tothe feedstock used, the upgrading processes for

Gas Production 87

synthesis gas preparation, and the gasificationprocess itself. Whereas high-temperature gasi-fication processes yield high concentrations ofcarbon monoxide and little methane, interest inSNG production is concentrated on gasificationprocesses that yield high methane content in therawgas. This is possible, for instance,with Lurgipressure gasification of coal, especially whengasification pressures of 8.0 – 10.0 MPa are ap-plied, which has been successfully tested (seeSection 4.4). Methanation processes with littlemethane in the raw gas suffer principally from(1) high exothermic heat release during metha-nation, (2) the very large quantities of synthe-sis gas that must be handled (four to five vol-umes of dry synthesis gas yield one volume ofmethane), and (3) the high proportion of steamformed during methane synthesis, which lim-its the directly achievable SNG quality in wetmethanation steps.

Synthesis gases for methane synthesis areoften classified by their stoichiometric number(SN), i.e., the ratio of the reactable hydrogenvolume to the pro rata volume of carbon oxides:

SN =η·vH2

3vCO+4vCO2

where η stands for an equivalent hydrogen con-sumption required for the hydrogenation of un-saturated and higher hydrocarbons in producingmethane.

BecauseSNGspecifications demandahydro-gen content of <10 vol%, a general classifica-tion may be SN ≤ 1.05. Because of restrictionsin carbon dioxide content as well, the synthesisgas must meet the condition 0.98 < SN < 1.03for stoichiometric methane synthesis, which issuitable for direct production of SNG. Such stoi-chiometric methane synthesis allows synthesisgas to bepreparedby shift conversion and carbondioxide removal. After methane synthesis, onlycompression to pipeline pressure (if the gasi-fication pressure is not high enough) and dry-ing (e.g., by a glycol system) are necessary. Thequality of SNG varies slightly because fluctua-tions of the stoichiometric number are inevitablycaused by variables and control functions of thegasification and upgrading processes.

In stoichiometric methane synthesis, littlerisk exists of carbon formation according to theBoudouard reaction, even in high-temperaturemethanation up to 700 ◦C. These problems be-comemore serious at lower SN numbers. Lower

stoichiometric numbers always require carbondioxide removal as the final step before drying.The allowable load of carbon monoxide for thestoichiometric number of the synthesis gas mustbe determined experimentally for each gasifica-tion process. Individual optimization is neces-sary to attain lowest product costs, because shiftconversionmaybe expensive if the necessary va-por is not provided with the raw gas. Among avast variety of process configurations developedin the 1970s, the following are reportedly avail-able for commercial operation. However, onlythe process with recycle has demonstrated relia-bility under the harsh conditions of commercialproduction, due to the lack of interest in SNG.

Besides the stoichiometric number, thehighly exothermic reaction generally creates aproblem for the design of methane synthesisplants: either the temperature increase must belimited by recycling of reacted gas or steamdilu-tion, or special techniques such as isothermal re-actors or fluidized beds, each with indirect cool-ing by evaporating water, must be used.

Recycle Methane Synthesis. Figure 38shows a schematic of a methane synthesisplant with hot recycle. The type shown is de-signed as a two-stage recycle system with a fi-nal countercurrent-cooled reactor. These plantswere originally designed for a recycle quantitysuch that 500 ◦C is not exceeded at the outletof the adiabatic reactors. Current systems, op-timized with regard to low secondary energyrequirements, are available for maximum tem-peratures up to 650 ◦C. Inlet temperatures areca. 300 ◦C, which is also the operating rangeof the recycle compressor. An advantage ofthe hot recycle is that water vapor formed dur-ing the methanation reaction is not condensedand, therefore, the stoichiometric number isnot changed for the stages processing feed gas.Recycle systems show a high flexibility for ac-cepting synthesis gases of very different natures[153, 154]. The waste-heat recovery system isnormally designed to generate saturated steamup to 10.0 MPa, but superheated steammay alsobe withdrawn. Replacing the recycle compres-sor by a gas ejector reportedly produces positiveresults [155].

The higher the temperature selected for thefirst stage of methane synthesis, the more stages

88 Gas Production

Figure 38. Recycle methane synthesisa) Recycle reactors; b) Steam drum; c) Countercurrent cooling reactor; d) Air-cooled cooler; e) Hot recycle compressor

must be applied to reach specification-gradeSNG, for thermodynamic reasons. Therefore,with high-temperature methanation in the recy-cle stage(s), an intermediate step without feedgas addition at ca. 300 – 460 ◦C is required up-stream of a final dry or wet methanation stage.The countercurrent cooling reactor in Figure 38,has a similar design to that described in Section5.1.3 (Fig. 35). It allows product quality to bereached in one final wet step; otherwise, morestages are required to achieve the same analysis.

Steam Quenching Methane Synthesis. Forsynthesis gases with low stoichiometric num-bers (low H2: CO ratio), steam must always beadded to prevent carbon formation. For optimumresults, the required quantity of steam is mixedwith a proportion of the feed gas in such a waythat the temperature increase in the first metha-nation stage is kept within allowable limits and,after heat removal, a further portion of the feedgas is admitted to the first intermediate productgas stream as quenching stream. The quenchingprocedure after cooling is repeated until all the

feed gas is consumed. Thus, a multistage con-cept results, depending on the inlet gas compo-sition, as shown in Figure 39 for the RM pro-cess [156]. With the addition of steam, the equi-librium composition of any intermediate prod-uct gas is not favorable enough to enable directproduction of specification-grade SNG in a wetmethanation step. The removal of carbon diox-ide, therefore, is followed by a final, dry metha-nation stage.

Fluidized-Bed Methanation. Intensive re-search and demonstration have been reportedfor the methanation of feed gases of unfavorablecomposition, with little or no addition of steam,in a cooled fluidized-bed reactor in one stepwithout an additional cleanup reaction (Com-flux process) [157]. The dangers of carbon for-mation, catalyst deactivation by high carbonmonoxide partial pressure, and excessive cata-lyst consumption by the fluidized-bed processare said to have been overcome, and operationup to 6.0 MPa has been demonstrated success-fully in the pilot plant.

Gas Production 89

Figure 39. Flow principle of multi stage quenching system – RM process

Isothermal Operation. By increasing theallowable temperatures for methanation cata-lysts, methane synthesis can be performed by aonce-through method in quasi-isothermal reac-tors cooled by evaporatingwater, generating sat-urated steam. Under favorable conditions, suchsystems produce specification-grade SNG inonly one catalytic step (IRMA process), but su-perheated steam can be produced only by com-bining isothermal and adiabatic reactors [158].Any traces of sulfur compounds must be elim-inated carefully, because poisoning of the cata-lyst will displace the temperature profile in thereactor and the reaction will not go to com-pletion as the catalyst lifetime expires. Testswith isothermal reactors have been performedfor methane synthesis within the ADAM – EVAsystems for energy transport [159].

The ADAM – EVA Concept (Long-DistanceEnergy Transport System). Generally, methanesynthesis is used to generate methane, the heatof reaction being an inevitable byproduct. How-ever, a proposal by the Julich Nuclear Re-search Center (KFA-Julich, Germany) and theRheinische Braunkohlenwerke (Cologne, Ger-many) uses the energy release of the methana-tion reaction as the essential product. The energydemand for the steam reforming of methane isprovided by a nuclear high-temperature reactor;the resulting synthesis gas is then retransformedinto methane a few hundred kilometers away,thereby providing thermal energy up to 650 ◦C.Themethane is routed back to the nuclear systemby a second pipeline to restart the cycle; thus, nofossil energy is used up.

CH4+H2O�CO+3H2

Endothermic ExothermicBy the combination of power generation and

energy distribution at a level of ca. 150 ◦C, anoverall efficiency of about 60% can be reachedfor such a chemical heat pipe system. A gen-eral flow diagram of this process is given in[160]. The use of high-temperature solar heat-ing instead of nuclear power has been proposedfor steam reforming of methane, however, so-lar ADAM – EVA systems may be restricted toa few favorable locations. Plans have been re-ported to provide energy that has been recoveredand converted by high-temperature reforming intheNegevDesert to the industrial zones of Israel.

Studies have been made on the replacementof steam in the reforming step by carbon dioxideas oxidant [161]:

CH4+CO2→2 CO+2H2

Although this reaction sequence significantlyreduces the load in steam generation and can im-prove overall efficiency, transport of the thermo-dynamically necessary excess carbon dioxide inpipelines for both directions reduces this advan-tage. In addition, the danger of carbon formationin both reforming and methane synthesis at rela-tively high pressure must be examined carefully.

Other Synthesis Processes.Methanation of Raw Gases. The dry product

gas volume from methanation is only 20 – 35%of the synthesis gas volume, depending on thegasification process and its yield of primarymethane. Because nickel catalysts require com-plete removal of poisonous sulfur compounds,

90 Gas Production

relatively large amounts of gas must be handledin the upstream treating units. The Gas ResearchInstitute in Chicago has proposed direct metha-nation of raw gases containing sulfur and car-bon dioxide on special catalysts, to reduce gasvolumes in acid gas removal. Savings in over-all steam balance and in capital investment areclaimed for such process routes. A final metha-nation step for cleanup in a sulfur-free atmo-sphere yields specification-grade SNG.

Liquid-Phase Methanation. A proposal byChem Systems, Hackensack, N.J., is based onconventional gas preparation with methanationin the liquid phase [162]. Similar to processesemployingfluidized-bedmethanation or isother-mal reactors, liquid-phase methanation appearsto be able to handle high carbon monoxide con-centrations and to yield high-quality methanein a single reactor. An inert liquid (aliphaticoil) containing dispersed catalyst is circulatedand the feed gas is passed through this. Theexothermic reaction heat is taken up by the liq-uid, mainly as sensible heat and partly by vapor-ization, resulting in excellent temperature con-trol.

5.4. Absorption Processes

5.4.1. General

Principles. Processes that use a liquid sol-vent to remove one or more compounds from agas stream are called absorption processes [140,163 – 165] (→ Absorption).

The main purpose of absorption processes ingas production is the removal of acid gas com-ponents. They can also be used successfully topurify gases containing low concentrations ofother components such as ammonia, higher hy-drocarbons, water, cyanide, and organic sulfurcompounds. Absorption processes are furtherapplied to the withdrawal of gas constituentspoisonous to catalysts or to the environment andalso for the recovery of liquid gas or gasolinecomponents from refinery gases.

In electric power generation from coal or liq-uid hydrocarbons by the combined power cycleprocess (see Section 7.3), absorption is used toselectively remove sulfur compounds and otherimpurities such as cyanide andmercury from thegas in order to meet gas turbine requirements

and environmental regulations. So-called zero-emission IGCCs are discussed in view of CO2emissions regulations and greenhouse gas abate-ment. Carbon of the feedstock is converted tocarbon dioxide, which is removed in the absorp-tion process and must be concentrated in a purecarbon dioxide stream for sequestration (see Fig.59).

In the absorber, gas and liquid are broughtinto contact countercurrently. The solvent re-moves one or more components from the rawgas,more or less selectively. The absorber is usu-ally equipped with trays or packing to provide alarge gas – liquid contact area. The types mostfrequently used are valve and sieve trays. Bubblecap or tunnel trays are used only in special cases(a comprehensive treatment of trays and pack-ings used in absorption is presented elsewhere;→ Absorption).

The solvent laden with the absorbed compo-nents is withdrawn from the bottom of the ab-sorber, routed to a regeneration system, whereit is freed of absorbed gas components, and re-turned as lean solvent to the absorber. In mostgas treating processes, absorption is reversible,and dissolved components are released chem-ically unchanged during regeneration. Liquid-phase oxidation processes for hydrogen sulfideremoval differ in this respect (see Section 5.4.3).

The different absorption processes are char-acterized as physical or chemical depending onwhether gas components are simply dissolvedphysically or are bound chemically to the sol-vent. Typical equilibrium lines for chemicaland physical absorption are shown in Figure 40.Loading capacity c (kilomoles per ton or per cu-bic meter of solvent) is represented as a functionof the partial pressure of the dissolved compo-nent. The loading in physical absorption, ini-tially following Henry’s law, is almost directlyproportional to the partial pressure in the gasphase, which results in a nearly linear equilib-rium line. The equilibrium line in chemical ab-sorption is bowed sharply during saturation ofthe chemically active solvent component. Aftersaturation of the chemical capacity of the solu-tion, only weak physical absorption in the sol-vent’s main component (usually water) is pos-sible which results in a steep linear equilibriumcurve.

At lowpartial pressure (e.g., p2 in Fig. 40) theabsorption capacity cch of the chemical solvent

Gas Production 91

is much higher than the absorption capacity cphof the physical solvent, whereas the physical sol-vent gives better results at high partial pressure,(e.g., p1).

Figure 40. Gas – liquid equilibrium lines for (a) chemicaland (b) physical absorption

The strong dependence of physical solubilityon partial pressure can be utilized efficiently forsolvent regeneration by means of simple flash-ing (i.e., pressure reduction). This improves theeconomy of operation considerably, particularlyin case of bulk removal of impurities at high op-erating pressure.

If the dissolved components are chemicallybound, less gas is released (∆cch< ∆cph) whenthe solvent is depressurized. Thus, regenerationin chemical absorption processes is almost al-ways accomplished by reboiling.

Fewer trays are generally required for chem-ical absorption than for physical absorption dueto the acceleration of mass transfer by chemi-cal reaction in the liquid phase and the low acidgas equilibrium pressure over the solvent at lowloading.

Selection of Solvents. Selection of the appro-priate solvent in a gas treating plant is mostimportant because the solvent circulation ratedetermines equipment size and thus capital andoperating costs. The following simplified rulesare applicable, if linear dependence of the load-ing capacity on partial pressure across the entireconcentration range in physical absorption and

saturation of a chemical solvent above a certainpartial pressure are assumed.

The solvent circulation rate required forphys-ical absorption is largely independent of theconcentration of the component to be removedbut proportional to the total quantity G of crudegas and inversely proportional to the total pres-sure p and the absorption factor αi . The mini-mum solvent circulation rate Lmin required toremove component i completely can be esti-mated from the component balance around theabsorber.

Lmin =G

αi·pThe solvent circulation rate for chemical ab-

sorption is proportional to the quantity of thecomponent to be removed and thus highly de-pendent on the concentration of this componentin the gas, but practically independent of pres-sure. The maximum allowable concentration ofthe absorbed component in chemical solvents isoften set by practical limits like increasing cor-rosion rate, increasing viscosity or tendency offoaming at higher loading.

In addition to these basic rules, the follow-ing points should be considered in selecting anabsorption process.

Achievable Purity. High product gas puritycan be obtainedwithmany absorption processes,but only a few are suitable from an economicviewpoint. The minimum concentration of thecomponent to be removed in the clean gas ismainly determined by the residual loading ofthat component in the lean solvent fed to theabsorber top. Physical solvents can easily be re-generated to low residual loadings by pressurereduction and mild reboiling, whereas chemicalsolvents require high energy input for reboilingto achieve low loading.

Selectivity. The selectivity of a solvent is itstendency to absorb primarily one or a group ofcomponents from a gas mixture. Selectivity isoften defined as the concentration ratio of com-ponent i to component j in the regenerator off-gas to the concentration ratio of component i tocomponent j in the raw gas.

S =yi,off−gasyj,off−gas

/yi,raw−gasyj,raw−gas

A high selectivity between the componentsto be removed and the product gas componentsis advantageous. High absorption rates of the

92 Gas Production

valuable product gas components reduce pro-cess economy and may affect treatment of theoff-gas. If two or more components are to be re-moved from the raw gas and these componentsmust be obtained separately, the solubility of thevarious components in the solvent should be dif-ferent so that the selectivity is high.

Utilities. For chemical absorption, cheaplow-pressure steam or waste heat with a suffi-ciently high temperature is generally required,whereas cheap electric power and cooling waterat low temperature are an advantage for physicalabsorption.

Crude Gas Impurities. Impurities such ascarbonyl sulfide, carbon disulfide, hydrogencyanide, ammonia, mercaptans, thiophenes,phenols, heavy hydrocarbons, metal carbonyls,chlorides, and inorganic components such asmercury directly influence selection of both sol-vent and process [166]. Some trace compoundsof the raw gas can react with the solvent or otherabsorbed compounds to form nonvolatile reac-tion products, which cannot be removed dur-ing regeneration. They accumulate in the sol-vent circuit and can lead to operational prob-lems like plugging, foaming, corrosion and tosolvent losses by degradation. Many commer-cial solvents are not capable of removing organicsulfur components and cyanide to the requiredlow levels tomeet emission regulations or down-stream process requirements. Raw gases fromcoal gasification containmercury whichmust betrapped. The alternative to the selection of a sol-vent which allows the removal of all these tracecomponents in the absorption process step is theinstallation of a process chain consisting of COShydrolysis and cyanide conversion reactors andadsorption guard beds.

Therefore trace components often determinethe suitability of a solvent.

Other Solvent Properties. If the raw gas hasto be dried or a low water dew point of the prod-uct gas has to be achieved, absorption processesusing non-aqueous organic solvents are favor-able. Aqueous chemical washes are preferablewhen upstream process stages produce water asa component of the treated gas mixture and thiswater is required in the product gas downstream.The laden and the lean solvent should be noncor-rosive to avoid expensive materials of construc-tion. High chemical and thermal stability reducesolvent losses due to degradation and avoid ex-

pensive solvent reclaiming. Toxicity affects sol-vent handling, storage, and safety equipment.The solvent should be easily available on themarket.

The selection of an appropriate solvent canbe supported by computer tools such as AB-SORPERT (Ruhr-University Bochum [167]).This program contains heuristic rules, thermo-dynamic properties of a large number of sol-vents and a databank of industrial absorptionprocesses. All the information is linked to pro-vide an expert system for a systematic solventselection and a basic process design based onthe information of the feed stream and the de-sired separation.

Regeneration Methods. Absorption pro-cesses with solvent recovery consist, in princi-ple, of absorption and regeneration (desorption)process steps. Several possible solvent regener-ation methods are available; these are shown inFigure 41.

The simplest and cheapest method is flash re-generation (Fig. 41 A), in which the pressurizedladen solvent, is depressurized in one or morestages. The residual content of dissolved com-ponents depends on the pressure in the last stage;therefore, this is often reduced to vacuum. Re-generation efficiency can be increased consid-erably by inert gas stripping (see Fig. 41 B). Ifthe inert gas is completely free of the gas to beremoved, the residual load of the solvent can bereduced as desired, provided the volume of strip-ping gas is large and the stripping rate is highenough. It is, however, unavoidable that the re-moved gas component is diluted by the inert gas.This may be desirable in some cases, whereasin others, further processing becomes economi-cally prohibitive. Extremely high purity can beobtained without using inert gas if the impuri-ties are removed in a hot regeneration stage (Fig.41 C). This method makes use of the fact that,in general, the solubility of all gases decreasessharply with increasing temperature. The sol-vent is stripped by its own vapor, which is thencondensed from the overhead gas and recycledas reflux. However, this method is rather costlyin terms of capital investment (1) because it re-quires heat exchangers for the lean and ladensolvent and a regeneration column with reboilerand condenser, and (2) because the utility de-mand for heating and cooling is rather high.

Gas Production 93

Figure 41. Absorption with different regeneration methodsA) Flashing; B) Stripping; C) Reboiling

5.4.2. Processes for Carbon Dioxide andSulfur Compound Absorption [168 – 171]

Numerous chemical and physical absorptionprocesses are available for removing carbondioxide and various sulfur compounds from rawgases. Hydrogen sulfide is the most widespreadsulfur component and, like carbon dioxide, is anacidic compound. Other sulfur impurities thatmay be present in a raw gas are carbonyl sulfide,carbon disulfide, mercaptans, and thiophenes. Asummary of chemical and physical absorptionprocesses is given in Table 28.

Physical absorption processes are used pre-dominantly to remove carbon dioxide and hy-drogen sulfide fromgaseswith high carbondiox-ide partial pressure, especially converted gasesproduced in partial oxidation plants based oncoal or residual oil. Here the partial pressureof carbon dioxide is above the level appropriatefor economic application of chemical absorptionsystems.

Chemical absorption process (normal aminebased) are used widely, especially for crude gascontaining carbon dioxide and hydrogen sulfideat low partial pressure. As shown above, variousamines are employed. The selection of amine isinfluenced by specific process conditions such asthe presence of further impurities in the raw gas,which may cause solvent degradation, corrosionand foaming. Low coabsorption of valuable gascomponents in the aqueous amine solution is anadvantage of all amine processes.

Table 28. Absorption processes for acid gas removal

Trade name Solvent LicensorChemical absorption with aqueous amine solutionsMEA 1 – 3 N monoethanolamine free processDEA 2 – 4 N diethanolamine Elf Aquitaine

and othersDGA/Econamine

4 – 6 N diglycolamine(2-(2-aminoethoxy)ethanol)

Fluor Daniel,Huntsman

DIPA 2 N diisopropanolamine Shell and othersMDEA 2 – 5 N methyldiethanolamine free processADIP DIPA or MDEA ShellAmineGuard/Ucarsol

formulated MDEA UOP/UnionCarbide

aMDEA MDEA + activator for enhancedCO2 absorption

BASF

Gas Spec formulated MDEA DOW ChemicalFlexsorb hindered amine Exxon

Chemical absorption with alkaline solutions (hot potassiumcarbonate)Benfield K2CO3 + activator (DEA) UOPCatacarb K2CO3 + catalyst EickmeyerGiammarco-Vetrocoke

K2CO3 + activator (arsenictrioxide)

Giammarco

Vacasulf K2CO3 + NaOH Krupp Uhde

Physical absorptionRectisol methanol Linde/LurgiSelexol polyethylene glycol dimethyl

ether (DMPEG)UOP

Purisol N-methyl-2-pyrrolidone (NMP) LurgiMorphysorb N-formylmorpholine (NFM) Krupp Uhde

Physical – chemical absorptionSulfinol D/M sulfolane (tetrahydrothiophene

oxide) + DIPA or MDEAShell

Amisol methanol + MEA, DEA or DETA Lurgi

Sulfur components like hydrogen sulfide,carbonyl sulfide, and mercaptans have a highersolubility in physical solvents than carbon diox-

94 Gas Production

ide. This facilitates production of a sufficientlyhydrogen sulfide rich gas for the sulfur recoveryprocess. Compared to aqueous amine solutions,selectivity with regard to valuable componentsis lower in organically based physical solvents,which leads to higher losses of these compo-nents.

Amine solutions do not show a selectivity ofhydrogen sulfide over carbon dioxide at equilib-rium. Therefore, very high sulfur purity can beobtained only if carbon dioxide is removed si-multaneously to a very low level. However, thereaction rate of certain amines with hydrogensulfide is much higher than with carbon dioxide.Hydrogen sulfide can therefore be selectively re-moved by utilizing this difference in reactionkinetics. In the absorber a large gas liquid inter-facial area has to be provided for a fast hydrogensulfide mass transfer to achieve the required lowsulfide concentration in the purified gas. Simul-taneously a short contact time between gas andliquid is required so that only a small portion ofthe absorbed carbon dioxide can react with theamine [172].

Figure 42. Equilibrium curves of CO2 in various solventsa) H2O (30 ◦C); b) N-methyl-2-pyrrollidone (40 ◦C);c) Methanol (− 15 ◦C); d) Methanol (− 30 ◦C); e) Hotpotassium carbonate solution (110 ◦C); f) Sulfinol solution(50 ◦C); g) 2.5 MDiethanolamine solution (50 ◦C); h) 3 MAmisol DETA solution

In an attempt to combine the advantagesof chemical and physical absorption, processesusing chemical absorbents mixed with or-ganic solvents have been developed. These pro-cesses, termed physical – chemical absorptionprocesses or mixed solvent processes, are suit-

able both for absorbing carbon dioxide and hy-drogen sulfide and for removing organic sulfurcompounds, giving high end purities. A surveyof the equilibrium behavior of carbon dioxide incommon solvents is shown in Figure 42.

Initial selection of a gas purification processbased on the partial pressure of the acid gas inthe raw gas versus the product gas specificationcan be done with the aid of Figure 43.

Figure 43. Selection of appropriate gas purification processfor simultaneous H2S/COS and CO2 removal

If the acid gas partial pressure in the raw gasis very low, dry-bed adsorption processes (e.g.,solid beds of zinc oxide, iron oxide, activatedcarbon, or molecular sieves) or nonregenerablescavenger solutions are used for gas cleaning.

5.4.2.1. Physical Absorption Processes

Water was an industrially important solvent forcarbon dioxide absorption, particularly in am-monia synthesis, until organic solvents were in-troduced. Organic solvents have a much highersolvent capacity than water, which improvesprocess economy.

Organic solvents are especially suitable fortreating converted raw gases obtained from par-tial oxidation of oil or coal. The carbon dioxidepartial pressure in these gases can be as highas 2.5 MPa. At this high acid gas partial pres-sure the solvent capacity, i.e. solvent loading,

Gas Production 95

of organic solvents exceeds that of chemical ab-sorbents considerably.

Rectisol Process. The Rectisol process, in-vented by Lurgi around 1950 and developedfurther together with Linde, is particularly suit-able for the specific requirements of high-pres-sure synthesis gas production [173].This processachieves syngas quality with very low total sul-fur content in the ppb range and removal of allimpurities and bulk CO2 in one process step.More than 60 Rectisol units were reported tobe operating worldwide in 2005 purifying morethan 210× 106 m3/d (STP) of synthesis gas,which is equivalent to more than 75% of the gasproduced from coal, oil, and waste gasification.

Originally designed and applied in conjunc-tion with Lurgi pressure gasification of coal, theRectisol process is now also used successfullyin oil-based synthesis gas plants and other coalgasification processes, as well as in special casesof natural gas treatment.

The Rectisol process uses cold methanol assolvent. The solubility of acid gas in metha-nol increases strongly with decreasing tempera-ture whereas the solubility of the valuable lightcomponents is almost temperature independent.Temperatures as low as − 75 ◦C are thereforeapplied in order tominimize the required solventcirculation rate and the loss of valuable gas com-ponents. The low operation temperature also re-duces solvent losses into the product streams bylowering the vapor pressure of methanol. Goodmass and heat transfer are provided bymethanoldue to its low viscosity. Operating the processat low temperature does not necessarily requirelarge refrigeration units. The major part of thecooling is done by internal heat exchange andby utilizing the temperature decrease by gas de-sorption and expansion. In Figure 44, the tem-perature dependence of the absorption coeffi-cient λ, which is defined by c = λp, is shownfor various gases at 1 bar. The absorption coef-ficient is not entirely pressure-independent butrises above the value measured at 1 bar as thepartial pressure of the gas approaches its satu-ration vapor pressure. The nonideal behavior ofthe gas phase at higher total pressure also mustbe considered to model the gas liquid equilibriacorrectly (→ Absorption).

Figure 44. Absorption coefficient λ of various gases inmethanol at low pressure (1 bar)

Standard Rectisol Process. In gas produc-tion, the Rectisol process is used in two basicconfigurations: the nonselective standard pro-cess and the selective process configuration. Allacid gas constituents are absorbed simultane-ously in the standard Rectisol process. The basicflow diagram of the standard process with twoabsorption sections is shown in Figure 45. Flash-regeneratedmethanol, fed to the lower part of theabsorber, largely removes hydrogen sulfide andcarbon dioxide at bottom temperatures between− 10 to − 30 ◦C. The remaining hydrogen sul-fide and carbon dioxide are removed from thegas to the required purity by heat-regeneratedcold methanol in the absorber top section.

Under these process conditions, solvent load-ing is very high at the bottom of the absorber,which results in a low solvent rate. The heatof absorption raises the temperature of the sol-vent; the absorber bottom temperaturemay be asmuch as 50 ◦C higher than the solvent inlet tem-perature. During flash regeneration of the sol-vent, this heat of absorption is given off. and thesolvent temperature drops accordingly. The ap-plied low absorption temperatures are generated

96 Gas Production

Figure 45. Standard Rectisol process for simultaneous removal of H2S and CO2a) Absorber; b) Flash tower; c) Hot regeneration; d) Condenser; e) Methanol – water column

in this desorption step and not by the refrigera-tion unit. This phenomenon is often referred toas the “autorefrigeration effect”. In the fine washsection of the absorber, only the residual acidgases are removed. Therefore, little heat of ab-sorption is released in the fine wash section andthe temperatures remain low. Moreover, there isa smaller quantity of gas to be treated in the finewash section. Therefore, solvent demand in thefine wash section is considerably lower than thatin the main section of the absorber. This is de-sirable in view of the high investment and utilitycosts for hot regeneration.

In Fischer – Tropsch synthesis, SNG, andtown gas plants, carbon dioxide must only be re-moved to a final content of 1 – 2%. This valuecan even be reached in the main wash part withflash-regenerated methanol. The fine wash partthen has to be designed only for removal of hy-drogen sulfide and carbonyl sulfide, which re-sults in a lower solvent rate, normally one-tenthof the total amount of solvent withdrawn fromthe bottom, due to the higher solubility of thesecomponents.

