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Eni S.p.A. Exploration & Production Division

00342.HTP.PRC.PRGRev.1 January 2005

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DESIGN CRITERIA

GAS SWEETENING PROCESSES

SELECTION CHARACTERISTICS AND CRITERIA

00342.HTP.PRC.PRG

Rev.1 January 2005

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TABLE OF CONTENTS

1-SCOPE OF WORK: AIM OF SWEETENING

2- PROCESS CHOICE CRITERIA2.1- Required depuration level and other possibile impurity

2.2- Delivery specifies and ecological discharge

2.3- Toxic problems

2.4- Corrosion allowance

2.5- Decision schemes and plant costs

3- CHEMICAL AND/OR PHYSICAL ABSORBING PROCESS FOR SOUR GAS

SWEETENING

3.1- CHEMICAL ABSORBING WITH DEA AND MDEA SOLUTIONS

3.1.1- Physical and chemicals properties of amines

3.1.2- Absorbing mechanisms and regeneration

3.1.3- Column design3.1.4- Other main items design

3.1.5- Construction materials

3.1.6- Foaming problems and amine loss

3.1.7- Classic absorbing scheme

3.1.8- Split-Flow scheme

3.1.9- Double stage Scheme

3.1.10- Plant simulations

3.2- SELEXOL PROCESS

3.2.1- Process description

3.2.2- Used schemes

3.2.3- Design and measuring criteria3.2.4- Construction materials

3.3- SULFINOL PROCESS

3.2.1- Process description

3.2.2- Used schemes

3.2.3- Design and measuring criteria


4- OXIDATIVES PROCESS FOR SULPHIDRIC ACID OXIDATION

4.1- LO-CAT process

4.1.1- Plant schemes

4.2- Monterotondo process with biological oxidation

4.3- CLAUS process and variations

4.3.1- Process description, involved reactions and used catalysts

4.3.2- Items

4.3.3- Use and treatment of produced sulphur

4.3.4- Base scheme plant

4.3.5- Process variations


4.3.7- Further process development

5- CRYOGENIC PROCESS FOR CARBON DIOXIDE REDUCTION

5.1- RYAN-HOLMES process

5.2- CRYOFRAC process5.2.1- Process description

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1. SCOPE OF THE WORK: AIMS OF SWEETENING

This document aims to improve the know-how of gas sweetening plant operators, to

help them understand application limits and choose the most appropriate process for

each specific application.

The study is also part of the ongoing process to standardize the types of plants most

frequently used by ENI E&P and its Subsidiaries.

Sweetening natural gas and associated gas consists in removing sour gases such as

H2S and CO2 from the gas stream. Sweetening is performed to obtain a product thatdoes not cause safety or environment pollution problems when used.

H2S is a highly toxic gas: exposure for a few minutes to just several hundred ppm of

this gas can be fatal.

During combustion the H2S contained in the natural gas transforms into SO2, a

product which is just as toxic and which can be discharged into the atmosphere only

in limited quantities.

In cases where natural gas has a high sour gas content, sweetening is necessary also

to avoid corrosion problems in the transport and distribution network.

Sometimes the content has to be reduced to control the heat value and/or to make

the gas produced interchangeable.

A regeneration type sweetening plant releases all the absorbed H2S and CO2 during

the regeneration stage; as , unlike , cannot be discharged into the atmosphere, it has

to be conveyed to a recovery plant for transformation into elementary sulphur.

Only in the case of very limited H2S rates is it possible to send the gas to an

incinerator for transformation in SO2 before being discharged into the atmosphere.

Sulphur recovery is of primary importance in gas sweetening processes both for

environmental protection purposes and because it has an economic value: thesulphur recovered from the sour gases is one of the main supply sources for the

chemical industry. Processes based on the Claus reaction are mainly used to recover

sulphur from H2S.

Finally, Tail Gas Clean Up plants which treat the gas discharged from sulphur

recovery plants are worthy of mention: in short, these units greatly improve the

sulphur recovery efficiency of Claus units, reducing the quantity of SO2 discharged

into the atmosphere and further limiting the environmental impact of natural gas

treatment systems.

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In the case of liquid gas production, the dew point of the water must drop to

between -150°C and -160°C, while the dew point of hydrocarbons must be between

-40°C and -50°C in all the pressure ranges between the maximum operating

pressure and the atmospheric pressure.

2.1.1. Gas containing only CO2.

Absorption processes in liquid phase generally allow the purification values

specified to be achieved.

The primary amines are able to reduce the CO2 content in the gas produced to

partial pressures of around 10 mm H2O abs. The activated potassium carbonates can

reduce CO2 partial pressure values in the treated gas to 200 mm H2O abs.

If the absorption pressure is sufficiently high, physical solvents can also be used to

achieve partial pressure values of around 1000 mm H2O abs.

2.1.2. Gas containing only H2O

Amine solutions in water or organic solvents are able to reach the purity level

required even at fairly low absorption pressures; see figure 4.2.1 for MEA solutions

(20% weight) and figure 6.1.5 for other amines.

A physical solution can be used if the gas to be treated is at a sufficiently high

pressure.

2.1.3 Gas containing both H2S and CO2

If the gas contains both H2S and CO2, it is often useful to use a selective process for

the H2S as this process will leave the amount of CO2 permitted by the specification

in the treated gas. This will result in savings because a smaller rate of solution will

be needed for absorption purposes.

Selective processes that can be used are: oxidation in liquid phase, tertiary amines

and physical or chemical-physical processes.

At low pressure, oxidation processes in liquid phase such as STRETFORD, LO-

CAT and SULFINT can be used. However, STRETFORD will only tolerate a low

CO2 content in the gas supplied.

Selective amines such as methyldiethanolamine (MDEA) and diisopropanolamine

(DIPA) can be used at medium pressures; EXXON has developed a selective wet

process that uses stereo hindered amines which, as such, have the same selectivity

characteristics as tertiary amines.

Physical solvents such as SELEXOL, which are suitable at high pressures, can be

used in plants designed ad hoc also to selectively absorb H2S.

The selectivity of the physical solvents is based on the fact that the relative

volatility of CO2 is very high compared to that of H2S.

Primary and secondary amines can be used to obtain the specified purification valueif the process does not need to be selective.

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In figure 4.2.1 the operating pressure, the H2S/CO2 ratio in the crude gas to be

treated and the H2S purification value that can be reached in gas scrubbed using an

MDEA solution (20% weight) as absorbing medium are correlated.

The curves are based on a consumption of steam for regeneration purposes of 120

Kg/m3

of solution; with a consumption of steam of 150 Kg/m3

of solution, theresults of the purification can be up to 70% higher than the values shown.

The use of MEA – DEG solutions in standard concentration (MEA 20% - DEG

75% - H2O 5%) can lead to an improvement in the purification result of 50%

compared to the values indicated.

The purification results that can be obtained with other amines are shown in figure

6.1.5.

Note that, when high quantities of sour gas have to be absorbed, it is possible to use

a process which offers a good result in terms of absorption (high pick-up), followed

by a process able to achieve the required purification level, even with low pick-up.

This procedure is used in the HI-PURE process: a solution of active potassiumcarbonate is used to absorb most of the sour gas (bulk removal); final absorption of

the sour gas to just a few ppm, which is not economically viable with an active

carbonate solution, is performed using an amine solution in a completely separate

unit installed downstream of the active carbonate absorption unit. In this way the

final result is a gas which fully respects the specifications using a lower total rate of

absorbing solutions than what would have been needed if just one of the two

procedures had been used.

2.2. Delivery specifications and environmental constraints of discharges

It is important to remember that almost none of the solutions used in gas sweetening

processes can be discharged directly into drains but must be treated at the very least

in a biological system.

Therefore, precautions must be taken when replacing or discharging absorbing

solutions and specialized companies used for the disposal of non-biodegradable or

harmful waste.

The environmental problems of the various plants will now be discussed.

2.2.1. Absorption processes

The effluent from these processes is a gas current containing all the sour gases that

have been removed from the crude gas.

If the sour gas contains H2S then a Claus sour gas treatment unit must be installed

downstream of the sweetening plant to recover the elementary sulphur; in this case

H2S rich currents (25% vol. min.) will have to be treated with a minimum potential

of not less than 500 Kg/day. Oxidation processes should be applied in cases of sour

gas potential of less than 500 Kg/day or in the case of very limited concentrations of

H2S.

If the H2S is present in the sour gas in extremely limited quantities, under currentlaws on atmospheric protection, it would be possible to convey the sour gas to an

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incinerator which oxidizes the H2S to SO2 before being released into the atmosphere

at a suitable height above ground level.

However, a Claus unit with final incinerator will always be envisaged for the

oxidation to SO2 of the residual sulphur compounds in the tail gas; an emergencyflare is also needed to burn the sour gas should the Claus unit not be operational.

2.2.2. Oxidation processes

Oxidation processes in liquid phase are able to offer high sulphur recovery factors

with very low emissions.

However, these processes are based on rather delicate oxidation-reduction reactions

and secondary reactions which slowly degrade the solution.

The production of the degradation must be constantly purged, creating ecological

problems due to the toxicity of most of the solutions used in these processes.

The new generation liquid phase oxidation processes (LO-CAT, SULFINT and to a

certain extent also STRETFORD) use absorbing solutions which are not poisonous

and which are in any case biodegradable so many of the barriers to the intensive

development of these processes have been broken down.

The sulphur produced in oxidation processes is generally of a much lower quality

than that produced in the Claus units. Indeed, it is difficult to find a market outlet

for it unless it is further purified.

2.2.3. Cryogenic processes

Cryogenic processes, often used for the decarbonation of gas, do not pose

environment related problems.

2.3. Toxicity issues

Due to the heightened awareness in recent times of safety and ecology related

questions, some processes, while being very efficient, are rarely used or have been

almost completely abandoned.

Solution toxicity problems are encountered in the following processes:

- processes which use active potassium carbonate with arsenite/arseniate.

- processes which use potassium carbonate with organic activator and

concentrated MEA (and sometimes DEA) processes: indeed the latter use

vanadic anhydride or antimonium salts as a corrosion inhibitor.

- oxidation processes that use Vanadium or Arsenic salts as oxygen carrier.

Blowdowns of the solutions used in these processes are not biodegradable and can

only be used in very small quantities in biological treatment systems and only for

some processes.

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The amines and solutions used in almost all the physical and chemical-physical

processes are not easily biodegradable and cannot therefore be discharged into

drains or treated biologically, except in very small quantities.

The relatively high cost of the solutions and the related ecological problems meanthat closed systems have to be installed in almost all sweetening processes which

envisage absorption with water solutions to collect the continuous and

discontinuous blowdown and recycle it in the operative units.

2.4. Corrosion and protection

Corrosion problems are mainly encountered in wet absorption plants.

Corrosion may be due in some parts of the plant to the action of wet CO 2 and H2S,

while stress-corrosion cracking phenomena may occur in other parts.

The possibility of corrosion due to the acid attack of CO2 and H2S is common to all

processes of this type.The processes which have created most corrosion problems are those which use

active potassium carbonate or amine solutions in concentrations of more than 20%.

Different corrosion phenomena have been observed in plants which use amine or

potassium solutions depending on which amine was used, its concentration, the

presence of degradation products or other impurities, the sour gas rate, as well as

the temperature and velocity at which the solution flows through the equipment and

pipes.

The main cause of corrosion is the presence of sour gas such as CO2 and H2S in a

water environment.

Figure 2.4.1 shows the corrosion effect on MEA-DEG and MDEA solutions.

It has generally been noted that corrosion takes place more rapidly in the presence

of CO2 alone as this can form soluble iron salts at high temperature.

Instead, the acid attack of generates insoluble salts which form a film which is not

sufficiently thick to protect against further corrosion.

Therefore, in the presence of small concentrations of H2S, the joint action of H2S

and CO2 sometimes acts as a corrosion retardant.

Besides sour gases, another important cause of corrosion are the products of the

degradation of the amines which form due to irreversible reaction of the amine and certain impurities present in the gas to be treated.

It has been noted that the products of the degradation of monoethonolamine are

more corrosive than those of diethanolamine while corrosion phenomena are more

limited in plants that use diisopropanolamine or methyldietanolamine.

Corrosion is also favoured by erosion caused by suspended solids in the solution

such as iron sulphide formed through the action of the H2S.

Stress-corrosion typical of alkaline attack has sometimes been noted; these

phenomena generally occur after the plants have been in operation for some years.

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However, the situation which most often favours corrosion is the high concentration

of acid gas along with high temperature.

This situation generally occurs in the solution/solution exchanger in the high

temperature zone, rich solution side and in the regenerator head.

The regeneration column is also subject to corrosion above all in the head zone

where the concentration of acid gas and the temperatures are higher.

The boiler and head condenser, pipe side, where the process fluid normally flows,

are also subject to a high risk of corrosion.

Measures can be taken to reduce corrosion to acceptable limits for efficient

management of the plant.

The main criterion to follow is to use corrosion resistant materials instead of carbonsteel in the areas most at risk of corrosion.

304 or 316 type stainless steel is the most commonly used material.

The following equipment is usually made of corrosion resistant alloys:

solution/solution exchanger, boiler, regenerator head condenser, some parts of the

regeneration column and a portion of the piping.

One measure which it is always advisable is to avoid zones in the exchangers with a

very high wall temperature. This implies that very hot heating fluids should not be

used in the boiler and that very high pressures should not be used in the

regeneration process.

It is also important to prevent oxygen, which provokes the degradation of the

solution, from coming into contact with said solution.

Therefore, blowdown collection tanks and solution storage tanks must be kept in an

inert atmosphere.

Special care must be taken to ensure that there are no suspended solids in the

solution; filters must therefore be used.

Another criterion for limiting corrosion is to use heavy metal based corrosioninhibitors such as Vanadium and Antimony.

These corrosion inhibitors are used above all in plants where potassium carbonates

and high concentration MEA and DEA solutions are used.

Along with corrosion inhibitors there should be continuous control of the oxy-

reducing potential of the solution.

There has been a growing tendency recently to subject all the carbon steel

equipment and pipes that come into contact with amine solutions operating a high

and low temperature to an annealing treatment. Recent studies have shown that,

although this has not completely eliminated the stress-corrosion cracking phenomena, annealing has greatly reduced its statistical incidence.

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Programmed inspections of equipment and pipes is recommended during annual

plant shut-downs; however, common investigation techniques such as ultrasound

and penetrating fluids are not always effective in highlighting corrosion; it is better

to used magnetic particle type investigation systems with fluorescent fluids.

In all cases where repairs have to be made, the welding must be annealed.

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C O R

R O S I O N P E N E T R A T I O N - m m

CONTACT TIME - hours

FIGURE 2.4.1 - CORROSION OF MEA/DEG AND MDEA SOLUTIONS

(A) MEA 15% pesoDEG 80% peso

H2O 5% peso

(B) MDEA 50% peso

ACCIAIO AL CARBONIO

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2.5. Decision-taking scheme and engineering costs

The sweetening processes considered have been applied in many plants and so the

related problems are well-known.

Plants which use dry adsorption processes generally cause fewer operating

problems.

Indeed, adsorption and regeneration cycles have been automated and so many

installations now operate unmanned.

Plants that use wet absorption processes are without doubt more common. These

plants are easy to run although the condition of the circulating solution has to be

kept under control.

One operating problem is caused by the use of absorbing solution reclaiming

systems.

Another operative problem can be caused by the installation of a Claus unit

downstream of the sweetening system even although nowadays reliable continuous

analyzers are available which greatly simplify operations.

The machines can incur normal maintenance problems; in chemical absorption

plants the machines comprise pumps for the circulation of solution and condensate

while in physical absorption plants besides pumps there are often compressors with

the purpose of recycling part of the hydrocarbons released during regeneration of

the solution or of extracting and recompressing the gases generated during the

various flash stages.

Wet oxidation plants are those which most often cause operating problems. Indeed,

they are based on a very delicate chemism which therefore requires precise

analytical control. They generally also entail the complex problem of the disposal of

solution blowdowns and the extraction and purification of the sulphur produced.

Plants which use non-regenerative processes are technologically simple plants

which therefore cause no problems in normal operating conditions.

However, it is not easy to remove and handle the rich adsorbent mass which must

then be disposed.

Case histories of cryogenic separation plants are too few to have statistics available

about the operating problems.

A typical operative problem of amine or activated carbonate absorption plants is the

formation of foam, often due to the presence of liquid hydrocarbons in the

circulating solution; the widespread presence of foams causes the choking-up of the

absorption and regeneration columns which can result in uncontrollable entrainment

of liquids by the gas.

The risk of the solutions foaming can be reduced greatly by taking measures when

designing the plants such as using scrubbing solutions at temperatures above dew

point in hydrocarbons of the crude gas to be treated, suitably insulating the bottom

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of the absorber, supercooling the crude gas to be treated and complete separation of

liquids upstream of the acid gas absorption column.

The use of activated carbon filters to retain liquid hydrocarbons or other surfactant

substances completes the list of precautions to take.

Economic evaluation of the investment and operating cost of a plant is the decisive

criterion for the choice; this evaluation requires a specific study which takes into

consideration the industrial context the plant will be built in. To choose the best

solution a study should be made of the process as a whole in order to assess the

engineering and operating costs of the alternatives chosen.

The fundamental factor of the cost of an absorption plant is the rate of circulating

fluid, while secondary cost factors include the operating pressure and the acid gas

pick-up.

The operating cost is sometimes affected by local situations which can change the

economic viability of one process compared to others with investment costs varying

by 10% to 30%.

The results of a feasibility study carried out for different sweetening cases are given

below for the sake of information.

2.6 Sweetening of gas containing only CO2.

2.6.1 Case 1 – Average pressure

Operating pressure 40 eff. bar

incoming CO2 43 % vol.

outgoing CO2 2 % vol.

Cost factors of the utilities considered:

Electrical energy (Kwh ) 2.1

Fuel gas (1000 Kcal ) 1.0

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Process MEA

Classical

MEA

With

additives

DGA BENFIELD

Concentrationsol. (% weight) 20 35 60 30

Net pick-up

(mol GA/mol sol.) 0.3 0.35 0.30 -

Solution rate 1 0.51 0.57 0.67

Investment 1 0.72 0.90 0.71

Total cost of

purification

1 0.48 0.75 0.39

2.6.2 Case 2 – High pressure

Operating pressure 72 eff. bar incoming CO2 13 % vol.

outgoing CO2 1.4 % vol.


Electrical energy (Kwh) 1.0

Fuel gas (1000 Kcal) 0.08

Process MEA DEA MDEA

Two-stage

Concentration of sol.

(% weight)

20 35 40

Circulating sol. (m3/h) 1 0.6 0.33/0.95

Hourly cost of utilities 1 0.63 0.68

Investment 1 0.72 0.78


Gas rate 93,000 Nm3/h

Operating pressure 53 bar abs.

incoming CO2 1.0 % vol.outgoing CO2 0.005 % vol.


Steam (kg) 0.25


Fuel gas (1000 Kcal) 0.275

Process MEA MOLECULAR

SIEVE

SULFINOL-D

Investment 1.0 1.87 0.925

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Total cost of

purification

1.0 0.84 0.45




incoming CO2 1.0 % vol.



Steam (Kg) 0.25


Process SULFINOL PHYSICALInvestment 1.0 1.9

Total cost of purification 1.0 1.3





outgoing CO2 1 % vol.

Process MDEA SULFINOL-M SELEXOL

Circulation of sol. 1.10 1 1.60

Consumption of

electrical energy

0.75 1 2.10




incoming CO2 2.17 % vol.outgoing CO2 100 ppm vol.

Process DEA-SNEA SULFINOL-D activated MDEA

Investment 1 1.08 1.40

Circulation of sol. 1 1.24 1.14

Steam consumption 1 0.90 0.97

Total operating cost 1 1.07 0.91


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outgoing CO2 100 ppm vol.

Process DEA-SNEA SULFINOL-D

Investment 1 1.14

2.6.8 Sweetening of gas containing only H2S.

2.6.8.1 Case 1 – High pressure


incoming H2S 1.5 % vol.

outgoing H2S 10 ppm vol.

Process DEA DIPA DIPA

Amine concentration (%

weight)

21 26.5 53

Circulation of solution 1 0.71 0.44

Regeneration steam 1 0.76 0.58

2.6.8.2 Case 2 – Low pressure


incoming H2S 17 % vol.