Downstream of hot regeneration, a metha-nol – water distillation column is used to sepa-rate water entering with the raw gas. To preventicing in the cooling section of the raw gas, asmall amount of methanol is injected upstreamof the heat exchanger. The mixture of methanoland water is withdrawn and fed directly to thedistillation column. A refrigeration unit is usedto cover cold losses in the plant due to surge

energy, incomplete heat exchange, unrecoveredheat of absorption, and insulation losses. An am-monia absorption refrigeration unit can be ad-vantageously applied if low temperature (130 –200 ◦C) waste heat is available.

The standard Rectisol process combines allabsorbed gases in a single acid gas stream. Thesulfur compounds are thus dilutedwith the entireabsorbed volume of carbon dioxide. (Referredto the raw gas specification in Table 24, only 0.9vol% of hydrogen sulfide [example 2] and only1.8 vol%, of hydrogen sulfide [example 4] in theoff-gas are typical of the standard process withcomplete removal of carbon dioxide.) The off-gas with this low sulfur content is not suitablefor sulfur recovery in a Claus process. Other sul-fur recovery processes must be applied such asliquid-phase oxidation (see 5.4.3) or adsorption.The standard Rectisol process is therefore ad-vantageously applied for gases from low-sulfurfeedstock.

Selective Rectisol Process. Initially the se-lective Rectisol process was introduced to pro-duce an off-gas rich in hydrogen sulfide whichcould be processed in a Claus plant for sulfurrecovery. Pure carbon dioxide is produced as asecond product with a total sulfur content in thelow ppm range. This stream can be vented to at-mosphere or utilized, for example in urea plants,which are often built in conjunction with ammo-nia plants, or recovered in food grade (see Sec-tion 6.3),or can be used for enhanced oil recov-

Gas Production 97

Figure 46. Selective Rectisol process for separate removal and recovery of H2S and CO2a) H2S absorption; b) CO2 absorption; c) Flash regeneration; d) Stripper; e) Reabsorber; f) Hot regeneration

ery and sequestration. The solvent loops shownin Figure 45 are interchanged for that purpose(see Fig. 46). The raw gas is first desulfurizedwith a small amount of heat-regenerated metha-nol. The carbon dioxide coabsorption is low be-cause of its low solubility and the presaturationof methanol in the carbon dioxide absorptioncolumn.

Most of the carbon dioxide and the remain-ing traces of sulfur compounds are removed inthe downstream carbon dioxide scrubber in twostages, by using flash-regenerated methanol inthe middle section and heat-regenerated metha-nol in the top section.

Depending on the requirements, carbon diox-ide may be removed down to a few percent oreven a few parts per million. Removal of bothcarbonyl and hydrogen sulfides as well as mer-captans to 0.1 ppm is ensured in all cases.

According to Figure 44, the selectivity bet-ween carbon dioxide and hydrogen sulfide forthe raw gas indicated in column 4 of Table24 is not sufficiently high to give an off-gasrich enough in hydrogen sulfide for the Clausprocess. Therefore, an additional concentration

stage is required before the solvent passes to hotregeneration. In the bottom section of the reab-sorber column (d in Fig. 46) carbon dioxide isstripped out of the laden solvent by simple pres-sure reduction or stripping with inert gas. In thetop section (e) the small amount of hydrogen sul-fide which is stripped together with the carbondioxide in the bottom section is reabsorbed withmethanol to yield a sulfur-free off-gas.

In the selective Rectisol process, much moresolvent must be treated by hot regenerationthan in the standard Rectisol process. Additionalmethanol is required in the reabsorber, andmoremethanol is necessary in the hydrogen sulfideabsorber because the entire gas rate must betreated at higher temperature.

To avoid the extra load on the regenerationunit, the two Rectisol wash steps can be sep-arated by removing the sulfur compounds up-streamof carbonmonoxide conversion,with car-bon dioxide removal located downstream of theconversion unit. The solvent loops of both ab-sorption steps are still combined. A classic ap-plication of this case is synthesis gas productionby residual oil gasification.

98 Gas Production

As shown in Table 24, example 3, the ratioof hydrogen sulfide to carbon dioxide upstreamof shift conversion is approximately ten timeshigher than downstream (example 4). Therefore,the natural selectivity ofmethanol is sufficient toproduce a sulfur-rich off-gas suitable for sulfurrecovery in a Claus process. In addition, the gasrate upstream of carbon monoxide conversionis only 70% of the rate downstream; therefore,less solvent is required for desulfurization. Thissimplification of the gas treating unit, combinedwith the advantage of passing a sulfur-free gasto the carbon monoxide conversion unit and thepossibility to apply low-temperature shift con-version for minimum residual carbon monoxidein the raw hydrogen, must be paid for by coolingthe gas twice.

Because the selective Rectisol process ismuch more complex than the standard process,a decision must be made in each individual caseas to which route gives the most economical so-lution in conjunction with the downstream plantand the sulfur recovery unit.

For Fischer – Tropsch synthesis and SNGplants based on coal with a low sulfur content(example 2 in Table 24; pure carbon dioxide isnot required in either case), the standardRectisolprocess has mostly been used, with the off-gasbeing treated in a special chemical system.

A two train Rectisol plant is shown in Figure47. Part of the raw gas from a residual oil gasifi-cation is shifted and purified in the first train toproduce pure hydrogen. The remaining raw gasis purified in the second train giving gas for oxosynthesis. The solvent of both trains is regener-ated in a common hot regenerator column.

Applications [174, 175]. In view of (1) thehigh flexibility with regard to carbon dioxide,hydrogen sulfide, and carbonyl sulfide removal,as well as the removal of various trace com-pounds, (2) the high purity of the treated gas,and (3) the absence of corrosion at low tempera-ture, theRectisol process is used in the followingcases:

Standard processes are used for raw and con-verted gas from coal gasification (in this case,combination with prewash systems for other im-purities is also possible).

Selective processes with two absorptionsteps immediately following are used for raw andconverted gas fromcoal or oil gasificationplants.

Figure 47. Two train selective Rectisol plant downstreamof residual oil gasification producing 40 000 m3/h (STP)gas for oxo synthesis from unshifted gas and 30 000 m3/h(STP) hydrogen from shifted raw gas (Courtesy of Celaneseand Lurgi AG)

The selective process is used with two sep-arate absorption steps for raw and convertedgas from oil and coal gasification by partial ox-idation processes. Absorber – stripper systems(Fig. 41) are used to remove small quantities ofcarbon dioxide and hydrogen sulfide as integralparts of low-temperature units.

Plants with a one train throughput of morethan 15.6 × 106 m3 (STP)/d are already oper-ating. The utility requirements for a Rectisolunit vary widely depending on process condi-tions. Typical consumption figures for selectiveremoval of hydrogen sulfide and carbon diox-ide from raw gas, yielded by partial oxidationrelative to 100 000 m3 (STP)/h (H2 and CO) are

Power for drive shaft(without powerrecovery)

2000 – 2500 kW

Steam (0.5 MPa) 5 – 7 t/hWaste heat above110 ◦C

33 – 42×106 kJ/h using absorption

Cooling water (∆t =10 ◦C)

1500 – 2000 m3/h refrigeration

Methanol 60 – 70 kg/h

Gas Production 99

The application of the Rectisol Process forhydrogen production and for Integrated Gasifi-cation Combined Power Cycles (IGCC) [176]isdescribed in 5.4.2.4(see Sections 7.2 and 7.3).

Processes with Other Physical Solvents.TheRectisol process needs a refrigeration unit tomaintain the low absorption temperatures. Otherprocesses employ organic solvents with highboiling points and low vapor pressures. Thesesolvents provide a high absorption capacity foracid gases at ambient or slightly subambienttemperature without appreciable vapor losses.

Although higher absorption temperature low-ers carbon dioxide solubility, at a high partialpressure of carbon dioxide the loading capacityof organic solvents is still attractive and simi-lar to the capacity of chemical solvents undercomparable conditions.

Processes using high-boiling organic sol-vents are

1) Selexol process developed by Allied Chemi-cal Corporation (now licensed by Norton andUOP) and the Genosorb process of Clariantand Krupp Uhde, all using a mixture of di-methyl ethers of polyethylene glycol (DM-PEG);

2) Purisol process, developed by Lurgi, usingN-methyl-2-pyrrolidone (NMP);

3) Fluor solvent process usingpropylene carbon-ate (PC);

4) Morphysorb process developed by KruppUhde using N-formylmorpholine (NFM);

5) Sepasolv process by BASF applying a mix-ture of oligoethylene glycol methyl isopropylethers (MPE); and

6) the Gaselane process using N-methylcapro-lactam (NMC) developed by VEB Leuna[178]

The normal boiling point of these solvents isbetween 200 and 350 ◦C; their relative molec-ular mass, between 100 and 320. The solubilityof carbon dioxide is similar for all solvents.

If selective removal of hydrogen sulfide fromgas containing carbon dioxide is required, NMPis superior to Selexol, Sepasolv, and propylenecarbonate. The solubility of carbonyl sulfide inall solvents is remarkably lower than the solu-bility of hydrogen sulfide.

This means that if these processes are used topurify gases containing both carbonyl and hy-

drogen sulfides to synthesis gas quality, a car-bonyl sulfide hydrolysis step is generally re-quired.With regard to removal of carbon dioxidefrom sulfur-free synthesis gas, the process flowsheet is basically the same for all processes, con-sisting of a simple absorber – stripper system. Inaddition to nitrogen or other inert gases, air is asuitable stripping medium.

Selexol Process [177]. This process was de-veloped by Allied Chemical Corporation andis now licensed by Norton and UOP. The Se-lexol solvent is a trademark of the DOW Chem-ical Company and consists of a mixture of di-methyl ethers of polyethylenglycol (DMPEG).A similar mixture of DMPEG is used as sol-vent in the Genosorb process of Krupp Uhdeand Clariant (solvent supplier). More than 50Selexol plants had been constructed by 2005purifying more than 110× 106 m3/d (STP) ofnatural gas and synthesis gas. Most commercialapplications of the Selexol process are for treat-ing natural and synthesis gas with high contentsof hydrogen sulfide and carbon dioxide or highconcentrations of carbon dioxide alone [175].To purify gases containing carbon dioxide andhydrogen sulfide, the process may be used foreither bulk removal of both components or se-lective removal of hydrogen sulfide, with the de-sired quantity of carbon dioxide being retainedin the lean gas. Whereas the solubility of hydro-gen sulfide is approximately nine times that ofcarbon dioxide, the solubility of carbonyl sulfideis only twice that of carbon dioxide.

This means that the difference in solubilitybetween hydrogen sulfide and carbon dioxideis large enough to provide a good basis for se-lective absorption of the former. The solubilitydifference between carbonyl sulfide and carbondioxide, on the other hand, is less favorable.

If carbonyl sulfide must be removed togetherwith hydrogen sulfide, the solvent picks upmorecarbon dioxide and the feed to theClaus unit willbe diluted with more carbon dioxide than for re-moval of hydrogen sulfide alone. Hydrolyzingcarbonyl sulfide upstream of the absorption pro-cess can circumvent this problem. Nearly totalconversion of carbonyl to hydrogen sulfide canbe attained by vapor-phase catalytic hydrolysis,leaving only hydrogen sulfide and carbon diox-ide as key components.

A certain amount of water in the Selexol sol-vent does not greatly decrease the solubility of

100 Gas Production

acid gases. Thus, in the case of hot regenerationby reboiling under normal pressure, the boilingpoint of the solvent can be decreased by additionof a suitable quantity of water.

A disadvantage of the Selexol solvent is thatits lower boiling components are distilled off.Care must be taken that solvent properties donot change in the course of time.

The Selexol process is applied for followingpurposes:

1) Selective removal of hydrogen sulfide fromnatural gas and synthesis gas

2) Removal of high carbon dioxide content to alow residual level from natural gas and syn-thesis gas

3) Purification of landfill gas

Purisol Process. The Purisol process appliesa mixture of NMP and water as solvent. It canbe operated either at ambient temperature or— in conjunction with a refrigeration unit— attemperatures down to ca. − 15 ◦C. The Purisolprocess has been operated successfully for sev-eral decades in selective natural gas desulfuriza-tion and bulk CO2 removal from converted syn-thesis gas. The process has the advantage of ahigher hydrogen sulfide – carbon dioxide selec-tivity than Selexol, permitting the production ofsulfur-rich Claus gas with about 50% hydrogensulfide, even if the ratio of carbon dioxide to hy-drogen sulfide in the raw gas reaches extremelyhigh values.

The solubility of carbonyl sulfide is only20%that of hydrogen sulfide, but the former is partlyhydrolyzed to the latter by the NMP solvent re-sulting in a 60 – 80% carbonyl sulfide removal.If the absorber is designed with carbonyl sul-fide reaction zones, COS can be removed quan-titatively. The absorption of valuable gas com-ponents is approximately 10% lower comparedto Selexol. Therefore a smaller amount of flashgases has to be recompressed.

For carbon dioxide removal, the NMP sol-vent can be regenerated with air, if nitrogen isnot available. The carbon dioxide concentrationin the purified gas can be reduced to 1000 ppmor less, with almost no steam demand for regen-eration.

A typical example of a Purisol unit for remov-ing carbon dioxide from converted gas is shownin Figure 48. Two identical units of this type op-

erate in the same plant, treating synthesis gas forhydrogen production.

Pre- and afterwash steps are shown for boththe absorber and the stripper in the flow sheet ofFigure 48. Initially, raw and stripping gases aredried with a small amount of NMP to keep thewater content of the solvent in the carbon diox-ide absorption section low. A water wash on topof the columns reduces NMP losses. In general,NMP recovery with water is not necessary if thePurisol process is operated at subambient tem-perature. If a low water dew point of the productgas is required, NMP can be recovered bymeansof a glycol afterwash step.

The water – NMP mixture is separated in athird column, which simultaneously recoversNMP from the carbon dioxide off-gas at the flashstage. Carbon dioxide itself reduces the boilingpoint of the NMP – water mixture, which is adesirable effect.

Applications. The Purisol process is excel-lently suited to fours typical applications forhigh-pressure gases:

1) Selective removal of hydrogen sulfide2) Removal of high carbon dioxide content to a

low residual level3) bulk removal of acidic compounds down to

moderate product purity by using a simplifiedflash regeneration system, and

4) Removal of organic sulfur components (mer-captans)

Fluor Solvent Process. In the late 1950s, theFluor solvent process using propylene carbonatewas commercialized by Fluor and El Paso [180].Removal of carbon dioxide from natural gas atmildly subambient temperature was the first ap-plication. The main criterion for selecting PCin the Fluor solvent process was its high carbondioxide solubility, together with a relatively lowsolubility of methane. Propylene carbonate haslimited water solubility, and slowly reacts irre-versibly with water and carbon dioxide around90 ◦C. Thus, it is unsuitable for controlling thewater content by distillation under atmosphericpressure. Pre- or afterwash steps are, therefore,not applicable.

Morphysorb [181, 182]. Krupp Uhde ap-plied their experience with N-formyl morpho-line (NFM) as solvent to develop the Mor-physorb process for selective sulfur compoundsremoval and bulk CO2 absorption. Lower meth-

Gas Production 101

Figure 48. Purisol process for CO2 removal from high pressure converted gasa) Absorber; b) Stripper; c) NMP-water separator

ane solubility compared to other commercialsolvents is claimed. NFM has a freezing pointof 23 ◦C. To take advantage of the higher acidgas solubility at lower temperatures a mixture ofNFM and N-acetyl-morpholine (NAM) is pro-posed.

5.4.2.2. Chemical Absorption Processes

Amine Processes. Aqueous amine solutionshave been used on an industrial scale for theabsorption of carbon dioxide and hydrogen sul-fide for a long time. They are especially suitedfor gases with low organic sulfur content andless stringent purity requirements regarding hy-drogen sulfide, as for example natural gas ofpipeline quality or fuel gases. Synthesis gasqualities with total sulfur contents below 1mL/m3 require an additional sulfur adsorptionguard bed downstream of the amine absorber.

Mono- and diethanolamine (MEA, DEA),as well as diglycolamine (DGA, 2-(2-aminoethoxy)ethanol), diisopropanolamine(DIPA), and methyldiethanolamine (MDEA),are well known in this field. Several so-calledformulated amines are available on the marketsuch as Ucarsol, Gas-Spec, and aMDEA (seeTable 28). These solvents contain specific addi-tives in order to allow higher amine concentra-tions, to enhance carbon dioxide absorption rateand to reduce corrosion potential and foaming

tendency. The selectivity of the solvent towardshydrogen sulfide can be improved by stericallyhindered amines such as Flexsorb SE.

Alkanolamines can be regarded as derivativesof ammonia in which one, two, or three hy-drogen atoms have been replaced by a CnH2n-OH group. According to the number of or-ganic groups attached to the nitrogen atom theyare classified as primary, secondary, and ter-tiary alkanolamines. The primary alkanolaminesapplied for gas treating contain one ethanolgroup (MEA) or one glycolic group (DGA)The secondary alkanolamines employed containtwo ethanol (DEA) or isopropyl alcohol groups(DIPA). MDEA as anN-alkylated alkanolaminehas two ethanol groups and one methyl groupattached to the nitrogen. The amino group pro-vides the necessary alkalinity in the water so-lution to absorb the acid gases. The hydroxylgroup in the alcohol moiety increases the watersolubility and reduces the amine vapor pressure.

In the amine processes the acid gas compo-nents react reversibly with the alkaline solvent.The reactions of hydrogen sulfide and carbondioxide with amines can be represented as fol-lows:

Reaction of amine and water (protonation):

R1R2R3N+H2O�R1R2R3NH++OH−

Sulfide formation:

H2S+R1R2R3N�R1R2R3NH++HS−

102 Gas Production

Carbamate formation:CO2+2R1R2R3N�R1R2NCOO−+R1R2R3NH+

Bicarbonate formation:

CO2+OH−�HCO−3

Hydrogen sulfide reacts instantaneously withall amines to form sulfide. Carbon dioxide canonly react with primary and secondary amines toform carbamate. The carbamate reaction rate hasa moderate velocity. The bicarbonate formationis slow. The different reaction rates of hydrogensulfide and carbon dioxide with the amines arethe basis for selective hydrogen sulfide absorp-tion. MDEA has a higher selectivity comparedto primary and secondary amines since no carb-amate is formed with tertiary amines and carbondioxide can only be absorbed via the slow bicar-bonate reaction. In order to utilize this selectiv-ity based on the reaction kinetics, the absorbermust provide sufficient mass transfer area forcomplete hydrogen sulfide removal in combina-tionwith short residence time to suppress carbondioxide absorption [183].

When the temperature of the solvent is in-creased and the pressure reduced, the chemicalequilibrium of the above equations is shifted tothe left side and the acid gases are released.

The flow diagrams for all amine process arebasically the same (Fig. 49). Feed gas is broughtin contact with the solvent in an absorber (a).Laden solvent flows through a heat exchanger(d) to the regeneration system (e), in which acidgases are desorbed from the solvent at reducedpressurewith vapor generated by reboiling.Acidgases leave the regenerator overhead after thestripping steam has condensed (f). Regeneratedsolvent is returned to the absorber by way ofthe solvent heat exchanger (d) and solvent cooler(b). Amine losses with the lean gas have to bereduced by a water wash step. If the absorptionoperates at elevated pressure it is often econom-ical to recover the absorbed valuable gases inan additional flash stage downstream of heat ex-changer (d).

The amine concentration determines the ca-pacity of the solvent and thus the solvent circula-tion rate. The maximum amine concentration ofthe solution is limited with regard to corrosivityof the rich solvent, viscosity, and other proper-ties. Alkanolamines are susceptible to degrada-tion by several trace components present in the

raw gas and by carbon dioxide. Carbon dioxidecatalyzes the degradation of MEA and DEA toform heat stable salts and inactive polymer reac-tion products, which increase solvent viscosityand foaming tendency. DIPA reacts with carbondioxide to form oxazolidone. The capacity of theprimary and secondary amines has to be main-tained by continuous solvent reclaiming.MDEAas tertiary amine and hindered amines, which donot form carbamates, are not affected by carbondioxide.

Figure 49. Basic flow diagram for amine processesa) Absorber; b) Solvent cooler; c) Solvent pump; d) Heatexchanger; e) Regenerator; f) Condenser; g) Reboiler

Carboxylic acids (e.g., formic acid, aceticacid, oxalic acid) are frequently found in aminesolvents. They are formed directly from cyanideand by degradation of the amine with oxy-gen. At high carbon monoxide partial pressureformic acid is directly produced in alkaline so-lutions. Carboxylic acids tie up amines as heatstable salts (HSS). Sulfur dioxide also reacts ir-reversibly with the amine to formHSS. Heat sta-ble salts reduce solvent capacity and are a majorsource of corrosivity [184]. Common practiceis the addition of caustic soda to liberate theamine and to restore solvent capacity. The ef-fect of caustic addition on the corrosivity of thesolvent is small because the acid anions remainin the solution [185]. The salts of the carboxyl-ic acids have to be removed by ion exchange ordistillation.

Primary Amines. Solutions of MEA wereused almost exclusively for many years to re-move hydrogen sulfide and carbon dioxide fromnatural gas and certain synthesis gases becauseof their high reactivity and the high purity ob-tainable.

Gas Production 103

The disadvantages of MEA are the forma-tion of nonregenerable and corrosive degrada-tion products with carbonyl sulfide and carbondisulfide. Because of the high vapor pressureof MEA a water wash step of the purified gasis normally required. The high heat of reactioncauses high steam demand for hot regeneration.Thus MEA units are rapidly being replaced byother systems. However, the MEA solution isstill preferred for gas streams containing smallamounts of hydrogen sulfide and carbon diox-ide with practically no carbonyl sulfide or car-bon disulfide (e.g., the removal of carbon diox-ide from reformed gas produced in a tubular re-former).

Newly developed corrosion inhibitors allowaqueous concentration of 30 wt % MEA, whichresults in a lower solvent circulation rate. An ex-ample is the Amine Guard process from UnionCarbide. The Amine Guard process is suitableonly for removal of carbon dioxide because sul-fur components react with the inhibitors.

DGA is used in Fluor’s Econamine process.The high DGA concentration (typical 40 – 60wt %) reduces the required solvent circulationrate compared to MEA. Carbonyl sulfide is par-tially removed from the gas and the degradationproducts of DGA can be reclaimed at elevatedtemperature. No water wash step is required dueto its lower vapor pressure [186].

Secondary amines DEA and DIPA arewidely used to treat refinery gases and highpressure natural gases.

DEA is applied in the 20 – 25 wt % concen-tration range for refinery applications. DEA issuperior to MEA for treating of refinery gases,which contain appreciable amounts of carbonylsulfide and carbondisulfide, becauseDEAreactswith these components to a much lesser extentand the reaction products are less corrosive. TheSNEA-DEA process for natural gases was devel-oped by Societe National Elf Aquitaine (SNEA)[187]. A 25 – 30 wt % DEA solution is used re-sulting in higher solvent capacity. Acid gas load-ings up to 0.7 – 1 mol/mol of DEA is possiblewith sufficiently high acid gas partial pressure.The heat of reaction of DEA and carbon dioxideis about 25% lower than that of MEA and car-bon dioxide. Therefore, DEA solutions requireless steam for hot regeneration.

The ADIP process of Shell International OilProducts, licensed by JGC, Stork Engineers &

Contractors, and Chiyoda [188], utilizes DIPAin concentrations up to 50 wt%. ADIP units aresuitable for sweetening natural gas, refinery gas,and Claus tail gas. MDEA is also used in ADIPunits if higher hydrogen sulfide selectivity is re-quired.DIPA is claimed to bemore stable againstdegrading and corrosive trace components likecyanide, oxygen, and sulfur dioxide thanMDEA[212]. Therefore in cases where selectivity is notcritical, DIPA is the preferred solvent for hydro-gen sulfide removal. Raw gas purification andClaus tail-gas treatment can be integrated intoone plant (Fig. 50). One single regeneration col-umn (d) provides lean amine solvent for the finewash section of the high pressure absorber (a)and the Claus tail-gas absorber (e). The tail-gasabsorber operates at low pressure and hydrogensulfide has to be removed selectively to avoid alarge carbon dioxide recycle through the Clausplant. Therefore the rich solvent leaving the tail-gas absorber is not completely saturated and canabsorb acid gas in the lower section of the highpressure absorber (a).

Hydrogen sulfide concentrations in the puri-fied gas below 5 ppm (mL/m3) are achievablewith DIPA since it is easier to regenerate withless steam than MEA and DEA. DIPA has alower tendency to form corrosive degradationproducts than DEA.

Tertiary Amines. MDEA is the most widelyused tertiary amine. It is suitable to purify a largevariety of gases. The carbon dioxide absorptionrate is slower in tertiary amines because they donot form carbamates This makes MDEA partic-ularly suited for selective hydrogen sulfide re-moval from gases with elevated carbon diox-ide contents. Many formulated MDEA basedamines are available on the market, which areclaimed to have improved hydrogen sulfide se-lectivity (e.g.,GasSpec,Ucarsol) [189]. The lowheat of reaction facilitates regeneration down tolow acid gas contents in the lean amine requiredto achieve hydrogen sulfide concentrations be-low 5 mL/m3 in the purified gas. MDEA is ap-plied in the 30 – 50 wt % range. Acid gas load-ings up to 0.8 mol/mol of MDEA is possible incarbon steel equipment. The high MDEA con-centration in conjunction with the high attain-able acid gas loading result in a high capacityof the solvent and a low solvent circulation rate.For this reason many amine units originally de-signed and operated with primary or secondary

104 Gas Production

Figure 50. Integration of high pressure gas purification with Claus tail-gas treatinga) High pressure absorber; b) Solvent heat exchanger; c) Flash drum; d) Regenerator; e) Claus tail-gas absorber; f) Solventcooler; g) Amine filter

amines were converted to use MDEA as solventin order to increase plant capacity and reduceutility consumption [190]. As a guideline, thesteam required for regeneration per volume sol-vent is approximately 80 kg/m3 for bulk carbondioxide removal and 120 kg/m3 for selective hy-drogen sulfide removal.

MDEA is widely used in refineries for hydro-gen sulfide absorption from various medium- tohigh-pressure off-gas streams from hydrotreat-ing units and the hydrocracker. MDEA is alsosuitable for selective absorption of hydrogen sul-fide from low-pressure gases and is therefore ap-plied in the Claus tail gas treating process.

The carbon dioxide absorption rate inMDEAsolutions is enhanced by addition of smallamounts of primary and secondary amines[191]. This effect is employed for example inthe aMDEA process of BASF and the activatedMDEA process of Elf Aquitaine. The carba-mate formedwith the activator component is hy-drolyzed, giving bicarbonate and recovering theactivator. The consecutive carbamate formationand hydrolysis reactions are much faster thandirect bicarbonate formation with water.

Activator-carbamate reaction:

CO2+R1R2NH�R1R2NCOO−+H+

Carbamate hydrolysis:

R1R2NCOO−+H2O�R1R2NH+HCO−3

Part of the absorbed carbon dioxide can beliberated from aqueousMDEA solution by pres-sure reduction. Flash regeneration is sufficientif only bulk carbon dioxide removal is required.Lower carbon dioxide concentrations in the pu-rified gas are achievable with an additional finewash section, which has to be fed only with asmall stream of steam regenerated solvent (Fig.51). This makes the activatedMDEA process at-tractive in terms of steam consumption. A typ-ical aMDEA plant for the removal of carbondioxide from steam reformer gas is shown inFigure 52.

The largest acid gas absorbers were built fora liquefied natural gas (LNG) terminal with asingle train capacity of 1.6× 106 m3(STP)/h.Production of LNG requires bulk removal of allsulfur and carbon dioxide to avoid freezing inthe liquefaction units. The aMDEA process ofBASF is used to remove simultaneously hydro-gen sulfide and carbon dioxide to the low ppmlevel. Additional molecular sieve adsorbers traporganic sulfur components such as mercaptans.

Degradation products of MDEA are sec-ondary and primary amines, which also serve asactivator for carbon dioxide absorption. There-fore in selective MDEA applications the carbondioxide absorption is often higher than expected.

Gas Production 105

Figure 51. Activated MDEA process for carbon dioxide removal with semi lean amine regenerated by flash a) High pressureabsorber; b) First flash drum; c) Second flash drum; d) Solvent heat exchanger; e) Regenerator; f) Solvent cooler; g) Aminefilter

Figure 52. aMDEA plant removing carbon dioxide from 15000 m3(STP)/h raw gas of a steam reformer (Courtesy ofBASF AG and Lurgi Oel Gas Chemie GmbH)

Sterically Hindered Amines. Solvents basedon sterically hindered amines are licensed byExxon Research and Engineering Co. under thetrade nameFlexsorb [192, 193]. These stericallyhindered amines are secondary aminemoleculesbearing one bulky side group. This steric hin-drance reduces the carbamate reaction rate bymore than an order of magnitude compared tosecondary amines. The basicity of this hinderedsecondary amine is higher than that of MDEAresulting in a higher hydrogen sulfide absorptioncapacity. Therefore, very high selectivity is at-tained in absorberswith short contact time [194].

Flexsorb SE solvent is a strongly hinderedamine designed for high hydrogen sulfide selec-tivity [195]. It is therefore best suited for gaseswith high carbon dioxide: hydrogen sulfide ratioand in applicationswhere lowcarbondioxide ab-sorption is required, such as Claus tail-gas treat-ing. The benefits of Flexsorb have been demon-strated in various low pressure retrofit appli-cations. Higher gas throughputs were achievedwith lower solvent circulation rates comparedto the original solvent. In addition the steam re-quirement for regeneration is reduced due to thelower solvent circulation rate.