Process DEA DIPA DIPA

Amine concentration (%

weight)

21 26.5 53

Circulation of solution 1 0.71 0.48

Regeneration steam 1 0.75 0.61

2.6.9 Sweetening of gas containing both CO2 and H2S

2.6.9.1 Case 1 – High pressure




incoming H2S 0.1 % vol.

outgoing H2S 1.5 ppm vol.

Cost factors considered

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Fuel gas (1000 Kcal) 1

Process SELEXOL SULFINOL

Solution rate 1 0.89Investment 1 0.87

Total cost of purification 1 0.35

2.6.9.2 Case 2 – Average pressure




incoming H2S 5 % vol.


Process DEA DGA SULFINOL-D

Concentration (%

weight)

25 60 45

Net pick-up

(mol/mol amine) 0.67 0.26 0.44

Solution rate 1 1.06 1.06

Consumption of

regeneration steam

1 2.03 0.62

Investment 1 1 1

2.6.10 Sweetening of gas containing CO2 with re-injection of CO2

Refer to figures 2.6.1, 2.6.2 and 2.6.3 taken from Hydrocarbon Processing, May

1982, “Pick treatment for high CO2 removal” by C.S. Goddin, which show energy

consumption and operating costs for the decarbonation of streams with a variable

CO2 content, taking into consideration the cost of different utilities.

The diagrams have been prepared assuming the following operating parameters:

Pressure of crude gas 19 eff. bar

Pressure of sweetened gas 43 eff. bar Pressure of produced CO2 30 eff. bar

CO2 incoming supply variable

CO2 in outgoing sweet gas 1.5 % vol.

H2S incoming supply variable from 0.1 to 0.2 % vol.

H2S in outgoing sweet gas 1.5 ppm vol.

H2S in the CO2 produced 100 ppm vol.

It has been envisaged that the gas arriving from cryogenic treatment and from the

membranes is brought into line with the specification by means of sweetening with

DEA.

Investment costs of the various processes:

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%CO2 in the crude gas DEA CRYOGENIC MEMBRANES

20 1 1.67 -

40 1 1.27 1.01

60 1 1.10 0.9180 1 0.78 0.70

90 1 0.71 0.58

In figures 2.6.2 and 2.6.3 the operating cost of the final CO2 sweetening using a

DEA solution must be added to the operating cost of the alternatives with cryogenic

plant and membranes; the cost of the alternative with DEA, as indicated, includes

the final sweetening of the CO2.

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E N E R G Y R E Q

U I R E D – k c a l / N m 3 o

f C O 2

CO2 IN THE GAS TO BE TREATED - % molar

CRIOGENICO

MEMBRANE

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FIGURE 2.6.1 –CO2 CONTENT / PURIFICATION ENERGY

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3. CHEMICAL AND/OR PHYSICAL ABSORPTION PROCESSES FOR ACID

GAS SWEETENING

3.1 Chemical absorption with DEA and MDEA solutions

3.1.1 Physical and chemical properties of amine

Processes which use amine solutions have been applied in the chemical,

petrochemical and natural gas treatment industries for many decades and have

been greatly improved over the years; indeed they are now the most commonly

used processes in all possible fields of application such as natural gas and

associated gas, purification of synthesis gas, purification of refinery gas,

production of technical gas etc..

The growing impact of ecological problems and the development of Claus tail gas

treatment units have encouraged the study of selective amines and thedevelopment of increasingly interesting processes also from the point of view of

operative consumption.

At present there are different amine processes on the market which allow for an

extremely targeted choice of solution also for very specific problems.

All the processes listed are currently used more or less intensively; some specific

characteristics of amine processes are indicated in to tables 3.1.1 – 3.1.2 – 3.1.3

and 3.1.4.

3.1.1.1 Diethanolamine (DEA)

Free process, commonly used in refineries.

DEA is an averagely reactive, poorly selective secondary amine suitable for

absorbing acid gases at a pressure of a few atm.

It is not decomposed by COS and CS2.

It partially absorbs mercaptans and organic sulphur.

SNEA holds a patent for the use of concentrated solutions on the basis of

experience on a very large scale at Lacq in France. With this process it is possible

to push the pick-up to the maximum compatible with the driving force; moreaccurate filtering is sufficient to avoid corrosion and remove the products of the

degradation of the solution.

There is another Union Carbide process (AMINE GUARD-ST)which uses

additives also for DEA that allow more economic performance in the sweetening

of gas containing both H2S and CO2 by using solutions with a concentration up to

55% in weight.

3.1.1.2 Methyldiethanolamine (MDEA)

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Free process, with the possibility of using the patented experience of SNEA,

Union Carbide, Parsons, Exxon etc..

Highly selective tertiary amine suitable for absorbing low pressure acid gases; itallows 40–60% of the CO2 contained in the crude gas to be treated not to be

absorbed while it absorbs H2S to a few ppm for gas at a pressure of 80 eff. bar

It does not absorb but is not degraded by CO2.

It partially absorbs mercaptans and organic sulphur.

The variant “ACTIVATED MDEA” produced by BASF uses MDEA at 48%

weight added in a two-stage plant with flash (also vacuum) of the rich solution.

This process is able to greatly limit the consumption of regeneration heat, suitable

for the absorption of large quantities of CO2. However, the activator is not used in

the case of selective absorption.

Another variant (Union Carbide’s UCARSOL, HS/CO2/H2S process) uses MDEA

with a concentration of 24% weight in a patented absorber with plates designed to

increase selectivity.

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PROCESS DEA MDEA

TYPICAL CONCENTRATION

(% weight)20÷40 25÷55

SELECTIVITY H2S ON CO2 LOW HIGHABSORPTION FACTOR CO2 40 10

MOLECULAR WEIGHT

(Kg/Kmol)105.14 119.17

CORROSION LIMIT

(mol AG/mol amine)> 0.8 0.8

TYPICAL PICK-UP

(mol AG/mol amine)0.5÷0.85 0.1÷0.8

TYPICAL PICK-UP

(Nm3 AG/m3 solution)22÷73 5÷83

HEAT OF H2S REACTION(Kcal/Nm3) 432 370

HEAT OF CO2 REACTION

(Kcal/Nm3)714 671

STEAM FOR REGENERATION

(Kg/m3 solution)100÷130 100

BOILING POINT AT ATM PRESSURE (C°) 268 247

SOLUBILITY OF HYDROCARBONS IN AMINE 1 < 1

SOLUBILITY OF AMINE IN HYDROCARBONS 1 /

TABLE 3.1.1 – CHARACTERISTICS OF AMINE PROCESSES (SHEET 1 OF

2)

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PROCESS DEA MDEA

SOLUBILITY OF AMINE IN WATER (% at 20°C) 96.4 100

OPTIMAL OPERATING PRESSURE (abs. bar) > 8 ALL

FREEZING POINT (°C) 0 < -40 NEED FOR RECLAIMING NO NO

RECLAIMING TEMPERATURE (°C) / /

THERMAL EXCHANGE COEFFICIENT FACTOR AT 65°C 0.87 0.85

BOILING CONC./TEMP. AT 1.8 ata (% weight / °C) 21/118 24/118

FREEZING CONC./TEMP. (% weight / °C) 21/-5 24/-6

VISCOSITY FACTOR 1.3 1.06

COST OF PURE AMINE FACTOR 1 2.5

COST OF THE SOLUTION AT STANDARD

CONCENTRATION FACTOR 1.4 4.0

EMPIRICAL FORMULA OF AMINE NC4H11O2 NC5H13O2

TABLE 3.1.1 – CHARACTERISTICS OF AMINE PROCESSES (SHEET 2 OF

2)

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TABLE 3.1.2 – STRUCTURAL FORMULAS OF AMINE

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PROCESS DEA

CONCENTRATION (% weight) 26

40°C (mol CO2 / mol amine) 0.39

40°C (Nm3 CO2 / m3 amine) 31

PROCESS DEA

CONCENTRATION (% weight) 40

50°C (mol CO2 / mol amine) 0.31

50°C (Nm3 CO2 / m3 amine) 51

TABLE 3.1.3 – THEORETICAL PICK-UP (85% EQUILIBRIUM) WITH

PARTIAL PRESSURE OF CO2 EQUAL TO 1 PSIA

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PROCESS MDEA DEA – SNEA

AMINE CONCENTRATION (% weight) 44 – 55 30 – 35

LEAN SOLUTION

(mol acid gas/mol amine)0.004 0.016

PICK-UP OF RICH SOLUTION

(mol acid gas/mol amine)0.1 – 0.9 0.7 – 1.1

STEAM AT BOILER (Kg/m3 solution) 100 120

PROCESS MDEA MDEA MDEA

AMINE CONCENTRATION

(% weight)50 50 50

STEAM AT BOILER

(Kg/m3 solution)80 100 120

LEAN SOLUTION

(mol acid gas/mol amine)0.006 0.004 0.003

TABLE 3.1.4 – REGENERATION OF MDEA AND DEA – SNEA

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3.1.2 Absorption and regeneration mechanisms

Amines have been used for gas sweetening since before 1930.

Refer to table 6.1.1 where some unique characteristics of the amines discussed

here are summarized.

When H2S and CO2 are dissolved in water they form a weak acid which combines

with the amines, which have a weak base, to form a salt which is then

decomposed by high temperatures: this absorption mechanism is common to all

amines.

During absorption the molecules of the acid gases must pass from the gaseous

phase, through the gas-liquid interface, to the liquid phase; the H2S reacts

immediately with the amine molecules in the interface: its absorption is therefore

limited by its transfer from the gas phase. Absorption of CO2 is instead limited

also by diffusion in liquid phase as its reaction with the amines is notinstantaneous.

The amines used in sweetening processes can be primary, secondary and tertiary

depending on the number of radicals bonded to the nitrogen.

The absorption capacity of acid gases decreases from the primary amines to the

tertiary ones because the alkalinity decreases proportionately.

Table 3.1.1 shows the empirical formulas of the amines dealt with here; the

relative structural formulas are given in table 3.1.2.

The reactions which take place during the acid gas absorption and regeneration

stages of amine solutions are listed below along with the reaction speed:

a) MEA and DGA

2RNH2 + H2S <===> (RNH3)2S INSTANTANEOUS

(RNH3)2S + H2S <===> 2RNH3HS INSTANTANEOUS

2RNH2 + H2O+ CO2 <===> (RNH3)2CO3 FAIR

(RNH3)2CO3 + H2O +CO2 <===> 2RNH3HCO3 FAIR

2RNH2 + CO2 <===> RNHCOONH3R FAIR

b) DEA and DIPA

2R 2 NH + H2S <===> (R 2 NH2)2S INSTANTANEOUS

(R 2 NH2)2S + H2S <===> 2R2NH2HS INSTANTANEOUS

2R 2 NH + H2O + CO <===> (R 2 NH2)2CO3 FAIR

(R 2 NH2)2CO3 + H2O + CO2 <===> 2R 2 NH2HCO3 FAIR

2R 2 NH + CO2 <===> R 2 NCOONH2R 2 FAIR

C) TEA and MDEA

2R 3 N + H2S <===> (R 3 NH)2S INSTANTANEOUS

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(R 3 NH)2S + H2S <===> 2R 3 NHHS INSTANTANEOUS

2R 3 N + H2O + CO2 <===> (R 3 NH)2CO3 SLOW

(R 3 NH)2CO3 + H2O + CO2 <===> 2R 3 NHHCO3 SLOW

Note that the maximum quantity of H2S that can be absorbed by the amines isequal to 1 mol/mol amine while the maximum quantity of CO2 that can be

absorbed is 0.5 mol/mol amine for carbamate to form (only for primary and

secondary amines) while it is equal to 1 mol/mol amine for the bicarbonate

reaction.

The reactions listed take place from left to right during the acid gas absorption

stage and from right to left during the solution regeneration stage. Absorption is

favoured by low temperature while regeneration takes place at the boiling

temperature corresponding to the minimum operating temperature needed for the

acid gas produced.

In the case of the presence of stronger acid gases than and such as thiosulphate

and thiocyanate, the reactions with the amines causes the formation of salts that

cannot be regenerated even at high temperature.

It has already been mentioned that the reaction of H2S with the amines is

controlled by the gaseous film with the consequence that the H2S absorption

capacity is similar for all the amines.

However, the reaction between CO2 and amine is more complex: CO2 reacts with

the primary and secondary amines to form carbamate as these amines have an

unstable hydrogen atom which favours the reaction while the tertiary amines,

which do not have a hydrogen atom, do not form carbamate.

When CO2 dissolves in water it is firstly hydrolyzed to H2CO3 and then reacts to

form bicarbonate; as the dissociation of carbonic acid is slow, the entire CO2

absorption kinetics are slow.

The above stated reactions of CO2 are exploited when necessary to make a

selection, above all using concentrated solutions with low water content which

allow the absorption of CO2 to be favourably limited, leaving the H2S absorption

capacity unchanged.

The primary amines have stronger bases and are not suitable for selectiveabsorption of ; secondary amines are moderately selective: however, DEA has

double the CO2 absorption capacity of DIPA which is therefore more selective.

Instead, the tertiary amines are selective as the CO2 does not react to form

carbamate. However, TEA has too weak a base to offer an acceptable pick-up, at

least with moderate partial pressures of H2S while MDEA allows fair pick-ups

even with low partial pressures.

Note that MDEA and DEA are stable amines with a boiling point and atmospheric

pressure of 170°C and 268°C respectively while TEA, which has a boiling point

and atmospheric pressure of 180°C, decomposes before reaching boiling point.

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This means that this particular amine is not widely used for industrial sweetening

purposes even if in some applications with high partial pressures of H2S the

absorbing capacity of TEA can be cost-effectively exploited.

The selectivity of the amines, above all DIPA and MDEA can be maximized by

acting on the time the solution remains in the absorption column and the flow rateof the circulating solution (short time in the column and low circulation rate

favour selectivity).

The short time in the column can be obtained by reducing the contact volume

while the low circulation rate can be obtained with a higher concentration of

amine.

Figure 3.1.2 shows the effect of the flow rate of the circulating solution and the

contact volume on the selectivity and degree of purification that can be obtained

with a selective amine.

One way of increasing selectivity is to use an absorber which can supply the lean

solution at different heights: the correct contact volume is generally decided when

the unit is started up.

The behaviour of the amines in relation to the absorption capacity of organic

sulphur differs greatly: MEA absorbs CO2 and CS2 but it is degraded and only

absorbs a few of the lighter mercaptans; DEA and MDEA partially absorb organic

compounds such as CO2 and CS2 and do not react in any way with these

substances: DGA easily absorbs CS2 but is also easily degraded by it; finally,

DIPA absorbs the COS but is degraded by CO2 and by COS and . In the case of

the presence of organic sulphur compounds in the gas to be treated, the choice of

amine must be carefully considered; if the organic sulphur compounds are present

in large doses and are to be removed it is probably better to use a physical

absorption system or other more suitable processes such as catalytic hydrolysis of

the crude gas before sweetening; on the market there are specific catalysts able to

ensure a high conversion of CO2 and CS2 in H2S at temperatures of between

100°C and 150°C in the presence of fairly limited quantities of steam (3%-5% in

volume).

The presence of NH3 in the crude gas is tolerated by all the amines that are able to

ensure their almost complete removal; HCN instead causes the degradation of all

the amines.

The absorption of H2S and CO2 by the amines takes place due to the so-called driving force of the gas phase to the liquid phase.

In the gaseous phase the partial pressure of the acid gas is equal to the total

pressure multiplied by the molar fraction of the gas.

In the liquid phase the acid gas absorbed exercises a vapour pressure on the

solution; the vapour pressure depends on the concentration of the acid gas and the

solution temperature. The driving force is the difference between the partial

pressure of the acid gas in the crude gas to be treated and the vapour pressure of

the acid gas in the absorbing solution; at equilibrium, the two pressures are equal

and so the driving force is annulled and the solution cannot absorb any more acid

gas.

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When designing a sweetening system note that only the partial pressure of the acid

gas is not modifiable: the vapour pressure of the acid gas depends instead on the

type of amine, temperature, concentration of the amine and concentration of the

acid gas in the amine; all these parameters can be modified by the engineer to

optimize the sweetening system.

The absorption capacity of the various amines as regards acid gas and the vapour

pressure of the acid gases on amines are the fundamental parameters for

understanding the possibility of using the different amines.

Refer to table 3.1.6 which indicates the absorbing capacity of the amines: despite

the fact that MEA and DGA are able to chemically absorb very high quantities of

acid gases, the table has been drawn up assuming a limit of 0.4 mol acid gas/mol

amine considered the maximum possible value if serious corrosion of the plant is

to be avoided. The table shows the net pick-up that can be obtained using a lean

solution at 20°C, and the consumption of steam for regeneration of the solution

for cases of absorption at 15 and 50 eff. bar Examination of the table shows that the best performance from a merely

quantitative point of view are those that can be obtained with concentrated DEA

and DIPA (MDEA has a similar performance to DIPA).

Figure 6.1.4 shows the adiabatic load curves at equilibrium of the various amines

assuming a temperature of the lean solution of 40°C, as a function of the partial

pressure of the crude gas to be treated; also in this case the maximum

concentration of acid gas was limited to 0.4 mol/mol amine for MEA and DGA.

The figure shows that the DGA is the most interesting amine on the low partial

pressure field for H2S while DIPA and DEA are the most effective amines as the

partial pressure rises.

An interesting comparison between MEA, DEA and DIPA is shown in figures

6.1.41 and 6.1.42 where the vapour pressures of H2S and CO2 on equimolar amine

solutions at a temperature of 40°C are shown.

For both acid gases the MEA is the most competitive MEA, followed by DEA and

DIPA.

Figure 3.1.2 shows the vapour pressures in that can be obtained with the different

amines depending on the consumption of steam for regeneration. The figures

show that the best results in terms of crude gas purification can be obtained in thefollowing order: TEA, DIPA, DEA, MEA and DGA. MDEA has a similar

behaviour to DEA.

As can be seen from the data in the figures and tables, it is not possible to identify

a single most competitive amine: the total pressure of the gas, the partial pressure

of the and of the in the crude gas to be treated, the presence or absence of organic

sulphur compounds, the degree of purification and selectivity desired can play a

decisive role in the choice of process. Only detailed study of the problem with

reference also to the consumption of utilities can allow for the most cost-effective

choice to be identified.

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The modern approach is to use secondary or tertiary amines when selectivity is

required and/or when large quantities of acid gas have to be treated; primary

amines are used for fairly small quantities of acid gases when selectivity is not

necessary and when the partial pressures of H2S and CO2 are low.

The use of anti-corrosion additives can reduce the heat needed for regeneration of

the MEA by around 30%; however, the use of additives implies the payment of a

license. Anti-corrosion additives are generally poisonous and not biodegradable;

despite the fact that their presence in solutions is limited to just 0.5 – 2% in

weight, the fact that they are poisonous has limited the development of these types

of processes even although performance in terms of cost-effective operations is

very high.

Another system used to improve the performance of MEA is to add diethylene

glycol (DEG) to the watery MEA solutions. At the same operating pressure the

addition of glycol provokes a rise in the boiling point of the solution with ensuingimprovement in performance in terms of final purification of H2S and CO2 (the

residual H2S is less than half that obtainable with MEA alone).

The physical characteristics of some lean amine solutions are indicated below.

DGA and TEA have not been included as they are not commonly used nor are

DIPA solutions which are subject to license: sufficient data is available in any

case in past studies (Kohl and Riesenfield - Gas Purification).

The physical characteristics of the rich solutions have not been included but are

also available from past studies.

The figures below give a series of physical data and vapour pressures of acid

gases on amine solutions:

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PROCESS ACTIVE AMINETYPICAL AMINE

CONCENTRATIONKmol/m3

CONCENTRATION% weight

SOLVENTDENSITY

AT25°CKg/m

BOILING

POINT AT1.8 ata °C

DEA(HOCH2CH2) 2 NHM=105

2 21 1020 118

DEA – CONC.

(HOCH2CH2) 2 NHM=105

3 31 1035 119

M = Molecular weight Kg/Kmol

TABLE 3.1.5 – AMINES – COMPOSITION AND PHYSICAL PROPERTIES

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15 eff. bar 50 eff. bar

AMINE

CONCENTR . % weight of

amine inwater

NET

PICK-

UP

Kmol

Acid gas

m3 solutio

n

REGENERATIO

N STEAM

SOLUTION

Kg steamm3 solution

SWEETENIN

G STEAM

ACID GAS

Kg steamKmol acid gas

NET

PICK-

UP

Kmol

Acid gas

m3 solutio

n

REGENERATIO

N STEAM

SOLUTION

Kg steamm3 solution

SWEETENIN

G STEAM

ACID GAS

Kg steamKmol acid gas

DEA 21 1,62 152 94 1,84 99 52

DEA 31 2,16 183 85 2,55 113 44

(*) – CONSIDERING A LIMIT OF 0.4 mol acid gas/mol amine

MAXIMUM TO LIMIT CORROSION

THE TABLE WAS DRAWN UP CONSIDERING AN H2S SWEETENING FROM 10%

vol. TO 4 ppm vol. WITH APPROACH TO EQUILIBRIUM OF 33% AT THE

ABSORBER HEAD AND OF 70% AT THE TAIL OF THE ABSORBER

TABLE 3.1.6 – AMINES – ABSORBING CAPACITY

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% H 2 S A B S O R B E D

S E L E C T I V I T Y

- - - - - - - - - - - - - - - - - - - - - - - - - - - -

% C O 2 A B S O R B E D

SOLVENT IN CIRCULATION

FIGURE 3.1.1 – AMINES – SELECTIVITY EFFECT

N = NUMERO DI PIATTI

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Of the amines, DEA and MDEA are those with least tendency to absorb

hydrocarbons; MEA and DIPA have greater solubility while DGA has even higher

solubility.

Table 3.1.22 shows the solubility of hydrocarbons in water, in DEA and in MEA;despite the fact that solubility depends not only on the temperature but also on the

partial pressure and on the concentration of amine and the pick-up of acid gases,

the values in the table should be considered fairly precise for the calculation of the

global performance of a sweetening plant.

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D

E N S I T Y – k g / d m 3

TEMPERATURE - °C

FIGURE 3.1.3 – DEA – DENSITY OF SOLUTIONS

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V I S C O S I T Y – c e n t i p o i s e

TEMPERATURE - °C

FIGURE 3.1.4 – DEA – VISCOSITY OF SOLUTIONS

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S P E C I F

I C H E A T – K c a l / K g • ° C

TEMPERATURE - °C

FIGURE 3.1.5 – DEA – SPECIFIC HEAT OF SOLUTIONS

ACQUA

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T O T A L P R E S S U R E – m m H g

TEMPERATURE - °C

FIGURE 3.1.6 – DEA – VAPOUR PRESSURE OF SOLUTIONS

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P I C K - U P S O L U T I O N – m o l C O 2 / m o l D E A

VAPOUR PRESSURE CO2 – mm Hg

FIGURE 3.1.7 – DEA – VAPOUR PRESSURE OF CO2

CONCENTRAZION DEA 30% peso

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P I C K - U P S O L U T I O N – m o l H 2 S / m o l D E A

VAPOUR PRESSURE H2S – mm Hg

FIGURE 3.1.8 – DEA – VAPOUR PRESSURE OF H2S

CONCENTRAZION DEA 30% peso

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V A P O U R P R E S S U R E O F C O 2 - a t a

PICK-UP OF SOLUTION – mol CO2/mol DEA

FIGURE 3.1.9 – DEA – VAPOUR PRESSURE OF CO2 – 50°C

PARAMETRO mol H2S / mol DEAIN FASE LIQUIDA

DEA 36% PESO

TEMPERATURA 50°C

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V A P O U R

P R E S S U R E O F H 2 S - a t a

PICK-UP OF SOLUTION – mol H2S/mol DEA

FIGURE 3.1.10 – DEA – VAPOUR PRESSURE OF H2S – 50°C

PARAMETRO mol CO2 / mol DEAIN FASE LIQUIDA

DEA 36% PESO

TEMPERATURA 50°C

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V A P O U R


PICK-UP OF SOLUTION – mol H2S/mol DEA

FIGURE 3.1.11 – DEA – VAPOUR PRESSURE OF H2S – 70°C


DEA 36% PESO

TEMPERATURA 70°C

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V A P O U R

P R E S S U R E O F C O 2 - a t a

PICK-UP OF SOLUTION – mol CO2/mol DEA

FIGURE 3.1.12 – DEA – VAPOUR PRESSURE OF CO2 – 70°C


DEA 36% PESO

TEMPERATURA 70°C

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D

E N S I T Y – K g / m 3

AVERAGE CONCENTRATION - % weight

FIGURE 3.1.13 – DEA – DENSITY OF SOLUTIONS

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V I S C O S I T Y – c e n t i p o i s e

TEMPERATURE - °C

FIGURE 3.1.14 – DEA – VISCOSITY OF SOLUTIONS

ACQUA

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S P E C I F I C H E A T – K c a l / K g ° C

TEMPERATURE - °C

FIGURE 3.1.15 – MDEA – SPECIFIC HEAT OF SOLUTIONS

(ACQUA PURA)

PUNTO DIEBOLLIZIONE

PUNTO DI

CONGELAMENTO

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V A P O U

R P R E S S U R E – m m H g

TEMPERATURE - °C

FIGURE 3.1.16 – VAPOUR PRESSURE OF PURE MDEA

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T H E R M A L C

O N D U C T I V I T Y K c a l / m • h • ° C

MDEA CONCENTRATION - % weight

FIGURE 3.1.17 – MDEA – THERMAL CONDUCTIVITY OF SOLUTIONS

TEMPERATURA 40°C

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V A P O U R


PICK-UP OF SOLUTION – mol CO2/mol MDEA

FIGURE 3.1.18 – MDEA – VAPOUR PRESSURE OF CO2 – 40°C


MDEA 40% PESO

TEMPERATURA 40°C

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V A P O U R


PICK-UP OF SOLUTION – mol CO2/mol MDEA

FIGURE 3.1.19 – MDEA – VAPOUR PRESSURE OF CO2 – 100°C


MDEA 40% PESO

TEMPERATURA 100°C

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V A P O U R


PICK-UP OF SOLUTION – mol H2S/mol MDEA

FIGURE 3.1.20 – MDEA – VAPOUR PRESSURE OF – 40°C


MDEA 40% PESO

TEMPERATURA 40°C

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V A P O U R


PICK-UP OF SOLUTION – mol H2S/mol MDEA

FIGURE 3.1.21 – MDEA – VAPOUR PRESSURE OF H2S – 100°C


MDEA 40% PESO

TEMPERATURA 100°C

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ABSORBING

FLUIDWATER

DEA

20% WEIGHT

CH4 0,021 0,014

C2H6 0,041 0,029C3H8 0,031 0,021

C4H10 0,020 0,013

C5H12 0,012 0,008

C6H14 0,0024 0,0018

C7H16 0,0007 0,0005

SOLUBILITIES ARE EXPRESSED IN Nm3/m3•ata AT AMBIENT TEMPERATURE

TABLE 3.1.22 – SOLUBILITY OF HYDROCARBONS IN H2O, MEA AND DEA

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3.1.3 Column design

Amine plant absorbers are usually of the plate type; packed columns are used

when load flexibility is required or when large quantities of acid gas have to

be absorbed: the packing is normally metallic or ceramic.

In both cases it is advisable during design to consider that the amines are

fluids which can cause foaming and so suitable dimensioning margins must be

applied for the plates which must also be suitably spaced, especially in the

case of MEA.

The typical number of plates in an absorption column is 18-20; even although

the number of plates must be calculated on a case by case basis depending on

the operating conditions, it is rare that fewer than 18 plates are used; on the

contrary, columns with more than 35 plates have been designed and

developed.

The absorption columns generally have a head demister to contain

entrainments and operate with the lowest head temperature compatibly with

the need to avoid hydrocarbon condensation during contact between the gas to

be treated and the absorbing solution.

Ways of avoiding entrainments and hydrocarbon condensation in the column

include supercooling the crude gas to be scrubbed, using effective systems to

separate droplets in the gas to be scrubbed, insulating the incoming gas lines

and the bottom of the column and, finally, working with a temperature of the

solution at the head which is more than 10°C higher than the dew point in the

hydrocarbons of the gas to be scrubbed.

The regenerators are usually plate columns except when large quantities of

acid gas have to be treated.

Typically 18–20 regeneration plates are used with 2-3 upper plates for

scrubbing the acid gases released through the reflux and replenishment

condensates.

The dimension of the regenerators also takes into account the specific use,with the risk of the formation of foam.

Metallic material is preferred for the packing instead of ceramic material, even

if ceramic is still fairly widely used; the decision to use metallic packing

material is due partly to the undoubted superior ultimate strength under load

conditions and during operations, and partly to the instability of some types of

ceramics to the alkaline attacks of the amines at the relatively high

temperatures of the regenerators with ensuing tendency towards porosity and

disintegration. The suitability of the ceramic packing material for use with

amines can however be checked in advance by means of simple tests.

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The solutions are regenerated almost exclusively using steam supplied by

kettle or radiator type boilers; only in very particular cases is direct steam

used; however, it is difficult to dispose of because amine plants generally have

a fairly limited water balance.

Regeneration of the solution with direct flame boilers is used above all in

industrial contexts which do not have a heating fluid such as steam or hot oil

available. Direct flame boilers should be used with extreme caution in order to

avoid thermal degradation of the solutions due to the high temperatures of the

heat exchange surfaces.

Another limit to the use of direct flame boilers derives from some mechanical

calculation codes that make the design impossible when design pressures of

0.5 eff. bar are exceeded.

3.1.4 Design of other main equipment

The solution/solution exchangers which in the past consisted of series of pipe

bundles are now more frequently plate exchangers which are more compact

and efficient; there is a tendency to apply plate exchangers to cool the solution

when cooling water is available and application is possible at a fairly low

temperature.

In any case, the new generation of plate exchangers can be applied also at

pressures of 20-25 eff. bar

In most cases however, the solution is cooled and the acid gas condensed

using air coolants.

Turbines to recover energy from the rich solution leaving the absorber are not

normally used in amine plants due to the low flow rate of the circulating

solution, at least compared to absorption processes for large quantities of acid

gas such as hot carbonate plants.

However, in the case of two-stage designs with high rates of MDEA or DEA,

the use of recovery turbines greatly improves the economic result of the plant.

It has already been mentioned that a flash tank for hydrocarbons is needed to

recover hydrocarbons dissolved in the solutions (above all DIPA) or in cases

where the acid gas has to be supplied to a Claus unit and must thereforecontain only minimum quantities of hydrocarbons.

The flash tank is dimensioned on the basis of the time of residence which

varies from 3 to 6 minutes depending on the hydrocarbon absorbing capacity

of the solution.

In some case special systems (skimmers) are used to remove gasoline floating

on the solution in the flash tank.

If the gas released into the flash tanks has to be recovered with low H2S

content, it is normal practice to place a small scrubbing tower on top of the

flash tank supplied by a small quantity of lean solution.

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A mechanical cartridge filter is generally installed on the cold lean solution as

well as an activated carbon filter to entrain hydrocarbons.

The optimal position of this second filter is the subject of debate: we tend to

install it on the rich solution leaving the absorber to block the hydrocarbons before they reach the regenerator. Although it is theoretically possible to

regenerate activated carbon, it is preferable to replace it with a new charge

when it is rich.

Depending on the amine and type of service, the rate of solution to send to the

filters varies from a minimum of 5% to a maximum of 100% of the circulating

solution. The mechanical filters are usually designed to entrain particles of 50

micron even if filtering to 10-20 micron is preferable and in some cases

necessary for the plant to function correctly.

Many plants are fitted with injection systems for anti-foam agents to use onlyin cases of real need. The systems range from manual load tanks to small

dosed injection stations.

The type of anti-foam to use is not easy to identify in advance but can depend

on the operating conditions of the plant: however, the injection of anti-foam

agents must be limited over time so as not to exasperate the phenomenon due

to the presence in the solution of the degradation substances of the anti-foam

agent.

The solidification point of pure amines is lower than 0°C; the amines used in

accordance with their typical concentrations have solidification points of

around -5°C/-6°C except for DGA which solidifies at -40°C/-50°C.

However, it is advisable to use heating coils inside the storage tanks or

blowdown collection tanks where the solution can remain for long periods of

time without agitation.

The need to protect the pipes against the risk of freezing (with steam or

electric marking) depends on the design temperature of the plant; generally,

no protection is needed in Italian coastal areas but the pipes are emptied

during the long winter shut-downs.

Plants with amine solutions always have an underground blowdown collection

tank fitted with booster pump and a solution storage tank. The storage tank

must be able to contain all the solution present in the equipment; the

blowdown tank generally has a minimum capacity of 2-3 m3 and a maximum

capacity of 10% that of the storage tank.

A pump conveys the solution from the storage tank to the absorption cycle; all

the discharges from the plant equipment are conveyed to the blowdown

collection tank.

The importance of collecting all solution blowdowns, including accidental

blowdowns, lies in the need to completely eliminate any leakage of aminesfrom the plant. Modern amine plants are fitted with solution blowdown

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collection systems and are built in such a way as to minimise the risk of

accidental leakages (for example, pump stuffing box sealing fluids are not

used unless strictly necessary for machines subject to great stress but pumps

with self-fluxed mechanical seals are preferred).

Plenum in the blowdown collection and storage tanks is obtained with inert or

fuel gas to avoid the degradation of the solution in contact with the oxygen:

this is preferred also when not strictly indispensable, for example for MDEA

and DEA.

3.1.5 Construction materials

Mention has already been made of the corrosion caused by MEA, which is

more accentuated as the pick-up of the acid gas increases; it has also been

mentioned that the pick-up of the MEA can be increased by around 30% by

using special patented anti-corrosion additives (manufactured by UnionCarbide).

Corrosion of carbon steel similar to that caused by MEA is common also to

DGA and DIPA while DEA and MDEA have a decidedly less aggressive

behaviour.

Despite the reduced aggressiveness of DEA and MDEA compared to other

amines, nowadays all amines tend to be considered equally corrosive and so

the same materials are generally used in amine services.

Even if some companies used carbon steel with significant corrosion over-

thickness for amine services in the past, they now prefer to use stainless steel

where there is a statistically high probability of corrosion.

However, it should be noted that the parts most subject to corrosion are those

in contact with the acid gases, with greater corrosion where only CO2 or

mainly CO2 with H2S of not more than 1% vol. is present, followed

immediately by the parts in contact with the hot rich solution (temperature of

more than approx. 80°C).

Note once again that present-day practice is to anneal the weld seams of allthe carbon steel parts of the equipment that are in contact with the amines in

continuous and dynamic service (therefore not the tanks).

The presence of chlorides in the replenishment water is another source of

severe corrosion which is difficult to control even when stainless steel is used.

The absorption and regeneration column are made of carbon steel with

stainless steel interiors (plates or packing). The regeneration columns are

generally lined with stainless steel from 2-3 inches below the upper tangency

line up to the head.

The bottom of the regeneration column is also sometimes lined with stainlesssteel.

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For small columns in contact with particularly aggressive solutions (MEA,

MEA-DEG, DGA), it is common practice to use only stainless steel for the

entire construction.

The solution/solution exchanger is made of stainless steel in the case of plateexchanges and sometimes also, but only for the hottest element, in the case of

multiple pipe bundle exchangers.

The solution coolant is generally made of carbon steel except for plate type

coolants where stainless steel is used.

The acid gas coolant is generally made of stainless steel as is the boiler and

reclaimer solution side, heads excluded.

The flash tank and amine storage and blowdown collection tanks are made of

carbon steel as is the body of the solution filters.

The reflux accumulator of the regenerator (acid gas separator) is instead made

of stainless steel.

The solution and acid condensate circulation pumps have stainless steel body

and rotors.

The other machines in discontinuous service, such as solution blowdown

pumps and solution loading pumps are made of carbon steel.

The solution pipes downstream of the solution-solution economizer at the

regenerator head are made of stainless steel; the hot and cold acid gas pipes

and the acid condensate pipes are also made of stainless steel.

In the case of acid gas at a temperature of not more than 50°C and with high

H2S content it is possible to use carbon steel with steam or electric marking to

prevent the steam from condensing.

All the other pipes are made of carbon steel even if sometimes the lean

solution pipes to the solution-solution exchangers are made of stainless steel

to offer better protection.

All the solution control valves are made of stainless steel as are the manual

regulation valves even if they are assembled on carbon steel pipelines.

3.1.6 Foaming and amine leakage problems

Accurate evaluation of the solvent rate and the global thermal balance of an

amine plant requires precision calculations above all in the case of

simultaneous absorption of H2S and CO2; this section describes valid, rapid

methods to make rough evaluations or feasibility assessments.

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Plants which use loads of more than 0.4 mol/mol must however be designed

with great care, above all increasing the filtration of the solution; the use of

inhibited solutions allows the load limit indicated above to be increased to

0.5—0.55 mol/mol.

Calculation with DEA

The performance obtained in terms of H2S purification can be inferred from

figure 3.1.2.

MDEA has an intermediate behaviour between TEA and DEA.

The quantity of regeneration vapour is usually between 80 and 120 Kg/m3 of

solution (see table 3.1.4).

The preliminary calculation can be set assuming the use of a 40% weight

solution and a temperature of 10-20°C higher than the temperature of the head solution.

Using figures 3.1.18, 3.1.19, 3.1.20 and 3.1.21 the H2S and CO2 vapour

pressures can be extrapolated after assuming an absorbing solution rate.

The assumed solution flow rate is considered acceptable when the approach to

equilibrium is 60%-75%; that is, when the H2S and CO2 vapour pressures in

solution are around 1.35 – 1.7 times the partial pressures of H2S and CO2 in

the crude gas.

In the case of selective absorption it is assumed that it is possible to obtain an

acid gas with an H2S/CO2 ratio which is four times more than the same ratio in

the crude gas.

Other elements used to evaluate the economics of sweetening are the

hydrocarbons that are dissolved and then lost in the absorbing solutions and

the losses of solvent through degradation, entrainment and evaporation with

the sweet gas and acid gas.

Hydrocarbon losses can be calculated for MEA and DEA using table 3.1.22.

The solubility of hydrocarbons in MDEA is even lower than that in DEA.

Losses of amine through degradation cannot be prevented in advance; if there

is a risk of degradation due to the presence of reactive products in the gas to

be treated, a process which uses non-degradable amine must be chosen.

Losses through entrainment are almost always due to the presence of foaming

agents (surfactants such as oils, liquid hydrocarbons) in the absorbing

solutions: in this case it is necessary to identify and eliminate the cause of the

formation of foams and to limit their development by injecting anti-foam

substances; however, also in this case the loss of solvent cannot be evaluated.

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Losses through evaporation can be calculated assuming that the sweet and

acid gas contain evaporated amine in equilibrium at the vapour pressure of the

amine on the absorbing solution under set operating conditions (amine

concentration, temperature and pressure).

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C CONCENTRATION OF DEA % weight

T TEMPERATURE AT THE BOTTOM OF THE ABSORBER °C

P H2S VAPOUR PRESSURE ata

R REFERENCE LINE

mol H2S / mol DEA

FIGURE 3.1.2 3 – H2S VAPOUR PRESSURE ON DEA, H2S AND CO2

SYSTEMS

PARAMETRO

CARICAMENTO CO2

mol CO2 / mol DEA

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C CONCENTRATION OF DEA % weight

T TEMPERATURE AT THE BOTTOM OF THE ABSORBER °C

P CO2 VAPOUR PRESSURE ata

R REFERENCE LINE

mol CO2 / mol DEA

FIGURE 3.1.2 3 – CO2 VAPOUR PRESSURE ON DEA, H2S AND CO2

SYSTEMS

PARAMETRO

CARICAMENTO H2S

mol H2S / mol DEA

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V A P O U R

P R E S S U R E – m m H g

TEMPERATURE - °C

FIGURE 3.1.25 – DEA VAPOUR PRESSURE

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V A P O U R


H2S LOAD – mol H2S/mol amine

FIGURE 3.1.26 –H2S VAPOUR PRESSURE ON AMINE SOLUTIONS

TEMPERATURA 40 °C

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V A P O U R P R E S S U R E O F C O 2 - a t a

CO2 LOAD – mol CO2/mol amine

FIGURE 3.1.27 – CO2 VAPOUR PRESSURE ON AMINE SOLUTIONS

TEMPERATURA 40 °C

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3.1.7 Traditional absorption lay-out

The traditional lay-out, shown in figure 3.1.28, consists in using a solution with

a single regeneration level conveyed to the head of the packed or plateabsorption column; the gas to be treated is scrubbed in counter-current; the rich

solution leaving the bottom of the absorber passes into a flash tank if the

pressure level of the absorber is higher than approx 10 eff. bar or if the lean

acid is to be sent to a Claus unit.