With Flexsorb SE Plus hydrogen sulfide con-centrations as low as 10 ppm (mL/m3) at atmo-spheric absorber pressure are possible at the costof 25% higher circulation rates relative to Flex-sorb SE. Acid buffer is added to the solution,which facilitates solvent regeneration.

Hot Potassium Carbonate Process. Thepotassium carbonate process has been success-fully applied for carbon dioxide removal since1960. The hot potassium carbonate process (HP)uses a potassium carbonate concentration inaqueous solution of 20 – 40 wt %. The reactionwith carbon dioxide that occurs in the processalso shows an equilibrium behavior that is fa-vorable to absorption even at elevated tempera-ture. As a result of the high absorber temperature(near the atmospheric boiling point of the solu-tion), relatively low heat input is required forregeneration by reboiling.

106 Gas Production

Figure 53. Typical flow diagrams of the HP process for CO2 removalA) Single stage; B) Single stage with splitflow; C) Two stage process

Investment costs are lower than forMEA andDEA because solvent heat exchangers are notrequired. Figure 53 shows three typical versionsof the process. The single stage type (Fig. 53 A)can be varied by cooling part of the solvent tolower the vapor pressure of carbon dioxide sothat a higher purity of the treated gas can be ob-tained (Fig. 53 B). A carbon dioxide content ofless than 0.5% in the treated gas can be obtainedeconomically only by using a two-stage design(Fig. 53 C).

The main application of the HP wash is re-moval of carbon dioxide from converted reform-ing gas, because process conditions are almostideal for this. The carbon dioxide partial pres-sure after conversion is normally in the rangeof 0.4 – 0.7 MPa and, therefore, in the optimumrange for equilibrium behavior of the solution(see Fig. 42, line e). In the production of hydro-gen and synthesis gas for ammonia, a methana-tion stage follows as a fine purification stage forconverting carbonmonoxide and carbon dioxideresidues. This step is operated at high tempera-ture and in the presence of water, as in carbonmonoxide conversion. This is in keeping withthe high absorption temperature, as well as theaqueous solvent basis of the HP process. In viewof the 0.25 – 0.4% carbon monoxide in the gasdownstreamof the shift conversion, the gas in theabsorption stage does not have to be cleaned toa carbon dioxide content of less than 0.1%. Theconverted gas contains enough steam to coverthe heat requirements of the HP process.

Figure 54 shows a characteristic working di-agram for a two-stage HP absorption process. Inaddition to the equilibrium curve, which is prac-tically identical for absorption and regeneration,operating lines of the absorber and the regener-ator, together with determination of the theoret-ical stages according to McCabe – Thiele, areshown. The slopes of the operating lines are de-termined by the split fraction of semi-lean andlean solvent. Optimum adjustment of the splitfraction results in the minimum number of the-oretical stages required.

The process is available in different versions.Modifications differ less in the basic principlesof the process than in the additives used as cor-rosion inhibitors or activators to increase masstransfer.

Modifications. The Benfield process usesmainly solutions with 25 – 30% potassium car-bonate and vanadate as an additive. Several de-velopments have led to its widespread commer-cial application, especially with regard to am-monia plants based on natural gas and naphtha[196].

The solution used in the Catacarb processcontains proprietary additives.Amine promotersenhancemass transfer rates and inhibit corrosionin applications such as purification of ammoniasynthesis gas, hydrogen production, and naturalgas. If no carbon dioxide is present a solutionof potassium borate is used for hydrogen sulfideremoval instead of potassium carbonate.

Gas Production 107

Figure 54. McCabe – Thiele diagram for a two-stage HPprocess

The Lurgi HP process uses borate as an ad-ditive, which accelerates absorption and is ef-fective as a corrosion-inhibiting compound. Inthis process modification, both carbonyl and hy-drogen sulfides can also be absorbed in the hotpotassium carbonate solution. In the high-purityversion, for example, hydrogen sulfide is re-moved to < 1 ppm. It can be absorbed selec-tively against carbon dioxide under certain con-ditions because of their different reaction veloc-ities. The selectivity, together with other proper-ties of the HP process (high absorption temper-ature, low investment and utility costs), is par-ticularly advantageous for purifying coal gasifi-cation gas to fuel gas quality. Because coal gasusually contains additional impurities that can-not be removed in an HP wash, a precleaningstep is required (see Section 5.4.4).

The regeneration operates under vacuum inthe Vacasulf process of Krupp Uhde. The pro-cess is mainly applied for low pressure desul-

furization of coke oven gas. Stringent lean gasqualities are achieved by a fine wash step withcaustic soda.

Absorption in Ammonia Solution andCaustic Soda. Desulfurization with ammoniasolution to purify gases containing ca. 500 mg ofhydrogen sulfide per cubic meter was formerlyvery common for treating coke oven gas andis still used in some cases. However, the pro-cess has not become established for cases withmore stringent purity requirements. On the otherhand, it is an integral part of new developmentsfor separating hydrogen sulfide – carbon diox-ide and ammonia for sulfur recovery from sourgases containing ammonia, and for recoveringammonia from gas condensates.

Absorption of carbon dioxide and hydrogensulfide in caustic soda produces bicarbonate andbisulfide the bonding of which is so strong thatthe vapor pressure of the sour components prac-tically disappears. Therefore, gas purities as lowas 0.1 mL/m3 (CO2 and H2S) or less can be at-tained. However, regeneration, which is possiblein principle but very complex, is dispensed with.Caustic soda is thus used only for fine purifica-tion.

5.4.2.3. Physical – Chemical AbsorptionProcesses

Processes utilizing a solvent mixture of a phys-ical organic component and a chemically actingcomponent are termed physical – chemical ab-sorption processes or mixed solvent processes.Mixed solvents provide a higher capacity thanthe pure physical solvent because of the addi-tional chemical bonding of the acid gases withthe chemical acting constituent. Fewer stages arenormally required in the absorber due to the ac-celeration of mass transfer by the chemical re-action. The physically acting constituent of themixed solvent provides absorption capacity fortrace components such as carbonyl sulfide andorganic sulfur components, which can not be re-moved with aqueous amines.

The best known processes are the Shell Sulfi-nol process and the Lurgi Amisol process.

Sulfinol Process. The Sulfinol solvent is athree-component solution containing the fol-lowing constituents: an alkanolamine, usually

108 Gas Production

DIPA (diisopropanolamine), sulfolane (tetra-hydrothiophene dioxide), and water. The ratioof the constituents varies according to the re-quirements of the application. The process, in-troduced in 1964, was used in its early yearsto remove hydrogen sulfide and carbon dioxidefrom gasification gases in steam reforming, aswell as from natural gases [197]. More than 180plants are reported to be in operation in 1998.

The process displays many properties andadvantages of both chemical and physical pro-cesses. Acid gas loading is higher at elevatedpressure and less steam for regeneration is re-quired compared to purely chemical solvents.The physical character of the solvent facilitatesto achieve hydrogen sulfide concentration as lowas 1 ppm (mL/m3) in the purified gas. Sulfinolis capable of removing other sulfur compoundsincluding mercaptans, carbonyl sulfide, and car-bon disulfide. The process flowsheet is very sim-ilar to chemical absorption processes (Fig. 49).Flashing of the rich solvent is often included asadditional step to recover valuable gases and toincrease the hydrogen sulfide concentration inthe Claus gas.

The Sulfinol-M process uses MDEA insteadof DIPA as chemical constituent in the solvent.This process canbedesigned for selective hydro-gen sulfide absorption while a large percentageof the carbon dioxide from the raw gas slips tothe treated gas.

Amisol Process. The Amisol solvent con-sists of an alkanolamine, normally MEA orDEA, in conjunction with methanol. This sol-vent mixture was successfully used in the firstcommercial Amisol plant in the early 1970s toremove carbon dioxide, hydrogen sulfide, andcarbonyl sulfide from raw gas produced in a par-tial oxidation plant handling vacuum residues[198]. A clean gas with very low total sul-fur content (hydrogen sulfide, carbonyl sulfide,mercaptans) of < 0.1 mL/m3, as required forsynthesis gas, can be obtained by the Amisolprocess. Good absorption of trace componentssuch as hydrogen cyanide, and higher hydrocar-bons is also achieved. Because of the low boil-ing point of methanol, absorption and regener-ation temperatures are not far apart; therefore,no lean – rich solvent heat exchangers are nor-mally required. Selective removal of both car-bonyl and hydrogen sulfides versus carbon diox-

ide is possible by using aliphatic alkylaminessuch as diisopropylamine (DIPAM) and diethyl-amine (DETA) [199].

The Amisol process is suitable for treatingvarious raw gases to absorb hydrogen sulfide,carbonyl sulfide, and carbon dioxide along withtrace contaminants or to selectively remove sul-fur compounds versus carbon dioxide. An exam-ple of a commercial Amisol plant using diethyl-amine involves purification of coal gas producedby an entrained-bed gasifier. A concept was de-veloped for clean fuel power stations.

Flexsorb PS. The Flexsorb PS solvent is amixture of amoderately hindered amine, a phys-ical solvent, and water. This solvent is primarilysuited for natural gas clean up. More than 95%of carbonyl sulfide and mercaptans can be re-moved from the gas.

5.4.2.4. Comparison of Various AbsorptionProcesses for Hydrogen Production andIGCC

Hydrogen Production Based on ResidueOil Gasification (see also 7.2). Refineries re-quire large amounts of hydrogen for hydrotreat-ing and hydrocracking to meet stringent fuelspecifications and toprocess heavier opportunitycrude oils. The residues from visbreaker or hy-drocracker units are an inexpensive feedstock forhydrogenproductionbygasification.The carbonmonoxide from the raw gas is converted to hy-drogen in a CO shift unit and the resulting rawhydrogen must be purified from acid gases andtrace impurities. The selection of the appropri-ate gas purification technology depends on therequired hydrogen purity, the value of the fuelgas and CO2 byproducts, and the plant emissionregulations.

Acid gasesmust be removed from the raw hy-drogen downstream of the sulfur-tolerant shiftunit. This can be achieved with various pro-cesses.

In the first case (see Fig. 55A), the Rectisolprocess produces an H2S-rich sour gas suitablefor theClaus unit. In addition, CO2 bulk removalis economical and a pure CO2 stream is gener-ated which can be used for sequestration or asindustrial gas, or vented without further treat-ment. A pressure swing adsorption (PSA) unit is

Gas Production 109

Figure 55. Gas purification options for hydrogen production from gasification of heavy residue oilA) Rectisol-based acid gas removal and PSA for hydrogen fine purification; B) Selective acid gas removal and PSA forhydrogen purification; C) Acid gas fine removal by Rectisol and methanation for hydrogen purification

used to free the hydrogen from impurities suchas methane, nitrogen, carbon monoxide, and ar-gon, which are concentrated in the PSA off-gas.The PSA off-gas is a high-BTU, low-sulfur gas,since Rectisol removes besides H2S also eco-nomically virtually all the CO2 and other traceacid components.

In the second case (see Fig. 55B) a selectiveacid gas removal (AGR) process is applied suchas Purisol or Selexol, which also produces anH2S-rich sour gas suitable for the Claus unit.Typically, 20–30%of theCO2 content of the rawgas is removed in a selective gas purification pro-cess. This CO2 is produced as a separate, impureCO2 stream with a significant sulfur concentra-tion. The sulfur impurity in the CO2 stream isnormally not a problem for sequestration pur-poses or enhanced oil recovery. Simple venting,however, is usually not permitted.

Since AGR is selective, a considerableamount of CO2 passes to the PSA unit in thiscase and is separated there. The PSA unit alsoseparates the remaining sulfur impurities. Dueto the high CO2 content the PSA off-gas hasa low heating value. It also contains a signifi-cant amount of sulfur because both Selexol andPurisol leave traces of H2S in the low ppm rangein the purified gas and remove carbonyl sulfide(COS) only partly.

In the third case (see Fig. 55C), if hydrogenproduction is to be maximized but high purityis not required and additional fuel gas is of lowvalue, the PSA unit can be replaced by a metha-nation reactor which converts CO and CO2 to

methane. Since all impurities end up in the prod-uct hydrogen, the gas purification unit must re-move all sulfur species to the ppb level and re-move bulkCO2 to the low ppm range. Rectisol isthe process of choice for this configuration. Hy-drogen purity exceeding 97% can be achieved,which is normally sufficient for hydrotreatingand hydrocracking applications, and no fuel gasis produced. This process configuration gives thehighest specific hydrogen production.

Gas Purification for Power Production byCombined Cycle IGCC. In an integrated gasi-fication combined power cycle (IGCC), fuel gasproduced by partial oxidation of coal or residualoil is used to power a gas turbine (see 7.3). Agasification process is advantageously appliedwhich recovers the energy of the hot gas leav-ing the reactor in a waste-heat boiler to producehigh-pressure steam. The heat recovered fromgasification and the exhaust gas of the gas tur-bine are utilized for a steamcycle (seeFig. 56). Ifthe IGCC is installed in a refinery, additional hy-drogen can easily be produced. Part of the desul-furized gas downstream of the acid gas removalis passed to a CO shift unit and the hydrogen ispurified in a PSA unit.

Due to mechanical restrictions gas turbinesusually operate at pressures below 30 bar,whereas gasification and gas cleaning oper-ate advantageously at pressures above 50 bar.Higher operating pressures have the benefit ofsmaller equipment size. Additionally the in-creased solubility of the sulfur compounds in the

110 Gas Production

Figure 56.Cogeneration of power (IGCC) and hydrogenWHB-waste-heat boiler, AGR-acid gas removal, HRSG-heat recoverysteam generation, PSA-pressure swing adsorption, MARS-metal/ash recovery system

absorbing solvent at higher pressure results in alower solvent circulation rate that reduces utilitydemand for regeneration and solvent cooling.

The pressure difference between gasificationand gas turbine can be utilized for additionalpower generation by expansion of the purifiedgas downstream of gas cleaning in a gas ex-pander (see Fig57c). During expansion withpower generation the gas temperature is reducedconsiderably. The expanded cold gas can be usedfor solvent cooling (d). The Rectisol process canbe advantageously used for fuel gas productiondue to its ability to remove all sulfur compo-nents, including carbonyl sulfide. The requiredcooling to low temperatures can be done com-pletely by the cold gas downstream of the expan-sion, if the pressure difference between gasifica-tion and gas turbine is high enough.

A comparison of a process with low-pressuregasification, gas purification, andgas turbine anda processwith high-pressure gasification and gascleaning, gas expansion, and low-pressure gasturbine is compiled in Table 29. The calculationis based on a residual oil gasification of 100 t/hand a net electrical power generation of 500MW.Capital expenditures only for the gas cleaningunit are increased by 17% by the gas expansionequipment, but for the complete IGCC plant thecapital expenditures are increased only by 2%.The utility demand is decreased by 13%, andthis leads to an overall increase of the electrical

power generation efficiency of the IGCC from45.4 to 46.2%.

Table 29. Comparison of low pressure IGCC and high pressureIGCC with gas expansion

High pressure gascleaning withexpansion

Low pressure gascleaning

Gasifier pressure 58 bar 24 barGas turbine pressure 22 bar 22 bar

Utility consumptionCooling water for gas

cleaning180 m3/h 1230 m3/h

Steam for regeneration 8.6 t/h 13 t/hElectrical power for gas

cleaning− 0.4 MW 8.1 MW

Electrical power for airseparation

32.6 MW 29.1 MW

Total electrical power 32.2 MW 37.2 MWCapital expenditures (relative)Gas cleaning unit 117 100Complete IGCC plant 102 100

Efficiency of electrical power generationOverall efficiency 46.2% 45.4%

An off-gas-free selective desulfurization forfuel gas production within an IGCC applyinga selective absorption process such as Selexoland Purisol is shown in Figure 58. This processconfiguration operates successfully in an IGCCplant in Italy based on heavy residue gasifica-tion. Undesired trace components of the gasifi-cation such as cyanides or higher hydrocarbonsare removed in a prewash stage (a). The rich sol-vent of the main wash is flashed in the reabsorp-

Gas Production 111

Figure 57. Rectisol for integrated gasification combined power cycle (IGCC) with gas expansion a) Crude gas cooling;b) H2S absorber; c) Gas expansion turbines; d) Solvent cooling by expanded gas; e) Solvent cooling by refrigerant, f) Flashregeneration, g) Hot regeneration

Figure 58. Off-gas free Purisol process for selective hydrogen sulfide removal in an integrated combined power cycle (IGCC)a) Prewash section, b) H2S absorber, c) Reabsorber, d) Flash regeneration, e) Hot regeneration, f) Claus tail gas recycle blower,g) Fuel gas recycle compressor (optional)

tion column (c). Carbon dioxide and valuablecomponents are recovered as fuel gas, whereashydrogen sulfide is reabsorbed in the solvent inorder to increase the hydrogen sulfide concen-tration in the sour gas to the Claus unit. Thefuel gas leaving the reabsorption column over-head can either be recompressed and added tothe high-pressure purified gas or can be usedas low-pressure fuel gas for the steam boiler ofthe IGCC. The typical sulfur recovery rate of atwo-stage Claus unit is 95 - 96%. The off-gasof the Claus unit contains beside unconvertedhydrogen sulfide considerable amounts of SO2and COS, which are converted to hydrogen sul-fide in a catalytic hydrogenation step. The hy-

drogenated off-gas is compressed and fed to thereabsorption step of the acid gas removal unit.

So-called zero-emission IGCC (Fig. 59) isdiscussed in view of the Kyoto protocol for re-ducing carbon dioxide emissions. In addition tototal sulfur removal, carbon from the feedstockis also captured and concentrated in a secondpure carbon dioxide product stream. The fuelgas to the gas turbine is pure hydrogen. Thegasification process converts the carbon of thefeedstock mainly to carbon monoxide, which isfurther converted to carbon dioxide and hydro-gen in a carbon monoxide conversion unit (→5.1). The carbon dioxide can easily be separatedfrom the hydrogen in the absorption process and

112 Gas Production

Figure 59. “Zero-emission” IGCC concept WHB-waste-heat boiler, AGR-acid gas removal, HRSG-heat recovery steamgeneration, PSA-pressure swing adsorption, MARS-metal/ash recovery system

a pure carbon dioxide stream is produced, whichcan be used for industrial purposes such as ureaproduction, for enhanced oil recovery, or can besequestered for carbon dioxide emission control.The economically achievable carbon recoveryrate and the purity of the carbon dioxide productstream depend on the selected absorption pro-cess. The Rectisol process allows the productionof ultralow-sulfur fuel gas (hydrogen) to the gasturbine and generation of very pure carbon diox-ide. Selective-solvent processes such as Selexoland Purisol leave more sulfur in the fuel hydro-gen. As expressed by the designation selective,this type of solvent requires greater effort for re-moval of bulk carbon dioxide. In addition, thecarbon dioxide product stream contains largeramounts of hydrogen sulfide. Chemical absorp-tion processes are not particularly suitable forremoving these large amounts of carbon diox-ide due to the large amount of steam requiredfor solvent regeneration.

5.4.3. Liquid-Phase Oxidation Processes

In liquid phase oxidation processes hydrogensulfide is absorbed from a gas stream in a liquidsolution and directly oxidized to sulfur in the so-lution. The overall reaction can be expressed asfollows:

H2S+1/2O2 → S+H2O

The oxidizing agent in the solvent is reducedand solid sulfur is formed in the gas liquid con-tactor. In the subsequent regeneration step thesolution is oxidized with air converting the oxi-dant to its oxidized state.

The main process steps are shown in Figure60. The feed gas is contactedwith the redox solu-tion in a suitable gas – liquid contactor such asa countercurrent packed bed absorber, a spraytower, a cocurrent sparged tower, or a venturiscrubber. The sulfur formation takes place in theabsorber and is completed in a downstream reac-tion vessel which may also serve as gas – liquidseparator if the absorber works in the cocurrentmode. In the oxidizer the active component isconverted from the reduced to the oxidized stateand the sulfur particles are separated either asfroth by flotation or as slurry by settling. Theregenerated solution is recycled to the gas – liq-uid contactor. The sulfur froth or slurry is eitherseparated in a filter unit or a sulfur melter.

Liquid-phase oxidation processes are espe-cially suitable for selective removal of hydrogensulfide versus carbon dioxide, because carbondioxide does not react with the redox solution.Essentially complete hydrogen sulfide removalwith concentrations below1 ppm(mL/m3) in thetreated gas is attainable. Contrary to physical orchemical absorption processes in which the off-gas of the regenerator must be generally treated

Gas Production 113

Figure 60. Schematic flow diagram of liquid phase oxidation processes

in an additional unit to produce sulfur or sulfu-ric acid, in liquid-phase oxidation processes theformation of elemental sulfur occurs within thepurification unit itself.

The main applications of liquid-phase oxida-tion processes are

1) Desulfurization of coke oven gas, natural gas,and synthesis gas with low hydrogen sulfideconcentration,

2) Treatment of gases that cannot be processedeconomically in a Claus plant (e.g., due tolow hydrogen sulfide concentration, small to-tal amount of sulfur),

3) Treatment of Claus tail gases,4) Replacement of dry purification processes,

and5) Treatment of very low raw gas rates (e.g.,

landfill gas).

Economical operation is limited to sulfur pro-duction rates less than 20 t/d due to high utilitycosts. Continuous solvent make-up is requiredbecause the solvent is susceptible to chemicaland thermal degradation and because of losseswith the sulfur filter cake. Continuous solventpurging is required due to the accumulation ofside products such as thiosulfate in the solventcirculation. Thiosulfate production can be keptbelow 1% of the hydrogen sulfide converted, ifoperating conditions such as temperature and so-lution pH are carefully controlled. A bufferingagent in the solution avoids large pH swings dueto carbon dioxide absorption.

Carbonyl sulfide and organic sulfur compo-nents are only partly removed in liquid oxidationprocesses. Hydrogen cyanide and sulfur dioxide

are almost quantitatively absorbed and react tothiocyanate and sulfate in the solution. Thesesalts accumulate in the circulating liquor and in-crease the required purge stream and the amountof hazardous waste byproducts generated. Am-monia is also absorbed and causes pH changes.Ammoniumpolysulfide is formedwith solid sul-fur, increasing the salt concentration in the so-lution. Ammonia is stripped in the oxidizer andmay cause odor problems in the vent stream.

Sulfur quality is often not crucial because ofthe small amount of sulfur produced. Residualsolvent impurities have to be removed from thefilter cake bywashingwith demineralized water.An additional sulfur washing and melting stepis required to achieve sulfur quality similar toClaus plant sulfur.

A drawback of all liquid-phase oxidation pro-cesses is the relatively low oxidation capacity ofthe solution for hydrogen sulfide. This necessi-tates large solvent circulation rates. The heat ofreaction generated by the oxidation of hydrogensulfide must be dissipated at low temperatureand cannot be used, as in the Claus process, togenerate valuable steam. Only part of the heatof reaction can be used for the heat demand ofthe process.

The development of liquid phase oxidationprocesses started with iron oxide suspended inaqueous alkaline solutions (e.g., Ferrox, Gluud,and Manchester). The second generation ofprocesses applied thioarsenate (e.g., Thylox,Giammarco–Vetrocoke [200]). They have lostmarket acceptance because of the toxicity ofthe solution. An aqueous mixture of vanadiumpentoxide, naphthoquinone, and alkali carbon-

114 Gas Production

ate is the basis of the widely used Stretford pro-cess [201], the Takahax process [202] and theSulfolin process. Presently, environmental con-cerns about the vanadium content in the purgestream have resulted in the selection of ironchelate-based processes such as Lo-Cat [203],SulFerox [204] and Sulfint [205] for most ap-plications. Processes using sulfur dioxide areTownsend [206] and IFP Clauspol [207]. In theend they have not been commercially successful.A new class of processes apply microorganismsto convert hydrogen sulfide to sulfur (e.g., Shell-Thiopaq [208] and Bio-SR).

Vanadium-Based Solutions. The Stretfordprocess, originally developed for the removal ofhydrogen sulfide from coke oven gas, has provedto be equally suitable for desulfurization of othergas streams such as natural, synthesis, and tailgases.

The Stretford process uses an aqueous solu-tion containing sodium carbonate, sodium bicar-bonate, anthraquinonedisulfonic acid (ADA),and sodium metavanadate. The process can beconsidered to proceed in the following steps:

1) Absorption of hydrogen sulfide in the alkalisolution,

H2S+Na2CO3 →NaHS+NaHCO3

2) Oxidation of hydrogen sulfide to sulfur bysodium metavanadate,

4 NaVO3+2NaHS+H2O→Na2V4O9+2 S

+4NaOH

3) Reoxidation of the reduced vanadate by ADA

Na2V4O9+2NaOH+H2O+2ADA→4 NaVO3+2ADA(reduced)

4) Oxidation of reduced ADA by air in the re-generator

2 ADA (reduced)+O2 → 2 ADA+2H2O

Figure 61 shows a simplified flow diagramof the Stretford process for selective desulfur-ization. The raw gas is brought in contact coun-tercurrently with the solution in the absorber (a)and purified to a level of < 1 ppm of hydrogensulfide. The solution flows from the absorberbottom to the reaction vessel (c) to complete theconversion of hydrogen sulfide to elemental sul-fur; then the solution is regenerated in the oxi-dizer (d) by contact with air.

Sulfur is separated by flotation and removedat the top of the oxidizer (d) as froth containing5 – 10% solids. The sulfur froth is collected ina vessel and subsequently processed in a filteror centrifuge. The sulfur has to be washed withwater to separate solution and sulfur and to pro-duce relatively pure sulfur. Though installed innumerous commercial plants, the Stretford pro-cess suffers from the drawbacks of formationof byproducts and environmental concern aboutvanadium in the effluent streams.

In the solvent used in the Sulfolin processADA is replaced by an organic nitrogen vana-dium promoter [209]. Like the Sulfolin processthe Unisulf process [210] claims a reduced for-mation of byproducts and lower corrosivity, re-sulting in lower chemical consumption.

Iron-based Solutions. Iron is kept in theaqueous solution by organic ligands or chelat-ing agents (L). The basic chemistry in the ironchelate solution can be summarized as follows:

Absorption:

2 Fe3+L+H2S→ 2 Fe2+L+S+2H+

Regeneration:

2 Fe2+L+1/2O2+2H+ → 2 Fe3+L+H2O

In iron chelate processes sulfur is generallyseparated by settling. The slurry containing 10 –15 wt % sulfur is further processed by filtrationin order to recover the solvent. Stainless steelequipment is required due to the high solventpower of the solution for iron.

The Lo-Cat process is licensed byU.S. Filter.The flow sheet of the conventional Lo-Cat pro-cess is similar to the SulFerox process shownin Figure 63. Oxidation of the rich solvent andsulfur settling take place in the same vessel. Thesolution contains up to 0.3 wt % iron resulting inlarge solvent circulation rates with high powerdemand for solvent pumping especially in highpressure gas purification.

In the Lo-Cat autocirculation design, hydro-gen sulfide absorption, the conversion to sul-fur, and the oxidation of the redox solution takeplace in one vessel in separate compartments(Fig. 62). The autocirculation design is appli-cable for low pressure gases such as acid gasesfrom amine units and Claus tail gases. Due to thecompact design hydrogen sulfide is not totally

Gas Production 115

Figure 61. Schematic of a Stretford plant for selective H2S removala) Absorber; b) Air compressor; c) Reaction vessel; d) Oxidizer; e) Solvent vessel; f) Sulfur filter

converted to sulfur before the solution is con-tacted with air. This causes an increase in thethiosulfate formation of up to 5% of the con-verted hydrogen sulfide.

Figure 62. Lo-Cat autocirculation design

The Lo-Cat process operates with the stoi-chiometric ratio of iron(III): hydrogen sulfide of2: 1. In the Lo-Cat II process the ratio of iron(III)to hydrogen sulfide is maintained at 1: 1. An im-proved solution chemistry enhances the reactionkinetics. The rich solvent from the absorber ismixed with oxidized solution upstream of theoxidizer. Residual HS− in the rich solvent isthus totally converted to sulfur by iron(III) be-fore the solution is contactedwith oxygen result-ing in thiosulfate formation below 1%. Due to

the substoichiometric operation less solvent hasto be circulated to the high pressure absorber, re-ducing the power requirement for pumping. TheLo-Cat II process is also available in the auto-circulation design.

The SulFerox processwas developed by Shelland is licensed by DOW and Shell (Fig. 63).Special feature of this process is the high ironconcentration of 2 to 4 wt %, reducing equip-ment sizes and pumping costs. The high ironconcentration accelerates the reaction kineticsand thus the hydrogen sulfide absorption rate.One possible design for the absorption sectionis the patented cocurrent pipeline contactor withproprietary mixing elements which makes useof the fast reaction kinetics. Low pressure drop,nearly infinite hydrogen sulfide turndown capa-bility, and low fouling tendencies are providedin a small diameter vessel. Due to the fast re-action kinetics all of the hydrogen sulfide ab-sorbed is converted to sulfur and no HS− ionsare present in the rich solutionwhen entering theoxidizer. Thiosulfate formation is claimed to bebelow 0.1%.

The pH of the SulFerox solution is kept be-low 8. Carbon dioxide solubility in the SulFeroxsolution is therefore similar to that in water, re-sulting in a very high hydrogen sulfide selec-tivity. The SulFerox solution is also capable ofremoving about 70% of mercaptans present inthe raw gas.

A comparison between one of the amine pro-cesses plus Claus plant, autocirculation Lo-Cat,direct Lo-Cat, Lo-Cat II, and the SulFerox pro-

116 Gas Production

Figure 63. Schematic flow diagram of the SulFerox process. A) Absorber, b) Gas-liquid separator, c) Flash drum, d) oxidizer,e) Sulfur settling vessel, f) Sulfur filter (Courtesy of Shell International Oil Products B.V.)

cess in technical and economical terms was per-formed by the Gas Research Institute (GRI)[213]. Lo-Cat II and SulFerox perform very sim-ilar for the two high pressure gas cleaning casesstudied. The two Lo-Cat modifications and anamine process plus Claus plant result in signifi-cantly higher sulfur production costs.

Biological Processes. Suitable bacteria con-vert hydrogen sulfide under certain conditions tosulfur. The processes consist of the same processsteps as conventional liquid phase oxidation pro-cesses (Fig. 60). In the Shell-Thiopaq processhydrogen sulfide is absorbed into an iron-freecaustic solution. The dissolved sulfide is sub-sequently oxidized to sulfur in the Thiopaq re-actor by bacteria of the genus Thiobacillus. TheBio-SR process of NKKCorporation uses a non-chelated iron(III) sulfate solution. The bacteriaact as catalyst in the oxidizer to enhance the ki-netic of regeneration from iron(II) to iron(III)sulfate. Both processes have demonstrated tobe capable for removing hydrogen sulfide fromhigh pressure natural gases.

5.4.4. Removal of Gas Impurities of LowConcentration

Raw gases contain numerous gaseous impuritiesbesides the acid gases. High boiling hydrocar-bons have to be removed because of dew pointspecifications in the purified gas. Organic sulfurcompounds are harmful tomost catalysts.Unsat-urated hydrocarbons tend to polymerize in thesolvent circulation. Acids, hydrogen cyanide,

and ammonia tend to form salts in the absorptionsolution. Gases from residual oil or coal gasifi-cation often contain volatile metal compoundssuch as mercury and metal carbonyls.

The absorption of impurities depends on thesolvent. Higher hydrocarbons will only be part-ly absorbed in water-based amine solvents andcause foaming problems. Low-boiling impuri-ties are trapped in organic solvents such as Se-lexol and Purisol and can not be removed bysolvent stripping.Rectisol has the advantage thatthe impurities absorbed are purged together withthe process water from the solvent loop. Al-though the concentration of impurities in the gasis often in the ppm range theymay accumulate inthe solvent circulation to levels that cause prob-lems such as fouling, foaming, corrosion, andsolvent degradation. Therefore they are often re-moved in a separate prewash process upstreamof the acid gas removal.

Water Wash Process. Even a simple waterprewash can remove acids, cyanide compounds,and ammonia very effectively from the gas, be-cause their water solubility (e.g., expressed bythe absorption coefficient) is very high (see Fig.64).

However, high solubility also means difficultregeneration, especially because purity require-ments of wastewater are stringent and fresh wa-ter is scarce. Therefore, the laden water must of-ten undergo complex stripping with steam andimpurities must then be incinerated.

Gas Production 117

Figure 64. Solubility of different gases in water at lowpressure 0.1 MPa

Oil Wash Process. Hydrocarbons, espe-cially the normally undesirable unsaturated onesthat tend to polymerize, as well as saturatedhigh-boiling ones, are best removed by using anoil wash, which generally is designed with a hotregeneration stage (see Fig. 41C). Occasionally,the vapor pressure of the washing agent in theabsorber is too high so that a higher boiling oilfraction must be used. If the boiling temperatureof this fraction is too high, it cannot be boiledindirectly in the regenerator with steam but onlydirectly with live steam or a fired reboiler.

In some types of raw gas—especially thatgenerated from coal— removal of the main gasimpurities can require a whole series of pre-wash steps before highly concentrated gas com-ponents such as carbondioxide are separated andchemically converted. Examples are prewasheswith water for hydrogen cyanide and ammonia,followed by absorption of hydrogen sulfide in

aqueous ammonia. In a final fine wash step ben-zene, naphthalene, gum formers, and other un-saturated hydrocarbons are absorbed in an oilwash.

Rectisol Prewash. The previously describedprewash systemswere particularly important forraw gas produced from coal pressure gasifica-tion, which contained almost all the impuritiesnamed, extending to gas naphtha. Therefore, at-tempts were made to find a simpler solution tothis problem, which succeeded in conjunctionwith the development of the Rectisol process:all impurities are removedwith only one solvent,methanol, which is also used downstream for re-moval of hydrogen sulfide and carbon dioxide.

Figure 65 shows a simplified scheme of sucha prewash in which only a small quantity of coldmethanol removeswater, hydrogen cyanide, am-monia, gum formers, gas naphtha, and other im-purities from the raw gas. Methanol is regen-erated by extraction with water and two-stagedistillation.

Condensation. Low boiling substances andwater have to be removed from natural gasesto meet pipeline dew point specification. This isoften accomplished by condensation at low tem-perature. The Ifpexol process (Fig. 66) of InstitutFrancais du Petrol (IFP) removes water and re-covers liquid hydrocarbons from raw natural gas[211]. Part of the raw gas is brought into con-tact with a methanol rich stream in the methanolstripper (a). The methanol concentration in thewater leaving the stripper as bottom product canbe reduced to 50 mg/kg. The methanol vapor inthe overhead raw gas prevents hydrate and iceformation in the gas chiller (b). The liquid con-densate is separated in a settler vessel (d) into theliquid hydrocarbon fraction and the methanol –water fraction. The methanol – water fraction isrecycled to themethanol stripper. The pretreatedgas can be further purified by any suitable pro-cess.

Mercury Removal. Mercury emissionsfrom coal-based power plants are of environ-mental concern. Conventional coal-fired powerstations have to treat a huge volume of flue gasin an activated-carbon adsorber. Power produc-tion based on coal gasification (IGCC) allowseasier removal of mercury. The Rectisol process

118 Gas Production

Figure 65. Rectisol prewash for raw gas from coala) Extraction system; b) Azeotropic distillation column; c) Methanol – water distillation column

Figure 66.Gas pretreating by condensationwith Ifpexol process. a) Methanol stripper, b) Gas chiller, c) Gas – liquid separator,d) Liquid – liquid separator

for acid gas removal also traps mercury withno additional adsorption step required. Otherabsorption processes require an additional acti-vated carbon guard bed, which has to treat thesmaller amount of fuel gas to the gas turbine.

5.5. Adsorption Processes

5.5.1. Fundamentals (→ Adsorption)

Three types of adsorbents are particularly suit-able for the adsorption of gaseous componentsin fixed beds on an industrial scale:

Gas Production 119

1) activated aluminum oxide or silica gel,2) molecular sieves, and3) activated carbon.

Apart from process conditions such as pres-sure and temperature, the degree of adsorp-tion on the above-mentioned solids is deter-mined also by the structure of the adsorbent andthe properties of gaseous components, such asvolatility and polarity. Adsorption of substancesthat are characterized by high volatility and lowpolarity— typical examples are hydrogen andhelium— is very poor or almost imperceptible,whereas water and carbon dioxide, for example,are well adsorbed because of their low volatilityand high polarity. If process conditions are nearthe dew point of a gas component, that compo-nent is adsorbed preferentially.

Adsorption Mechanisms. The qualities ofadsorbents are attributable in part to their char-acteristic material structures. Activated carbon(→ Carbon, Chap. 5), for instance, traps and re-tains gas molecules in minute pores measuringonly a few molecular diameters.

In addition to this minipore effect, watermolecules are adsorbed on silica gel or activatedaluminum oxide in the form of water of crystal-lization, filling vacancies in the crystal lattice. Inthe case of molecular sieves (→ Zeolites), elec-trostatic forces come into play in addition to theminipores and surface-enlargingmicropores, in-creasing their adsorption capacity, especially forpolar and polarizable gases.

Adsorption Conditions. A look at adsorp-tion processes in the molecular range suggeststhat these effects are improved by increasingprocess pressure and decreasing temperature.

For a particular gas and a selected adsorbent,adsorption capacity B is shown in Figure 67 as afunction of pressure p at varying temperature T.This group of curves, which is characteristic ofeach adsorbate and adsorbent, cannot be calcu-lated quantitatively but must be determined byempirical means.

Basic Modes of Operation in AdsorbentBeds. Figure 67 also shows the basic modesof operation possible for a fixed-bed adsorber.Three possibilities exist:

Figure 67. Adsorption capacity B of adsorbents as afunction of pressure and temperature

First Basic Mode of Operation: Path 1 – 2.At T0 and p0, the adsorbent is loaded to the de-gree B1. If the temperature T0 is kept constantand the pressure is increased from p0 to p1, theadsorbent can collect the higher load B2. Thedifference B2 – B1 is the adsorption capacity,which can be used for pressure-swing adsorp-tion (PSA).

Second Basic Mode of Operation: Path 3 –4. While the pressure is kept constant at p1, thetemperature is reduced from T2 to T14. This op-eration forms the basis of a temperature-swingadsorption (TSA) process.

Third Basic Mode of Operation: Path 4 – 5.This classic method of adsorption and regenera-tion uses the low temperature T1 and high pres-sure p1 for adsorption, and then regenerates thesystem by decreasing the pressure to p0 and in-creasing the temperature to T2 by heating. Theadsorption capacity, described by the differenceB4– B5, is clearly the highest of the three cases.

Possible Applications. For gas condition-ing, the adsorption of gases on fixed beds hasthree possible applications:

1) elimination of undesirable trace impuritiesfrom gas mixtures,

2) removal of contaminating gaseous compo-nents to yield a highly pure finished product,and

3) separation of substances.

120 Gas Production

Industrially useful applications can be derivedonly from an appropriate combination of gascomponents to be adsorbed and effective adsor-bents. The following text describes a few typicalprocesses used in practice.

5.5.2. The “Classic” Method

The classic method (third mode of application,above) may be illustrated by means of the re-moval of water and carbon dioxide traces.

If gas mixtures such as hydrogen – car-bon monoxide, hydrogen – methane, nitrogen –methane, or air are separated into their individ-ual components in cryogenic processes, tracesof water and carbon dioxide will lead to icingand clogging of equipment. For this reason, suchsubstances must be removed completely fromthe gas mixture at ambient temperature.

High carbon dioxide contents in the gas mix-ture are normally reduced by chemical or phys-ical wash processes (see Section 5.4) to a leveldetermined by the optimum arrangement of thewash process with a view to efficient subsequentadsorption.

The moisture content of the gas—as shownin the example in Figure 68—can be signifi-cantly decreased by lowering the temperatureto near the freezing point (ca. + 5 ◦C). The wa-ter and carbon dioxide traces remaining in thegas mixture can then be adsorbed efficiently andcompletely on a fixed bed.

Process Description.As shown in Figure 68,most carbon dioxide is removed from the gasmixture in a carbon dioxide scrubber, and thegas is cooled to ca. +5 ◦C. Condensed water isthen removed in a downstream gas separator.

Loading Period. The gas mixture containinga residual quantity of 100 – 150 ppm of carbondioxide and water is fed to adsorber (a), flowingthrough it from bottom to top. For water adsorp-tion, adsorbers (a) and (b) may be filled with ac-tivated aluminum oxide, silica gel, or molecularsieves. The latter have been found most effec-tive for adsorbing carbon dioxide. After carbondioxide and water have been removed, the gas isfed to the cryogenic section.

Partial Regeneration by Flashing. At thesame time, the pressurized adsorber (b) is iso-lated from the gas flow at working temperatureand depressurized to regeneration pressure. The

gas in the adsorber is lost to the process. Ad-sorbed water and carbon dioxide, which are lib-erated from the adsorbent as the system is de-pressurized, are removed from the adsorptionbed by the regeneration gas flowing in the oppo-site direction to the gas flow during adsorption.

Fine Regeneration by Reheating. To com-plete regeneration, the adsorber bed is purgedwith hot regeneration gas containing neither wa-ter nor carbon dioxide. In the process, the ad-sorbent is heated to ca. 250 – 300 ◦C. It is thenrecooled with cold regeneration gas in the finalstage. Regenerated adsorber (b) is pressurizedagain with clean process gas and can be used foranother adsorption cycle.

Changeover. When adsorber (a) has beenfully loaded, the gas flow is switched to adsorber(b) while adsorber (a) is regenerated.

Process Conditions. The gas compositionand the length of the adsorption – regenerationcycles determine the size of the adsorption unit.The following conditions in particular must bemet: (1) The total quantity of impurities in thegas mixture must not be too large so that theadsorbent volume—and, along with it, the sizeof the adsorbers—as well as the recycled re-generation gas rate are kept within reasonablelimits. (2) The adsorbent quantity must be ad-equate for the adsorption stage to ensure thatthe heating and cooling time is sufficient for therequired amount of regeneration gas during thesimultaneous regeneration stage. With this ar-rangement, the adsorption and regeneration cy-cles normally require 4, 6, or 8 h.

Other Applications of the Classic Adsor-ber Loop. The system is also used to adsorbtraces of water from gases before they are sentthrough pipeline systems. Moisture is an unde-sirable ballast, leading to corrosion along withcarbon dioxide, hydrate formation with meth-ane, and freezing when cooled. The maximumallowable moisture content, normally specifiedfor water dew points from −5 to −30 ◦C, canbe achieved for small or moderate gas flow ratesat reasonable cost with adsorber units that useactivated aluminum oxide, silica gel, or molec-ular sieves in long adsorption cycles, as well asthe classic method of regeneration. Natural ortown gas may be dried in this way before be-ing fed to the gas grid. Similarly, adsorbers are

Gas Production 121

Figure 68. Reversible adsorber unit to eliminate residual carbon dioxide and watera) Adsorber; b) Adsorber; c) CO2 scrubber; d) Chiller; e) Gas separator; f) Cooler; g) Heater; h) Regenerator gas; i) Cryogenicprocessused to adjust the residual moisture content ofthe control air, instrument air, and compressedair used in plants to a dew point normally of−25to −30 ◦C.

5.5.3. Pressure-Swing Adsorption (→Adsorption, Chap. 7.2)

The principle of the PSA process can be illus-trated by the recovery of highly pure hydro-gen. If a gasmixture containing hydrogen passesthrough a bed of molecular sieves at elevatedpressure and ambient temperature, these sievespreferentially adsorb the high-density and polargas components. Hydrogen, being the compo-nent least amenable to adsorption, remains alonein the gas phase, provided the adsorber bed islarge enough for the gas rate flowing through it.

Process Description. At least two adsorbersare required to ensure continuous production ofhighly pure hydrogen. While one adsorber is inthe adsorption stage, the other is regenerated bydepressurization.

To purge the adsorbent bed, the flashed gasesare discharged from the bed in the opposite di-rection to the flowduring the adsorption stage. Inthis way, the pure gas released toward the end ofthe regeneration stage expels from the adsorberthose components that are adsorbed first duringthe adsorption stage. The regenerated adsorbentbed is repressurized with pure product gas or

with the gas mixture to prepare it for the nextadsorption stage.

Process Conditions. Depressurizing and re-pressurizing adsorbers requires less time thanheating and cooling the adsorbent bed duringclassic regeneration. Therefore, keeping the ad-sorber mass, and hence also the vessels, smallby using short adsorption – regeneration cyclesis cost-effective. Cycles between 5 and 10 minare efficient in practical applications. However,a drawback of the process is that one vessel vol-ume of recoverable hydrogen is lost during eachregeneration phase when the adsorber is flashed.This is why yields of highly pure hydrogen aregenerally poor relative to the recoverable hydro-gen content of the gas mixture.

Special Process Loops. To improve theyield, the number of adsorbers may be increasedfrom two to four, six, or ten.

In principle, these adsorbers or adsorbergroups simultaneously operate in the followingfour modes (→ Adsorption):

1) adsorption in the first adsorber,2) depressurizing the second adsorber,3) purging the third adsorber at lowpressure, and4) pressurizing the fourth adsorber.

With this loop, losses of recoverable hydrogenare reduced by using the gas released when thesecond adsorber is depressurized to repressurizethe fourth adsorber to medium pressure.

122 Gas Production

Product Purity. A four-adsorber loop canyield hydrogen of 99.999% purity and producesa uniform product flow, achieving yields as highas 72 – 78% of the hydrogen in the gas mixture.The types and quantities of contaminants in thegas mixture influence the size of the adsorberbeds and, to a certain extent, the yields. How-ever, they do not influence the function of thepressure-swing adsorption unit or the purity ofthe recovered hydrogen. Contaminants are dis-charged at ambient pressure together with thehydrogen used for purging. In the last ten years,this process has been used in several plants de-signed for an output up to 100 000 m3 (STP)/hof highly pure hydrogen.

Multicomponent Separation [214]. Fur-ther developments of current adsorption pro-cesses aim, on the one hand, at improving theadsorbents and optimizing process loops and ad-sorption – regeneration cycles and, on the other,at utilizing various desorption effects dependingon the respective components. In the past, ad-sorption processes were tailored to recover oneproduct in a highly pure form (e.g., H2 from gasmixtures, O2 or N2 from air). Modern develop-ments take advantage of the different propertiesof individual components during desorption toincrease the content of certain components inthe outgoing flows. To this end, the origin of thedesorption flash gas is considered, for example,raw gas from the inactive volume of the vesseland the intergranular volume, raw gas from thepore volume, and desorbed gas from the innersurfaces of the adsorbent at various desorptionpressures (medium pressure, normal pressure,vacuum). The upgrading effect is reinforced bysuitable process arrangements during desorption(e.g., cocurrent or countercurrent desorption), incombination with the passing of adsorbed gasesto other adsorbent beds.

This process arrangement can be used suc-cessfully to separate gas mixtures consisting ofhydrogen, carbon monoxide, and carbon diox-ide into highly pure product hydrogen, a carbonmonoxide fraction, and a residual gas fractionwith an increased carbon dioxide content. Anal-ogously, a mixture consisting of nitrogen, meth-ane, and carbon dioxide can be separated into ni-trogen waste gas, a high-methane fraction, and acarbon dioxide fraction. In both cases, the com-ponents participating in the process behave very

differently because of their properties, not onlyduring individual adsorption cycles but also dur-ing desorption cycles.

5.5.4. Adsorption on Activated Carbon

Activated carbon (→ Carbon, Chap. 5) shows astrong adsorptive effect because of its extensivepore system. Suitable preparation of carbona-ceousmaterials can produce different pore struc-tures, so that the distribution of pore radii maybe matched to specific adsorption requirements.A variety of pore sizes down to micropores andminipores with diameters on the order of smallgas molecules can be produced.

Applications. For gas purification, adsorp-tion on activated carbon is used to protect down-stream process stages or systems from tracesof compressor lubrication oil or other low-volatility hydrocarbons, such as toluene andwashing oil, which may have entered the gasflow through a special wash process. These ad-sorbers are characterized by high adsorptivity,especially for substances that are near their dewpoint or have been entrained with the gas flowin the form of minute condensate droplets. Acti-vated carbon adsorbers can usually be regener-ated by heating them and simultaneously purg-ing the adsorbent bed with hot steam. Coking ofthe adsorbent, however, cannot be avoided en-tirely, so that the adsorption capacity of the ac-tivated carbon adsorber decreases with each re-generation cycle. Frequently, greater economyresults from providing a single activated car-bon bed of sufficient size, operating it to break-through, and refilling it with fresh activated car-bon during a brief shutdown.

5.6. Cryogenic Processes (→ CryogenicTechnology)

Cryogenic processes are used to recover puregas fractions from gas mixtures. Typical exam-ples are carbon monoxide and hydrogen. The re-quired purity of both products depends on theirintended use. The accompanying substances andcontaminants are more or less the same in allcases: hydrogen, carbonmonoxide, carbon diox-ide, residual methane, nitrogen, oxygen, argon,

Gas Production 123

steam, contaminants such as oils, saturated andunsaturated higher hydrocarbons, sulfur com-pounds, nitrogen oxides, and acetylene. Theirconcentrations vary according to the methodused to produce the synthesis gas. Because thegas mixture is further processed, admixturessuch as hydrogen or carbon monoxide, carbondioxide or steam may lead to undesirable sec-ondary reactions. Especially when unused gasesare recycled to the process, nitrogen, argon andmethane will lead to a build-up of inert com-ponents and thus impair the process conditions.Contaminants especially in the ppm range suchas sulfur compounds, steam and acetylene, butalso carbon monoxide in hydrogen, frequentlypoison the catalyst or produce secondary reac-tions at the catalyst.

In practice, a carbon monoxide product gascontaminated with not more than 1 – 3% hydro-gen, methane, or nitrogen meets the quality re-quirements for most downstream chemical pro-cesses. Hydrogen used for hydrogenation or am-monia synthesis must meet even more stringentpurity requirements.

Sulfur compounds, acetylene, carbon diox-ide, and water are removed from synthesis gasby washing, adsorption, and absorption. This re-duces the problem to the recovery of pure hy-drogen or carbon monoxide from a gas mixtureessentially containing hydrogen, carbonmonox-ide, and methane. Cryogenic separation is apromising route, in addition to selective absorp-tion and adsorption, for separating the compo-nents of this gas mixture.

Plant engineering has benefited from severalcryogenic processes, and various process combi-nations have been used for the respective appli-cations. Two principal process loops seem to bepromising for carbon monoxide recovery fromsynthesis gas mixtures: (1) partial condensationand (2) the liquidmethane wash process. A thirdprocess— the liquid nitrogenwash process— isused to recover hydrogen for ammonia synthe-sis.

5.6.1. Partial Condensation

The first stage of the partial condensation pro-cess for recovering hydrogen or carbon monox-ide from synthesis gas mixtures is the lique-faction of higher boiling gas components such

as carbon monoxide and methane. Dependingon the pressure applied, liquefaction begins atca. −100 to −180 ◦C, and ends when the sub-limation temperatures of methane and carbonmonoxide are reached.

The first components to liquefy in a hydro-gen atmosphere aremethane and carbonmonox-ide. Separation of the gaseous and liquid phasesalone yields relatively pure product flows.

Because carbon monoxide recovery requirestemperatures as low as−200 ◦C, for a cryogenicprocess to work properly all components thatmay freeze and clog the pipingmust be removedcompletely.

In partial condensation, carbon monoxide iscondensed along “cold walls.” This leaves aquantity of carbon monoxide, corresponding tothe carbon monoxide vapor pressure, uncon-densed. The “cold wall” is produced in a heatexchanger by flashing liquid carbon monoxideto the lowest possible pressure and evaporating itin countercurrent. This defines the coldest pointof the partial condensation process: it is the tem-perature of liquid carbon monoxide evaporatingat the lowest possible pressure [215, 216].

If the carbon monoxide condensate containsan unacceptably high percentage of methane,this must be removed by distillation. The car-bonmonoxide yield is 75 – 80%. Losses are dueto uncondensables in the hydrogen flow and tocarbon monoxide removed along with methane.However, recovery of only part of the carbonmonoxidemay be desirable if hydrogen gas con-taining a certain proportion of carbon monoxideis required.

Process Description. Figure 69 shows a typ-ical process flow diagram of a partial conden-sation unit. After all contaminants that mightfreeze and clog the piping have been removedfrom in a pretreatment stage, the gas mixtureis cooled in heat exchangers (a) in countercur-rent to the cold products coming from the plant.Part of the carbonmonoxide andmethane is con-densed, and the gas and liquid are separated ina separator (b). Hydrogen gas, which still con-tains some carbon monoxide, is reheated. Theliquid containing traces of dissolved hydrogenand methane is depressurized; dissolved hydro-gen boils off and is removed in the downstreamseparator (c). The liquid, which no longer con-tains any hydrogen, is fed to a column (d) wherethe carbon monoxide – methane mixture is sep-

124 Gas Production

arated into pure carbonmonoxide and amethanefraction containing a certain percentage of car-bon monoxide. All cold products are evaporatedat ambient pressure and heated in countercur-rent to the hot incoming gas mixture. The car-bon monoxide yield may be increased by modi-fying the process by additional process steps andrecycling of carbon monoxide containing flashgases.

Figure 69. Cryogenic partial condensation processa) Heat exchanger; b) Separator; c) Degasser; d) CO – CH4separator; e) Expander

In this process, hydrogen is obtained at itsfull pressure, which is not always required ina carbon monoxide recovery unit. The hydro-gen produced can, therefore, also be used forflash refrigeration in an expander (e). This pro-cess element enables the coldest temperature tobe lowered by ca. 20 to −200 ◦C, so that theamount of carbon monoxide remaining in thegaseous phase is greatly decreased and the car-bon monoxide yield increased [215, 216]. Pro-cess results are given in Table 30.

Table 30. Typical analyses for the process flow diagram shown inFigure 69

Feed gas H2fraction

COproduct

Residualgas

Without expanderstageAnalysis, mol%

H2 70 88 1.5 40CO 27 12 98.5 50CH4 3 10

Volume parts 1000 650 50 300With expander stageAnalysis, mol%

H2 70 96.5 1.5 40CO 27 3.5 98.5 50CH4 3 10

Volume parts 1000 600 100 300

5.6.2. Liquid Methane Wash Process

Liquid methane is used to wash a pressurizedhydrogen flow slightly above the freezing pointof methane (i.e., ca. −180 ◦C). During this pro-cess, carbon monoxide carried along with hy-drogen is dissolved in methane; 100% of car-bon monoxide in the feed is recovered [217] sothat hydrogen leaving the column overhead isfree of carbon monoxide. For calculation meth-ods applicable to isothermal wash processes, see[218].

Spent methane is regenerated by distillation,expelling carbon monoxide overhead. By us-ing the heat pump principle and employing aflash refrigerator, carbon monoxide can be con-veniently recycled and used to maintain the re-frigeration balance.

Process Description (see Fig. 70). Pre-treated synthesis gas is cooled in a heat ex-changer (a) to the washing temperature of−180 ◦C. High-boiling components are thuscondensed, and only the gaseous carbonmonox-ide in the hydrogen gas is washed out withmeth-ane in column (b). This washingmethane, meth-ane from the gas mixture, and carbon monox-ide that has been condensed or washed out arefed to the carbonmonoxide – methane separator(c). There, the carbon monoxide product is dis-charged overhead while the washing methaneremains as the bottom product. A liquid meth-ane pump (d) delivers the washing methane viaa supercooler (e) to the top of the wash column(b), fromwhich carbonmonoxide-free hydrogenis discharged overhead. The product streams ofcarbon monoxide, hydrogen, and residual meth-

Gas Production 125

Figure 70. Liquid methane wash processa) Heat exchanger; b) Liquid methane wash column; c) CO – CH4 separator; d) Liquid methane pump; e) Liquid methanesubcooler; f) Recycled CO expander; g) Recycled CO compressor

ane are heated in the exchanger (a), flowing incountercurrent to the feed gas, and recovered in-dividually.

The carbon monoxide compressor (g) closesthe refrigeration loop with the expander (f). Thesame compressor may also be used to adjust thecarbon monoxide product to the required bat-tery limit pressure. Process results are given inTable 31.

Table 31. Typical analyses for the process flow diagram shown inFigure 70

Feed gas H2fraction

COproduct

Residualgas

Analysis, mol%H2 70 98.5 1.7CO 27 98.5CH4 2 1.5 0.1 100

Volume parts 1 000 705 275 20

5.6.3. Liquid Nitrogen Wash Process

A finishing treatment in a liquid nitrogen washunit is the last process stage in the recovery ofhydrogen for ammonia synthesis from hydro-carbons. In preceding stages, carbon monoxideis converted to carbon dioxide, producing addi-tional hydrogen, and carbondioxide is then elim-inated completely from the synthesis gas. Hy-drogen gas entering the liquid nitrogenwash unit

contains the following undesirable substancesand contaminants:

1) traces of unreacted oxygen from the gasifica-tion process,

2) residual methane from unreacted hydrocar-bons,

3) some unconverted carbon monoxide, and4) argon that entered the system together with

the oxygen required for gasification

Whereas argon and methane in the synthesisgas are inerts and would unnecessarily burdenthe ammonia synthesis loop, the residual oxy-gen and carbon monoxide are poisonous to theammonia synthesis catalyst. Except for an un-avoidable residue of argon, the above-mentionedcontaminants are all eliminated by means of aliquid nitrogen wash unit.

The refrigeration needed for this process isobtained by flashing the required amount of ni-trogen for washing the synthesis gas into the pu-rified hydrogen already in the cryogenic unit. Bymixing the nitrogen with the hydrogen stream,the pressure of the cryogenic nitrogen is loweredfrom the full process pressure at which the purenitrogen is cooled to the nitrogen partial pres-sure in the hydrogen stream, which more or lessreflects the proportion of nitrogen (24 – 25%) inthe synthesis gas. Essentially, the positive excessenthalpy ∆H occurring during this process pro-

126 Gas Production

vides the refrigeration necessary for the entireprocess [219].