The flash tanks usually operate at a pressure of 4-6 eff. bar, which is sufficient

to send the rich solution to the regenerator head without the aid of a pump. The

gas dissolved in the absorber and released into the flash tank is conveyed to the

fuel gas network, if envisaged, or is recovered; if necessary the flash gas is

scrubbed again with a small amount of lean solution.

The solution leaving the flash tank is heated in an economizer, cooling the leansolution, and then enters at the regenerator head at a temperature of between 80

and 110°C.

The solution is regenerated by means of indirect heating; the heat required is

supplied by a boiler assembled at the base of the regeneration column; this can

be a packed or plate column.

The boiler can be supplied by steam, hot gas, hot oil or pressurized water;

sometimes, in the case of small units, boilers with direct flame in which a gas is

burnt in a combustion chamber which directly heats the solution to be

regenerated are used. The boiler can be kettle, horizontal or vertical

thermosiphon or once through type.

Typical regeneration temperatures are between 105°C and 150°C depending on

the amine used and pressure reached (typically 0.5-1 eff. bar).

The acid gases released at the regenerator head are scrubbed on three scrubbing

plates and then sent to a condenser where they are cooled to 35°C-60°C; after

being separated from the condensate, the acid gases are sent for further

treatment (flare, incinerator, Claus unit).

The acid condensate separated in the reflux accumulator is pumped to the

regenerator head scrubbing plates; generally the process requires a water make-

up prepared with demineralised water supplied to the reflux accumulator or

directly to the regenerator. Whenever the system has excess water, the excess

acid condensate is collected at the reflux pump delivery.

The lean solution taken from the bottom of the regenerator is sent to the heat

recovery with the solution to be regenerated and is then pumped to the head of

the absorber after being cooled to 40°C-60°C.

If necessary a booster pump is installed upstream of the main pump.

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The temperature of the head solution is kept at least 10°C higher than the

hydrocarbon dew point of the gas to be treated to avoid gasoline condensation

in the absorber.

A mechanical filter is used for the lean solution and an activated carbon"chemical" filter able to retain the hydrocarbons and organic substances on the

rich solution.

Moreover, an underground blowdown collection tank with vertical booster

pump and storage pump with booster pump are normally envisaged; in both

tanks plenum is provided with inert gas or fuel gas; both tanks are fitted with

heating coil.

The storage tank can be included in the absorption cycle between a booster

pump and the main pump.

Amines are usually sold in 160 litre drums; due to the high freezing point, if aDEA solution is used, a heating system should be envisaged for the amine

drums to cut the loading and top up times during winter.

Solution reclaiming plants are envisaged when amines are used in services with

a high risk of degradation.

The reclaimer functions with heating fluid at high temperature; the reclaiming

operation consists in the continuous distillation of the solution with ensuing

concentration of high-boiling products of degradation of the amine which will

then be evacuated by the system (generally manually). Sometimes sodium has

to be used to neutralize any acid products of degradation.

The vacuum reclaimer is much more complicated in engineering terms and is

operated in continuous mode. Vacuum reclaiming systems are needed for

amines such as DGA and DIPA even if the gas to be treated does not contain

impurities such as CO2-CS2 or organic sulphur.

An important aspect for the efficient performance of the amine processes is to

avoid gasoline condensation in the absorber: it is therefore important to use gas

which is not saturated with gasoline or at suitable temperature in order not to

cause condensation.

If the gas to be treated is saturated in hydrocarbons, it is advisable to insulatethe bottom of the absorber and maintain the head solution at a higher

temperature than that of the gas; moreover, in these cases systems have to be

installed upstream of the absorber to separate the gasoline from the crude gas.

3.1.8 Split-Flow Lay-out

The split-flow lay-out, shown in figure 3.1.30, is very similar to the traditional

lay-out. The difference is that part of the lean solution is conveyed to the head

of the absorber at a lower temperature than that of the solution conveyed to themiddle of the absorber.

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FIGURE 3.1.29 – AMINES – VACUUM PRESSURE RECLAIMER

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FIGURE 3.1.31 – AMINES – TWO-STAGE LAY-OUT (SPLIT – STREAM FLOW)

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FIGURE 3.1.32 – AMINES – EQUITHERMAL TWO-STAGE LAY-OUT

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3.2 Physical absorption – Selexol Process

3.2.1 Description of the process

The SELEXOL process was developed by Allied Chemical and is now under Norton license.

The solution used in the SELEXOL process typically comprises 5% weight of

H2O and 95% weight of polyethylene dimethylether with molecular weight of

280 Kg/Kmol; the solution’s freezing point is between -22°C and -29°C and its

flash point is 151°C.

The typical characteristics of the solution are shown in figures 3.2.1 (vapour

pressure), 3.2.2 (density), 3.2.3 (specific heat), 3.2.4 (viscosity) and 3.2.5

(thermal conductivity). Table 3.2.6 shows the solubility of Selexol in relation to

a number of compounds using the solubility of methane gas as reference parameter.

The main characteristics of the SELEXOL process are:

- low regeneration cost, this is common to all physical absorption processes.

- limited circulation flow rate due to the good load that can be obtained in

high partial pressure conditions.

- no corrosion problems.

- limited losses of solvent through vaporization or degradation.

- good dehydrating capacity.

Like all physical absorption processes, SELEXOL can be profitably used when

the partial pressure of the acid gas is sufficiently high. The solution is conveyed

to the absorber also at very low temperature after being cooled.

If the total pressure of the gas to treat were high as well, the process would

guarantee a high degree of purification. SELEXOL is also suitable in cases

where selective absorption of H2S compared to CO2 is required and is also

suitable in cases where the gas to be sweetened contains CO2, CS2, mercaptans

or other organic compounds.

The process is used for sweetening natural gas and for treating synthesis gas.However, it must be borne in mind that, like all physical solvents, SELEXOL

absorbs noticeable quantities of hydrocarbons and this can be considered a

disadvantage.

As in all absorption procedures that use amines, Selexol is also highly sensitive

to the presence of liquid hydrocarbons in the gas to be treated; indeed, gasoline

provokes severe foaming of the solution with losses of solvent through

entrainment.

A particular characteristic of the SELEXOL procedure is the possibility of

regenerating the absorbing solution for subsequent flashes, followed by

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stripping with air or steam in cases where a high level of purification is required

as will be described in detail in the following section.

3.2.2 Lay-outs used

Figures 3.2.7, 3.2.8, 3.2.9 and 3.2.10 show four typical lay-outs for the

SELEXOL process.

Each process can is suitable for a specific sweetening needs.

The lay-out in figure 3.2.7 can be used to sweeten natural gas with high CO2

content.

Regeneration takes place for flashes only and so does not reach very high levels

of regeneration. However, this is acceptable as the sweetened gas does not

usually require very high decarbonation levels.

Small quantities of H2S are also allowed in the gas to be treated (10-20 ppm)

because in this case the CO2 flash has an adequate stripping effect.Final regeneration of the solution with air would be possible in the case of

crude gas containing only CO2 (integrating the lay-out in figure 3.2.7) if a fairly

high CO2 purification level is required but in the complete absence of H2S in the

gas to be sweetened.

Regeneration with air of SELEXOL solutions in cases of absorption of gas

containing is not possible for the transformation of into sulphur by the oxygen

present in the air and the consequent precipitation of metallic sulphur in

solution.

In cases where crude gas containing is to be treated, the final regeneration could

therefore be performed using steam or inert gases.

The lay-out in figure 3.2.8 is suitable for sweetening natural gas with a low CO2

content. Indeed, by exploiting the selectivity of the solvent with reference to

H2S (see table 3.2.6) all the CO2 acceptable for the specification can be left in

the treated gas while the H2S is completely absorbed.

A high level of regeneration of the solvent provides a sweetened gas of the

required purity in terms of H2S.

The solution envisaged in the lay-out in figure 3.2.9 sweetens natural gas with a

high CO2 content and marked H2S content. The H2S is absorbed selectively with

a part of the solvent deriving from the absorption of CO2. The solution leaving the H2S absorbed is regenerated separately from the rest of

the solvent, obtaining in this way an acid gas which is rich in H2S and which

can be conveyed to a sulphur recovery unit.

The lay-out in figure 3.2.10 is an example of selective absorption and

regeneration using stripping gas instead of steam.

Note that the stripping gas is generally an inert gas without oxidizing agents to

avoid oxidation of H2S to elementary sulphur.

Final regeneration with steam would however be possible even in cases of

treatment of crude gas containing H2S.

3.2.3 Design and dimensioning criteria

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The absorber generally has a packed column, even if plate columns have been

used in some exceptional cases.

SELEXOL has not shown a tendency to form foam if used to scrub synthesisgas while foaming phenomena have been note in the case of natural gas

sweetening.

In this case filters and gasoline elimination systems must be installed at the

absorber inlet. The plates or packing material must also be dimensioned with

care. The bottom of the absorber is usually designed in such a way as to ensure

a residence time of the absorbing solution of less than a few minutes to favour

complete absorption of the acid gases by the solution.

The correct dimensioning of the absorber bottom is essential to reach the

absorption approach to equilibrium set when calculating the solvent flow rate.

The regeneration column is generally a packed column. A typical packingconsists of 1” or 2” Pall metal rings. The packing height, in some cases where a

high level of purification is requested, can be more than 20 metres.

The heaters used to heat and cool the solvent can be pipe bundle type. However,

it is more cost effective to use plate exchangers for which it is possible to have

a pure counter-current flow.

The flash tanks are usually very large with residence times of 5-6 minutes to

allow good gas/liquid separation.

A cartridge solution filter is always used in the absorption cycle, able to treat 2-

10% of the circulating solution and able to entrain particles of up to 5 micron.

Plenum for the stored SELEXOL solution is not necessary providing that it has

been regenerated first; in general solution storage sections are not planned but,

during maintenance jobs, the solution is transferred from one flash tank to

another.

A collection tank must be envisaged for all continual and accidental blowdowns

with a booster pump to convey the blowdowns to the absorption system not

only to avoid pollution of the sewers but also because of the high cost of the

solution.


SELEXOL is a non corrosive solvent which also acts as a corrosion inhibitor

when working in a highly corrosive environment like that of a wet gas with

high CO2 concentration.

Indeed, decarbonation plants are built in carbon steel as are H2S elimination

ones.

However, some equipment in the regeneration zone where materials come into

contact with wet CO2 and H2S must be made of stainless steel.

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This equipment includes the regeneration column (even if in the past they were

made of carbon steel with high corrosion reinforcement) with relative packing,

the head condenser and the reflux accumulator.

It is advisable to use stainless steel also for the lines which convey the vapoursarriving from the boiler; a kettle type boiler can be made with carbon steel pipes

and skirt.

The SELEXOL solution originally contains around 10 ppm of chlorides; it is

therefore extremely important to control the chlorides especially if frequent top

ups are necessary.

It is advisable to use stainless steel for the pipes that convey the preheated

solution to the regeneration column head.

To assess the effectiveness of a process lay-out first of all it is necessary todefine the global energy balance. That is, the thermal energy needed to

regenerate the solvent, the energy needed for circulation and the energy needed

to cool the solvent must be evaluated.

Indeed, gas can be sweetened in certain plants with a wide range of process

parameters.

For example, the flow rate of solvent can be reduced by increasing the

regeneration heat and lowering the temperature of the solvent conveyed to the

absorption.

See the graphs in figures 3.2.11 and 3.2.12 for a rough assessment of the rate of

solvent needed to sweeten a given flow of gas.

The graph in figure 6.3.11 can be used to estimate the flow rate of solvent

required to absorb the CO2, while the graph in figure 3.2.12 can be used to

estimate the flow rate of solvent needed to absorb the H2S.

Not taking into consideration the interaction between CO2 and H2S it can be

said that the flow rate to assume is the higher value of the two values indicated

on the respective graphs.

The graph in figure 3.2.11 is used with the molar fraction of the CO2 in the gas

to be treated and its pressure, plotted in quadrant number 1.

With the calculated parameter move up to quadrant number 2 until intersecting

with the value the approach to equilibrium on the bottom of the absorber value;

typical approach to equilibrium values for CO2 are 85 % - 90 %.

Then move left to the third quadrant until intersecting with the temperature of

the solution at the bottom of the absorber and then down to quadrant number 4

until intersecting with the difference in concentration of the CO2 between the

gas to be treated and the sweetened gas. This point represents the flow rate of

solvent needed in m3/h for a gas rate of 100,000 Nm3/h.

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Evaluation of hydrocarbon absorption

As stated above, hydrocarbons are soluble in the SELEXOL process solution.

Since releasing hydrocarbons along with the acid gas during regeneration of the

solvent could represent excessive loss of energy, the hydrocarbons are oftenrecycled in the absorber and separated by flashing the solvent.

It is essential to choose the pressure at which the flash is to take place with

great care. Indeed, if the flash takes place at high pressure few acid gases would

be released and there would be a saving of hydrocarbon recompression energy;

instead, if the flash takes place at low pressure, the hydrocarbons would not be

lost during the subsequent regeneration stages but many acid gases which

should be recycled in the absorber would be lost, with high recompression

costs.

A pressure is usually chosen where the separated gases can be recycled in theabsorber with a single compression stage. For example, if absorption takes

place at 70 bar the flash of hydrocarbons could take place at 20 - 25 bar.

To make a rough estimate of the amount of hydrocarbons dissolved in the

SELEXOL solution the relative solubility data given in figure 3.2.6 can be

used, taking as reference data the solubility of the acid gas obtainable from

figures 3.2.11 and 3.2.12.

The results that can be obtained from decarbonation of a gas at a pressure of 70

abs. bar and a temperature at the bottom of the absorber of 0°C with approach

to equilibrium of 85%; a flash of the solution at a pressure of 21 abs. bar has

been estimated with recycling of the flash gases in the absorber; the rate of the

Selexol solution is 690 m3/h.

Crude gas Sweet gas Absorbed gas

CO2 (Nm3/h) 20,000 1,300 18,700

CH4 (Nm3/h) 60,000 58,570 1,430

C2H6 (Nm3/h) 10 ,000 4,795 5,205

C3H8 (Nm3/h) 6,000 -- 6,000

C4H10 (Nm3/h) 4 ,000 -- 4 ,000

Total (Nm3/h) 100,000 64,665 35,335

Note the large quantities of hydrocarbons absorbed with CO2, above all ethane;

this result represents the limit of applicability of the Selexol process.

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V A P O U

R P R E S S U R E – m m H g

TEMPERATURE - °C

FIGURE 3.2.1 – SELEXOL – VAPOUR PRESSURE

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D E N S I T Y – K g / d m 3

TEMPERATURE - °C

FIGURE 3.2.2 – SELEXOL – DENSITY

SELEXOL 100%

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S P E C

I F I C H E A T – K c a l / K g • ° C

TEMPERATURE - °C

FIGURE 3.2.3 – SELEXOL – SPECIFIC HEAT

SELEXOL 100%

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V

I S C O S I T Y - c e n t i p o i s e

TEMPERATURE - °C

FIGURE 3.2.4 – SELEXOL – VISCOSITY

SELEXOL 100%

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T H E R M A L

C O N D U C T I V I T Y – K c a l / m • h • ° C

TEMPERATURE - °C

FIGURE 3.2.5 – SELEXOL – THERMAL CONDUCTIVITY

SELEXOL 95% PESOH20 5% PESO

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ALL THE SOLUBILITY VALUES REFER TO A TEMPERATURE OF 20 °C, AT A

PRESSURE OF 70 bar, WITH MOLAR CONCENTRATION APPROACHING ZERO

AND ARE EXPRESSED AS A PARAMETER OF THE SOLUBILITY OF METHANE

COMPONENT

H2 _______________ 0.20

CO _______________ 0.43

C1 _______________ 1.0

C2 _______________ 6.4

C2H4 _______________ 7.3

CO2 _______________ 15.0

C3 _______________ 15.3

iC4 _______________ 28.0

nC4 _______________ 35.0COS _______________ 35.0

iC5 _______________ 67.0

H2S _______________ 83.0

H2S _______________ 134.0

C6 _______________ 165.0

CH3SH _______________ 340.0

C7 _______________ 360.0

CS2 _______________ 360.0

C4H4S _______________ 8100.0

H2O _______________ 11000.0

FIGURE 3.2.6 - SELEXOL – RELATIVE SOLUBILITY OF GAS

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FIGURE 3.2.7 – NON SELECTIVE SELEXOL– GAS WITH HIGH CO2 CONTENT AND TRACE

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FIGURE 3.2.8 – SELECTIVE SELEXOL FOR H2S – GAS WITH LOW CO2 CONCENTRATION

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FIGURE 3.2.10 – SELECTIVE SELEXOL FOR H2S – GAS WITH HIGH H2S AND LOW C

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TEMPERATURE OF SOLUTION AT THE BOTTOM OF THE ABSORBER APPROACH TO EQUILIBRIUM

RATE OF SELEXOL FOR 100000 Nm3/H OF CRUDE GAS CRUDE GAS PRESSURE

INLET-OUTLET MOLAR FRACTION OF CO2 MOLAR FRACTION OF CO2 IN THE CRUDE GAS

FIGURE 3.2.11 – SELEXOL – EVALUATION OF SOLVENT RATE - CO2

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Page. 100 of 175 TEMPERATURE OF SOLUTION AT THE BOTTOM OF THE ABSORBER APPROACH TO EQUILIBRIUM

RATE OF SELEXOL FOR 100000 Nm3/H OF CRUDE GAS CRUDE GAS PRESSURE

INLET-OUTLET MOLAR FRACTION OF CO2 MOLAR FRACTION OF CO2 IN THE CRUDE GAS

FIGURE 3.2.12 – SELEXOL – EVALUATION OF SOLVENT RATE - H2S

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3.3 Chemical-physical absorption - SULFINOL Process

3.3.1 Description of the process

Shell’s SULFINOL process is a chemical-physical process: The first plant wentinto operation in 1964 and there are now some 150 plants throughout the world.

The process is used for the purification of natural gas (around 50% of the

operations) and for the purification of synthesis gas (the other 50%).

The fluid used for the absorption of the acid gases is a watery solution of

physical solvent (tetrahydrothiophene 1-1 dioxide or Sulfolan) and a chemical

solvent such as diisopropanolamine (DIPA); other amines can be used such as

MDEA in mixture with Sulfolan, as happens in some cases in more modern

plants.

Sulfolan acts as a physical solvent while the amine’s typical absorbing capacityis the result of a chemical reaction; the typical concentrations of the

SULFINOL solution are:

Sulfolan 40 % weight

DIPA 45 “

H2O 15 “

The SULFINOL solution is ideal in the cases of absorption of COS, CS2 and

mercaptans for which its absorbing capacity is extremely high (at 70 abs. bar

less than 10 ppm vol. of mercaptans can be reached in the sweet gas); the

presence of a physical solvent along with DIPA does not however allow

selectivity as regards CO2; if selectivity is necessary than MDEA-Sulfolan

mixtures must be used .

The optimal field of use for this process is gas with an H2S/ CO2 ratio of more

than 1.

Consequently, the presence of a physical solvent means that the solution has a

high absorbing capacity as regards hydrocarbons which provokes greater

absorption of the amines but which is still less than that of the traditional

physical processes; a flash tank between the absorber and regenerator istherefore needed to produce an acid gas with hydrocarbon content which is

compatible with supply to a Claus unit.

As opposed to physical absorbents, SULFINOL solutions are not able to dry the

sweet gas except for a slight reduction of the dew point in H2O of not more than

2 - 5°C.

Figure 3.3.1 shows the theoretical absorption capacity of Sulfolan, water, MEA

20% weight and typical SULFINOL solutions as a function of the partial

pressure of H2S; it is clear how advantageous this solution is as the partial

pressure increases.

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Some physical characteristics of Sulfulan:

Molecular weight 120.17 Kg/Kmol

Boiling point 285 °C

Freezing point 100% weight95% weight

0% weight

27.56

0

°C°C

°C

Flash point 177 °C

Critical temperature 528 °C

Critical pressure 52.2 atm abs

Heat of fusion 2.7 Kcal/Kg

Heat of vaporization 123

113.9

Kcal/Kg

Kcal/Kg

Specific heat 0.31

0.35

0.40

Kcal/Kg.°C

Kcal/Kg.°C

Kcal/Kg.°CVapour pressure 0.00002

0.00210

0.126

Kg/cm2 abs

Kg/cm2 abs

Kg/cm2 abs

Figures 3.3.2, 3.3.3, 3.3.4 and 3.3.5 show some physical characteristics of

standard lean SULFINOL solutions: density, viscosity, specific heat and

thermal conductivity.