Figure 71. Liquid nitrogen wash unita) Heat exchanger; b) Wash column; c) Valve for wash ni-trogen; d) Refrigeration valve; e) Trim valve

Process Description (see Fig. 71). Hydro-gen gas containing contaminants such as oxy-gen, argon, carbon monoxide, and methane iscooled to ca. −180 ◦C in a heat exchanger (a).At a temperature around its dew point, the gas isfed to the bottom of the wash column (b). Liquidnitrogen trickling down countercurrently to thehydrogen gas dissolves the contaminants and isdischargedwith them from the bottomof the col-umn. The overhead product is purified hydrogenwith an nitrogen content that reflects the equi-librium stage of a hydrogen – nitrogen mixtureat the of temperature and pressure prevailing atthe top of the column. Nitrogen required in thewash unit is compressed to the process pressure,cooled in an exchanger (a), and fed to the washcolumn through a valve (c). Nitrogen needed forrefrigeration is cooled in the same way, flashedthrough a valve (d) into the purified hydrogen,and evaporated in the exchanger (a) in counter-current to the uncooled and untreated hydrogenand the nitrogen flow. The valve (e) is used toadd nitrogen to establish the hydrogen – nitro-gen ratio (3: 1) required for synthesis.

Wash nitrogen containing the dissolved con-taminants is discharged from the columnbottom,flashed to a lower pressure, evaporated, heated,and discharged as residual gas (for process re-sults, see Table 32).

Table 32. Typical analyses for the process flow diagram shown inFigure 71

H2 gas N2 gas Synthesis Residualgas gas

H2 94.11 mol% 75 mol% 8.31 mol%N2 0.05 mol% 100 mol% 25 mol% 26 mol%O2 ppm ppm 2 ppm ppmAr 0.60 mol% 50 ppm 7.14 mol%CO 4.44 mol% 5 ppm 52.68 mol%CH4 0.50 mol% 10 ppm 5.87 mol%Volumeparts 1000 337 1253 84

5.7. Gas Separation by Membranes (→Membranes and Membrane SeparationProcesses)

Fundamentals. Certain materials, whenproduced in the form of nonporous membranesof defined thickness, absorb gas componentsin their molecular structure at different ratesand allow them to diffuse through the mem-brane with a certain pressure drop. The diffu-sion rate depends on (1) the type of material,(2) the size of the molecules, (3) the pressuredifferential across the membrane, and (4) thediffusion length (membrane thickness) and ac-tive area of the membrane. The following typesof membrane can be distinguished:

Polymer membranesCarbon membranes with microporesLiquid membranesMetallic membranesCeramic membranes

Whereas membranes based on ceramics, metals,and liquids have been known for a long time andhave been developed to an industrial standard forsome applications (e.g., palladium ceramics forthe production of highly pure H2), the develop-ment of polymer membranes for gas separation[220 – 225] has been accelerated since 1977 bythe rapid development of reverse osmosis mem-branes for seawater desalination [226].

Several gas separation membrane systems,which compete in materials and design of themodules, can be found on the market:

Gas Production 127

1) Prism separator (Monsanto), an asymmetricporous membrane on a polysulfone carrierwith a nonporous silicon-based coating, in theform of hollow-fiber bundles;

2) Delsep (Delta Engineering Corporation), anasymmetric porous membrane based on cel-lulose acetate with a nonporous celluloseacetate-based coating, in the form of a coil-type module with a flat membrane;

3) Separex (Air Products Company); and4) ROmembranes (DuPont), asymmetric porous

membranes based on aromatic polyamide orcellulose acetate with a nonporous aramid-based coating, in the formofhollow-fiber bun-dles.

Applications. Hydrogen and helium havethe best sorption and diffusion behavior of allgas components, and their separation efficiencyis very good in comparison to nitrogen, meth-ane, carbon dioxide, or carbon monoxide. Suchgas separation units are, therefore, employed torecover hydrogen from reaction gas mixtures byusing the same process principle as shown inFigure 72 in all cases.

Process Description. Case 1: Gas Separa-tion Membrane installed in a Slipstream. Purgegas which must be extracted continuously froma synthesis gas loop (as installed for the ammo-nia production) to keep down the level of un-desirable ballasts and nonreacting componentsessentially contains the same valuables as theloop (e.g., H2, N2, or CO). To improve overallefficiency of the synthesis loop, these valuablescan be recovered by means of a gas separationmembrane and recycled at low pressure to thesuction side of the raw gas compressor. Thosegas components that do not permeate the mem-brane (e.g., CH4, Ar, CO, higher hydrocarbons)remain in thewaste gas. Typical gas compositionis shown in Table 33.

Table 33. Typical gas composition for purge gas treatment

Raw gas Permeate Waste gasH2, vol% 74 96 57CO, vol% 26 4 43Flow rate, kmol/h 100 43.6 56.4Yield H2: 56% CO: 93%

Case 2: Gas Separation Membrane in theMain Stream. Hydrogen sulfide and carbon

dioxide exhibit much better sorption and dif-fusion behavior than the product gases carbonmonoxide and methane. This favorable separa-tion efficiency can be used in a pretreatmentstage to reduce the size of downstream conven-tional fine treatment (see Fig. 73). Most of thesour gases are removed by the gas separationmembrane, although some valuable componentsare lost as well. Such losses of valuable com-ponents can be minimized by splitting the pro-cess into several stages and recycling the gases— a method which, of course, requires addi-tional energy. Being preceded by a membrane,the absorption unit is turned into a fine treatmentstage, whose dimensions and complexity can bereduced considerably.

5.8. Addition of Inerts or OtherSubstances

Normally, the quality of a gas is described notonly by its composition but also by such charac-teristics as its calorific value and Wobbe index(see Chap. 8).

These characteristics are adjusted as neces-sary, either by adding inerts such as nitrogen orair to decrease the calorific value and increasethe density of a gas, or by adding high-calorificsubstances such as liquefied gas, butane, or gaso-line to increase the calorific value. Proper mix-ing is achieved by installing flow turbulators,providing spray injection of liquids into the gasstream, or adding densifiers.

If air is added, the oxygen level in the gasmust be watched closely. Explosive air – gasmixtures, which may occur briefly in areaswhere air is injected into the gas, can be elimi-nated by flow turbulators.

If liquefied natural gas is used as feedstock, itcan be conditioned to the required calorific valueby adding liquefied nitrogen in the cryogenic liq-uid phase. Proper mixing is ensured as the gasevaporates at discharge pressure. This condition-ing method saves energy because only a mini-mum of energy is required to pump the liquefiednatural gas and liquefied nitrogen, rather thannecessitating installation of compressor stationsfor natural gas, nitrogen, or the gas mixture.

Substances with a manifest odor—usuallyhighermercaptans—are added tonormally odor-less gases to warn of any leaks.

128 Gas Production

Figure 72. Principle of H2 recovery from reaction gas mixtures by means of gas separation membranesa) Gas separation membrane; b) Raw gas compressor; c) Recycle gas compressor

Figure 73. Pretreatment by means of gas separation membranea) Gas separation membrane (pretreated stage); b) Sour gas treatment; c) Absorption (fine treatment)

6. Handling of Byproducts

Cooling the crudeproduct gas leads to the forma-tion of aqueous and hydrocarbon condensates.Usually, water-insoluble hydrocarbons are re-moved from the aqueous phase and handled sep-arately. Treatment of the condensates and thecrude gas also results in gaseous byproducts.

6.1. Aqueous Condensates

Rates and Components. All productionmethods for gases containing carbon monoxideand hydrogen primarily generate a hot crude gasmixedwith water vapor, regardless of feedstock.The water content of the crude gas originates

Gas Production 129

from the feedstock moisture (coal), unconvertedgasification steam, or reaction water (oxidationof part of the feedstock hydrogen). Cooling thecrude gas produces aqueous condensates ladento a greater or lesser extent with organic or in-organic substances, depending on the type offeedstock and the gasification method. Table 34shows typical condensate rates and contents ofdissolved or suspended components for somefeedstocks and selected gasification methods.The numbers may change depending on gasi-fication conditions. As a rule, condensate rateand impurity content increase with decreasinggasification temperature and increasing feed-stock complexity [227 – 231]. In addition tothe classical feedstocks nowadays a variety ofwaste materials are used. These have to be pre-pared, e.g., by pelletizing using suitable binderssuch as bituminous coal and/or by mixing withcoal in certain ratios between 50% and 85%.The waste includes solid and liquid materialslike plastics, sewage sludge, rubber, fluff, con-taminated wood, residues of paint, householdwastes, industrial waste, etc., as well as tar/oilfrom fixed bed gasifiers, used oils, solventsetc. [232]. This changes the composition of thecondensates slightly. The modern practice is toselect gasification types and conditions in sucha way that the condensates obtained are as cleanas possible and/or show low specific rates, inorder to reduce the wastewater treatment cost(see Chap. 4).

Typical Specifications of Treated Waste-water. Condensates are recycled as far as possi-ble within the gas production plant. Only excesscondensate must be specially treated. Treatmentmethods depend on the final destination of con-densates:

1) Use as boiler feedwater (liquid disposal),2) Use as cooling water make-up (gaseous dis-

posal),3) Burning of contents and evaporation of water

(gaseous disposal), or4) Release to receiving water (liquid disposal).

The first two are possible only if the con-densates contain easily removable components(e.g., using steam reforming condensate asboiler feedwater after degassing) or are cleanedby various treatment processes (e.g., using coalpressure gasification condensate as cooling wa-

ter after tar – dust separation, filtration, extrac-tion, stripping, adsorption, biological oxidation,etc.) [233, 234]. Burning is limited by theamount of water to be evaporated, the environ-mental impact of sulfur and nitrogen oxides,and economic aspects (proportion of recover-able heat). In practice, incineration of gas con-densates has been considered only for small rates[235] or for the concentrated residues after treat-ment [236].

The last method is the most common: gascondensates are treated tomeetwastewater spec-ifications (local authority regulations) so that theeffluent can be discharged into receiving water(river, canal, lake, sea).

No universal wastewater disposal standardsexist. Specifications must be negotiated withthe local water control authority, which usuallyconsiders the plant location, type of feedstock,conversion process, and type of receiving water[237, 238]. Typicalwastewater specifications forauthorized effluent disposal are listed in the fol-lowingmaterial [BOD5 = biological oxygen de-mand (five days)], indicating the targets for var-ious wastewater treatment processes describedin the following sections.

Temperature max. 30 ◦CpH 6.5 – 8.5Suspended solids max. 30 mg/LBOD5 max. 30 mg/LCOD max. 500 mg/LHydrogen sulfide max. 0.1 mg/LAmmonia (measured as nitrogen) max. 200 mg/LFree cyanide max. 0.5 mg/LCNS− max. 15 mg/LMonohydric phenols max. 0.1 mg/LTotal phenols max. 10 mg/LSubstances extractable withpetroleum ether max. 20 mg/L

6.1.1. Mechanical Treatment

Gravitational Separation. Suspendedsolids such as soot, char, or ash and entrained taror oil are removed mechanically from gas con-densates. In coal pressure gasification plants,the first cleaning step is usually carried out intar separators, often followed by additional oilseparators. In oil gasification plants, soot is usu-ally extracted with oil or naphtha and recycledto the process.

130 Gas Production

Table 34. Specific rates and composition of gas condensates

Process examples Lurgi coal pressure gasifier British Gas – Lurgi slagginggasifier

Texaco, Shell,Lurgi MPG Steam reforming

Type moving bed moving bed partial oxidation catalyticFeedstock bituminous coal bituminous coal heavy vacuum residue naphthaSpecific rate of feed, m3/t 0.7 – 1.2 0.15 – 0.2 0.4 0.5 – 0.7COD, mg/L 25 000 2400BOD5, mg/L 8000 1000Composition, mg/L

Suspended solids 100 120Suspended oil 1000 20Free ammonia 15 000 20 000 1500Fixed ammonia 350 3 000 150CO2 30 000 40 000 2400 400H2S 700 600 1600HCN 50 100 300Monohydric phenols 4500 4000 <1Total phenols 6000 5000 <1Dissolved organic

substances500

Halides (Cl−, F−) 250 4000Fatty acids 800 500 400Nickel 0 2 – 5Iron 4 6 – 15Vanadium 0 3 – 8

pH 8 – 9 8 – 9 8 – 9 5 – 7

Primary tar separators may be simple tanksprovided with a feed nozzle, an expansion gasexit to release dissolved gases, an oil skimmer,and drawoffs for tar, solids, and cleaned water.Oil has a lower density, and tar – dust a higherdensity, than water so they form top and bottomlayers that arewithdrawnwhilewater is obtainedin the intermediate layer. The settling area of tarand oil separators is usually designed to removeall particles having a settlingvelocity higher than1 – 3 m/h from the aqueous phase. Remainingparticles are removed in the following filtrationstep.

Part of the mixture of tar and dust (heavy tar)is recycled to the gasifiers by positive displace-ment pumps. Excess tar sometimes undergoesfurther dust removal and is sold as a product.After primary tar separation, the gas condensatestill contains up to 100 mg/L of solids and 500– 5000 mg/L of tar and oil [239].

The gas condensate is cleaned further in oilseparators down to a content of oil and solidsin the range of 20 – 50 mg/L. The efficiencyof these separators can be enhanced by fittinga large number of parallel corrugated plates,which not only increases the settling area butalso achieves laminar flow conditions, resultingin low Reynolds numbers. An example of thisseparator type is the Shell CPI [240].

Filtration. The following treatment steps(e.g., extraction, stripping, adsorption) requirethat tar and solids content be reduced to 10 mg/Lor lower. This can be done in simple gravel fil-ters operated at low specific loads of ca. 5 – 7m3 m−2 h−1.Modern dual- ormultimedia filterscan operate up to 20 m3 m−2 h−1. In dual-mediafilters, the gas condensate flows first through anupper layer of lower density, coarse filter media,which allows finer solids to penetrate deeply intothe filter bed, and then through a lower layer offiner filter media to ensure final cleaning.

Filters operate batchwise. Filter beds loadedwith tar and solids are regenerated by backflush-ing with (hot) water. Usually, the bed is alsoloosened by blowing gas through it. After clean-ing, filter bed particles are classified accordingto Stokes’ law by means of an upward streamof water at a defined rate. The filters not onlyremove solids to < 10 mg/L but, at low temper-atures and low velocities, remove tar and oil aswell. These partly stick to the filter media andpartly collect in the bottom of the filter, fromwhich they can be removed prior to backflush-ing. The backflushing liquid (mud liquor) is nor-mally recycled to the primary tar separator.

Gas Production 131

Phenol 45Cresols 30Xylenols 6Polyhydric and other phenols 10Pyridine bases 1Neutral oil, pitch 5Water 3

6.1.2. Extraction and Adsorption of OrganicSubstances

Filtered gas condensates containing high levelsof dissolved organic substances cannot directlybe treated biologically. First, the phenol contentmust be reduced to an acceptable level, usually< 700 mg/L. At low feed rate or concentration,this is done by dilution. At high concentration ofsteam volatile phenols, the condensate is dephe-nolized prior to stripping the dissolved gases.This is carried out either by liquid – liquid ex-traction using organic solvents—e.g., benzene,methyl isobutyl ketone (MIBK), butyl acetate,or diisopropyl ether (DIPE)—or by adsorptionon activated carbon. Combinations of these pro-cesses are also possible [241 – 243]. Adsorp-tion on activated carbon is mainly used for post-treatment after dephenolization and/or biologi-cal treatment of condensates in gas productionplants [244 – 246].

Phenol Recovery by Solvent Extraction.Figure 74 shows a typical phenol extraction pro-

cess used in gasification plants, aswell as in cokeoven and carbonization plants, in the phenol in-dustry, and in coal hydrogenation plants. In thePhenosolvan process the filtered and cooled phe-nolic effluent is treated countercurrently with asuitable solvent such as DIPE in amultistage ex-tractor. The extract is separated by fractional dis-tillation into pure solvent (overheadproduct) andcrude phenols (bottoms product). The solvent isrecycled to the extractor. The raffinate still con-tains a small amount of solvent which is reco-vered by strippingwith recycled gas. The solventis then removed from the gas by absorption incooled and recycled crude phenols, from whichit is subsequently recovered in the stripping sec-tionof the fractionator. ThePhenosolvanprocesscan be operatedwith different feedstocks, differ-ent solvents or solvent mixtures (e.g., benzeneandDIPEorDIPEplusMIBK) [247], and differ-ent solvent recovery systems (e.g., steam, gas, orammonia stripping). All steam volatile phenolsand neutral oils can be recovered almost com-pletely, but only partial recovery of pyridines,fatty acids, and dihydric phenols is possible. Thesuitability of this process for hard-to-extract or-ganics is limited only by economics, becauseresidual phenols and other organics can be re-moved more cheaply by adsorption or biologi-cal treatment. The intermediate phenol contentbetween extraction and final treatment is bestdetermined by means of an economic optimiza-tion study, an example of which is shown in Fig-

Figure 74. Lurgi Phenosolvan process for phenol recovery (with steam-saving solvent recovery by gas recycle)a) Extractor (mixer – settler); b) Fractionator; c) Solvent absorber; d) Solvent stripper

132 Gas Production

ure 75. An advantage of solvent extraction isthe recovery of crude phenol by fractionatingthe extract which contains valuable byproducts.Typical composition of a Phenosolvan extractfrom coal pressure gasification is given below,in percent by weight:

Figure 75.Economical phenol recovery point between phe-nol extraction and final treatment (typical, e.g., for coal gasi-fication condensates containing ca. 4500 mg/L monohydricand ca. 1500 mg/L polyhydric phenols)a) Extraction; b) Biological treatment; c) Total cost;d) Biological and activated carbon treatment

If large gas condensate flows must bedephenolized, extraction must be performed inmixer settlers, exhibiting high stage efficiencies.Lower flow rates can be treated in extractioncolumns by using many stages to compensatefor low efficiency.

Recovery by High-Pressure Extraction.Organics can also be recovered by using super-critical solvents [248]. This method might beeconomical for small condensate rates with highfeed concentrations of phenols and other valu-able organics, but it is not yet being used in gasproduction plants [249]

6.1.3. Removal and Recovery of Ammoniaand Sulfur

Before gas condensates rich in dissolved gases(NH3, CO2, HCN, H2S) can be treated in bio-logical oxidation or in heavy-metal precipitationplants, theymust be strippedwith gas or steam todecrease free ammonia to< 50 mg/L, hydrogensulfide to < 1 mg/L or hydrogen cyanide to <10 mg/L, or to lower the pH (previously 8 – 9)to 6. This is done in the total stripper commonto all removal or recovery processes. Its primarypurpose is to produce stripped gas condensate asbottom product. The overheads contain all thefree gases and water vapor. Processes differ inthe method of treating this gas mixture.

Total Stripper. Depending on themajor feedcontaminant, this stripper has different namessuch as ammonia still and sour water, hydro-gen sulfide, or hydrogen cyanide stripper [250].It usually operates at normal atmospheric pres-sure, but the pressuremay be adjusted dependingon the destination of the top gases. Typical steamconsumption is 140 – 250 kg of steam per cubicmeter of gas condensate.

The stripper is usually equipped with a topcondenser to reduce the water content of thegases from > 90 to < 50%.

Incineration, Catalytic Oxidation and Sul-fur Recovery. The mixture of ammonia, sourgases, andwater can beflared off only if the envi-ronmental impact (e.g., of SO2, NOx ) is accept-able.Otherwise, a special incineration plantwithflue gas treatment is required.Ammonia can alsobe removed by catalytic oxidation [251].

If a Claus plant (→ Sulfur, Chap. 6) is avail-able, the strip gas can be mixed with the feedfor sulfur recovery and ammonia oxidation. Ifthe ammonia flow is too high compared to theClaus sulfur stream, the strip gasmust be dividedinto an ammonia-free sour gas and an ammo-nia stream low in hydrogen sulfide [252]. Bothstreams are fed separately to theClaus plant; firstthe ammonia stream is oxidized above stoichio-metric conditions and then mixed with the sourgas stream at below stoichiometric conditions toproduce sulfur. Any NOx formed in the first stepreacts with sulfur produced in the second step;thus, both sulfur dioxide and NOx emissions areeliminated [253].

Gas Production 133

Recovery of Ammonia. If the ammoniaquantity exceeds a certain level, incinerationis unacceptable but recovery becomes economi-cal compared to (catalytic) oxidation. Formerly,the strip gases were washed with sulfuric acidto obtain ammonium sulfate, and the remainingacid gaseswere fed to sulfur recovery plants. Be-cause ammonium sulfate sales today are limited,the recovery of anhydrous or aqueous ammo-nia is preferred in most cases. Generally, twodifferent methods are available for the difficultseparation of basic and acid gases from aqueoussolution:

1) Chemical methods using agents that selec-tively absorb ammonia under certain condi-tions, but release ammonia under regenerationconditions [254]; and

2) Physical methods (without agents), based ondifferent vapor pressures of ammonia and acidgases at different concentrations in aqueoussolution and at different temperatures or pres-sures [242, 255, 256][257].

The first method is represented by thePhosam-W process (U.S. Steel, USX Engineers& Consultants, Aristech Phosam-W), and thesecond by the CLL ammonia recovery process(Chemie Linz/Lurgi) or by the WWT process(Chevron); both types of processes are usedin coal pressure gasification plants [258 – 261].The largest plant worldwide for the recovery ofclean liquid ammonia from coal pressure gasifi-cation wastewater operates in Secunda at SasolII and III since 1980. Both CLL trains togetherhave a total capacity of 1000 t NH3 per dayformed as a byproduct.

Phosam-W Process (see Fig. 76). The totalstripper (a) overheads are treated in the ammoniaabsorber (b) with lean ammoniumphosphate so-lution, which selectively absorbs ammonia un-der prescribed conditions. The wet overheadacid gases are dried (c) and sent to a sulfur recov-ery unit. The rich ammoniumphosphate solutionis regenerated in the Phosam stripper (d) at hightemperature and recycled to the absorber. Thestripper overheads (mainly ammonia and water)are condensed, treated with sodium hydroxideto bind coabsorbed acid gases, and fractionatedto yield pure anhydrous ammonia as final prod-uct. Hot (475 K) fractionator bottom water andsodium hydroxide are flashed into the total strip-per.

Figure 76. Phosam-W process for chemical separation ofaqueous ammonia – acid gas mixturea) Total stripper; b) Phosam ammonia absorber; c) Acid gasdrier; d) Phosam ammonia stripper

Clean Liquid Ammonia Recovery Pro-cess (CLL; see Fig. 77). The total stripper(a) overheads are treated in the acid gas absorber(b) and ammonia stripper (c) column. Ammoniais cleaned by fractional condensation and fed to aliquefaction unit, to yield anhydrous ammonia asfinal product. Alternatively, ammonia can be ab-sorbed in water or fed into the ammonia pipelineof the factory (Chemie Linz). The acid gas ab-sorber bottoms are concentrated in the ammoniastripper (c) and fed to the deacidifier (d) whereclean carbon dioxide and hydrogen sulfide areobtained as top product, suitable for use directlyas sulfur recovery feed gas.

Deacidification is enhanced by dilution withwater or elevation of temperature and pressure.Deacidified gas condensate is fed to the totalstripper, closing the CLL loop.

The gas condensate is fed either to the deacid-ifier if it contains more acid gases and/or inertgases and/or traces of solvent from the dephe-nolization step or to the total stripper if it con-tains more ammonia. Diluted condensates arealso fed to the total stripper.

Both processes described yield ammoniaqualities comparable to synthetic ammonia (Ta-ble 35); hence, process selection depends oneconomic considerations.

134 Gas Production

Figure 77. CLL Ammonia recovery process for physicalseparation of aqueous ammonia – acid gas mixturesa) Total stripper; b) CLL ammonia stripper; c) CLL acid gasabsorber; d) Deacidifier

Table 35. Typical composition of anhydrous ammonia

Clean liquidammoniarecovered fromgas condensates

Synthetic ammonia

Water (added), wt % 0.01 – 0.2 0.2Ammonia, wt % 99.99 – 99.8 99.8Oil, ppm <2 <5Organic substances, ppm <8 not determinedH2S, ppm <1 not determinedCO2, ppm <50 not determined

Economic Considerations. Chemical ab-sorption – regeneration plants usually have lowcapital costs but high operating costs becauseof the use of stronger agents and the need forhigh-pressure steam to regenerate the chemicals.Chemical absorption plants are less flexible withrespect to overload, since the agents have a dis-tinct limit of loading capacity.Physical absorber– stripper plants have higher capital costs butlower operating costs because of the use ofcheaper steam. The higher the gas condensateflow rate or the ammonia feed concentration, themore economical are physical methods; chemi-cal methods are more economical in the case oflow flow rates or low ammonia feed concentra-tions.

6.1.4. Biological and Final Treatment

After extraction and stripping, gas condensatesstill contain organic and inorganic substances

such as fatty acids, dihydric phenols, and an-ions (e.g., chlorides and fluorides). Usually, theCOD and BOD of pretreated gas condensatesare decreased biologically. Aerobicmicroorgan-isms break down and metabolize most of the or-ganics to carbon dioxide and produce sludge.

Essentially, two methods exist for aerobictreatment of gas condensates. The activated-sludge process is most common because it is lesssensitive to variations in feed BOD. The tricklebed process, in which a film of microorganismsoxidizes the organics, is seldom used for gascondensates [262 – 264]. SometimesPAC (pow-dered activated carbon) is used, mainly in acti-vated sludge processes [265 – 269].

If the biotreated gas condensate is to be (part-ly) reused (e.g., as boiler feedwater), final filtra-tion for removal of suspended solids and, if nec-essary, ion exchange (→ Ion Exchangers), re-verse osmosis, or multiple flash evaporation forremoval of strong anions may be employed aswell.

Activated Sludge Process. For aftertreat-ment of gas condensates, a single stage con-sisting of an activated sludge basin equippedwith aerators, to provide the required oxygenand effect mixing of the basin contents, and aclarifier for separation and thickening of theactivated sludge is usually adequate. Nutrients(phosphorus and nitrogen compounds) for themicroorganisms are needed in the ratio of BOD:N: P = 100: 5: 1. Gas condensates normally con-tain the required nitrogen in the form of fixedammonia; phosphorus must usually be added inthe form of phosphoric acid.

Typical design figures for biological after-treatment of gas condensates to achieve a resid-ual BOD5 of 20 ppm are listed in the fol-lowing material (MLSS = mixed liquor – sus-pended solids, water-free; BOD5 = biologicaloxygen demand (five days); DO = dissolvedoxygen).

Activated sludge basin:Specific volumetric load,kg COD m−3 d−1 ca. 1.00

Specific sludge load,kg BOD5 kg−1 MLSS d−1 ca. 0.25

Specific sludge content, kg MLSS/m3 ca. 4.00Oxygen content, ppm DO 1 – 2Oxygen demand, kg O2/kg BOD5 ca. 1.50Specific oxygen transfer, kg O2 kW−1 h−1 ca. 2.50

Gas Production 135

Postclarification basin:Retention time, h ca. 2.50Rise rate (water), m3 m−2 h−1 ca. 0.80Specific area load, kg MLSS m−2 h−1 4 – 5

Addition of PAC to the activated sludge basinimproves removal efficiency in some cases, be-cause of various favorable properties of PAC:(1) substances toxic or inhibitory to biodegra-dation are adsorbed; (2) settling, thickening, anddewatering characteristics of the sludge are im-proved; and (3) the sludge age can be increased,resulting in higher retention times of difficultlybiodegradable organics. Typical PAC consump-tion is 0.1 – 0.3 kg per cubic meter of gas con-densate.

6.1.5. Removal of Heavy Metals

Wastewater from different processes (refining,coking,metallurgy, fossil fuel conversion, powergeneration, and others) contains greater or lesseramounts of heavy metals. The type and con-tent depend mainly on feedstock composition[270, 271] and the conditions under which prod-ucts come in contact with water (process con-densates, quench water, cleaning liquid, etc.).Some of the wastewater must be treated forremoval or recovery of heavy metals. Typicalmethods are liquid – liquid ion extraction, fixed-bed ion exchange, and precipitation by addingsuitable chemicals such as metal salts, lye, oracid [272 – 274].

In the majority of all cases, gas condensatescontain only small quantities of heavy metals.However, a considerable nickel and vanadiumcontent is found, for instance, in gas conden-sates from partial oxidation of heavy fuel or vac-uum residues. These metals are environmentallyharmful. The newLurgiMulti Purpose Gasifica-tion (MPG) process provides a special treatmentof the heavymetals (MARS,Metals Ash Recov-ery System, see Sections 3.2.2 and 7.2). Thisreduces the metals concentration in the effluentsconsiderably.

Precipitation of Vanadium andNickel. Af-ter removal of hydrogen cyanide, carbon diox-ide, hydrogen sulfide, and ammonia, the strippedcondensate typically has a pH of ca. 6. It iscooled to 40 ◦C and fed to the heavy-metal re-moval plant, which consists mainly of precip-

itation steps, sludge dewatering and filter cakestorage.

Precipitation of vanadium and nickel can becarried out in one step at pH 9 if all metals areto be removed in one mixture [273]. This mix-ture is clarified in a Sedimat (an automatic clar-ifier) [275], equipped with a central flocculationchamber to which a coagulant (polyelectrolyte)is added. The flocculated sludge settles to thethickening section in the bottom of the Sedimat,from which it is discharged.