The SULFINOL solution in small quantities is biodegradable in the typical

refinery biological treatments; the toxicity of the solution for fish is fairly low

(more than 4 - 5 g/litre).

The cost of the SULFINOL solution is fairly high; compared to primary and

secondary amines the unit cost is around 3 times more; considering the high

concentrations of the solutions used it is therefore necessary to note that the

investment for the first charge is around 10-15 times higher than that of amines

such as MEA and DEA.

3.3.2 Lay-outs used

The typical lay-out for the SULFINOL process is the classical one for amine processes shown in figure 3.1.28.

However, a vacuum reclaimer must be installed as shown in figure 5.1.2, able

to remove the products of degradation such as DIPA-oxozolidone from the

solution.

The predicted consumption of DIPA due to secondary reactions is fairly

important; a plant with a fairly low circulation (30 - 40 m3/h) which treats gas

containing H2S, CO2 and other sulphur compounds has an estimated

consumption of approx. 1 Kg/h of solution.

3.3.3 Design and dimensioning criteria

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The absorber is generally a plate columns which can sometimes have more than

30 plates as in the case of high level purification of H2S and other sulphur organic

substances.

The tendency of the solution to foam is limited, so there is no need for high

margins in the dimensioning of the columns. Like for most physical solvents, the

residence time of the solution must be approx. 6 minutes at the bottom of the

absorber and at the bottom of the regenerator in order to approach equilibrium of

physical absorption and to move away from equilibrium during regeneration.

The bottom of the absorber has a thin layer of packing material to favour the

absorption of acid gases.

Despite the fact that Sulfulan and DIPA do not have high vapour pressure, it is

preferable, to limit consumption, to install a gas cooling or scrubbing system atthe head of both columns; the ideal temperature of the sweet gas leaving the

absorber or of the acid gas leaving the regenerator is less than 35°C.

The tendency of the solution to become dirty is fairly limited; if plate exchangers

are used to recover heat between solution/solution it is preferable to increase the

capacity of the mechanical filter; in the case of transformation to SULFINOL of a

plant previously operating with another solvent is it necessary to clean the plant

chemically to prevent the SULFINOL from acting as a pickling agent and

removing all the metal scales present in the equipment involved in the circulation.

Unlike other amine solutions, SULFINOL does not expand when it solidifies so

no special anti-freeze measures are needed.

The need for a flash tank between the absorber and regenerator to separate most of

the hydrocarbons present in the rich solution; the residence time of the liquid is

around 6 minutes.

If necessary, an absorption tower can be used to absorb the H2S in the gas leaving

the flash tank to limit the H2S concentration; a small quantity of fresh solution is

sufficient to ensure good purification.

Considering that the SULFINOL solution tends to absorb heavy aromatichydrocarbons, if a Claus unit is installed downstream suitable acid gas treatment

techniques must be used to achieve full combustion (see section 4.3.).

The regeneration column is generally a plate column with 15-20 regeneration

plates plus 2-3 acid gas scrubbing plates at the head.

In general, threaded connection should be avoided, as generally required for

physical solvents, to prevent leakage of the solution; however, Teflon seals are

suitable for the threaded parts.

The seals of pipes and equipment are made of graphite impregnated asbestos;

rubber or elastomer seals should be avoided. A mechanical filter is usuallyinstalled to entrain particles of up to 50-100 micron; during the pre start-up stage

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of the plant it is common practice to filter the solution through cartridges able to

entrain particles of 4-5 micron; after having stabilized the concentration of solids

in the solution, the cartridges are replaced with normal ones able to filter up to 50

-100 micron.

Plenum in the storage and blowdown collection tanks must be provided using

inert gas.

The dimensioning of a SULFINOL plant requires detailed calculations and the

preparation of a global thermal balance; for a rapid evaluation of the flow rate it is

however possible to use the method described below which allows several

simplifications but which is sufficiently accurate for a rough feasibility study.

The net pick-up of acid gases in the SULFINOL solution varies depending on the

concentration of acid gas and on its partial pressure.

In general the load that can be obtained is more favourable for H2S than CO2; the

pick-ups indicated in figure 3.3.1 refer to H2S but can be considered valid for

preliminary dimensioning calculations also in the case of CO2.

The pick-ups indicated in the figure are calculated at equilibrium; the flow rate of

solution is generally dimensioned by calculating 60-70% of approach to

equilibrium. The partial pressure to consider for the calculation is the total of H2S

+ CO2.

The execution of preliminary calculations need not take account of the incoming

concentration of sulphur products such as mercaptans, COS and CS2, which are

taken into consideration using laws of interaction which cannot be simplified in

formulas or diagrams.

The result obtainable in terms of purification is however always less than 1% in

volume for CO2 (generally 500 ppm) and 20 ppm for total sulphur products when

operating at a pressure of approx 70 abs. bar; proportional results can be obtained

at different pressures.

Refer to figure 3.3.6 for H2S; this figure correlates the consumption of vapour

used to regenerate the solution with the partial pressure of H2S at the equilibriumof the lean solution regenerated at 40°C (allowing a difference in temperature of

the solution between the regenerator head and tail of not more than 20°C).

The consumption of regeneration vapour is however usually not less than 80

Kg/m3 of solution; the approach to equilibrium to be considered in this case is 25-

35% of the partial pressure of H2S at the equilibrium indicated in figure 3.3.6.

An example of use of the diagram.

To determine the H2S that can be obtained in an absorber operating at 50 abs. bar,

consuming 80 Kg/m3

of regeneration steam and considering an approach toequilibrium of 25%; the allowable parts per million in volume of H2S are:

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16 x 10-6 x 106

ppm H2S = ------------------- = 1,28 ppm vol.

50 x 0,25

A steam consumption of 80 Kg/m3 of solution is typical when the gas to be treated

has a moderate presence of COS and mercaptans; in the case of the presence of

COS in quantities of more than 50 ppm in the supply gas, the consumption of

steam must be increased even up to 150 Kg/m3 of solution.

The quantity of C1-C2-C3 hydrocarbons absorbed by the standard SULFINOL

solution is less than half that absorbed by a physical solvent; it is similar to that

absorbed by a physical solvent for C4 hydrocarbons; instead, aromatic

hydrocarbons are absorbed in a much more massive manner.


The solution used in the SULFINOL process has similar corrosion characteristics

to primary amines; the level of corrosion is however higher if the gas to be treated

contains CO2 with H2S of less than 1%.

Cast iron must be not be used.

In general, when temperatures are below 80°C, carbon steel is used with corrosion

reinforcement of 3mm; the carbon steel parts are generally subjected to annealing

to prevent the typical stress cracking phenomena of amines.

The absorber is made of carbon steel killed with stainless steel plates; if the

bottom temperature is high, the lower part of the absorber can be cladded with

stainless steel.

It is advisable to avoid the accumulation of chlorides in the solution as these could

cause severe corrosion probably because they destroy the protective film created

by the SULFINOL solution in the steel; the chlorides are almost always

introduced into the system with replenishment water which is not fully

demineralised: their content should never exceed 50 ppm.

The packing material at the bottom of the absorber is AISI 321 o 347.

The solution/solution exchanger is made of stainless steel, except in cases where

the temperature is less than 80-90°C; if pipe bundles with several bodies are used,

it is more cost-effective to use carbon steel for the pipe side of the first body,

which operates at lower temperature.

The flash tank, which generally operates at a temperature of less than 80°C, is

made of carbon steel.

The regenerator is made of carbon steel with stainless steel plates.

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The parts of the boiler in contact with the solution are made of carbon steel if a

heating fluid at a temperature of less than 150°C is used, otherwise stainless steel

is used.

The material used for the acid gas coolant, acid gas separator and relative pipesdepends in the type of acid gas: in the past carbon steel was used for H2S and

stainless steel for CO2; nowadays stainless steel is preferred for both gases to

avoid severe corrosion event after just a few months of operation.

The solution circulation pumps are made of stainless steel; carbon steel would

have a very short life (even less than one year): stainless steel is used also for the

solution recycling pumps from the reclaimer and for the acid condensate pumps.

The reclaimer is made of carbon steel, except for the boiler pipes which are

generally made of stainless steel.

Mention has already been made of rubber seals which must be avoided in

operations with SULFINOL solution while the only acceptable elastomero is

Ethylene-Propylene-Diene-Monomer EPDM which must be applied without

greases.

As a rule, carbon steel should be used for solution pipes up to 80°C and stainless

steel for higher temperatures; the regulation valves should be made of stainless

steel.

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D

E N S I T Y – K g / m 3

TEMPERATURE - °C

FIGURE 3.3.2 – SULFINOL – DENSITY OF LEAN SOLUTION

SOLUZIONE SULFINOL STANDARDDIPA 45% PESO

SULFOLANO 40% PESOH20 15% PESO

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V I S C

O S I T Y - c e n t i p o i s e

TEMPERATURE - °C

FIGURE 3.3.3 – SULFINOL – VISCOSITY OF LEAN SOLUTION



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S P E C I F

I C H E A T – K c a l / K g • ° C

TEMPERATURE - °C

FIGURE 3.3.4 – SULFINOL – SPECIFIC HEAT OF LEAN SOLUTION

SOLUZIONE SULFINOL STANDARD

DIPA 45% PESOSULFOLANO 40% PESO

H20 15% PESO

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T H E R M A L C O

N D U C T I V I T Y – K c a l / m • h • ° C

TEMPERATURE - °C

FIGURE 3.3.5 – SULFINOL – THERMAL CONDUCTIVITY OF LEAN SOLUTION



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P A R T

I A L P R E S S U R E H 2 S I N T H E

S O L U T I O N R E G E N E R A T E D A T 4 0 ° C - a t a • 1 0 - 6

STEAM CONSUMPTION FOR REGENERATION – Kg/m3

FIGURE 3.3.6 – SULFINOL – STEAM FOR REGENERATION OF SOLUTION

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4. OXIDATION PROCESSES TO REDUCE HYDROGEN SULPHIDE

This classification includes all the oxidation absorption processes in liquid phase

and the oxidation processes in gaseous phase; the first include the STRETFORD,

SULFINT and LO-CAT processes; the second include the CLAUS units.

The processes in liquid phase are selective and directly transform the H2S content in

the gas to be treated into sulphur; they can operate at high or low pressure and are

limited by the sulphur potential which ranges from several Kg/h to a maximum of

around 10 t/d, above which they are no longer economically advantageous.

4.1 SULFINT and LO-CAT processes

Refer to figure 4.4.1 for the SULFINT process.

The H2S is absorbed by the solution (complex of iron chelates) in a Venturi or anabsorption column; the rich solution is oxidised with air in an oxidizer: the finely

divided solid sulphur formed during oxidation precipitates to the bottom of a settler.

The clear solution collected at the head of the settler is recycled in the absorber.

The sulphur mud, at a concentration of 5% in weight, is further concentrated

through centrifugation and is then conveyed at a concentration of 30% in weight,

for subsequent purification, for example through fusion, to obtain a more easily

marketable product.

The settled solution, containing secondary compounds formed during the various

stages of the process, (mainly sulphates), is sent to permeable membranes which

separate the large molecules (iron chelates) from the small molecules (sulphates);

the regenerated solution is recycled in the oxidizer while the by-products, which are

not harmful, are discharged into a effluent treatment system.

In small plants mud treatment can be greatly simplified if the sulphur produced can

be discharged directly.

If the quantity of H2S in the gas to be treated is high several absorption columns

would be needed.

LO-CAT Process

Refer to figure 4.1.2 for the LO-CAT.

The idea is similar to that of the SULFINT process; the solution used is iron chelate

based buffered at pH 8 with Na2CO3 and KOH; the H2S is oxidized to sulphur by

means of a special non-toxic ethylendiamine tetra-acid based catalyst.

The H2S can be absorbed and the solution is oxidized in a single tank when the

procedure is used for low pressure exhaust gas; otherwise, a separate packing or

bubble absorption column preceded or not be a Venturi can be used.

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The sulphur produced can be sent for further treatment if a high purity product is

required; a part of the solution must be evacuated to keep under control the

concentration of secondary compounds in the solution.

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FIGURE 4.1.2 – OXIDATION PROCESSES IN LIQUID PHASE – LO-CAT

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4.2 Monterotondo process with biological oxidation

The development of fields with a high rate of H2S has contributed to the study of

innovative processes to reduce this dangerous pollutant, with competitive

engineering costs compared to traditional technologies.

For this reason EniTecnologie is developing a redox system based on a chemical

and biological approach in liquid phase, characterized by an almost zero

environmental impact considering that the process has been classified “zero

discharge”.

Redox technologies in liquid phase are advantageous also in reducing H2S in

gaseous current with low pollutant content.

These technologies are based on two main steps: in the first step the H2S is oxidised

to elementary sulphur through reduction of metal ions solubilized in water in free or

complex organ; in the second step the reduced ions are re-oxidised generally usingthe oxygen in the air. The high demand for redox technologies has led to the

development of new plants mainly based on the using the Fe2+/Fe3+ [2] redox

couple.

In recent years much space has been gained by biotechnological processes based on

the exploitation of the capacity of Thiobacilli to directly oxidise sulphides to

elementary sulphur [3].

An interesting way of combining efficiency of the oxidation reaction of sulphides

with the ferrous ions and the oxidation capacity in Thiobacilli acid environment is

based on the following pairs of reaction:

H2S + Fe2(SO4)3 -> S + 2 FeSO4 + H2SO4

2 FeSO4 + H2SO4 + 0,5 O2 -> Fe2(SO4)3 + H2O_____________________________

H2S + 0,5 O2 -> S + H2O

The first reaction takes place spontaneously with high kinetics; the second is

catalyzed by Thiobacillus ferrooxidans able to react in a pH range pH (1.4-1.8) in

which the iron ion does not precipitate, increasing the natural oxidation kinetics of

the ferrous iron some 500,000 fold [4].

The difficulties involved in this process regard the choice of a suitable reactor for

the precipitation and separation of the elementary sulphur, the low oxidation

efficiency of the biological systems and the limited stability over time of thealignment between the two main stages of the process. During the experimentation

stage the data needed to design a preliminary bench-scale plant was collected. In

particular, for the chemical section of the process, the mechanisms involved in the

formation of crystalline forms of sulphur (more easily separable from the reaction

mixture) as a function of the liquid-gas contact system, reaction temperature and

reactor configuration were studied. As regards the biological section of the process,

aspects related to the multiplication of the micro-organisms were studied in more

detail, making the nutritional needs compatible with the need to avoid the formation

of nutrient bouillon very low solubility. Moreover, the oxidation capacities with

respect to the ferrous ions of a micro organism cultivated in dispersed or adhesive

form on inert supports.

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The continuative management of the bench plant for around one year firstly allowed

us to define the conceptual diagram of the process, as shown in figure 4.1.3.

BIOLOGICAL AND CHEMICAL REDUCTION OF H2S IN GASEOUS STREAMSCONCEPTUAL DIAGRAM OF THE ENITECNOLOGIE PROCESS

FIGURE 4.1.3 - DIAGRAM OF THE ENITECNOLOGIE CHEMICAL

BIOLOGICAL REDUCTION PLANT.

Subsequently the know-how needed to design a prototype plant and to make the

first economic estimations was developed. The study also continued with the aim of identifying supports suitable for the biological substratum. The prototype is a skid

mounted type with a maximum capacity of 10 kg/day.

As can be seen in the figure, the acid gas to be treated is sent to the reaction column

into which the acid solution of ferric ions is introduced in a concentration of 0.3 M.

The slurry obtained in this way (density of 1.2-1.4 g/l) is conveyed to a filtering

system, which separates the ferrous sulphates from the sulphur. Then, this current is

fluxed with water in a fluxing reactor and then vacuum filtered, to obtain a cake

with 40-60% in weight of sulphur.

The filtered current is then sent to the biological reactor following replenishment of

the ferric ion solution and the introduction in the system of the nutrients needed for the growth of the bacteria. The redox reaction to restore the original solution takes

place in the reactor. The bacteria possess the energy needed for the growth from

oxidation of the ferrous ions.

As stated, continuative operations in conditions which simulate those existing in

production plants has verified that the new process can be considered “zero

discharge” and does not generate effluent or waste of any type. The use of an

autotrophic micro-organism such as T. ferrooxidans allows even a small portion of

carbon dioxide (15 kg per ton of sulphur produced) to be eliminated.

The elementary sulphur produced contains good commercial characteristics, and is

particularly suited for use in agriculture in terms of purity (~ 99%), dispersion inwater, iron ion content and requires modest consumption of chemicals and energy.

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Gaseous currents of different origin (excluding those containing SO2 in a

concentration > 2 mM) can be treated with hydrogen sulphide at concentrations

ranging from just a few ppm up to more than 70% in volume, demonstrating

marked operative flexibility.

It should also be noted that the plant operates at pressures and temperatures close toenvironmental values.

Comparative evaluations to date show how the new chemical-biological process is

competitive an economic point of view and as regards environmental impact

compare to competitors’ processes based exclusively on chemical approaches.

These considerations have led to the inclusion of this process among the future

hydrogen sulphide reduction options.

4.3. Claus process and variants

The CLAUS process which transforms H2S directly into sulphur is applied to acid

gases produced in the sweetening plants and to the gaseous effluents arriving from production units containing high concentrations of H2S or other substances,

sulphides such as CO2, CS2, mercaptans or other sulphur compounds.

4.3.1 General description of the process, reactions involved and catalysts involved

The CLAUS process is based on a technology used since 1883 and which consisted

in burning the H2S with air in the presence of a catalyst to obtain sulphur.

Already in 1932 the IG CLAUS process was based on the combustion with SO2 of a

third acid gas to be treated; the products of combustion of the acid gas are combined

with the remaining part of the gas and sent to a catalytic converter which operated

at 200 - 300 °C; the sulphur was therefore produced only in the catalytic conversion

stage (see conversion reactions on the following pages) with fairly modest

performance.

In 1936 K. Braus developed the so-called MODIFIED CLAUS process which

envisaged the combustion of all the H2S contained in the acid gas with the

stoichiometric air for the transformation of the H2S into sulphur; this discovery

allowed for marked simplification of the technology and above all allowed high

levels of H2S /sulphur conversion to be reached through exploitation of the heat

conversion reactions during the combustion stage (see below) besides the catalyticreaction which took place downstream of the combustion of the acid gas.

At present the technology uses the MODIFIED CLAUS process as it was developed

by several US and European companies immediately after the war and its ongoing

improvements made possible by in-depth investigation of various aspects of the

process.

The best materials available for the construction and above all the use of particular

instrumental components have resulted in the very high level of reliability of these

processes above all if compared to the operative and maintenance situations in the

early Seventies which were certainly not optimal.

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The cost of CLAUS units has dropped greatly since the “package” concept was

introduced on an industrial scale at the end of the Fifties; this concept envisages

gathering together the equipment needed for the process in a few multi-purpose

components (2 or 3) with a marked saving in cost and without operative and

management constraints.

The concept of “package” construction is economically advantageous for units with

dimensions of up to 60 t/d of produced sulphur even it is applied in part also for

higher potential.

The basic criterion of the MODIFIED CLAUS process therefore implies the full

combustion with SO2 of a third of the H2S contained in the supply acid gas.

The combustion reaction takes place in a single stage, burning the acid gas sub-

stoichiometrically with the quantity of air needed for the transformation of H2S into

sulphur; that is, using a half mole of oxygen per mole of H2S.

The SO2 produced in the combustion reacts with the H2S which has not reacted

forming sulphur in the form of vapour and water vapour according to the following

simplified reactions:

H2S + ½O2 === H2O + SO2 oxidation

2 H2S + SO2 === 3S + 2H2O conversion

The sum of the two reactions can be expressed as

3 H2S + 1½O2 === 3S + 3H2O

or better, the simplified reaction can be expressed as

H2S + ½ O2 === S + 3 H2S

It has already been mentioned that the various reactions take place in different

stages: first of all there is a thermal reaction in which the H2S is burned in an

environment with very small oxygen content; then there are two or more stages of

catalytic conversion in which the process gas is subject to subsequent conversion

reactions in such a way as to obtain the almost complete conversion of the H2S into

sulphur.

The catalytic conversion reaction is exothermal so greater efficiency can be

obtained operating at lower temperature, compatibly with the dew point, of the

sulphur produced in the reacted gas.

Any condensation of the sulphur on the catalyst would drastically limit the activity

of the catalyst itself; in this case it would be lost most of the contact surface that

would be blocked by the sulphur.

Each conversion stage, whether thermal of catalytic, is followed by a cooling stage

of the process gas to condense the sulphur produced which is separated from the gas

and recovered in liquid form; the elimination of the sulphur produced during previous the conversion stage obviously maximizes conversion in the subsequent

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In any case, other treatment plants (Tail Gas Clean Up) are available for installation

downstream of the CLAUS units which are able to increase the efficiency of global

recovery of the supplied sulphur up to 99.9%; the cost of these plants can vary from

30% to more than 100% of the cost of the CLAUS unit, depending on the efficiency

requested.

The potential of the CLAUS units can vary from less than 1 t/d up to 1500 and more

t/d per line; in Italy units are installed with a potential of between 1 and 300 t/d per

line; the largest units are installed in France and in Canada downstream of the large

natural gas desulphurization units.

The sulphur produced in the CLAUS units is of excellent quality with purity of up

to 99.95% in weight; the sulphur is generally stored in liquid form and transported

in truck-tanks if the destination is within a radius of a few hundred kilometres from

the site of production; however, it is possible to solidify the liquid sulphur for

shipment in flakes or capsules, both in bulk and in sacks.The sulphur produced by the CLAUS units is an excellent raw material for the

production of sulphuric acid and for agricultural and pharmaceutical uses.

The tail gas leaving the CLAUS plants is incinerated by means of combustion at a

temperature of 650-900°C to obtain full oxidation of the sulphide compounds; the

combustion fumes are discharged into the atmosphere through chimneys of a

suitable height.

The treatment of gas with very low H2S contents (less than 20%) is possible taking

measures to maintain the required flame temperature; these measures range from

preheating the combustion air and the acid gas to be treated to the use of oxygen

enriched air.

Direct catalytic oxidation processes allow acid gases with H2S content of less than

5% to be treated.

Other engineering solutions such as supported of the acid gas allow the acid gases

to be treated in the entire concentration range of 5% to 20% of H2S; however very

low concentrations of H2S correspond to limited efficiency with recoveries that in

the worst case scenario may not even exceed 50%.

Supported combustion of acid gas allows the required flame temperature to be

reached, burning the fuel case with more equivalent air than that required in

stoichiometric combustion of the acid gas the subsequently injecting the acid gas in

the combustion products of the fuel gas; this technique is applicable in the case of

acid gas with a wide variation in H2S content in low concentration ranges (5 - 20%

in volume).

The CLAUS process is based on the combustion of H2S in sub-stoichiometric

conditions in order to maximize the transformation of H2S into sulphur.

The conversion of the H2S takes place in two distinct stages: the first is essentiallythermal and takes place during the combustion of the H2S, the second is catalytic

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and takes place in two or three serial reactors with cooling of the gas between each

reactor and removal of the sulphur produced in the previous stage.

The main reactions taking place during thermal conversion of H2S are:

H2S ==== H2 + ½ S2 - 905 Kcal/Nm3 H2S

H2S + ½O2 ==== H20 + ½ S2 + 1674 Kcal/Nm3 H2S

H2S + 1½O2 ==== H2O + SO2 + 5531 Kcal/Nm3 H2S

During thermal conversion the hydrocarbons present in the acid react as follows:

CH4 + 1½O2 === CO + 2 H2O + 5538 Kcal/Nm3 CH4

C2H6 + 2½O2 === 2CO + 3 H2O + 9190 Kcal/Nm3 C2H6



The reactions of the other components of the acid gas are:

H2 + ½O2 === H2O + 2579 Kcal/Nm3 H2

2NH3 + 1½O2 === N2 + 3H2O + 3379 Kcal/Nm3 NH3

CO2 + H2S === CO2 + H2O - 321 Kcal/Nm3 H2S

CO2 + 2H2S === CS2 + 2H2O - 359 Kcal/Nm3 H2S

The main reaction which takes place during catalytic conversion is:

2H2S + SO2 === 2H2O + 3S + 557 Kcal/Nm3 H2S

Catalytic conversion takes place in a temperature range of 220 to 260°C for the first

converter, 200 to 240°C for the second and 180 to 220°C for the third converter.

In the first converter, if the reaction temperature exceeds 250°C, hydrolysis

reactions of CO2 and CS2 also take place with the formation of H2S; the hydrolysis

of CO2 and CS2 is favoured by the high temperature.

Note that all the sulphur produced in the thermal and catalytic conversion stages is

present in gaseous form as S2-S4-S6 and S8 according to equilibriums which depend

on the temperature; the sulphur is liquefied as S1 in the condensers located downstream of the single reaction stages.

The conversion reactions are accelerated on synthetic alumina based catalysts; in

the past bauxite was widely used but its performance from the point of view of load

losses was less constant due to the load losses given the inconstant particle size of

the natural product.

Catalysts with well defined grain size and characteristics are available on the

market: from typical alumina catalysts specifically for the CLAUS reaction those

with activators to favour hydrolysis of CO2 to CS2, to those most suitable for

working in slightly oxidizing atmospheres.

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Catalysts for Claus units are manufactured by European companies such as Rhone

Poulenc and US companies such as Kaiser and ALCOA.

Despite the differences claimed by the manufacturers the performance of

homogeneous types of catalysts are fairly similar and in any case are not decisive inappreciably modifying the result of global conversion.

4.3.2 Equipment

Thermal conversion takes place by means of combustion of H2S with air in special

burners, studied ad hoc, reaching temperatures in the reaction products which

depend on the composition of the acid gas and the presence or absence of products

such as NH3 or HCN; the development of precise residence times of the products of

combustion is essential to achieve good levels of conversion.

The burners are generally of a combined type, able to operate with acid gas duringnormal operations and with fuel gas during plant start-up and shut-down.

In the case of the presence in the acid gas of ammonia, HCN, mercaptans, large

quantities of hydrocarbons or in case for high potentials (more than 60 - 70 t/d of

sulphur) it is necessary to have residence times of the combustion products at the

flame temperature in a refractory muffle to ensure complete combustion of all the

components of the acid gas, an essential condition if the plant is to operate

correctly.

If the acid gas is NH3, HCN and mercaptan free and contains limited quantities of

hydrocarbons combustion is possible in a submerged or fire tube combustion

chamber, thus obtaining the required residence time of the flame at a lower

temperature.

The submerged or fire tube combustion chamber allows a more compact and

economical plant to be used with shorter start-up and shut-down times even of the

result of the global conversion into sulphur is slightly less than plants with

refractory muffle.

The various systems used to heat the gas sent to catalytic reactors are worthy of

mention.

The most efficient systems in terms of management and of the result of the

conversion into sulphur are the indirect heating types (with steam, hot oil, heating

elements): however, it is not always possible to have fluids at the required heating

temperature (preferably 300°C and more for the first reactor) and therefore direct

heating systems have to be used with hot process gas (reheat gas) taken from

strategic positions downstream of the thermal conversion equipment (see 4.3.4.1) or

burning fuel gas or part of the acid gas in inline heaters (see 4.3.4.2).

Management of the direct heating systems, above all inline heaters, makes the

investment higher and plant management fairly complex; the use of inline heaters

should therefore be avoided when alternative systems can be used.

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Another possible way of heating the gas at the reactors is to use indirect systems

which use the hot gases leaving the first reactor to heat the gas entering the

following reactors: these systems (gas/gas exchangers) are fairly cheap even if they

lengthen the plant start-up and shut-down times, due to the high inertia during the

reactor heating stages.

Incineration of tail gas

The tail gas leaving the CLAUS units is incinerated before being discharged into

the atmosphere; the most commonly adopted practice is to have a thermal

incinerator which operates at high temperature (650-900°C) and which generates

fumes with a residual oxygen content of 1.5% - 2% volume.

Almost total oxidation of the H2S still contained in the tail gas to a residual content

of even less than 10 ppm is possible with high flame temperatures.

Alternatively, catalytic incinerators which operate with striking temperature of around 300°C and which have much lower fuel gas consumption can be used; the

management of this equipment is more expensive and more complex especially in

the case of deviation from normal the normal operating parameters of the CLAUS

unit; other negative factors of catalytic incinerators are the additional load losses

and the need to operate with close control of the excess of combustion oxygen, with

ensuing risk of forming noticeable quantities of SO3 in the case of great excess and

incomplete combustion of the H2S in the case of scarcity.

The use of catalytic incinerators also rules out the possibility of direct incineration

of part of the acid gas as sometimes required for thermal incinerators.

The catalysts used in the past for catalytic incineration were based on bauxite (the

same used for the CLAUS reaction): nowadays there are specific catalysts with

special activators; these catalysts have a high activity and more suitable mechanical

characteristics for the service.

Heating stages

A final aspect to be taken into consideration in order to fully understand how a

CLAUS unit works regards the heating of the unit before it begins operating with

acid gas and after the shut-down of the acid gas.

Indeed, before introducing the acid gas and start producing sulphur, all theequipment which comes into contact with process gas must be heated to a

temperature of more than 120°C; i.e. the temperature at sulphur solidifies.

To this end fuel gas is burned and the hot products of the combustion are circulated

through the unit until they reach the necessary heat level; the fuel gas is burned in

the same burner by the acid gas is a special nozzle.

Immediately after the gas shut-down all the sulphur in the unit (as liquid absorbed

by the catalyst or accumulated at some points of the unit) has to be eliminated; this

is done by burning the fuel gas in the combined burner and evaporating the sulphur

with the hot products of the combustion.

This is the most delicate stage of operations using a CLAUS unit; indeed, it isnecessary to have stoichiometric combustion of the fuel gas to avoid the formation

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of gas black in the case of scarcity of oxygen as well as uncontrolled combustion of

sulphur present in the unit in the case of excess of oxygen in the products of

combustion which pass through all the equipment of the unit.

In the case of units with muffle, a further complication during heating operations isthe need to send a cooling fluid (inert gas or steam) to the products of combustion to

lower the flame temperature to values which are compatible with the refractory

materials used.

4.3.3 Treatment and use of the sulphur produced

The sulphur produced in the various condensers located downstream of the thermal

and catalytic conversion is sent to an underground cement tank and from there it is

transferred to tanker trucks, to other storage tanks or for solidification.

The modern plants are fitted with degassing sections for the sulphur produced toeliminate the dissolved H2S, which can be dangerous especially during liquid

transport; the H2S contained in the sulphur tends to degas during transport, creating

a poisonous atmosphere which is potentially explosive in the vapour phase above

the level of the liquid in the tankers.

Sulphur degassing processes are able to reduce the H2S to 10 - 15 ppm in weight

and can be either catalytic (SNEA – EXXON processes) or mechanical (SHELL –

AMOCO processes); however, in all cases the H2S is removed in an air current

which is sent to the incinerator along with the tail gas for oxidation of the H2S

content.

In the CLAUS units the equipment containing liquid sulphur is heated using steam

powered coils while the pipes and valves are lined with vapour; even if electrical

heating is theoretically acceptable, it is rarely used, above all in Europe, in liquid

sulphur lines.

The steam generally used to heat the lines and equipment containing liquid sulphur

must have a pressure of 2.5 - 5 eff. bar; higher pressures are not acceptable because

they can generate high temperatures of the liquid sulphur such as to result in very

high viscosity of sulphur with ensuing difficulty in its transport, as shown in figure

4.3.1.; it is generally preferred to use saturated steam at a pressure of 2.5 – 3.5 eff. bar to heat the sulphur.

4.3.4 Basic layout of a plant

It has already been mentioned that CLAUS units can be "straight through" and

"split flow" depending on the H2S content in the acid gas.

Figures 4.3.2 and 4.3.4 show the typical layouts of the two processes.

Refer to figure 6.6.2.

The acid gas arriving from the regeneration section of the sweetening unit, with anH2S concentration of more than 60% in volume, is sent to a burner positioned in

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The "split flow" process is applied to acid gas with an H2S content of between

approx 20 and 60%; part of the acid gas is supplied to the muffle burner with all the

air needed for the conversion of all the H2S present in the acid gas into sulphur.

The products of the combustion, which must have a minimum flame temperature of 1000°C to avoid problems of instability, after recovering the heat in the boiler, are

sent to the first condenser for condensation of the sulphur produced during the

thermal reaction; in cases of the treatment of acid gas with low H2S concentration

(20 - 40 % vol.), the condenser is not needed because the formation of sulphur in

the thermal reaction is very limited or even zero.

The remaining acid gas is mixed with the process gas arriving from the boiler or

from the first condenser and the mixture is sent to the first reactor; from this point

on the process is the same as that described for the "once through" case; the

differences are simply due to the different thermal level of the gas entering the

catalytic reactors needed to maximize the conversion; the thermal level depends onthe concentration of H2S concentration of the acid gas and on the presence or

absence of impurities in the acid gas.

However, an acid gas with a high content of C2+ hydrocarbons cannot generally be

treated in a "split flow" plant because with this lay-out total combustion of the

hydrocarbons is not possible with the risk of coking on the catalytic reactors and the

ensuing formation of plugs which prevent the correct operations of the plant.

One variant of the "split-flow" lay-out is to send the acid gas in by-pass to the

muffle downstream of the combustion; in this case it is possible to partly mitigate

the problem of the presence of hydrocarbons in the acid gas: this stratagem can only

be applied in the case of acid gas which can reach a final adiabatic temperature

downstream of combustion of not less than 950 - 1000°C.

The combustion temperature can be increased by re-heating the combustion air and

the acid gas supplied to the burner.

This chapter briefly describes the most commonly used lay-outs for "straight

through" processes only.

"Split-flow" processes with all variants needed to treat acid gas with a lowconcentration of H2S, are not considered of interest for the purposes of this report.

The wider use of package plants allows grouping together in a single apparatus the

acid gas combustion chamber (fire tube), the heat recovery system from the

combustion fumes and sulphur condensers with relative separation chambers; the

catalytic reactors are grouped together in a single shell.

This layout is applicable in its entirety up to a potential of 50 - 60 t/d of sulphur,

provided that the acid gas does not contain NH3 or HCN or large quantities of

hydrocarbons.

For higher potential it is necessary to have a combustion chamber with relative heatrecovery system separate from the sulphur condensers.

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The lay-out cannot be applied for potential of approx 100 t/d and a combustion

chamber separate from the fume heat recovery system.

Figure 4.3.5, which shows a package plant with direct heating of the gas at the firstcatalytic reactor and indirect heating of the gas at the second reactor; the minimum

flexibility of a plant of this type is around 4 to 1; note that in the case of direct

heating of the gas at both catalytic reactors, the minimum flexibility would drop to

3 to 1.

The acid gas from the sweetening plant’s regeneration sector passes into the V-1

separator where it releases any condensation which are recycled to the absorption

unit (through gravity, with a pump or blowing with inert material or steam), and

then enters the combined burner assembled on the front of the recovery boiler; the

rate of acid gas is regulated automatically; the entire gas load arriving from the

sweetening section is sent to the CLAUS plant.

The combustion air is supplied by the K-1 blower – centrifugal or volumetric

depending on the potential. The air rate is controlled in two ways: approximate

regulation depending on the flow rate of the acid gas while fine regulation is based

on the analysis of H2S/SO2 in the tail gas. Both regulations have a dedicated control

valve.

The products of combustion remain for the time needed to complete the reaction in

the combustion chamber submerged in boiler B-1 and then cooled to around 600 -

700°C in a second passage through the boiler.

Boiler B-1 has 5 passages and produces steam at a pressure of 1.5 - 5 eff bar; the

supply water flow rate is controlled automatically by a level meter; a pressure metre

regulates the rate of steam produced.

A small part of the hot gas at 600 - 700°C (reheat gas) leaving the second passage

of the boiler is used to heat the gas sent to the fires catalytic reactor (about 3-5% of

the total flow rate), while the rest of the hot gas is sent to the third passage of the

boiler which acts as first sulphur condenser; most of the sulphur contained in the

process gas in gaseous phase is condensed and separated by the gas in a preliminary

separation chamber fitted on the front of the boiler.

The liquid sulphur produced is sent to the TK-1 storage trench using a hydraulic V-

2 Liveley seal; the cooled gas, at a temperature of 160 - 200°C is mixed with the hot

gas leaving the first passage of the boiler so as to obtain the desired catalyst

temperature, which is generally 220 - 260°C, and is then introduced into the first

catalytic reactor R-1.

Catalytic conversion, taking place on alumina, results in an increase in temperature

of 40-80°C depending on the concentration of acid gas; the converted hot gas is sent

to the E-1 gas/gas exchanger where it is cooled, heating the supply gas to the second

reactor; the gas leaving the E-1 enters the fourth passage from boiler B-1 (second condenser) where most of the sulphur contained in the process gas is condensed;

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after separation in a chamber on the back of boiler B-1, the sulphur is sent to the

TH-1 storage trench using the V-3 Liveley seal .

The gas leaving the fourth passage of the B-1 boiler at a temperature of 150 - 180°C

is heated in E-1 to catalyzing temperature, generally 200 - 240°C and is thensupplied to the second reactor R-2.

Catalytic conversion generates an increase in temperature of 20 - 40°C; the

converted gas enters the final passage of the boiler B-1 (third condenser) where

most of the sulphur contained in the gas is condensed. After separation of the

sulphur in a chamber on the back of boiler B-1, the sulphur is sent to the TK-1

storage trench using a V-4 Liveley seal.

The tail gas leaving the final passage of boiler B-1 is sent for thermal or catalytic

incineration; the incineration reaction is supported by the combustion of fuel gas.

The sulphur produced, collected in the TK-1 trench, is loaded onto tanker trucks or

sent for solidification using the upright pump P-1 assembled above the trench; the

trench operates with a slight vacuum, preferably directly connected to the

combustion chamber of the thermal incinerator to avoid leakage of sulphur vapours

of the H2S dissolved in it.

Before sending the acid gas to the plant the entire plant has to be heated to set

minimum temperatures: for this purpose the fuel gas is burnt in the combined

burner on the front of boiler B-1. The products of hot combustion ensure the

necessary heating and are then discharged into the atmosphere through the

incinerator.

In plants with fire tube it is not necessary to lower the flame temperature of the fuel

gas during the unit heating stage; the heat exchange between the reaction tube and

water in the boiler is sufficient to obtain a temperature of the metal which is not

dangerous.

The combined burner has a pilot light with electric spark ignition; some plants have

graphite ignitions for the direct ignition of the fuel gas burner; in all cases one or

two photoelectric cells monitor the presence of the flame.

A logic (relay, solid state or with PLC) governs the combined burner ignition;ignition is based on semi-automatic sequences.

The plant is protected by a safety system which intervenes in the absence of the

flame, low boiler level, high acid gas separator level, high pressure on the boiler

front, low rate of acid gas and combustion air, low (and sometimes high) pressure of

fuel gas to the pilot light when envisaged and no flame at the incinerator; of the

safety system intervenes the supply of acid gas, combustion air and main fuel gas

(and pilot light when envisaged) will be interrupted.

The air blowers have an independent automatic protection system and an anti-

pumping system for centrifugal machines or, for volumetric machines, machines todischarge the excess air into the atmosphere.

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Figure 4.3.6 shows a plant with a single zone muffle with three recovery boilers and

indirect heating of the gas at the reactors; the proposed lay-out can treat gas with an

ammonia content of less than 1% in vol. with a flexibility of 5 to 1.

The plant description is similar to the one in the previous chapter; the difference

consists in the development of the heat reaction in the H-1 refractory muffle, able to

develop the residence time needed to complete the combustion of the acid gas at a

very high temperature and therefore with greater H2S conversion into sulphur than

that obtainable with combustion in “fire tube”.

The steam can be produced in the B-1 recovery boiler at a pressure of more than 60

eff. bar; a production of 10-20 eff. bar is generally required and the process gas

leaves the boiler at a temperature of 280 - 350°C.

The first and second sulphur condenser are installed in a second boiler which produces steam at 3-5 eff. bar, while the third condenser is a low pressure steam

generator able to cool the tail gas to 135°C, thus allowing an increase in conversion

efficiency of 0.8 – 1.5% compared to a steam production boiler at 3 - 5 eff. bar.

The gas to the reactors is heated using easy operating electric heaters with continual

thyristor regulation; in Italy electric heaters have been used up to a potential of 50

t/d of sulphur.