The metal-free effluent is filtered to removesuspended solids, neutralized with sulfuric acid,and discharged to receiving water.

Thickened sludge streams from the clarifiersare filtered in order to reduce the water contentto 70%.

6.1.6. Example of an Industrial Application

Gas condensate from a typical Lurgi moving-bed coal pressure gasificationwith the followingapproximate composition, in grams per liter:

Tar (suspended portion only) 11.0Oil (suspended portion only) 3.5Suspended solids 1.5Steam volatile phenols 4.6Dihydric phenols 0.9Fatty acids as acetic acid 0.7Soluble organic substances 0.4Free ammonia 15.6Fixed ammonia 0.7Carbon dioxide 32.0Hydrogen sulfide 0.4Free cyanide 0.1Water balance

and a flow rate of 500 m3/h is to be treatedso that 20% can be released to receiving wa-ter, 75% used as cooling tower make-up water,and 5% used as boiler feedwater. A typical pro-cess sequence is shown in Figure 78. Interme-diate battery limit conditions should be selectedin such a way that total treatment cost is mini-mized whereas final conditions are still satisfied[276]. This leads to the following typical designcriteria, in milligrams per liter, for intermediateproduct streams (after treatment):

Tar separationTotal suspended matter 1000

CPI oil separatorsTotal suspended matter 350

136 Gas Production

Dual media filters (for 500 m3/h)Suspended matter 10

Phenosolvan plantSteam volatile phenols 10Dihydric phenols 280

Ammonia recoveryFree ammonia 50Hydrogen sulfide 1

Activated sludge treatmentBOD5 reduction from

700 to20

Final dual media filtrationActivated sludge content 20

Activated carbon posttreatment (using GAC, GranularActivated Carbon; [277])(for 400 m3/h) Biologically refractory substance 5Ion exchanger (for 25 m3/h)

Strong ions 10

In addition, the aqueous condensate treat-ment units will produce the following valuablebyproducts, in tons per hour:

Tar 5.4Oil 1.7Crude phenol 2.7Sulfur (via Claus plant) 0.1Clean liquid ammonia 7.9

6.2. Hydrocarbon Condensates

The yield and composition of hydrocarbon con-densates depend very strongly on the type ofcoal and the gasification conditions. In a coun-tercurrent gasification system (e.g., Lurgi pres-sure gasification), coal travels through a retort-ing zone where volatile components are vapor-ized and leave the gasifier along with the gas.When the product gas is cooled in subsequentheat exchangers and coolers, different conden-sate fractions are recovered. When the temper-ature drops below the dew point of water, wa-ter condenses along with the hydrocarbons andforms an aqueous phase in the final separators(see Section 6.1).

Two fractions are recovered from Lurgi pres-sure gasification: a tar fraction containing thehigh boilers and an oil fraction boiling in themedium range. A third light naphtha fractionis obtained when the gas stream is cooled fur-ther for low-temperature purification in a Recti-sol plant.

Figure 78. Treatment of aqueous condensates frommovingbed coal gasifier (example given in the text)* These processes are optional

Whereas the naphtha fraction can be treateddirectly in downstream units, the tar and oil frac-tions must be pretreated by conventional meth-ods to remove suspended solidmaterial (ash) andundissolved water. Typical properties and com-positions of tar and oil are shown in Table 36.

To obtain products with suitable propertiesfor use as fuel oil and gasoline, the combinedpretreated medium- and high-boiling conden-sates are separated into their respective frac-tions by distillation [278]. Distillate boiling inthe gasoline range is then refined along with thelight naphtha fraction by hydrogenation to re-move unsaturated compounds (which are poten-tial gum formers) and to eliminate sulfur and ni-

Gas Production 137

trogen.Recovery of chemically pure compoundssuch as naphthalene or cresols (isomeric mix-ture) is usually uneconomic because of the smallamounts present in the condensates [279].

Table 36. Typical composition of condensates from Lurgi pressuregasification

Tar OilRate, wt % of total 60 40Total water, wt % 10 – 15 15 – 20Suspended solids, wt % 15 – 25Properties after removal of solidsand undissolved waterSpecific gravity, d30 1.05 – 1.1 0.94 – 0.96Solidification temperature, ◦C + 15 <− 20Flash point, ◦C + 60 + 19Water, wt % 4 2Nitrogen, wt % 1 1Sulfur, wt % 0.4 0.4C6– C9 aromatics, wt % 6 16Naphthalene, wt % 3 5Methylnaphthalene, wt % 3 4Phenolic compounds as cresol,wt % 21 19Initial boiling point, ◦C 98 96Portion distilled at 360 ◦C, wt % 60 75Distillation residue, wt % 40 25

Properties of distillation residueSoftening point, ◦C 67 63Portion insoluble in toluene,wt % 10 7Portion insoluble in quinoline,wt % 2 0.1Coke (Conradson coking test),wt % 29 28

6.2.1. Distillation

In a first dewatering column, the water content islowered by distillation. The dehydrated hydro-carbon condensate is then distilled in an atmo-spheric column. A typical distillation unit yieldssix fractions (see Table 37). Light and heavynaphtha are obtained as top and first side frac-tions; they are combined with the Rectisol naph-tha and refined in a hydrogenation unit (seeTable38).

Two more side fractions, heavy and mediumcreosote, are recovered from the atmosphericcolumn. The bottoms fraction is flashed into adrum under vacuum where it separates into va-porous residue oil and pitch.

Fractions boiling above 210 ◦Care used (e.g.,as fuel or impregnation oil for wood). Pitch issuitable as a binder for briquetting of fine coal

or may be fed to a partial oxidation plant to pro-duce synthesis gas.

Table 37. Typical fractions from a coal tar distillation unit

Approximatetotal, wt %

Approximate boilingrange, ◦C

Light naphtha 4 ≤170Heavy naphtha 7 160 – 210Medium creosote 25 190 – 290Heavy creosote 21 260 – 400Residual oil 13 max. 5 vol%

≤300Pitch * 30 max. 15 vol%

≤350

* Softening point of pitch is ca. 70 ◦C; nonvolatile residue ofpitch, max. 30 wt %.

Table 38. Typical feedstocks for hydrorefining (approximatefigures)

Heavynaphtha

Lightnaphthafrom tardistillation

LightnaphthafromRectisolunit

Totalblend

Rate, wt % of total 15 25 65 100Specific gravity, d20 0.93 0.84 0.83 0.85Bromine consumption,g/100 mL

114 33 37 70

Total sulfur, wt % 0.35 0.49 0.42 0.4Mercaptans, mg/kg 20 70 80 68Phenolic oil, vol% 25 10Alkaline nitrogen, mg/kg 10 100 7600 270 5000Naphthalene, wt % 10 1.5Initial boiling point, ◦C 130 – 140 90 – 95 50 – 55 55Boiling end point at 95wt %, ◦C

200 – 210 150 – 160 140 – 145 220

Chemical hydrogenconsumption, m3 (STP)/t

360 135 60 120

6.2.2. Hydrorefining

Hydrorefining is used (1) to saturate olefins anddiolefins to their respective alkanes; (2) to con-vert sulfur in mercaptans and thiophenes to hy-drogen sulfide; (3) to convert nitrogen in pyri-dines and amines to ammonia; (4) to convert or-ganic chlorine to hydrogen chloride; and (5) toconvert oxygen in phenols to water. The feed-stock is a blend of the various naphtha streamsrecovered from condensates (see Table 38).

The Benzoraffin process, developed around1960 jointly by BASF, Veba-Chemie, and Lurgi[280, 281], is commercially successful in refin-ing the pretreated naphtha fractions derived fromcoal gasification. It is a fixed-bed, catalytic, gas-phase hydrogenation process using a cobalt –molybdenumcatalyst operating at ca. 5 MPa and350 – 420 ◦C (Fig. 79).

138 Gas Production

Figure 79. Benzoraffin process for hydrorefining of naphtha fractionsa) Feed pump; b) Vaporizer; c) Mixer; d) Prereactor; e) Heat exchanger; f) Reactor; g) Cooler; h) Separator; i) Recycle com-pressor

Modern alternatives to this process (e.g.,by BASF, Lurgi-Sudchemie, IFP et al.)— alsoused for the preparation of feedstocks forextractive distillation of toluene (→ Toluene,Chap. 3.1)—also work with two reactors, butthe first reactor contains a palladium catalystand operates in the liquid phase at low tempera-tures, preventing gum formation from diolefinsor styrene.

The feed naphtha contains highly reactivechemical compounds with a strong tendency toform resinous deposits in heaters and vaporizers.A special vaporizer developed by Lurgi avoidsfouling and frequent shutdowns for cleaning.

Process Description. Feed naphtha is vapor-ized under pressure and mixed with hydrogen-rich gas in a specially designed vaporizer (b)[281]. A small stream is withdrawn as residuefrom the bottom and returned to the tar distilla-tion unit. The mixture of gas and naphtha vapor

passes through a prereactor (d) filled with cobalt– molybdenum catalyst. Here, the most reac-tive unsaturated compounds are hydrogenated toprevent fouling in the heat exchanger (e). Reac-tion occurs at 200 – 250 ◦C. The outlet temper-ature is ca. 5 – 10 ◦C higher due to the exother-mic reaction.After further heating in the heat ex-changer (e) to about 350 ◦C, the prereactor efflu-ent enters the main reactor (f). This reactor hastwo catalyst beds and a facility for injecting re-cycle gas to adjust the reaction temperature. Thefirst bed contains cobalt – molybdenum catalystpellets; the second bed is filled with a molybde-num catalyst. The exothermic reaction leads toan effluent temperature> 400 ◦C. The gas – va-por mixture is cooled to 30 – 40 ◦C in a heat ex-changer (e) and cooler (g). Water should be in-jected into the cooler to prevent salt deposits, be-cause hydrogen chloride and ammonia formed

Gas Production 139

during the reaction tend to solidify at lower tem-perature in the form of ammonium chloride.

The condensed raffinate is separated from thegas in a separator (h). After depressurization toca. 1 MPa, the raffinate is stripped in a strip-per column to remove hydrogen sulfide, ammo-nia, carbon dioxide, and water. If it still containssome phenolic compounds, a caustic wash anda subsequent water wash are recommended. Aportion of the separated gas is withdrawn to re-move sufficient light byproducts and inerts forthe partial pressure of hydrogen to be adjustedproperly. The main stream is compressed andrecycled to the vaporizer after addition of freshhydrogen.

A typical specification for a hydroraffinatefrom the naphtha blend in Table 38 is shownbelow.

Appearance clear and free of suspended solidsColor, Saybolt 0Water max. 0.1 wt %Total sulfur max. 0.1 wt %Existent gum max. 4 mg per 100 mLAromatics ca. 75 wt %

Because of its high aromatic content, this raf-finate is a valuable blending component for gaso-line. Alternatively, pure benzene, toluene, andother aromatics may be recovered from the raf-finate by appropriate separation processes. Sol-vents with specified boiling ranges for variousapplications in the chemical industry can be ob-tained by fractionation.

6.3. Gaseous Byproducts

Sulfur Compounds (→ Hydrogen Sulfide,see → Sulfur). Hydrogen sulfide constitutes themain part of the gaseous sulfur components. Itusually accounts for 90 – 95% of the sum ofall sulfur compounds. The remainder is primar-ily COS; CS2, SO2 and mercaptans are usuallypresent in trace amounts only.

These sulfur components occur in variouswaste gases and aremore or less strongly dilutedwith CO2. Hence, processing depends totally ontheir concentration and usually aims at recover-ing elemental sulfur. In special cases, the directproduction of sulfuric acid can also be of inter-est and even low concentrated H2S gases can beused for this purpose.

There are various methods available for therecovery of elemental sulfur from low concen-trated waste gases (e.g., 1%H2S, 98%CO2 and1% inert gases). Thesemethods include concen-tration by selective absorption or by adsorption,and sometimes intermediate steps with (partial)combustion to SO2 are also involved. Processesof this type have been occasionally employed.

From a theoretical standpoint, the most eco-nomical method should be direct oxidation, ei-ther in the gaseous or in the liquid phase as per-formed in an oxidation wash, e.g., the Lo-Catprocess (see Section 5.4.3). However, it shouldbe taken into consideration that organic sulfurcompounds in the gas, primarily COS and CS2,remain practically unchanged and this couldrepresent an undue environmental burden. Onthe other hand, precautionary measures must betaken if thewaste gas contains trace components,such as HCN, that could cause solvent degrada-tion. Therefore, low investment cost are oftencounterbalanced by high operating cost for sol-vent reclaiming and/or make-up of chemicals.Today oxidation wash systems are mainly usedfor low-pressure off-gases with low sulfur con-tents, with limitations in capacity.

If hydrogen sulfide is present in concentratedform, i.e., in concentrations of more than 20%(in some cases 10% are sufficient) sulfur can berecoveredmore economically by using theClausprocess. However, the Claus tail gas almost al-ways requires an additional treating step; it usu-ally still contains 3 to 6% of the sulfur, mainlyin the form of H2S and SO2. A modern solu-tion is the integration of a three stage Doxosul-freen process (→ Sulfur, Chap. 7.2.1.1) shownschematically in Figure 80. Here, the first twocatalytic reactors after the combustion chamberare continuously operated according to the clas-sical Claus principle. After condensing the ele-mental sulfur formed in the Claus reactors, COSand CS2 are converted by catalytic hydrolysisat elevated temperature. The following reactors—each containing a Sulfreen stage and a directoxidation stage—are operated in cycles: in one,the Claus reaction is first, at a slight surplus ofH2S, continued below the sulfur dew point withadsorption of the newly formed sulfur, and then,at further reduced temperature, the residual H2Sis almost completely converted by catalytic ox-idation with air. In the second reactor, operated

140 Gas Production

Figure 80. Two-stage Claus plant with internal hot gas bypass, integrated Doxosulfreen plant and thermal incinerationa) Claus reactors; b) Sulfur condensers; c) Separators; d) Heater; e) Hydrolysis reactor; f) Cooler; g) Sulfreen and h) Directoxidation stages; i) Thermal incineration

in parallel, the sulfur is desorbed by acid recy-cle gas and is subsequently, as in the first twosteps, separated in liquid form in coolers, withthe generation of LP steam [282].

Carbon Dioxide. Today, there are basicallyfour different ways of releasing sulfur-free car-bon dioxide:

– Release into the atmosphere (there might berestrictions in the future),

– Use for enhanced oil recovery (EOR) or– As chemical raw material, and– Recovery as pure CO2 for the food and phar-maceutical industry.

Usually a large part of the CO2 waste gas isreleased into the atmosphere for lack of otheruses, although it should be taken into consider-ation that even sulfur-free CO2, in larger con-centrations, is dangerous (in Germany theMAKvalue is 5000 ppm). However, in the case of fur-naces, a sufficiently high flue-gas temperaturenormally ensures adequate dilution in the atmo-sphere; special precautions are required for coldprocess waste gases and the high carbon diox-ide density relative to air. If introduction intoa flue gas is not possible (dew point!), coolingtowers can be used provided that the danger of

corrosion is taken into account, otherwise, onlyspecial stacks may be employed.

In connection with the general discussionabout the various factors influencing the earth’sclimate and the necessity to decrease the CO2emission into the atmosphere, several solutionshave been considered theoretically, e.g., the de-positing of CO2 in absorbed form in deep layersof the oceans. However, none of these sugges-tions have been realized so far on a technicalscale.

One of the more modern applications of CO2is in enhanced oil recovery. The specification forthe CO2 to be injected in aging oil wells (to re-duce the oil’s viscosity) is usually less stringent:CO2 concentration 98%, C1/C3+ hydrocarbonsup to 0.3%, total sulfur in the 1% range. Thispurity may often be accomplished simply by de-hydration of the raw CO2 [283].

Various requirements have to be met for theuse of carbon dioxide as a chemical raw mate-rial: in the production of carbon monoxide orsynthesis gas from natural gas, which has thelowest C/H ratio of all raw materials, carbondioxide is frequently removed from the processgas by absorption and recycled to the reformer.Care must be taken that the catalyst used is not

Gas Production 141

damaged by solvent vapors or other trace com-ponents. Waste gases from hot potash, amine,and physical washes operated with sulfur-freesolvents generally cause no problems, as longas effective separators are employed which areusually followed by a re-compression step. Thiscompression step also contributes to residualseparation.

Carbon dioxide used in the synthesis of ureahas tomeetmore stringent requirements. Indeed,a maximum concentration of inert gases, usu-ally 1 – 2% of hydrogen and carbon monox-ide, should not be exceeded and sometimes eventraces of solvents can cause problems. However,these small amounts can generally be separatedby a simple adsorption step with an appropriatemolecular sieve.

The highest demands are made on the purityof carbon dioxide used in the pharmaceutical orfood industry, e.g., in the formof dry ice for cool-ing or packaging purposes or in compressed orliquid form for the production of beverages. De-pending on the impurities to be separated (evena few ppm of odorous substances are often suf-ficient to cause an unpleasant smell), usuallyseveral different, usually adsorptive, purificationsteps are necessary. More modern installationsinclude also distillation stages (to separate “sourCO2” together with heavier hydrocarbons andlight inerts) as well as oxidation stages to con-vert the reminder of undesirable hydrocarbons(Fig. 81) [283, 284].

7. Typical Examples of Complex GasProduction Plants

7.1. Methanol Production from NaturalGas

The production of methanol (→ Methanol)with natural gas (→ Natural Gas) as feedstockis described by the following process steps:(1) gas production, (2) methanol synthesis, and(3) distillation. Details of the three single pro-cess steps are described in other chapters; thischapter deals with specific features of the over-all arrangement.

Basically two different gas production tech-nologies, steam reforming and catalytic au-tothermal reforming, or a combination of both

are used to produce methanol from natural gas,depending on plant size and availability of car-bon dioxide from outside battery limits.

Synthesis gas produced by steam reformingis characterized by a relatively low pressure anda surplus of hydrogen. By addition of carbondioxide the composition of the synthesis gas canbe adjusted to be more favorable for methanolproduction.

Pure catalytic autothermal reforming resultsin a synthesis gas with a relatively high pressurebut with a deficiency in hydrogen for methanolsynthesis. A carbon dioxide removal unit is re-quired or additional hydrogen has to be providedto adjust the appropriate gas composition. Suchhydrogenmay be obtained either from themeth-anol synthesis purge gas and from additionallyproduced reformed gas or from a conditionedmethanol synthesis purge gas.

The combinationof both technologies, the so-called combined reforming process, has the ad-vantage of an optimum composition and a highpressure of the produced synthesis gas.

For the methanol synthesis itself a widerange of technologies are available, which differmainly in the design of the methanol converter.In principle, three different types of convertersare used in the industry; the quench-cooled, thegas-cooled, and thewater-cooled reactor. For op-timization of the reaction and for capacity in-crease also combinations of these technologiesare applied.

In the distillation section the raw methanolproduced in the synthesis section is purified tomeet the required specification. Depending onplant capacity and optimization with regard tocost or energy savings, a two or three columnconcept can be used.

At locationswhere no carbon dioxide is avail-able most of the methanol plants are based onthe following gas production technologies, de-pending on their capacities: steam reforming forcapacities up to 2000 t/d, combined reformingfrom 1800 to 2500 t/d [285, 286].

Erection of plants with a capacity of 5000 oreven 10 000 t/d is also discussed nowadays, e.g.,for minimizing flaring of associated gas in oilfields by producing methanol as a “liquid stor-age”. For this size of plants pure catalytic au-tothermal reforming in combination with a highefficient conversion of synthesis gas to metha-nol is the most economical way. The process is

142 Gas Production

Figure 81. Purification of CO2 for the production of beveragesa) Fractionation; b) Adsorptive sulfur guard; c) Catalytic oxidation; d) Adsorptive dryers

designed in such a way that conditioned purgegas with the required hydrogen content is ob-tained to adjust the reformed gas composition.This technology is illustrated in the simplifiedprocess flow diagram in Figures 82 and 83 anddescribed in Section 7.1.1.

Steam reforming technology and combinedreforming technology are discussed in Section7.1.2.

7.1.1. Methanol Production Based onCatalytic Autothermal Reforming (Figs. 82and 83)

Gas production, methanol synthesis, and subse-quent distillation are the primary process steps.The relevant units are combined with an inte-grated steam system usually on several pressurelevels. Steam is generated and conditioned inthe synthesis, in the reformed gas and flue gaswaste-heat system, and in the auxiliary boiler.

Natural gas is preheated anddesulfurized.Af-ter desulfurization the gas is saturated with amixture of preheated process water from the dis-tillation section, process condensate, and make-up water in the saturator (d). The gas is furtherpreheated and mixed with process steam as re-quired for the prereforming process. In the pre-reformer (f) the gas is converted to hydrogen,carbon dioxide, and methane. Final preheatingof the gas is achieved in the fired heater (b). Inthe autothermal reactor (g) the gas is reformedwith steam and oxygen, the product gas con-tains hydrogen, carbon monoxide, carbon diox-ide, and a very small amount of non-convertedmethane and inerts together with undecomposedsteam. The reformed gas leaving the autother-

mal reactor possesses a considerable amount ofheat, which is recovered in form of high pres-sure steam, preheating energy, and energy forproviding heat for the reboilers in the distilla-tion section.

The reformed gas is mixed with hydrogenfrom the pressure swing adsorption unit to adjustthe synthesis gas composition. Synthesis gas ispressurized to 5 – 10 MPa by the synthesis gascompressor (o). Subsequently, this compressedgas is mixed with recycle gas from the synthesisloop. With the recycle gas compressor (t), therecycle:make-up gas ratio can be adjusted. Theinlet gas to the reactor system consists mainly ofhydrogen, carbonmonoxide, and carbon dioxideas reactive components, alongwithmethane andnitrogen as inerts. This gas mixture is convertedto methanol according to the following exother-mic reactions:

CO+2H2�CH3OH

CO2+3H2�CH3OH+H2O

Details of the reaction mechanism are de-scribed elsewhere (→ Methanol). The synthesisreactor system (q) is the key section of the loop.Here superheating of the catalyst is avoided bygeneration of medium pressure steam and thereaction parameters are selected in such a waythat properly conditioned purge gas is obtained.

In the pressure swing adsorption unit hydro-gen is separated from the purge gas and routedto the reformed gas. The remaining gas is usedas fuel in the fired heater.

The distillation unit is a three-columnenergy-saving system. In the light-ends column (u),low-boiling impurities are removed overhead.

Gas Production 143

Pure methanol distillation is divided into twosections. The first section (v) operates at ele-vated pressure, and ca. 50% of the methanolis distilled overhead. In the second section (w),which operates at ambient pressure, the remain-ing methanol is separated from water and high-boiling impurities.

Whereas the first section is heated with low-pressure steam, heat for the second section isprovided by condensing methanol vapor fromthe first section.

7.1.2. Comparison of Conventional SteamReforming and Combined ReformingProcesses for Methanol Production

The main reason for applying the combined re-forming process in methanol production is ad-justment of the stoichiometric number (SN) asdefined in Table 39, closer to the theoreticalvalue of 2.0. Whereas the practical SN is bet-ween 2.6 and 2.9 for conventional steam reform-ing, it is decreased to 2.03 in the combined re-forming process.

Table 39. Conventional steam reforming versus combined processfor methanol production

Conventionalsteamreforming

Combinedprocess

Process natural gas feed to tubularreformer,%

100 40 – 60

Process natural gas feed toautothermal reactor,%

60 – 40

Outlet pressure tubular reformer,MPa

1.9 3.7

Outlet pressure autothermalreactor, MPa

3.5

Outlet temperature tubularreformer, ◦C

875 780

Outlet temperature autothermalreactor, ◦C

950

Hydrocarbons converted in tubularreformer,%

83 22

Stoichiometric number, SN =(CH2−CCO2

)/

(CCO+CCO2

)2.6 – 2.9 2.03

Total natural gas lower heatingvalue demand, GJ *

31.6 29.5

Thereof fuel consumption ofreformer, GJ *

12.9 3.1

Discharge – suction pressure ratio(synthesis make-up gascompressor)

4.4 2.2

Instrument air ** separate unit byproductPlant air ** separate unit byproductNitrogen ** separate unit byproduct

* Per t of methanol.** With oxygen generation in an air separation plant.

Besides reducing the SN, placement of an au-tothermal reactor downstream of the tubular re-former means that the tubular reformer can bedrastically reduced in size for two reasons: first,it has to treat only about half of the natural gasfeed while the other half is by-passed directly tothe autothermal reactor and, second, it does nothave to convert all of the methane but leaves anessential part for the secondary reforming step.So the fired steam reformer may be ca. 75%smaller, can operate at lower temperature, andtherefore, can operate at higher pressure thana steam reformer in the conventional steam re-forming process.

Because of the lower SN, less surplus hydro-gen must be compressed and introduced into thesynthesis loop, which results in a significant re-duction of equipment size and compression en-ergy. Table 39 shows the main operating data ofthe combined reforming process in comparisonwith the conventional steam reforming process.

7.2. Hydrogen Production Based onHeavy Residues

Economic and environmental considerationshave led petroleum refiners to intensify effortstomaximize production of high-quality productsfrom crude oil. Tomeet this goal, a number of re-finery processes have been developed to upgradeheavy residues to, for instance, naphtha and gasoil. Some of these conversion processes (e.g.,hydrotreating or hydrocracking) require consid-erable quantities of hydrogen. Todaymost of thehydrogen is produced by steam reforming of nat-ural gas (83% of total production) or naphtha(14%). Only 3% is obtained by gasification ofheavy residues [287]. However, though the cap-ital investment for such a plant is considerablyhigher than for a unit based on steam reformingof natural or refinery gas, gasification of residuescan still be advantageous, particularlywhere nat-ural gas or other light hydrocarbons are not read-ily available or expensive.

In some refining schemes, particularly thosefor treating heavy crudes, the hydrogen bal-ance of the refinery may require utilization ofthe residues for hydrogen production. In addi-tion, partial oxidation of heavy residues can con-tribute to reducing the size of the fuel oil pool

144 Gas Production

Figure 82. Methanol from natural gas, catalytic autothermal reforming route—production of synthesis gasa) Natural gas interchanger; b) Fired heater; c) Desulfurizer; d) Saturator; e) Natural gas preheater; f) Prereformer;g) Autothermal reactor; h) Waste-heat boiler; i) Circulation water heater; j) Distillation reboilers; k) Final cooler; l) Finalseparator; m) Oxygen preheater

Figure 83. Methanol from natural gas, catalytic autothermal reforming route—methanol synthesisn) Synthesis make-up gas compressor turbine; o) Synthesis make-up gas compressor; p) Heat exchanger; q) Synthesis reactorsystem; r) Cooler; s) Crude methanol separator; t) Recycle gas compressor; u) Light ends column; v) Pure methanol pressurecolumn; w) Pure methanol atmospheric column

and increasing its quality by consuming some ofthe highest sulfur- andmetal-containing compo-nents of the pool [288].

A typical hydrogen specification for a hydro-cracker is

Gas Production 145

Hydrogen ≥98%Carbon monoxide and carbon dioxide ≤10 – 50 ppmOxygen ≤100 ppmInerts (N2, Ar, CH4) balance

The raw gas supplied by a partial oxidationunit fed with a heavy residue contains nearly50% carbon monoxide, as well as other com-ponents (CO2, H2S, COS) which must be con-verted or removed in a number of gas treat-ment stages to yield acceptable product qual-ity (see Table 24, example 3). In practice, gasplants based on partial oxidation are often mul-tiproduct plants, which in the refinery environ-ment could imply cogeneration of clean fuel gasor, for example, methanol for methyl-tert-butylether (MTBE) (→ MethylTert-ButylEther) pro-duction.

Selection of gas treating processes must bemade carefully, because ca. 20 – 30% of the to-tal investment for an overall hydrogen plant isconcentrated in these units. Plant size and possi-ble cogeneration of other products also influencethe final choice of process.

For the partial oxidation unit, either of thethree commercially proven processes (Shell,Texaco, Lurgi) is suitable.

Conventional Gasification-Based Hydro-gen Production. The principal process steps forhydrogen production based on heavy residues asfeedstock are as follows (see Figure 55):– Gasification– Acid gas removal (H2S, CO2), in one or two

separate units– CO shift conversion– Final purification

The byproducts have to be handledwithin therefinery (steam, fuel gas) with the appropriatespecifications or should be salable (sulfur, metalash).

As outlined in Section 3.2.2, the gasifica-tion process itself can be applied in two differ-ent configurations followed by different down-stream treatmentmodes of the raw synthesis gas:1) Gasification in boiler configuration with

downstream CO shift of desulfurized syngas.2) Gasification in quench configurationwith raw

gas shift (see Section 5.1) prior to desulfur-ization (and CO2 removal).

In the boiler configuration, the sensible heat ofthe hot raw gas leaving the gasification reactor

is utilized to generate high-pressure steam in awaste-heat boiler before the gas enters the scrub-ber. After the subsequent desulfurization and aclean gas CO shift conversion, CO2 must be re-moved from the syngas by a second washingstep. This route has been chosen for the ShellPernis plant [289]. The disadvantages of thisconcept are high capital costs due to the ex-pensive gasification stage (waste-heat boiler),CO shift conversion (with cooler-saturator cir-cuit) and the two-stage acid gas removal unitalong with limited feedstock flexibility due tothe boiler specification. An asset is a rather highthermal efficiency.