During the plant heating stage before going into operation, the flame temperature

reached in the stoichiometric combustion of the fuel gas would be such as to

damage the refractory materials used in the muffle: to limit its value, quench steam

is added to the combustion products in quantities which lower the flame

temperature to 1300 - 1350°C.

The heating times of a plant with muffle are more than double those of a plant with

fire tube, because the increase in temperature of the muffle’s refractory materials

has to be gradual.

The temperature of the muffle, between 1000 and 1400°C, is measured using

special Platinum-Rhodium thermo-couples or optical pyrometers which are more

reliable and last longer; it is not necessary to measure the flame temperature in firetube plants.

If acid gas containing ammonia is to be treated (typical example is the gas from the

Sour Water Stripper in refineries) two-zone muffles have to be used, adopting the

split flow lay-out principle.

See figure 4.3.7.

The acid gas containing ammonia is supplied to the burner along with all the air

needed for the Claus reaction to take place and the part of the acid gas not

containing ammonia needed to reach a flame temperature of approx. 1400°C; the

rest of the acid gas not containing ammonia is sent to the second zone of the muffle

where it reacts with the reaction products arriving from the first zone.

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The effect of the high temperature and the presence of a high concentration of SO2

in the first zone, along with an adequate residence time of the fumes, results in the

total destruction of the ammonia.

Incomplete combustion of ammonia could cause the precipitation of ammonium

salts in the colder parts of the plant with the formation of plugs which are difficultto remove.

In units with muffles, in the case of emergency shut-down, the unit must be fluxed

for a few minutes with inert gas to prevent damage caused by thermal radiation

from the hot refractory materials.

The use of a two-zone muffles allows very high, almost unlimited, turn downs,

designed with a single burner supplied with fuel gas needed to produce fumes in

suitable quantities to guarantee the minimum mass velocities to avoid the formation

of sulphur mists; the combustion air of the acid gas is supplied along with the fuel

gas in the first zone of the muffle.The acid gas is instead applied to the second zone where it reacts with the air

contained in the fumes of the first zone.

The gas at the reactors is heated using an indirect method using external heating

fluids.

Of course, for operations with a load of more than 20% higher than the nominal

load, the acid gas is sent only to the first zone which operates in a traditional way

while the second zone is not supplied.

The two zone muffle technique with very high turn downs cannot however be

applied for acid gas containing NH3 in a concentration of more than 1% Vol. unless

controls are made on a case by case basis.

Another way of heating the gas at the reactors for muffle plants and fire tube plants

consists in installing inline heaters; that is, combustion chambers with burner

supplied with fuel gas or acid gas installed in a container where the gas is heated to

the desired temperature, mixing it with the combustion products; the inline heaters

allow flexibilities of 8 - 10 to 1 but are more expensive than the other systems

described and more difficult to manage; moreover, only in the case of heaters

supplied with fuel gas, they can cause operating problems such as poisoning of the

catalysts through plugging with lampblack in the case of combustion with scarcity

of oxygen or sulphation in the case of incomplete combustion and the presence of oxygen in the products of combustion.

The use of inline heaters should therefore be limited to cases where alternatives are

not viable.


The main construction material in CLAUS units is carbon steel; the muffles, boilers,

condensers, reactors and heaters are made of carbon steel; a corrosion reinforcement

of 3 mm is generally envisaged for all parts in contact with process fluids and this is

sufficient for a long industrial life of the plant.

All the parts in contact with the liquid sulphur are made of carbon steel, includingthe heating coils immerged in the storage tank.

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V I S C O S I T Y - c e n t i p o i s e

TEMPERATURE - °C

FIGURE 4.3.1 – VISCOSITY OF LIQUID SULPHUR

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FIGURE 4.3.2 – CLAUS: STRAIGHT THROUGH LAY-OUT – THREE REACTORS

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V A

P O U R P R E S S U R E – a t a

TEMPERATURE - °C

FIGURE 4.3.3 – VAPOUR PRESSURE OF LIQUID SULPHUR AS S1

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GURE 4.3.6 – CLAUS – LAY-OUT WITH SINGLE ZONE MUFFLE

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FIGURE 4.3.7 – CLAUS – LAY-OUT WITH TWO-ZONE MUFFLE

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4.3.7 Further developments of the process

Processes below dew-point. This class of processes extends the Claus reaction

below the dew point of the sulphur. Sulphur recovery it increased by the more

favourable equilibrium reaction. One of these processes is Amoco’s Cold Bed Adsorption (CBA), comprising a pair of catalytic reactors operating at a

temperature below the sulphur dew point. Considering that the sulphur condenses

on the catalyst, regeneration is necessary to avoid a reduction in the catalytic

activity. The catalyst is restored by injecting hot tail gas in the reactor to vaporize

the condensate (this sulphur is then recovered in a condenser).

Two catalytic reactors in parallel are required, one in active mode and the other in

regeneration mode. The sulphur recovery with this system is 98.5-99%.

Along these lines other processes have been developed with three or four Claus

reactors and CBA lay-out, or completing eliminating the incoming water vapour

(the water reduces the catalytic activity as it adsorbs).

Another type of plant belonging to this class is the Claus processes. In these plants

the outgoing gas of a normal Claus process is contacted in a fixed bed with

polyethylene glycol at around 120°C. Sulphur dioxide and hydrogen sulphide are

absorbed by the organic solvent and react in a reactor with a sodium organic salt

based catalyst. The liquid glycol and glycol are practically insoluble; the two stages

are then separated.

The sodium salt is however low-boiling and its evaporation provokes a reaction

with the anhydride the precipitation of insoluble sulphate and the formation of

organic acid. The salt tends to accumulate in the column; this means that the plant

has to be shut-down periodically for cleaning operations. In this way sulphur

recovery reaches up to 99.7%.

SuperClaus process. This is a selective oxidation process comprising a normal two-

stage Claus unit followed by a SuperClaus catalytic stage. However, in this plan the

Claus unit operates in a slightly different way to the standard one. In fact, the

supply air is injected to have an H2S/SO2 ratio of more than 2 and to have a total

consumption of anhydride in the two stages of the process. The acid still has a

concentration of 09%.

The gas leaving the second stage is heated to around 300°C and mixed with a pre-

heated current of air. The supply obtained in this way is sent to the SuperClaus

reactor where a patented catalyst is used to oxidize the hydrogen sulphide. Anoxidation level of 85% can be reached. Total sulphur recovery is 99%.

Other processes. Besides the above, other processes have been patented which

allow sulphur purification and recovery of more than 99.99%. These processes are

often very expensive (even more than double the costs of a traditional dual stage

Claus unit) and for specialist use with cases limited to contingent cases. For further

information see the report OKIOC - PROCESS STUDY (Sulphur recovery and

disposal).

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5. CRYOGENIC PROCESSES TO REDUCE THE CARBON DIOXIDE

The cryogenic processes are based on the use of an additive such as oil or butane or

LNG for the separation by means of distillation of CH4 and CO2 from the other

hydrocarbons; the separation columns operate at temperatures of between -50°C and -90°C with pressures of around 40 - 80 abs. bar.

5.1 Ryan-Holmes process

Figure 5.1.1 shows the base lay-out of an acid gas separation plant according to the

Ryan-Holmes process applied to EOR.

The crude gas to be treated, containing H2S and up to even more than 90% of CO2,

is compressed, dehydrated, cooled and sent to the head of the CH4-CO2 separation

column demethanizer. In this case the additive is the recovered LNG.

The gas leaving the head of the column can be sent to the gas pipeline, after the heathas been recovered.

The LNG - CO2 solution and the heavier hydrocarbons leaving the bottom of the

demethanizer is sent to the CO2, separation column at a higher temperature and

lower pressure; the additive LNG which allows separation of the CO2 from the

hydrocarbons and the H2S is supplied also to this column.

The CO2 leaving the head contains up to around 100 ppm of H2S depending on the

concentration of H2S in the crude gas to be treated; the solution leaving the bottom

is sent for recovery of the additive by means of adjustment.

The additive is recycled to the columns while the light hydrocarbons leaving the

head of the last column are sent to a sweetening system with amine for final

conditioning.

The acid gas separated in the amine regenerator is sent to a Claus unit for sulphur

recovery.

The diagram does not indicate the various heat recoveries that are usually activated

for the energy optimization of the process.

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FIGURE 5.1.1 – CRYOGENIC - RYAN HOLMES PROCESSES

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5.2 CRYOFRAC Process

5.2.1 Process description

The CRYOFRAC process was developed by SnamProgetti to allow natural gas witha high CO2 content (more than 30 - 35% vol.) to be produced.

The procedure is based in the absorption at low temperature with a selective

physical solvent (Cryosol) for acid gas.

The typical working range is -30 – -50°C at a pressure of around 40 bar.

At present the process has been intensively and positively tested in a pilot plant but

operative industrial plants are still not available.

Natural gas with a high CO2 content can be purified using a cryogenic separation process; however, the thermodynamic behaviour of methane-CO2-H2S mixtures at

low temperature is complicated by the fact that the CO2 can solidify and the H2S can

involve immiscibility problems.

In the phase diagram of a methane- CO2 system (Figure 5.2.1) two zones can be

distinguished:

a) a zone above 49 abs. bar where the CO2 cannot crystallize as at this pressure,

greater than the critical pressure of the methane (46.4 abs. bar), it is not

possible to obtain pure methane.

b) a zone below 49 abs. bar where the formation of a solid phase is possible in a

wide concentration range.

This behaviour is also complicated by at least another two phenomena:

- Nitrogen, a compound frequently present in natural gas, widens the field of

composition in which the CO2 can solidify.

- The methane-H2S system has a wide field of immiscibility.

From the above it is clear that cryogenic distillation provides natural gas which is

almost completely free of acid gas but with a plant with great regulation difficulties

to avoid crystallization risks.

The solvent used in the CRYOFRAC process, called Cryosol, is a mixture of polar

organic compounds with patented compositions and property and of which it is

possible to give only qualitative information.

The main characteristics of Cryosol are:

- Stability and non-corrosivity under operating conditions

- Low fusion point (lower than -100°C)

- Low vapour pressure under operating conditions.

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In the natural gas-CO2 system, the solvent has good selectivity characteristics

during the acid gas absorption stage without causing problems of crystallization of

CO2 and immiscibility of H2S.

From a purely qualitative point of view, it can be said that, conditions being equal,Cryosol absorbs 400% of the CO2 absorbed by the methanol and 10% less of

methane.

The possibility of having a high load of solvent allows a flat vaporization curve in

wide range of enthalpy so most of the cooling is obtained at a (low) almost constant

temperature; see figure 5.2.2

The regeneration of the solvent through flash represents one of the unique aspects

of the process with the possibility of recovering the cooling.

The low thermal level of the cooling also allows the surface to be reduced as well as

the cost of the exchangers.

5.2.2 Lay-outs used

The CRYOFRAC process uses the typical lay-out of the physical absorption

processes; the acid gas is absorbed at low temperature while regeneration takes

place in flash stages in series, with the final flash in vacuum.

If it is also necessary to remove the H2S to the levels required by the natural gas

specifications, the solution should be regenerated also by means of distillation.

Two diagrams are shown below of the application of the process on gas containing

H2S and CO2 and CO2 alone.

Figure 5.2.3 shows the most common case of application, assuming the treatment of

a natural gas containing H2S and CO2.

Before treatment with the CRYOFRAC, the gas must be dehydrated and separated

from the heavy hydrocarbons which should be almost completely removed from the

solution.

The first acid gas absorbed in the CRYOFRAC process is H2S, given its selectivity

compared to CO2; the current without H2S is then cooled, part of the CO2 iscondensed while the remaining part is absorbed in a second absorption column.

The two acid gases are absorbed with a completely regenerated solution for the CO2

and a partially regenerated solution; that is, still containing CO2 and methane for the

H2S.

The solvent is regenerated by means of multi-stage flashes the last of which is

generally under vacuum while the current which absorbs the H2S requires further

regeneration by means of distillation. The H2S rich current (distillation head) can be

sent to a sulphur recovery plant.

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As the absorption of CO2 takes place at a very low temperature (-30 – -50°C),

cooling recovery has to be optimized. This takes place recovering the cooling

developed by the flash of the CO2 rich solution during cooling: if the content of CO2

in the gas is more than 35%, the cooling produced is such as to make the process

self-cooling and consumption is consequently reduced. A small scale cooling cyclewill be sufficient to cool the plant.

The following product specifications can be obtained with this type of process:

Scrubbed gas

H2S 1-4 ppm volume.

CO2 50 ppm vol. - 3% volume.

Recovered CO2:

H2S 1 - 100 ppm volume.

The process described can be simplified if the gas to be treated does not contain

H2S, or if the H2S in the scrubbed gas is not very severe. In this case absorption can

take place in a single column and final distillation to eliminate the H2S from the

solvent can be avoided.

Figure 5.2.4. shows a process in which only CO2 is removed from the already

dehydrated and degasolinized natural gas containing a CO2 concentration of 35%

vol.

The gas to be treated is cooled in the crude gas cooler E-2 to approx. -60°C by a

current of solvent arriving from the second V-2 flash tank and enters the C-1 plate

absorption column which operates at around 40 abs. bar.

The Cryosol solution leaving the bottom of the column is regenerated in four flash

stages, one of which under vacuum.

The gas freed from the first flash tank V-1 at approx. 10 abs. bar is still rich in

methane and is then recycled at the absorption column by the K-1 compressor.

The gas freed during the next three flashes are rich in CO2, and are the coolants

obtained from the expansion of the freed CO2 which are then used to make the process self-cooling.

Indeed, the solvent leaving the second flash tank V-2 transfers coolants to the gas

scrubbed in the E-4 cooling system and to the already mentioned coolant gas to be

treated E-2.

The gas leaving the third flash tank V-3, which operates at atmospheric pressure,

cools the recycled methane at the absorption column C-1, in the coolant E-1.

From the vacuum flash tank V-4, the solvent regenerated at a temperature of -50°C

is pumped by the P-1 recirculation pump and, mixing with the scrubbed gas leaving

the head of the absorption column C-1, scrubs the gas of the last traces of CO2. The

pressure in the vacuum flash tank V-4 is maintained by the vacuum pump K-2.

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The sweet gas/solution mixture is separated in the separator V-5, from whose

bottom the solvent s pumped by pump P-2 to the head of the absorption column C-

1.

If necessary a part of the solvent is conveyed to a reclaimer to avoid accumulationsof water or light hydrocarbons still present in the crude gas to be treated.

The figure shows the main controls carried out on the various currents; these

controls are usual also in other physical absorption processes.

5.2.3 Design criteria of the main equipment

The design criteria of the equipment are similar to those adopted for other physical

processes.

The absorber is preferably a plate column, which must however be preceded byunits upstream able to remove the humidity of the gas to be treated and the fractions

of heavy hydrocarbons.

The various flash tanks must be dimensioned with residence times which censure

complete separation of the gas from the liquid.


From an engineering point of view, Cryosol being a non-corrosive solvent, the

materials are chosen on the basis of the design temperatures of the various pieces of

equipment; i.e.:

- carbon steel killed for temperatures to -50°C, as in the case of the hotter flash

separators.

- alloys at 3.5% Ni for temperatures to -100°C, for example for the absorption

column, the circulation pumps and the colder separators.

- alloys at 9% Ni and stainless steel for lower temperatures.

5.2.5 Other applications of the process

An interesting application of CRYOFRAC consists in using it in Enhanced Oil

Recovery (EOR) processes where CO2 is used a miscible agent for the recovery of

crudes from fields nearing depletion. Indeed, CO2, modifies the flow characteristics

of the oil, facilitating production.

In this application the CO2 has to be separated from the methane and the same CO2

has to be reinjected into the reservoir at pressure of between 60 and 200 bar.

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If the CO2 content in the methane is high, and in some cases it can reach 80% vol.

of CO2, it would be advisable to combine the CRYOFRAC process with low

temperature distillation upstream to separate a certain quantity of CO2, which will

separate in liquid state at the bottom of the fractioning column while at the head the

methane and the rest of the CO2 to be supplied to the CRYOFRAC would beobtained at the head.

Also in this case the global energy efficiency of the process (cryogenic distillation +

CRYOFRAC) is interesting as it is self-cooling. Indeed, the liquid CO2 from the

bottom of the distillation column, expanding, provides coolant to the head

condenser and is then compressed before being reinjected into the reservoir.

Figure 6.5.5. shows the percentage of CO2 typically removed from a natural gas

current by means of cryogenic distillation.

However, the economic feasibility of this installation has to be carefully assessed in

relation to the price of oil.

Another application of CRYOFRAC could be to remove acid gas in the ammonia

synthesis gas or oxo-gas preparation line starting from carbon or oil fractions when

the acid gases are in a sufficiently high concentration.

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P R E S S U R E – a b s . b a r

TEMPERATURE - °C

FIGURE 5.2.1 – METHANE – CO2 SYSTEM

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T E M P E R A T U R E

ENTHALPY

FIGURE 5.2.2 – CRYOFRAC – CHARACTERISTICS OF THE SOLVENT

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FIGURE 5.2.3 – CRYOFRAC – ABSORPTION OF CO2 ALONE

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FIGURE 5.2.4 – CRYOFRAC – ABSORPTION OF H2S + CO2

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C O 2 R E M O V E D T H R

O U G H D I S T I L L A T I O N A T L

O W T

E M P E R A T U R E - % o f

t h e t o t a l

CONCENTRATION OF CO2 IN THE GAS TO BE TREATED - % vol.

FIGURE 5.2.5 – CRYOFRAC – EFFECT OF CO2 REMOVAL

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6 APPENDIX- PROCESSES NO LONGER USED

In the past processes were used on an industrial scale which have now become

obsolete for different reasons such as environmental protection, efficiency,

management or maintenance.This paragraph gives a brief overview of these processes.

6.1 Non-regenerative processes

These processes are only selective in terms of hydrogen sulphide, and can offer a

high level of purification. They were used with small quantities of pollutant and can

be divided into two categories: with solid adsorbent (iron and zinc oxides) or with

solutions (permanganates, dichromate and lead acetates). In this type of procedure

the operating pressure does not play a decisive role.

Solid adsorption plants usually had two reactors which were used in alternation: one

operating, one in maintenance with the removal of the degraded product and itssubstitution or in some cases, its regeneration. For example, with iron oxides the

sulphides formed were re-oxidized at high temperature with air.

Liquid absorption plants were used for small production units and were usually with

closed cycle.

These processes were gradually abandoned over time due to environmental

problems due to the disposal of the solid mass or the rich solution.

6.2 Chemical absorption with alkaline carbonates

These processes are based on similar mechanisms to those discussed for processes

with amine solutions. One of the most important at an industrial level was the

BENFIELD process belonging to Union Carbide.

The absorbing solution was a water solution, with K 2CO3 at 20-40% in weight,

DEA at 3% in weight and traces of V2O5. Carbonate was the absorbing solution and

amine the activator of the process. Vanadic anhydride was used as corrosion

inhibitor.

The reactions (all at equilibrium point) can be represented as

K 2CO3 + H2O + CO2 -» 2 KHCO3

K 2CO3 + H2S -> KHCO3 + KHS

The reactions are exothermic but their kinetics are favoured at high temperatures;

therefore, even if the conversion is depressed, it is preferable to work in this way.

The pressure favours the reaction kinetics.

DEA has the role of favouring absorption kinetics and of modifying the liquid-

vapour equilibrium.

During regeneration the operating temperature must be as low as possible, so as to

shift the equilibrium of the above-stated reactions to the left. Determining the

pressure logically establishes also the temperature of the regeneration column.

The reactions of the process are highly selective as regards acid gases.

The schemes allow for outgoing impurities of 20 ppm for CO2 and 1 ppm for H2S.

The heat required by the process is around half that required for a normal amine process.

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The lay-outs used are similar to those used in DEA and MDEA plants (classical,

split-flow, two-stage...).

The main problem of this type of plant is the danger factor of vanadic anhydride;

moreover, the major corrosion problems caused by the carbonates make the use of amine solutions more recommendable.

6.3 Absorption with MEA and MEA-DEG solutions

This type of plant (to which the Cupello plant belongs) has been abandoned due to

the excessive foaming problems of the solutions.

6.4 Other processes

A range of other processes have been used and then abandoned for the reasonsillustrated above. These plants belong to all types of process such as amine,

absorption with ammonia solutions, physical absorption, absorption on fluid bed

and oxidation processes in liquid phase.