In the quench configuration, the raw gas leav-ing the gasifier is shock-cooled by direct injec-tion and subsequent scrubbing with water. It canthus be routed directly to the sulfided CO shiftconversion (raw gas shift). The steam content ofthe quenched gas is normally sufficient for theshift reaction, which leads to a quite simple lay-out of the shift conversion plant. Furthermore,trace components like COS and HCN are de-stroyed in the shift reactor, which facilitates thesubsequent acid gas removal. This concept hasbeen proposed by Texaco [290] and Lurgi [291,292]. The advantage of this concept is lower cap-ital cost for both the gasification unit and theshift. A disadvantage is the loss in thermal effi-ciency.

Improved Route for Hydrogen Produc-tion. In order to further increase the economicsof a gasification-based hydrogen production unitcompared to a further steam reforming step, theabove described quench configuration route hasbeen improved by Lurgi introducing the MultiPurpose Gasification together with an advancedash handling system and an optimized CO shiftconversion. The new concept consists of the fol-lowing process units (Fig. 84):

– Air separation and oxygen compression,– Multi Purpose Gasification (MPG) in quenchconfiguration,

– Metals Ash Recovery System (MARS),– Sulfided CO-Shift (single-stage raw gasshift),

– Sulfur removal (Rectisol process),– Pressure swing adsorption (for remainingCO2 and low concentrated impurities),

146 Gas Production

Figure 84. Block flow diagram of the MPG based hydrogen productionASU = Air separation unit;MPG = Multi PurposeGasification;MARS = MetalsAshRecovery System; PSA = Pressure swingadsorption; SDA = Solvent deasphalting; BFW = Boiler feed water; CW = Cooling water

– Sulfur recovery unit (Claus process).

Two main features make this concept mosteconomic compared to conventional processroutes:

1) TheMPGprocesswith its high feedstock flex-ibility allows the simultaneous processing ofeven the heaviest refinery residues andwastes,which can reduce the costs for feedstock sup-ply (see below).

2) In the sulfided CO shift, the raw gas is con-verted in a single-stage reactor which resultsin an increased CO content in the convertedgas. On the one hand, this leads to a lower H2yield and therefore requiresmore feedstock tobe processed in gasification. The advantages,however, are:– A self-burning PSA off-gas which can di-rectly be used as fuel gas in an auxiliaryboiler or sent to the refinery’s fuel gas sys-tem. There is no need for fuel import.

– The capital costs for the overall plant arereduced because of the cheaper CO shiftconversion and the cheaper CO2 removal.

A simplifiedprocess flowdiagramof the gasi-fication and shift conversion section is shown inFigure 85. The raw gas is produced by partialoxidation of heavy residues using Lurgi’s MultiPurpose Gasification Process in quench config-uration at 6.0 MPa. The water-saturated raw gasfrom the quench vessel enters the scrubber at ca.240 ◦C, where residual soot/ash traces are re-moved by injection of additional quench water.No further raw gas cooling occurs in the scrub-ber. The water leaving the quench section con-tains most of the soot and ash from the raw gasand is further treated in theMetalsAshRecoverySystem (see Section 3.2).

The raw gas from the MPG unit is saturatedwith water and routed to the CO shift conversionunit. After preheating of the raw gas, its CO con-tent is reduced in a single-stage reactor to a resid-ual 6 to 9%, depending on the composition ofthe feedstock.No further steam injection into theshift feed is necessary to maintain the shift reac-tion. The sensible heat of the shifted gas is reco-vered by producing steam of different pressurelevels and by preheating several water streams(quench water, demineralized water). The con-

Gas Production 147

Figure 85. Simplified process flow diagram for the gasification (quench configuration) and the CO shift section in Figure 84

densate from the cooled shift product gas is sentto the MPG unit as quench and scrubbing water.

Both carbonyl and hydrogen sulfides arewashed out of the gas to a residual 1 ppm in aone—stage Rectisol unit. The factors determin-ing the selection of the Rectisol process insteadof a chemical absorption system are– Low steam consumption,– High selectivity for H2S versus CO2 for the

generation of a rich acid gas to facilitate thesulfur recovery.(see 5.4.2.4) for detailed dis-cussion on acid-gas removal options).The acid gas, which contains 25 – 60% hy-

drogen sulfide, depending on the sulfur contentof the feedstock, is normally discharged to a cen-tral Claus plant that also treats other sulfur con-taining gas streams originating in the refinery toproduce elemental sulfur.

In the final purification of the raw hydrogenin the pressure-swing adsorption unit CO2 andother impurities are reduced to 10 ppmof carbondioxide, 10 ppm of carbon monoxide, and 0.1%(maximum) of inerts, yielding hydrogen with apurity > 99.9% (see Section 5.5) at a pressurebetween 4.7 and 4.8 MPa.

Plant Performance. For easier comparison ofthe improved concept with conventional processroutes, three cases are considered for the eco-nomic evaluation:

Case 1: =MPG (quench) + single-stage rawgas shift + H2S absorption + PSA

Case 2: =MPG (quench) + two-stage raw gasshift + acid gas absorption (H2S, CO2) +PSA (for remaining impurities only)

Case 3: = Texaco Gasification (TGP, quench) +two-stage rawgas shift + acid gas absorp-tion (H2S, CO2) + PSA (based on [290],scaled to hydrogen production withoutpower generation)

The characteristics of the respective plantperformance are summarized in Table 40.

To compensate for the lower H2 yield of theMPG/single-stage CO shift process variant, theplant capacitymust be by6%higherwhich leadsto higher variable operating costs. On the otherhand the simple plant configuration of this routeresults in lower capital costs and fewer main-tenance costs and therefore reduced fixed costswhich makes this concept the most economicalone.

7.3. Combined Cycle Power SystemBased on Coal

7.3.1. Introduction

Combined cycle power plants are an attractiveoption to raise the efficiency of electric powergeneration. This power plant concept combinesa gas turbine and a conventional steam turbine

148 Gas Production

Table 40. Performance of improved-concept plants for hydrogen production from heavy residue

Case 1 Case 2 Case 3MPG, single-stage CO shift MPG, two-stage CO shift TGP [290]

Inputs SDA pitch (S: 5 wt %, ash: 0.07wt %)

SDA pitch (S: 5 wt %, ash: 0.07wt %)

Asphalted vacuum residue

Feed oil, t/d 960 905 1050Oxygen (99.5%), t/d 925 875 1120Process steam to gasifier, t/d 430 405 475Intermediate productsFuel gas from PSA (LHV, kJ/m3 (STP) > 4500 ca. 8000 ca. 1000Acid gas from Rectisol unit (vol% H2S)< 60 < 60 not availableProductsHydrogen, t/d 188.4 188.4 218Byproducts, t/dSulfur 47.8 45.1 58Metals, ash 0.6 (76 wt % V2O5) 0.56 (76 wt % V2O5) 8.5 (filter cake, dry)Steam export 1920 1810 2110 (calculated)

cycle. Based on available industrial gas turbinesearly plant concepts were developed and intro-duced in the early 1950s, these plants were fu-eled by conventional oil-derived gas or almostanykind of fuel gas.Althoughbasic process con-cepts are well established, novel concept varia-tions are still under development [293].

Due to its relatively low price as well as re-gional availability, coal is the major fuel forfossil-fueled power plants. To use combined cy-cle technology, coal must be transformed intoa clean fuel applicable to gas turbine operation(the same applies to residual oil, petroleumcoke,or wood). Impurities, which must be removed,are ash, alkali, halogens, and sulfur. Typical gasturbine specifications are given below:

Inlet temperature, ◦C (according to ISO) 1190Na + K, mg/kg 0.08V, mg/kg 0.1Pb, mg/kg 0.25C, mg/kg 2.5Cl, mg/kg 2.45Dust < 2 µm, mg/kg 4.4

The only way to meet these requirements isgasification because

1) Ash remains in the gasification residue,2) Alkali is trapped within the ash or removed in

one of the purification steps, and3) Halogens and sulfur are removed by dry or

wet absorption.

The combined effect of high efficiency andlow pollution makes the combined cycle withupstreamcoal gasification an attractive approachfor power generation from coal. In principle,two designs are possible: integrated and nonin-

tegrated plants. In the latter, the gasification andgas purification process units are installed andoperated separately from the power plant. Thismakes power plant operation easier, but an in-tegrated plant (IGCC) allows for more sophisti-cated energy management and higher efficiency[294]. Hence, most engineering and economicstudies and first-of-its-kind plants refer to thisdesign principle. In this section only the princi-ples of the integrated combined cycle based onthe gasification of solid fuels are discussed [295,296].

Environmental Aspects. The emissions ofcombined power cycles are inherently lowerthan that of coal combustion plants [297 – 299].This is due mainly to the following characteris-tics:1) Desulfurization of the fuel gas can be accom-

plished to any extent necessary at a relativelylow additional cost. For applicable processes,see Section 5.4.

2) The byproduct of desulfurization is elementalsulfur, which can be sold or easily disposedof, in contrast to flue gas scrubber residues.

3) LowNOx emissions are achievable due to thetightly controllable combustion conditions ina gas turbine.

4) Extensive particulate removal is essential forgas turbine operation, resulting in dust emis-sions far below existing requirements.

7.3.2. Fundamentals

Thermodynamics [300, 301]. Electricpower is conventionally generated by com-bustion to produce high-temperature thermal

Gas Production 149

energy and converting that thermal energy intowork to drive a generator. The working fluid canbe the combustion gas itself or a special workingmedium (steam, air) to which heat is transferredfrom the combustion gas.

The upper limit to the proportion of thermalenergy converted to work is set by Carnot’s law

η≈1− t2

t1

where t1 and t2 are the upper and the lower ab-solute temperatures of the working fluid and η isthe conversion efficiency. For a steam turbine cy-cle constructionmaterials usually limit the uppertemperature to ca. 560 ◦C, the lower temperaturedepends on the available cooling water tempera-ture. With sea water cooling lower temperaturesof ca. 30 ◦C are achievable. For an open cyclegas turbine (stationary application), t1 may beas high as 1050 ◦C, with t2 being ca. 500 ◦Cdepending on the pressure ratio.

When gas and steam turbines are combinedby using the thermal energy of the gas turbineexhaust gas to produce and superheat the steamfrom the steam cycle, t1 may be as high as1050 ◦C and t2 as low as 30 ◦C. This, in simpli-fied terms, is the basis for the efficiency increasewhen a dual cycle combining gas and steam tur-bines is used instead of a single working fluidcycle. Some losses occur along the route fromcoal to electrical energy. For simplification theoverall efficiency of the process can be writtenas the product of individual efficiencies:

η = ηCG·ηGC·ηCCwhere ηCG is the efficiency of coal gasification;ηGC, the efficiency of gas cleaning; and ηCC,the efficiency of the combined cycle.

Efficiency of Coal Gasification ηCG. Coalgasification processes suffer losses of useful en-ergy due to the following mechanisms:

1) The heat of evaporation of gasification steam,coal moisture or slurry water cannot be reco-vered.

2) The bottom or fly ash inevitably contains un-converted carbonwhich is lost for the process.The same applies for the sensible heat of theash.

3) Heat is lost to the environment, e.g., throughthe walls of apparatus and pipes.

4) The operation of drives, pumps and compres-sors consumes electric energy, which can berecovered only partly.

5) The gasification reactions require elevatedtemperatures. During cleaning of the hot rawgases, according to the requirements of thegas turbine operation and of the environment,the raw gas has to be cooled down by heatexchange which involves a certain degree ofirreversibility.

Each of the mechanisms leads to a differentamount of losses for the individual gasificationprocesses. Reference figures for ηGC may betaken from Section 4.4, 4.5, 4.6, 4.7 (Tables 18,20, 21, 23).

Efficiency of Gas Cleaning ηGC. Thelosses of useful energy during gas cleaning canmainly be attributed to

1) Nonreversible heat exchange during coolingof the raw gas and reheating of the clean gas,

2) The energy required for the regeneration ofthe solvent or adsorbent, which in most in-stances is lost,

3) The removal of sulfuric components and theco-removal of other combustible constituentsof the raw gas which, therefore, cannot con-tribute to the combined gas/steam turbine op-eration,

4) The removal of inerts such as H2O and CO2which reduces the mass flow through the gas-turbine and hence its energy output, and

5) The consumption of electric energy.

In most cases ηGC reaches 96 – 98%.

Efficiency of Combined Cycle ηCC. Forthermodynamic reasons the efficiency of theconversion of thermal energy intowork and con-sequently electric power is limited as has beenexplained earlier. In addition to that, other lossesof useful energy must be considered, e.g.:

1) The heat of evaporation of boiler feed wateris only partly used for generation of power.

2) The nonideal operation of gas and steam tur-bines increases the entropy of theworking flu-ids.

3) The flue gases leave the stack with a tempera-ture above 100 ◦C, their thermal energy is notused for the process.

4) Heat losses through the insulation.

150 Gas Production

Based on natural gas, a combined cycle witha gas turbine inlet temperature of 1100 ◦C (ac-cording to ISO)may reach 48% efficiency [299,302].

At a coal gasification efficiency of ca. 90% acombined cycle power generation unit based oncoal gasification achieves an overall efficiency η> 42%.

General Plant Arrangement. In principle,any high-pressure coal gasification processmen-tioned in Chapter 4 can be used as the front-end process. The selection of a gasification pro-cess depends on several boundary conditions,the most important of which are

1) Type of available coal (rank, particle size, ash,and moisture content);

2) Gasifier efficiency; and3) Maturity of the process and its proven relia-

bility.

Coal gasification can be integrated with com-bined cycle power generation in numerousways.Modifications differ in the degree of carbon con-version (partial or total gasification), the gasi-fication medium (air or oxygen), the measurestaken to adjust the turbine inlet temperature(excess air or steam production), and the gen-eral arrangement (stand-alone plant or repow-ering). Advantages and limitations of the indi-vidual variants are described in [294, 297], and[303 – 307].

7.3.3. Installations and Design Studies

Installations. The first commercial-scaleIGCCplants have been built at Lunen bySteag inthe Federal Republic of Germany [308] and nearBarstow, California (The CoolWater plant). TheLunen plant with a design output of 170 MWoperated between 1972 and 1977; the Cool Wa-ter plant (100 MW) began operation in 1984. Atpresent both plants are no longer in operation.

Lunen IGC Plant. The general arrangementof the Lunen plant is shown in Figure 86. Gasi-fication takes place in 4 + 1 Lurgi moving bedgasifiers (see Section 4.4).High volatile subbitu-minous coal is gasified with steam and air coun-tercurrently at a pressure of 2 MPa (c). The rawgas leaves the gasifierswith a temperature of 500– 600 ◦C. It is cleaned in a scrubbing unit (d)

where dusty and tarry compounds are removedfrom the gas. Simultaneously the gas is cooled toca. 160 – 170 ◦C and saturated with water. Thedusty tar removed from the gas is recycled in to-tal to the gasifiers after separation from the washwater. The clean gas is passed through a first gasturbine (e) which drives the booster compressor(b) for the gasification air. Here the pressure ofthe clean gas is reduced from 2 to ca. 1 MPa. Abypass stream of 1% of the clean gas is desul-furized with a hot potassium wash solution forpilot test purposes.

Figure 86. Simplified flow diagram of the IGCC Lunenplanta) Air compressor; b) Gasification air compressor; c) Lurgigasifiers (in all five); d) Wet gas cleaning; e) Expansionturbine; f) Supercharged boilers (in all two); g) Gas tur-bine; h), i) Boiler feed water preheaters; j) Steam turbine;k) Condensor; l) Boiler feed water pump; m) Generator

The clean gas is combusted in two super-charged boilers (f) at 1 MPa which is a partic-ular feature of this plant. In these boilers 340 t/hsteam are generated (13 MPa, 525 ◦C) power-ing a steam turbine (j) with an output of 96 MW.

Gas Production 151

The flue gases leaving the supercharged boilersdrive a second gas turbine (g) with an output of74 MW. The compressor for the total plant airis connected to the same shaft. In the gas tur-bine the flue gas enthalpy is reduced from 1.0Pa, 810 ◦C down to 0.1 MPa and 400 ◦C. Theresidual thermal energy of the exhaust gases isused for boiler feedwater heating (i, k), the stackinlet temperature is 170 ◦C [297, 309].

As a first-of-its-kind plant, the Lunen plantdid not operate continuously; altogether it pro-duced electrical energy for the grid for ca. 10000 h. During this time, the data in Table 41were obtained.

Table 41. Performance data, in percent, of the Lunen plant at 70%partial load

Efficiency Partial load 100% load

Design Operation Design

Designcoal

Coal 1 * Coal 2 ** Design coal

ηCG 91.7 90.2 86.1 92.9ηCC 35.8 35.2 35.0 39.4η 32.8 31.7 30.5 36.7

* Coal 1 = design.** Coal 2 = low grade (40% < 5 mm, 28% ash).

The plant generally operated at partial loadand with low-grade coal. Table 41 also showsestimated full load design data which were notachieved because only 70% of the necessarygasification air could be supplied due to limi-tations in air compression.

The second commercial-scale installation isthe Cool Water plant [310, 311]. The plant is ofcommercial size, with a net output of ca. 100MW.

In this plant, a Texaco entrained-flow gasifier(see Section 4.6)was used,which gasified a coal– water slurry with oxygen. According to pre-vailing emission standards, the gasification rawgas was purified by a physical absorption pro-cess (Selexol, see Section 5.4), which reducesthe content of carbonyl and hydrogen sulfide inthe purified gas to 50 ppm. Besides demonstrat-ing the IGCC process principles the Cool Waterplant was used to test construction materials andgeneral control strategies.

Design Studies. The two plants have notbeen designed to exploit the full potential ofIGCC. Their main purpose was the demonstra-

tion of the technical feasibility and operabilityof plants of this kind.

Hence the overall efficiency of these earlyIGCC plants is limited compared to modernpower stations based on the Rankine process.

The successful introduction of natural gasfired IGCC plants stimulated numerous designstudies based on gasification of solid fuels.Gasification processes which were consideredare the Lurgi fixed-bed gasification [297], theBritish Gas/Lurgi slagging gasifier [304, 312,313], entrained-flow gasification [304, 305], andfluidized-bed gasification [305, 314 – 316]. De-pending on constraints like available fuel qual-ity and sophisticated heat recovery systems effi-ciencies η � 40% can be expected.

The most recent commercial scale IGCCpower plants based on coal gasification havebeen commissioned at Buggenum (NL) (250MW) by DEMKOLEC B.V. [317 – 319],at Puertollano (Spain) (320 MW) [320];and Tampa (USA) [322]. The DEMKOLECBuggenum plant applies entrained-flow gasifi-cation for turbine fuel production; standard fuelis conventional hard coal (ca. 2000 t/d). Com-mercial operation started in 1994. In a singletrain 250 MW of electric power are generated;based on the experience gained so far, capac-ities of 500 MW are regarded as feasible. Theblock diagram (Fig. 87) depicts the main build-ing blocks of a typical IGCC layout followingthe DEMCOLEC design; Figure 88 presents anaerial view of the total plant . A major partof the setup is taken up by the gas cleaningand desulfurization systems. A conventionalClaus – SCOT plant converts H2S to elementalsulfur. The IGCC plant at Puertollano uses anentrained-flow gasifier to convert low grade sub-bituminous coal into fuel gas for a gas turbine[320, 321]. This project is jointly executed byseveral companies of different European coun-tries and funded by the European commission.The Tampa IGCC plant was commissioned in1996, using entrained-flow gasification; in ad-dition to hard coal refinery residues can be usedfor gasification [322].

Besides these plants in operation, numer-ous studies investigated the suitability of IGCCplants for a wide range of feedstocks [314, 316,323]. IGCCconcepts have also been investigatedwith regard to the gasification of unconventionalfuels [319, 323 – 325]. The use of biomass de-

152 Gas Production

Figure 87. Principle layout of an IGCC plant

rived fuels opens the possibility for a reductionof CO2 emissions whilst retaining high powergenerating efficiencies.

Figure 88. Aerial view of the DEMKOLEC plant

8. Analysis and Quality Control

Fuel gases [336] are sold principally as a sourceof energy, but in units described by volume.Con-sequently, considerable demand exists for spec-ification of the thermal content per unit volume.The range of fuel gases that will burn satisfacto-rily on a particular appliance is limited and bestdescribed by the Wobbe index of the gas. TheWobbe index is a measure of the heat input to anappliance at constant pressure, and is expressedas the ratio of the calorific value to the square

root of the relative density. This property shouldalso be specified. Safety requirements relate tothe flammable nature of the fuels, and engineer-ing considerations relate to their distribution andstorage.

Specifications for fuel gases cover variousproperties and may be devised or imposed bydifferent authorities. Requirements for the spec-ification of and means of ensuring satisfactoryodorization are commonly a responsibility ofgovernment. Control ofWobbe index and of par-ticular components that may have operationalimplications, such as water and condensablehydrocarbons, are usually the responsibility ofthe supply and distribution company. Differentcountries have different specifications [337], al-though common interests ensure a degree of con-sistency.

If a fuel gas is to be used as a chemical feed-stock, quite different specifications will be used;total carbon content and the quantity and typesof sulfur compounds present might be of muchgreater importance.

8.1. Quality Specifications

Natural gas supplies by far the greater part ofthe fuel gas market. The quality specifications

Gas Production 153

Table 42. Classification of fuel gases according to gas family (DVGW a Rule G 260) [340]

Symbol Unit First Family (town gas) Second family (natural gas) Third family(LPG)

Group A Group B Group L b Group H c

Wobbe index Ws

Total range MJ/m3

(STP)23.0 – 28.1 28.1 – 33.5 37.8 – 46.8 46.1 – 56.5 requirements

forNominal range MJ/m3

(STP)44.6 54.0 propane and

Range of variation inlocation distribution

Ws MJ/m3

(STP)+ 0.6 to − 1.4 + 0.7 to − 1.4 propane –

butanemixtures (≥ 60wt %

Gross calorific value Hs C4hydrocarbons)

Total range MJ/m3

(STP)16.6 – 19.8 18.0 – 21.2 according to

DINNominal range MJ/m3

(STP)17.6 19.8 51 622

Range of variation inlocation distribution

Hs MJ/m3

(STP)± 0.3 ± 0.3

Relative density d 0.40 – 0.60 0.32 – 0.55 0.55 – 0.704Appliance inlet pressurepeTotal range kPa 0.75 – 1.5 0.75 – 1.5 1.8 – 2.4 4.25 – 5.75Nominal range kPa 0.8 0.8 2.0 5.0

a DVGW = Deutscher Verein des Gas- und Wasserfaches e.V.b L = low.c H = high.

of natural gas have been reviewed in both Euro-pean and North American contexts [338, 339].

8.1.1. Combustion Characteristics

Fuel gases are categorized in one of three fami-lies: (1) manufactured gas (high hydrogen con-tent), (2) natural gas and (3) liquefied petroleumgas [340]. Table 42 shows these families and thesubdivisions for manufactured and natural gas,within which interchangeability [341, 342] maybe assumed. Test gases for evaluation of appli-ances within these families are described [343].

Other combustion properties are flame lift[344 – 346], incomplete combustion (yellow tip-ping) and flashback [347].

8.1.2. Minor Constituents

Specifications from several (mainly European)countries have been compared [337]. Typical na-tional requirements for maximum levels of im-purities (Table 43) [340] and odorization [348]are described.

Table 43. Recommended limits of gas impurities according toDVGW * Rule G 260 [340]

Gas family

First SecondBenzene hydrocarbons,g/m3 (STP) 10

Naphthalene, mg/m3 (STP) 50/p (p in bar)Hydrocarbon dew point, ◦C soil temperature atWater dew point, ◦C pipeline pressureFog, dust, liquids technically pureO2 (dry distribution grid),vol% 3 3

O2 (wet distribution grid), vol% 0.5 0.5NO, cm3/m3 (STP) 0.2Total sulfur, mg/m3 (STP) 200 120

Short-term value,mg/m3 (STP) 150

Mercaptan sulfur, mg/m3 (STP) 6Short-term value, mg/m3 (STP) 16

H2S, mg/m3 (STP) 2 5Short-term value, mg/m3 (STP) 10

NH3, mg/m3 (STP) 3CN, mg/m3 (STP) 150CO, vol% 6

* DVGW = Deutscher Verein des Gas- und Wasserfaches e.V.

154 Gas Production

8.2. Test Methods

8.2.1. Determination of CombustionCharacteristics

8.2.1.1. Calorimetry

The calorific value (heating value) H of a fuelgas is defined as the amount of heat released bycomplete combustion in air, at constant pressurep1, with the combustion products being returnedto the same temperature t1 as that of the reac-tants prior to combustion. Normally the superiorcalorific value HS (synonymous higher heat-ing value, HHV) is experimentally determinedwith all the water formed by combustion beingcondensed to the liquid state at t1. The inferiorcalorific value HI (synonymous lower heatingvalue, LHV) is defined as the amount of heatreleased with all the water formed remaining inthe vapor phase at t1 and p1.

The calorific value H can be measured or in-ferred in three ways: (1) direct combustion withmeasurement of heat released, (2) stoichiomet-ric combustion with measurement of gas – airratio, and (3) calculation from composition.

Direct Measurement. The calorific valuecan be measured directly in a referencecalorimeter, in which the heat liberated from adiscrete amount of gas ismeasured thermometri-cally [349]. More frequently, however, continu-ous recording calorimeters are used [350 – 354].The principle of continuous recording calorime-ters is combustion of a controlled stream of gas,followed by heat exchange of the products ofcombustion with another fluid, whose tempera-ture rise is recorded.

Stoichiometric Combustion. Stoichiomet-ric combustion relies upon the constancy of theratio of calorific value to air – gas ratio for sat-urated hydrocarbons. For this reason, it can beused for natural gas, but not manufactured gas[355]. Stoichiometric combustion devices havebeen used as part of an energy flow meter [356],where the rate of supply of gas to the meter iskept exactly proportional to the flow in the mainstream from which it is taken.

Calculation from Composition. The com-position of fuel gases can be defined very well

by gas chromatographic measurement. Knowl-edge of calorific values for the pure componentsand an appropriate calculation routine [357, 358]lead to the calorific valuesHS andHI of themix-ture.

8.2.1.2. Density

Gas Density Balance. The gas density (mass per unit volume) or relative density d (ra-tio of the density of the fuel gas to the density ofdry air) [359] is most commonly determined bya gas density balance [360, 361]. The method isbasedon the buoyancyof a nitrogen-filled spherein the test gas. Instruments of this type also allowcontinuous monitoring of gas density.

Oscillating Device. Another method for de-termining density employs an oscillating device(e.g., a tuning fork). The frequency of oscillationis a function of the density of the gas surroundingit [362].

Both types of density recorders are checkedand calibrated by standard gas mixtures ofknown (e.g., gravimetrically determined) den-sity [363]. Density and relative density may alsobe calculated from gas analysis [357].

8.2.1.3. Wobbe Index

The Wobbe index W is defined as the ratio ofcalorific value H to the square root of relativedensity at the same specified metering referenceconditions. As with the calorific value it mustbe distinguished between superior Wobbe indexWS and inferior Wobbe index W I. Normally, theWobbe index W is not measured directly, al-though it is one of the most important combus-tion characteristics for describing gas quality.

Calculation of Wobbe Index. The indexcan be calculated from the calorific value and therelative density more accurately than by directdetermination. The Wobbe index may also beobtained by combining the signals of a calorime-ter and a densimeter electronically, according toits definition equation.

Gas Production 155

Direct Measurement of Wobbe Index. Di-rect measurement of the Wobbe index with rel-atively simple and rapid instruments is recom-mended only for mixing gases of different qual-ity to yield a mixture of constant Wobbe index.TheseWobbe recorders are automatic calorime-ters whose gas flow rate is adjusted so that it isproportional to 1/

√d.

8.2.2. Analytical Methods

The analytical methods listed in the follow-ing sections are (1) general methods (i.e., ca-pable of measuring several components, Sec-tion 8.2.2.1); (2) specific methods for individualcomponents (Section 8.2.2.2); or (3) methodssuitable forminor components (Section 8.2.2.3).

Sampling is the first, and a most impor-tant, stage in gas analysis. Therefore, great careshould be taken to obtain representative sam-ples for determining gas properties and com-position [364]. Continuous automatic samplingtechniques have also been developed [365].

8.2.2.1. General Methods for Determinationof Several Components

OrsatMethod. The main constituents of manu-factured gases (CO2, O2, CO, unsaturated hy-drocarbons, H2, saturated hydrocarbons andN2)were determined volumetrically by the classicOrsat method involving successive absorptionof the constituents in specific reagents and se-lective combustion of hydrogen and saturatedhydrocarbons [366]. A simplified Orsat appara-tus is used at present only for flue gas analysis orin combination with the chromatographic Janakmethod for carbon dioxide, oxygen, and unsat-urated hydrocarbons.

JanakMethod. The Janak procedure, whichis suitable for manufactured gases as well as fornatural gas components up to C4H10, is a simplechromatographic method using carbon dioxideas carrier gas [367]. The carrier gas emergingfrom the column passes to a burette filled withconcentrated potassium hydroxide solution, inwhich carbon dioxide is absorbed and the re-maining components can be measured consecu-tively by volume. The apparatus available com-mercially registers the volume ratio of the dif-ferent components with an automatic burette.

Gas Chromatography (GC). Besides theJanak method, GC [368 – 370] with a ther-mal conductivity or flame ionization detectoris most often recommended for fuel gas anal-ysis. A large number of standardized methodsexist for the analysis of manufactured gas [371]and natural gas [372 – 374] by using an ad-sorption column (e.g., molecular sieve) for theseparation of hydrogen, helium, oxygen, argon,nitrogen, methane, and carbon monoxide, and apartition column (e.g., organic stationary phaseon a solid support) for the separation of higherhydrocarbons. For the analysis of natural gas,determination of hydrocarbons up to pentane[372] or octane [373, 374] is normally suffi-cient. For specific needs, methods have beendeveloped for analysis up to hexadecane [375].

For onstream applications, these analyseshave been completely automated (process GC):this includes periodic calibration of the total sys-tem with a calibration gas mixture [376] andevaluation of the chromatogram by an electronicintegrator. Gas analyses can also be performedbymass spectrometry [377]. The results of theseanalyses can be used to calculate the physical,thermodynamic, and combustion characteristicsof the gas (HS, HI, WS, W I, , etc.) [357, 358].

8.2.2.2. Specific Methods for Determinationof Individual Components

Hydrogen. In binary or quasibinary gasmix-tures, the hydrogen content can be monitoredeasily by measuring thermal conductivity, be-cause the conductivity of hydrogen is muchhigher than that of all other likely components.

Oxygen. Besides volumetric [366], chemical[378], and GC methods [373], oxygen (concen-tration >0.1 vol%) may be determined withoutsignificant interference from other componentsby using instruments based on the measurementof magnetic susceptibility [379].

Carbon Dioxide. The carbon dioxide con-centration [380] of fuel and flue gases is nor-mally recorded by nondispersive infrared instru-ments [381, 382]. The carbon dioxide absorptionband is overlapped only slightly by the carbon

156 Gas Production

monoxide band, and for a wide range of concen-tration ratios this interference can be eliminatedby using a filter cell containing carbon monox-ide. Several instruments are available that usesingle- or double-beam techniques and differenttypes of detectors. For discontinuous determi-nation, aforementioned volumetric [366] or GCmethods [371, 373] may be used.

Carbon Monoxide. Infrared absorption isalso the recommended method for monitoringthe carbon monoxide concentration of gases[383]. Interference from carbon dioxide must befiltered out in the manner described above. Gaschromatography [384] and detector tubes [385]may be used for consecutive measurements orfor a single check. Carbon monoxide detectortubes are subject to interference from propaneand higher hydrocarbons,whichmaybe found innatural gas. This interference can be eliminatedby first leading the gas through an adsorptiontube filled with activated carbon.

8.2.2.3. Methods for Determination ofMinor Components

Water Content. For discontinuous gravi-metric determination of water, a drying tubecontaining a suitable desiccant (e.g., magnesiumperchlorate) can be used [386]. Alcohols suchas methanol or glycols are also absorbed. Waterand methanol (concentration > 20 mg/m3) canbe determined by GC [387].

TheKarl Fischer method [388], in whichwa-ter reacts with iodine in an alcoholic solution ofsulfur dioxide andpyridine, is suitable for the de-termination of low water content of the order ofmilligrams per cubicmeter. Newer equipment inwhich iodine is generated coulometrically [389]is particularly suitable for this application.

A variety of hygrometers use different detec-tors [390, 391], such as electrolytic cells withphosphorus pentoxide as absorbent, conductiv-ity cells with glycerol [392], a capacitor withaluminum oxide as the moisture-sensitive di-electric medium [393], silicon chips, and an os-cillating quartz crystal. By absorption or adsorp-tion of water from the surrounding gas, the elec-trolytic current, conductivity, dielectric proper-ties, or frequency of oscillation is changed. Bet-ter instruments use a sophisticated purge and cal-

ibration technique to overcome contaminationproblems caused by hydrocarbon condensatesand other impurities such as glycols or etha-nolamines derived from gas treatment processes[394]. Some of these instruments measure watercontent in parts per million (mL/m3 or mg/m3);others indicate dew point temperature.

Although water content and dew point can berelated theoretically, some uncertainty exists asto whether this can be appliedwith any degree ofaccuracy to uel gases.Measurement fromany in-strument is more safely related to the techniqueused for calibration.

Water Dew Point. Because of the risk of hy-drate formation [395], gas temperature shouldnot drop below the dew point. The water dewpoint can be determined directlywhen a constantstreamof gas passes over a polishedmetalmirrorthat is cooled slowly [396]. The dew point corre-sponds to the temperature at which the first signsof condensation appear on the mirror. The mea-surement can be performed under pressures upto 10 MPa. Completely automated instrumentsemploying photocells are also used [390]. If thegas contains methanol or glycols, the dew pointcorresponds to the simultaneous condensationof water and alcohol. If the methanol content isknown, a correction can be applied to the ob-served dew point [396]. Condensation of hy-drocarbons may also interfere, when the hydro-carbon dew point is higher than that of water.This interference can be reduced significantly byfirst passing the gas through high-boiling liquidparaffin.

Sulfur Compounds.Total Sulfur. The combustion method is ap-

plicable to all fuel gases. A measured volume ofgas is burned with excess air or oxygen in a spe-cially designed apparatus (e.g., Wickbold [397],lamp [398], Lingener method [399]). Combus-tion products are absorbed in an oxidizing solu-tion. The sulfuric acid formed is determined byvolumetric, gravimetric, or turbidimetric meth-ods [400]. Sulfur dioxide can also be titratedcoulometrically [401].

In the hydrogenation method, sulfur com-pounds are catalytically hydrogenated to hydro-gen sulfide at high temperature over platinumgauze in the presence of excess hydrogen and

Gas Production 157

water vapor. Subsequent determination of hy-drogen sulfide by the methylene blue procedureis very sensitive [402]. The lead acetate papermethod in combination with hydrogenation isless precise but can be automated for continu-ous photometric recording [403].

Simultaneous Determination of HydrogenSulfide, Carbonyl Sulfide, and Mercaptan Sul-fur. Ameasured volume of gas is passed throughtwo sets of bubblers in series. The first set con-tains concentrated potassiumhydroxide solutionin which hydrogen sulfide and mercaptans areabsorbed; the second contains alcoholic mono-ethanolamine solution, which absorbs carbonylsulfide. Potentiometric titration with silver ni-trate measures each species [404].

Individual Sulfur Compounds. The preferredmethod for determining individual organic andinorganic sulfur compounds is GC. Althoughsulfur compounds can be detected by ther-mal conductivity or flame ionization detectors(the latter for organic sulfur compounds only),sulfur-specific detectors are preferred for theanalysis of naturally occurring or added odor-iferous sulfur compounds to avoid interferencefrom other components. The flame photometricdetector [368, 405] measures the light emittedat 396 nmwhen sulfur compounds are burned ina hydrogen-rich flame.

The electrochemical detector widely used forthe control of gas odorization registers the po-tential change of an electrolytic cell when chro-matographically separated sulfur compoundsare oxidized by concentrated chromium triox-ide solution [406]. Carbonyl sulfide,which is notindicated by this detector in its normal mode ofoperation, may also be detected when potassiumhydroxide solution is used instead of chromiumtrioxide.

The Hall detector measures the change inconductivity caused by absorption of sulfurdioxide in an electrolyte after the effluent of thechromatographic column has been burned withexcess oxygen by passing the gas stream througha narrow nickel tube at high temperature [407].The change in conductivity is directly propor-tional to the sulfur content of the effluent. Thesameprinciple has also been used formonitoringthe total sulfur content of fuel gases.

Hydrogen sulfide can be determined by theformation of brown or black stains of lead sul-fide on lead acetate paper. This can be a simple

semiquantitative test, a manual quantitative de-termination [408], or a completely automatedcontrol [403]. For spot measurements, detectortubes [385] may also be sufficient. Wet chemi-cal determinations are based on iodometric titra-tion [409] or the very sensitive Methylene Bluemethod [410]. The coulometric method uses theoxidation of hydrogen sulfide by electrochemi-cally generatedbromine. Interference fromotheroxidizable sulfur compounds can be eliminatedby using a difference method with selective ab-sorption solutions (e.g., CdSO4 for H2S andKOH for H2S and mercaptans) [411]. Specificelectrochemical cells are used for monitoringthe hydrogen sulfide concentration in air or fuelgases.

Carbon disulfide is normally determined byGC [368], but colorimetricmethods and detectortubes [385] may also be used.

Besides the aforementioned GC and poten-tiometric methods, carbonyl sulfide can be hy-drolyzed in dilute potassiumhydroxide solution.The resulting sulfide is determined colorimetri-cally by the Methylene Blue method.

Nitrogen Compounds.Ammonia. Ammonia in manufactured gas is

determined by absorption in standard sulfuricacid and titration of the excess acid [412]. Lowconcentrations are indicated by reaction withNessler’s reagent, which is also used for colori-metric determination. For spot measurements,gas detector tubes are available [385].

Hydrogen Cyanide. Hydrogen cyanide incoal gas with low hydrogen sulfide content canbe determined after absorption in a nickel –sodium carbonate suspension [413] or in an am-moniacal solution of zinc sulfate by titrationwith silver nitrate. Gases with a higher con-tent of hydrogen sulfide can be analyzed by thecyanogen bromide method [414].

Nitrogen Oxides. Nitrogen dioxide forms ared dye with Saltzman solution [415]. Nitric ox-ide present in coal gas, therefore, must first beoxidized by passing the gas through a sodiumpermanganate solution. Both oxides can be de-termined together, or nitric oxide alone, after se-lective absorption of nitrogen dioxide.

For the analysis of flue gases, the chemilumi-nescent reaction between nitric oxide and ozoneis used for NOx determination. If nitrogen diox-ide is included, the gas must pass through a con-

158 Gas Production

verter where nitrogen dioxide thermally decom-poses to nitric oxide and oxygen [416].

Condensable Hydrocarbons. Natural ormanufactured gases may contain amounts ofcondensable hydrocarbons that can cause prob-lems by forming gas condensates. In dry gas,the hydrocarbon dew point may be determinedby the mirror method [396]. The potential hy-drocarbon liquid content in natural gas can beevaluated either by weighing or volumetrically[417].

Gas benzene (i.e., BTX aromatics and higherboiling hydrocarbons) was formerly determinedby adsorption on activated carbon and subse-quent displacement by steam. Today, chromato-graphic methods are preferred.

Naphthalene and other condensed aromaticcomponents can be determined chromatograph-ically or in total by precipitation of the corre-sponding picrates [418].

8.2.2.4. Determination of Trace Constituents

For the determination of trace constituent (e.g.,nickel and iron carbonyl, chlorine [419], andmercury [420]), several methods have been de-veloped, using atomic absorption spectrometry(AAS) to determine heavy metals.

9. Acknowledgement

The entire topic was coordinated by GerhardHochgesand.

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282. S. Sevin, J. Nougayrede, W. Willing, G.Bandel, Hydrocarbon Eng. 3 (1998) no. 4,54/56 and 80.

283. J. E. Nobles, J. W. Stancik, Oil Gas J. (1983)Dec. 26, 156 – 158.

284. Leuna-Werke, DE 43 05 386 A1, 1994 (P.Szargan et al.).

285. E. Supp, W. Hilsebein, CEER Chem. Econ.Eng. Rev. 17 (1985) no. 5, (no. 188) 18 – 21.

286. E. Supp, Pres. at AIChE Spring Meeting,Anaheim, Ca. May 20 – 23, 1984.

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287. “Upgrading Heavy Crude Oils and Residues toTransportation Fuels: Technology, Economicsand Outlook,” multisponsored reprint Phase 6SFA Pacific, Inc., Mountain View, CA,February 1999.

288. R. K. Olson, V. A. Gembicki: MaximizingDistillate from Reduced Crude, NationalPetroleum Refiner’s Association, San Antonio,Texas 1984.

289. L. O. M. Koenders, S. A. Posthuma, P. L.Zuideveld: “The Shell Gasification Process forConversion of Heavy Residues to Hydrogenand Power,” Gasification Technol. Conf., SanFrancisco, Oct. 1996.

290. E. O. Gerstbrein, J. S. Falsetti: “RefineryUpgrading using Texaco RefiningTechnologies and the Texaco GasificationProcess,” 2nd Int. Conf. on the RefineryIndustry, Vienna, Dec. 1994.

291. H. Fellner, W. Liebner: “Multi PurposeGasification (MPG), Lurgi’s ProvenTechnology for Refinery BottomsManagement,” 4th Forum on Advances in theRefining Industry , Mexico City, Aug. 1998.

292. W. Liebner: “MPG – Lurgi/SVZ MultiPurpose Gasification, a New and ProvenTechnology,” Gasification—A Gateway to aCleaner Future, Dresden, Sept. 1998.

293. H. Edelmann, F. Stuhlmuller: “EFCC—Einzukunftiges Konzept furKohle-Kombikraftwerke?,,” VGBKraftwerkstechnik 77 (1997) no. 7, 537 – 543.

294. R. D. Lessard et al., 27th ASME Int. GasTurbine Conf., London, UK, April 1982.

295. S. Kjaer, H. Koetzer, J. van Liere, I.Rasmussen: “Neue Konzepte furkohlebefeuerte Kraftwerke Vergleich vonWirkungsgrad, Wirtschaftlichkeit, Umwelt-und Betriebsaspekten,” VGBKraftwerkstechnik 74 (1994) no. 7, 561 – 568.

296. K. Weinzierl: “Kohlekraftwerke der Zukunft,”VGB Kraftwerkstechnik 74 (1994) no. 2, 109 –114.

297. U. Dorstewitz, H. Wien, Conf. CoalGasification in Energy Technology, Dortmund,FR Germany, March 1979.

298. H. Maghon, B. Becker, VGBKraftwerkstechnik 64 (1984) 999 – 1004.

299. B. Becker, A. Kreutzer, K. Goebel, 1983 TokyoInt. Gas Turbine Congress, Tokyo, Oct. 1983,pp. 791 – 798.

300. I. J. Hartill, J. A. Lacey, The Inst. of GasEngineers, 123rd Ann. General Meeting,Torquay, UK, May 1986.

301. B. Robson, Energy Res. Abstr. 1 (1977) 157 –177.

302. J. S. Joyce, W. E. Zoerner, 5th CEPSI Meeting,Manila, Phillipines, Nov. 1984.

303. K. Weinzierl, VGB Kraftwerkstechnik 62(1982) 365 – 379.

304. D. R. Simbeck, R. L. Dickenson, E. D. Oliver,EPRI-Rep. AP-3109, Research Project 2207,June 1983.

305. C. W. Schwartz, Energy Progr. 2 (1982) 207 –212.

306. J. van Liere, M. van der Burgh: “Gasification—an alternative to natural gas,” Chem. Eng.603 (1995) 15 – 18.

307. R. P. Wieber, J. S. Halow: “Advanced IGCCPower Systems for the United States,” EnergyProgress 7 (1987) no. 2, 119 – 125.

308. H. Meyer-Kahrweg, Energie 29 (1977) 227 –231.

309. R. Reimert, CIMAC, 14th Int. Congress onComb. Engines, Helsinki, FI, 1981, pp. 19 –27/ESC-DF.

310. N. A. Holt: 5th Int. Conf. and Exhibition onCoal Utilization and Trade, vol. 1, London,UK, Dec. 1985, pp. 153 – 177.

311. D. R. Plumley: “Cool Water coal gasification—A progress report,” J. Eng. Gas TurbinesPower 107 (1985) 856 – 860.

312. J. Veen, D. Ceelen, G. Kreutzkamp, BMFTForschungsbericht T-86–047, 1986.

313. H. Vierrath et al., Int. Gas Turbine Conf. andExhibition, Dusseldorf, FR Germany, 1986,86-GT-294.

314. J. Keller, W. Adlhoch, H. Vierrath: “HighTemperature Winkler Coal GasificationProcess, A proven Technology for Productionof Syngas or Fuel Gas,” Int. Gas Union 9thWorld Gas Conf., Milan 1994.

315. J. Keller, W. Adlhoch, H. Vierrath, R. Rizzi:“Low rank Fuel HTW Gasification as anIntegral Part of Combined Cycle PowerStations,” EERC 17th Biannual Low RankFuel Symposium, St. Louis, MO, 1993.

316. K. Shippers et al.: “Cobra will demonstrateHigh Temperature Winkler IGCC,” MPSModern Power Systems 1993 Feb., 41 – 50.

317. G. D. Zon, H. H. J. de Winter, W. Willebroer:“Kohlevergasung im Dienste derStromerzeugung, die Demo- KV STEG,” VGBKraftwerkstechnik 74 (1994) no. 6, 436 – 441.

318. D. Borgmann, B. Schotter:“Kohlegas-Kraftwerk Buggenum,” VGBKraftwerkstechnik 74 (1994) no. 6, 532 – 536.

319. J. Th. G. M. Eurlings: “The future for theIGCC: scale, fuels and products,” ModernPower Systems 1997, Nov., 39 – 44.

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320. R. Cortes, W. Schellberg: “The PuertollanoIGCC Plant, Commissioning status andpreoperational records,” EPRI GasificationTechnology Conference, San Francisco 1998.

321. W. Schellberg, R. Wetzel:“Entwicklugstendenzen in Kombikraftwerkenmit integrierter Flugstromvergasung,Wirkungsgrade, Umweltdaten, Kosten,”Brennstoff Warme Kraft 47 (1995) no. 9, 354 –358.

322. J. Mc Daniel, Ch. Shelnut, T. Berry: “TampaElectric Company Polk Power Station IGCCProject, Project status,” EPRI GasificationTechnology Conference, San Francisco 1998.

323. B. Gericke, J. C. Loffler, M. A. Percavec:“Biomassenverstromung durch Vergasung undintegrierte Gasturbinenprozesse,” VGBKraftwerkstechnik 74 (1994) no. 7, 595 – 604.

324. J. Albrecht, C. Greil: “Fuel Gas Productionfrom Biomass for Industrial Plants and EnergyGeneration,” 10th European ConferenceBiomass for Energy and Industry, Wurzburg1998.

325. J. Albrecht: “Experiences from CFBGasification with Biomass and Waste Fuel,”Joint EU- VTT Seminar Power Productionfrom Biomass, Espoo, Finland 1998.

General References326. International Gas Union, Sub-Committee B:

Methods of testing gases and raw materials,14th World Gas Conference, Toronto 1979.

327. British Gas Corporation: Analytical Methods,London Research Station, 1970 ff.

328. British Standard (BS) 3156 Part 1 – 5 and 10 –12.

329. Gaz de France: Essai de Chemie Gaziere, Paris1954, 1964.

330. DVGW-Regelwerk, Arbeitsblatt G 261:Prufung der Gasbeschaffenheit, ZfGW-Verlag,Frankfurt 1985, under revision 1997.

331. G. J. van Rossum (ed.): Gas quality;proceedings of the congress of gas quality –specification and measurement of physical andchemical properties of natural gas, Elsevier,Amsterdam, Oxford, New York, Tokyo 1986.

332. H. Guerin: Traite de manipulation et d’analysedes gaz, 2nd ed., Masson, Paris 1981.

333. H. Seebaum, K. Wienke: AnalytischeMethoden fur die Untersuchung von Ferngas,4th ed., Vulkan-Verlag, Essen 1965.

334. J. Hengstenberg, B. Sturm, O. Winkler:Messen, Steuern und Regeln in der chemischenTechnik, 3rd ed., vol. 2, Springer-Verlag,Berlin – Heidelberg – New York 1980.

335. P. Profos: Handbuch der IndustriellenMesstechnik, 2nd ed., Vulkan-Verlag, Essen1978.

Specific References336. DIN 1340, ed. Dec. 1990.337. International Gas Union: Committee A,

Lausanne 1982, pp. 14 – 24.338. K. Hedden, Th. Heike in [331], pp. 761 – 782.339. G. F. Steinmetz in [331], pp. 11 – 25.340. DVGW-Regelwerk: Arbeitsblatt G 260,

Gasbeschaffenheit, draft, Wirtschafts- undVerlagsgesellschaft Gas und Wasser mbH,Bonn 1997;DIN 51622, ed. Dec. 1985.

341. J. A. Rijnaarts in [331], pp. 27 – 35.342. D. Vernet, Y. Deschamps, in [331], pp. 791 –

808.343. DIN EN 26, ed. Jan. 1981.

DIN EN 437, ed. March 1993.344. A. G. A. Testing Laboratories, Research

Bulletin 36.345. E. R. Weaver, J. Res. Nat. Bur. Stand. U.S. 46

(1951) no. 3, 213 – 245.346. B. C. Dutton, E. Grymziesky, J. Inst. of Energy

56 (1983) no. 427, 107 – 108.347. B. C. Dutton, Gas Commun. Inst. Gas Eng.

1246 (1984).348. DVGW-Regelwerk, Arbeitsblatt G 280:

Gasodorierung, Wirtschafts- undVerlagsgesellschaft Gas und Wasser mbH,Bonn 1990.

349. P. Cowan, A. E. Humphreys in [331], pp. 169– 179.

350. ASTM D 1826–94.351. G. J. van Rossum, G. J. Benes, GWF Gas

Wasserfach Gas/Erdgas 122 (1981) 12 – 19.352. L. J. Mondeil, F. M. Robert in [331], pp. 111 –

119.353. D. Buwalda, M. W. C. de Jonge, J. Reints in

[331], pp. 153 – 161.354. A. E. Humphreys, MESUCORA Conggress,

Session 11, Paris 1982, pp. 95 – 112.355. H. F. A. Trimbos, G. J. van Rossum, W. R.

Schierbeek in [331], pp. 139 – 151.356. C. H. Griffis et al. in [331], pp. 121 – 127.357. ISO 6976- 1995.

AFNOR X 20–522, ed. June 1982.ASTM D 3588–91. BS 3156 Part 10.8;DIN 51857, ed. March 1997.

358. G. J. van Rossum, H. Koning in [331], pp. 809– 830.

359. DIN 1871, Draft April 1997.360. ASTM D 1070–85.361. Verein Deutscher Ingenieure: VDI-Handbuch

Reinhaltung der Luft, 3490, Part 15, DraftNov. 1985.

168 Gas Production

362. M. Jaescke, GWF Gas Wasserfach Gas/Erdgas126 (1985) 93 – 102.

363. DIN 51870, ed. April 1981.364. ISO 10715- 1997.

AFNOR NF X 20–251, ed. Dec. 1982.ASTM D 5287–93.DIN 51853, ed. Nov. 1995.

365. Th. F. Welker, G. J. De Nobel, D. Stirnberg, K.Homann in [331], pp. 415 – 442.

366. BS 1756 Part 2.DIN 51872 Part 2, ed. July 1980, withdrawn1997.

367. DIN 51872 Part 3, ed. Feb. 1978, withdrawn1997.

368. C. J. Cowper, A. J. de Rose: The analysis ofgases by Chromatography, Pergamon Press,Oxford 1983.C. J. Cowper, P. A. Wallis, B. R. A. Wybrow in[331], pp. 481 – 493.

369. J. Tranchant (ed.): Manuel pratique dechromatographie en phase gazeuse, 3rd ed.,Masson, Paris 1982.

370. DIN 51405, ed. May 1987.371. ASTM D 1946–90.372. ISO 6568- 1981.

AFNOR X 20–303, ed. Nov. 1974.373. ISO 6974- 1984;

ISO/DIS 6974 Part 1–6, 1997.AFNOR NF X 20–501, ed. July 1985.ASTM D 1945–91.DIN 51872 Part 4, ed. June 1990,DIN 51872 Part 5, ed. Aug. 1996.BS 3156 Part 11.1.2.

374. L. Huber, Erdol Kohle, Erdgas, Petrochem. 37(1984) 249 – 257.R. Keuter, M. Struis, A. L. C. Smit, ProcessControl Quality 1 (1991) 127 – 136;GWF Gas Wasserfach Gas/Erdgas 134 (1993)21 – 27.C. Busch, GWF Gas Wasserfach Gas/Erdgas136 (1995) 23 – 30.

375. ISO 6975- 1997.BS 3156 Part 11.1.2.

376. ISO 6142- 1981,ISO/DIS 13275- 1995,DIN-ISO 6142, ed. May 1985.AFNOR NF X 20–202, ed. Dec. 1979.R. Keuter, R. J. Beks, H. J. Bardewel, A. L. C.Smit, GWF Gas Wasserfach Gas/Erdgas 135(1994) 164 – 170.

377. ASTM D 2650–93.AFNOR X 20–302, ed. Oct 1974.

378. DIN 51856, ed. June 1979, withdrawn 1997.379. AFNOR X 20–377, ed. Dec. 1980.

380. ISO/TR 6567/1- 1980.AFNOR X 20–380, ed. Nov. 1979.

381. AFNOR X 20–301, ed. Jan. 1978. BS 4314Part 1.

382. H. Karthaus, H. Engelhardt: Grundlagen derphysikalischen Gasanalyse, Hartmann undBraun, Frankfurt 1971.

383. AFNOR X 20–361, ed. Sept. 1978. VereinDeutscher Ingenieure: VDI-HandbuchReinhaltung der Luft 2459, Part 6, ed. Nov.1980.

384. AFNOR X 20–363, ed. Sept. 1976.385. K. Leichnitz: Prufrohrchen-Meßtechnik, (La

technique de mesure par tubes reactifs),Ecomed-Verlag, Landsberg 1981.Drager-Rohrchen Handbuch 11th ed.,Dragerwerk, Lubeck 1997.Gas Processors Association 2377–77.

386. ISO 11541- 1997.BS 1756 Part 4.

387. AFNOR NF X 20–530, ed. June 1984.388. ASTM E 700–79.

BS 3156 Part 10.4; DIN 51777 Part 1, ed.March 1983, Part 2, ed. Sept. 1974.

389. ISO 10101, Part 1–3-1987; DIN 51869, ed.May 1986.

390. M. A. Berliner: Feuchtemessung, VEB VerlagTechnik, Berlin-Ost 1980.

391. A. W. Jamieson, H. J. Sikkenga in [331], pp.289 – 299.

392. J. Rotter in [331], pp. 311 – 325.393. P. L. Cremonesi in [331], pp. 301 – 310.394. M. Mathews, Erdol Kohle, Erdgas, Petrochem.

37 (1984) 95 – 96.C. B. Blakemore, J. C. Steichen, G. Dallas in[331], pp. 327 – 335.

395. DVGW-Regelwerk: Arbeitsblatt G 285,Hydratinhibierung in Erdgasen mit Methanol,ZfGW-Verlag, Frankfurt 1974.

396. ISO 6327- 1981.AFNOR NF X 20–521, ed. Dec. 1981.ASTM D 1142- 95.BS 3156 Part 10.3. DIN 51871, ed. Dec. 1985.

397. ISO 4260- 1987.DIN-EN 24260, ed. May 1994.

398. ASTM D 1072–90.399. ISO 6326 Part 5–1989; DIN 51400 Part 4, ed.

Oct. 1990.400. DIN 51400 Part 1, ed. Dec. 1992,

DIN 51400 Part 2, ed. Feb. 1978.401. ASTM D 3246–92.

DIN 51400 Part 7, ed. June 1992.402. ASTM D 4468- 85.403. L. Below, GWF Gas Wasserfach Gas/Erdgas

120 (1979) 59 – 65.

Gas Production 169

404. ISO 6326, Part 3- 1989.DIN 51855 Part 6, ed. March 1989.

405. P. A. Gibbons, Chromatographia 19 (1984)254 – 256.ISO 6326 Part 4- 1994.DIN 51855 Part 8, ed. 1997.

406. ISO 6326 Part 2- 1981.AFNOR NF X 20–511, ed. Jan. 1982.DIN 51855 Part 7, ed. Dec. 1986.

407. C. A. Risk, R. A. Hall, J. Chromatogr. Sci. 15(1977) 156 – 159.DIN 51855 Part 8, ed. 1997.

408. ASTM D 2420–91.ASTM D 4084–94.DIN 51855 Part 3, ed. June 1979, withdrawn1991.

409. ASTM D 2385–81.DIN 51855 Part 4, ed. June 1995.I. P. 103/70.

410. ASTM D 2725–87.DIN 51855 Part 4, ed. June 1995.

411. G. Herbst, A. Romanski, GWF GasWasserfach Gas/Erdgas 121 (1980) 105 – 108.

412. DIN 51854, ed. Sept. 1993.413. DIN 51863 Part 1, ed. Sept. 1983.414. DIN 51863 Part 2, ed. Sept. 1983.415. DIN 51864, ed. Sept. 1986.416. AFNOR NF X 43–018, ed. March 1983.

Verein Deutscher Ingenieure: VDI-HandbuchReinhaltung der Luft 2456, Part 5 + 6., ed.May 1978.

417. ISO 6570 Part 1 – 3 1983/84.DIN 51859, ed. June 1990.BS 3156 Part 11.2.

418. DIN 51862, ed. Sept. 1982.419. DIN 51408 Part 1, ed. June 1983.420. W. Lommerzheim, GWF Gas Wasserfach

Gas/Erdgas 117 (1976) 430 – 432.