Mention should be made of the molecular sieve processes to purify solely CO2, used

more recently but which have not given satisfactory results.

These types of plants are no longer designed but they can still be found in plant

already in construction.

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7 GLOSSARY

SWEETENING – Process aimed at removing acid gas such as CO2 and H2S from

gaseous or liquid currents.

APPROACH TO EQUILIBRIUM – Percentage ratio between the vapour pressure

of the acid gas in the absorbing solution and the partial pressure of the acid gas in

the gas to be treated.

ABSORPTION – Operation aimed at removing the acid gas from the crude gas to

be treated.

PICK-UP – Specific quantity of acid gas absorbed chemically or dissolved in the

absorption fluid; it is expressed in Kmol of Acid gas/Kmol of solution, in Nm3

GA/m3 Solution or in other suitable engineering units.

NET PICK-UP – Difference in pick-up between the lean solution and rich solution;

it expresses the actual absorbing capacity of the solution.

DEGRADATION OF THE SOLUTION – Irreversible transformation of one or

more components of the absorbing solution with the formation of secondary

reaction products.

DRIVING FORCE – Difference between the partial pressure of H2S (or CO2) in the

crude gas and the vapour pressure of H2S (or CO2) in the absorbing solution. The

driving force measures the tendency of the acid gas to be absorbed by the solution.

ACID GAS - H2S and CO2 to be removed or H2S and CO2 freed during the

regeneration stage of the regeneration procedures.

SWEET GAS - Gas produced by a sweetening procedure.

CRUDE GAS - Gas entering the sweetening unit.

TECHNICAL GAS - Gas of not high purity used industrially as inert, oxidizer,

reducer or for other uses; it is sold in cylinders or other pressurised containers.

Technical gases are oxygen, hydrogen, nitrogen, helium, argon, carbon dioxide and acetylene.

INCINERATOR – Equipment in which sulphide components contained in a

gaseous current are completely oxidized to SO2; it can be thermal and thus based on

oxidation at a temperature of 600 - 1000°C or catalytic with an operating

temperature of 300 - 450°C.

HYDROLYSIS – Reaction of CO2 and CS2 with H2O with the formation of CO2

and H2S; hydrolysis can take place in liquid and gaseous phase. In the gaseous

phase the hydrolysis must be carried out on specific catalysts which operate at

temperatures of between 100 and 400°C.

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SOLID BED – Absorbing mass, or molecular or activated carbon sieve used in the

sweetening process.

MAKE-UP – Replenishment; top up: refers to the absorbing fluid lost through

degradation, vapour pressure, entrainment, etc., and to the solution water of theabsorbing fluid lost through evaporation with the sweet gas or with the freed acid

gas.

In some cases the excess condensate has to be removed when the sweetening plant

is supplied with gas at high temperature and saturated with water vapour.

PERMEABLE MEMBRANES – These are permeable fibres able to operate at very

high pressures and able to separate, through permeation, some compounds of from

gaseous currents; the permeated gas (acid gas) is rendered at low pressure while the

non-permeated gas (sweet gas) is rendered at high pressure.

MUFFLE – Combustion chamber internally lined with refractory material where thesub-stoichiometric combustion of the acid gas takes place in Claus units; it operates

at high temperature (1000-1420°C) and can have one or two reaction zones

depending on the composition of the acid gas to be treated and on the flexibility of

the plant. In two-zone muffles there is an internal wall which separated the zones.

OPTICAL PYROMETER – Instrument used to measure temperature base on the

measurement of the intensity of electromagnetic intensity (or radiant energy); it is

used to measure high temperatures (1000-1600°C), especially in ovens which treat

extremely aggressive fluids.

PARTIAL PRESSURE – The product between total absolute pressure of a gas and

the molar fraction of a component.

ABSORPTION PROCEDURE – Based on the polar affinity between the solid

absorbing bed and the acid gas to be removed.

CHEMICAL ABSORPTION PROCEDURE – Based on the reaction of the

absorbing fluid with the acid gas to be removed.

CHEMICAL-PHYSICAL ABSORPTION PROCEDURE – Based on the use of

absorbing fluid able to absorb the acid gas through chemical reaction or physicalsolubility.

PHYSICAL ABSORPTION PROCEDURE – Based on the solubility of the

absorbing fluid with regard to the acid gas to be removed.

CLAUS PROCEDURE – Oxidation process in gaseous stage to transform H2S into

sulphur which is recovered as a liquid with high purity.

CRYOGENIC PROCEDURE – Based on fractioning the gas at low temperature.

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NON-REGENERATIVE PROCEDURE – Based on the reaction of the absorbing

fluid with the acid gas to be removed; the product formed must be expelled from the

system.

OXIDATION PROCEDURE – Based on the direct transformation of H2S intosulphur; it can operate in gaseous phase through sub-stoichiometric combustion of

H2S or in liquid phase using oxidation air of the absorbing solution with an oxygen

carrying agent; in the latter case it is also a regenerative procedure.

REGENERATIVE PROCEDURE – Procedure based on the continual regeneration

of the absorbing fluid with release of the absorbed acid gas or removal of the

produced sulphur.

SELECTIVE PROCEDURE – Able to preferentially absorb H2S instead of CO2.

SECONDARY REACTION PRODUCTS – These are sub-products which areinevitable formed by reaction of absorbent fluids with some gaseous components

present in the crude gas to be treated (degradation of the solution).

RECLAIMING – Operation to condition the degraded solution: it consists in

removing the secondary products of degradation with the re-use of the solution in

the acid gas absorption cycle.

CONVERSION YIELD – Expresses the percentage of H2S present in the crude gas

to be treated which has been transformed into sulphur.

THERMOSIPHON BOILER – This is based on the circulation of the absorbing

solution, pipe side or sleeve side, due to the difference in specific weight between

the inlet and outlet of the boiler; the circulation ratios are higher for upright boilers

than for horizontal ones; the mixed phase freed at the outlet of the boiler is

conveyed to the solution regeneration column.

KETTLE BOILER – This is based on the passage of all the solution to be

regenerated in the boiler; the vapour freed is supplied to the solution regeneration

column while the lean solution is conveyed to the column itself or directly to the

circulation pump.

ONCE THROUGH BOILER – This is based on the passage of all the solution to be

regenerated through the boiler; the mixed phase freed (solution and vapour) is

supplied to the bottom of the solution regeneration column.

REGENERATION – Operation aimed at regenerating the rich absorbing solution; it

is performed through flashes, stripping with transport gas, stripping with direct

vapour, indirect heating, etc., depending on the process.

TWO-STAGE ABSORPTION LAY-OUT – Envisages the use of solutions with

different regeneration degree, besides different temperature; the regenerator is two-

stage: the lean solution leaves the regenerator at the bottom and is sent to the head of the absorber.

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This lay-out allows savings of regeneration heat.

CLASSICAL ABSORPTION LAY-OUT – This envisages using a lean solution

supplied at the head of the absorber; the regenerator is single-stage.

SPLIT-FLOW ABSORPTION LAY-OUT – This envisages the use of lean solution

supplied at the head of the absorber at different temperature levels; the regenerator

is single-stage.

SULPHATION – Reaction of the transformation of a salt into sulphate; in Claus

catalysts sulphation provokes the transformation of alumina (A1203) into

aluminium sulphate through reaction of the catalyst with a gas containing SO2 and

oxygen or containing only several tens of ppm of SO3; a sulphate catalyst loses its

activity proportionately to the entity of the sulphation.

RICH SOLUTION – Solution produced by absorption and therefore rich in acid gas.

REGENERATED OR LEAN SOLUTION – Solution used for absorption and

therefore with little acid gas.

VAPOUR PRESSURE – This is the absolute pressure applied by a component of a

solution in set temperature and concentration conditions.

THYRISTOR – Electronic continuous modulator of energy flow; a typical

application of a thyristor is to continuously regulate the temperature of the outgoing

fluids to be heated using electrical resistances.

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8. BIBLIOGRAPHY

The name of the author, the name of the publication, the title of the article or

publication or notes to identify the contents are given for each bibliographic

reference.

1926 Speer

Gas Age Record

Ferrox-Koppers Process

1927 Gluud-Schoenfelder

Chem. & Met. Eng. (12)

Gluud Process

1931 Mueller

Gas U. Wasserfach (28)Fisher Electrolytic procedure

1933 Bottoms

US Patent 1.783.901

Using MEA

1934 Gollmar

Ind. Eng. Chem. (2)

Thylox-Koppers Procedure

1935 Baehr-Mengdehl

US Patent 1.990.217

Alkazid Process

1938 Baehr

Chem. Fabrik (11)

Catasulf Process

1939 Hutchinson

Us Patent 2.177.901

Using MEA-TEG

1945 Gollmar-Lowry-Wiley

Chemistry of Goal Utilization

Oxidation in liquid phase

1945 Shreve

The Chemical Process Industries/Edit. McGraw Hill

Phenolate procedure

1946 Pieters-Van Krevelen

Elsevier AmsterdamStaatsmijnen/Otto Procedure

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Page. 164 of 175

1948 Sands-Wainwright-Schmidt

Ind. Eng. Chem (4)

Soda Iron Process

1951 Kohl

Petrol Processing – January

Using MDEA

1952 Sherwood-Pigford

Absorption and Extraction - McGraw Hill

General interest

1952 Griffith

Industrial Engineering Chemistry (5)

Tames North Board Gas Process (Sulphite oxidation)

1953 Terres

Gas U. Wasserfach (9)

Polythionate solutions

1954 Benson-Field-Jimeson

Chemical Engineering Progress - July

CO2 Absorption with Hot Potassium Carbonate Solutions.

1954 Jegorov-Dimitriev-Sikov

Hoepli

Processes with arseniate for desulphuration

1955 Blohm - Riesenfeld

US Patent 2.712.901

Using DGA

1955 Polderman-Dillon-Steele

Oil Gas Journal - May - 16

Degradation of MEA in Natural Gas Treating Service

1956 Benson-Field-Haynes

Chemical Engineering Progress - 52

Improved Process for CO2 Absorption uses Hot Carbonate Solutions

1956 Polderman-Steele

Oil Gas Journal - July - 30

Degradation of DEA in Gas Treating Service

1957 Guntermamm-Schnurer


Oxidation towers

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1958 Davis-Mills-Rydem

Gas Council/GB

Methylene blue procedure

1958 Reid-Towsend Oil Gas Journal - October - 13

Oxidation with SO2 in TEG solution

1958 Palo

Petroleum Refinery - December

Alkaline solutions – foaming agents

1958 Liedholm

Shell Develop. C. - February

Shell K3P04 Process

1959 Riesenfeld-Mullowney

Petroleum Refinery - May

Giammarco-Vetrocoke Process

1959 Powdrill

The Institution of Gas Engineers

Operation of Liquid and Gastechink Purification Plants at Cardiff

1959 Leuhddemann-Noddes-Schwartz

Oil Gas Journal - August - 30

Alkazid Process

1959 Sexauer


Silos system with regeneration

1960 Benson-Field

Petroleum Refinery - April

Benfield Process

1961 Bienstock-Field Corrosion - July - 17

Alkaline solutions - corrosion

1962 Eickmeyer

Chemical Engineering Progress - April - 22

Catacarb Procedure

1963 Nicklin-Holland

European Symposium on Cleaning Coke Oven Gas - Saarbrucken

Stretford Procedure

1963 Dingman

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Animine – concentrated DEA – SNPA Process

1970 Franckowiak

Hydrocarbon Processing - March

Estasolvan: New Gas Treating Process

1970 Hasebe

Chem. Econ. & Eng. Review (J) - March

Takahax Procedure

1971 Guyot-Martin

Canadian Gas Processors Association Meeting

Sulfreen Process

1971 Hawkes - Mago

Hydrocarbon Processing - AugustStop MEA CO2 Corrosion

1971 Ludberg

64th Meeting of Air Pollution Control Association

Removal H2S from Coke Oven Gas by Stretford Process

1971 The Benfield Corporation

The Way to Low Cost Scrubbing of CO2 and H2 S from Industrial Gas

1971 Barthel-Bistri-Renault

Hydrocarbon Processing - May

IFP Process: Oxidation with SO2 in solvent at high temperature

1971 Livingston

Hydrocarbon Processing – January

Pretreat Syn Gas Feeds (ZnO)

1972 Nonhebel

Edit. Newness-Butterworths-London

Gas Purification Processes for Air Pollution Control

1972 Barry

Hydrocarbon Processing - April

Reduce Claus Sulphur Emission

1973 Martin


Plant Peaks Sulphur Recovery

1973 Butwell-Hawkes- Mago

Chemical Engineering Progress - February

Corrosion Control in CO2 Removal Systems

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1973 Ranke

Linde Reports (18)

Rectisol Procedure

1973 RuziskaChemical Engineering Progress - 69

Packings for Hot Carbonate Systems

1973 Scheirman

Hydrocarbon Processing - August

Filter DEA Treating Solutions

1974 Moyes-Wilkinson

Chemical Engineering – February

Development at the Holmes - Stretford Process

1974 Kohl-Riesenfeld

Gas Purification

General Information

1974 Goar

Hydrocarbon Processing - July

Impure Feeds Cause Claus Plant Problem

1974 Bratzler-Doeges


Amisol Procedure

1974 Maddox

Gas and Liquid Sweetening

General information

1975 Kasay

Hydrocarbon Processing - February

Konox Process Removes H2S

1975 Baker Hydrocarbon Processing - April

Corrosion Free Gas Sweetening

1975 Strelzoff

Chemical Engineering - September

Choosing the Optimum CO2 Removal System utilization

1976 Anon

Natural Gas Processing Conference - Dublin

SULFINOL PROCESS

1976 Kent-Eisendef

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Hydrocarbon Processing – February

Better Data for Ammine

1976 Schianni

Natural Gas Processing & Utilisation Conference - DublinCryogenic Removal of Carbon Oxide from Natural Gas

1976 Raney


Remove CO2 with Selexol

1977 Archibald

Hydrocarbon Processing - March

Process Sour Gas Safely

1977 Vidaurri-KahreHydrocarbon Processing - November

Recover H2S Selectively from Sour Gas Streams

1977 Tennyson-Schaaf

Oil Gas Journal – January - 10

Guidelines Can Help Choose Proper Process Gas Treating

1977 Kelly


How Ammine Guard Saves Energy

1978 Vasan

Oil Gas Journal - 1978

Holmes-Stretford Process Offers Economic H2S Removal

1978 Judd


SELEXOL Procedure

1978 Klein

Oil & Gas International - September - 10DIPA/Sulfinol Processes

1978 Christensen-Stupin

Hydrocarbon Processing – February

Merits of Acid Gas Removal Processes

1978 Rayan-Loiselle

Hydrocarbon Processing - November

Make Sulphur from Lean Acid Gas

1978 Ouwerkerk Hydrocarbon Processing – April

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Design for Selective H2S Absorption

1978 Judd


SELEXOL Unit Saves Energy

1978 Palm-Caruthers

Oil Gas Journal - November

Guidelines Aid Control of SRU Tail-gas Emissions

1978 Pagani-Guerreri-Peri

USA Patent 4.097.257

Natural Gas Having High Content of Acidic Gases Purification Method

1978 Grancher

Hydrocarbon Processing - September Advances in Claus Technology

1978 Gas Processors Association


H2S Removal with MDEA

1978 Wright-Strauge

Oil Gas Journal - February

Modified Sulphur Recovery Process Meets Air Quality regulation

1979 Manning

Oil Gas Journal - October - 15

Chemsweet Process

1980 Kennard-Meisen


Control DEA Degradation

1980 Henderson-Hudson-Kimble

SPE of AIME Conference - September

Jay and Flowation Fields Sulphur Plant Operations

1980 Holmes-Ryan-Price-Styring

GPA

Pilot Test Prove Ryan/Holmes Cryogenic Acid Gas/Hydrocarbon

Separation

1980 Meissner

World Oil - October

Purifying CO2 for use in EOR

1981 Blanc-Elgue-Lallemand Hydrocarbon Processing - August

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MDEA Process Selects H2S

1981 Keaton

Hydrocarbon Processing – August

Activated Carbon System Cuts Foaming and Amine Losses

1981 Parnell


General information

1981 Sigmund-Butwell-Wussler


Ucarsol Process

1981 Zawacky-Duncan-Macriss

Hydrocarbon Processing - AprilProcess Optimized for High Pressure Gas Clean-up

1981 Kresse-Lindsay

Oil Gas Journal - January - 12

Stretford Process

1981 Vaz-Mains


Ethanolamines Process Simulated by Rigorous Calculation

1981 Huval-Van De Venne

Oil Gas Journal - August - 17

Fluor Econamine/DGA Process

1982 Meisen-Kennard

Hydrocarbon Processing - October

DEA Degradation Mechanism

1982 Holmes-Ryan-Price-Styring


Process Improves Acid Gas Separation

1982 Blanc-Grall-Demarais

University Oklahoma Conference

Degradation Products and Corrosion of Plants Using DEA/MEA

1982 Goodin


Pick Treatment for High CO2 Removal

1982 Mackinger-Rossati-Schmidt

Hydrocarbon Processing - MarchSulfint Process

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1982 Kaplan

Chemical Engineering - November

Cost-saving Process Recovers CO2 from Power-plant Flue Gas

1982 Brown

Unido Tech. Conference-Peking

Giammarco (with glycine)/Criteria for Selecting CO2 Removal Processes

1982 Hass-Fenton-Gowdy-Bingham

Sulphur Conference

Selectox and Unisulf: New Technologies for Sulphur Recovery

1982 Butwel-Kubek-Sigmund

Hydrocarbon Processing – March

Alkanolamine Treating

1982 Woelfer

Hydrocarbon Processing - November

Sepasolv Process

1983 Villa

Air Pollution Seminar-Kuwait

IGI Experience of Modified Claus Process

1983 Jarett


Fundamentals of Acid Gas Fractionation

1983 Nicholas-Wilkins-Li

Hydrocarbon Processing - September

Optimize Acid Gas Removal

1983 Russel

Hydrocarbon Processing-August

Field Tests for Delsep Permeators (Membrane)

1984 Goldstein-Edelmann-Beisner-Ruziska


Flexsorb Process/Hindered Amines Yield Improved Gas Treating

1984 Parro

Oil Gas Journal - September - 24

Membrane CO2 Separation

1984 Stanbridge-Hefner

AICHE Meeting Anaheim-California

Recent Developments in BASF Activated MDEA Process

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Page. 173 of 175

1984 Sulphur - November - December

Catalyst and the Claus Process

1984 Price

Oil Gas Journal - December Looking at CO2 Recovery in Enhanced Oil Recovery Projects

1984 Gazzi-D’Ambra-Rescalli


Cryofrac Procedure

1984 Mortko

Hydrocarbon Processing - June

Selexol/Remove H2S Selectively

1984 Bucklin SchendelEnergy Progress - September

Comparison of Fluor Solvent and Selexol

1984 Anon

Chemical Processing - November - 15

Non Toxic Catalytic Reagent Converts H2S in Sulphur

1984 Daviet-Sundermann-Donnelly


Switch to MDEA Raises Capacity

1985 Clem

35th Annual Gas Condition Conference

Flexsorb Process

1985 Pearce-Du Part


Amine Inhibiting

1985 Netzer

Hydrocarbon Processing - AprilRandall Process: Process Designed to Recycle CO2

1985 Thomason


Reclaim Gas Treating Solvent

1985 Byeseda-Deetz-Manning

Oil Gas Journal - June - 10

Optisol: a New Gas Sweetening Solvent

1985 HardisonHydrocarbon Processing - April

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Page. 174 of 175

Lo-Cat Process

1986 Schillmoller

Hydrocarbon Processing - June

Amine Stress Cracking: Causes and Cures

1986 Lath-Herbert


Amisol Solvent

1986 Riesenfeld-Brocoff

Oil Gas Journal - September - 29

Proc. MDEA/Tertiary Ethanoloammines More Economical for H2 S/CO2

Removal

1986 King-StanbridgeOil Gas Journal - September - 8

Rigorous Screening Selects Sour Gas Plant Process

1986 Maddox-Morgan


Select E.O.R. Processes for CO2

1986 Bradley

Oil Gas Journal - March - 17

CO2 E.O.R. Requires Corrosion Control Program in Gas Gathering

1986 Bianco

Oil Gas Journal - August

Calculation Methods Simulates LPG H2S MEA

1987 Carnell


New Fixed-bed Adsorbent for Gas Sweetening

1987 Heisel-Masold

Hydrocarbon Processing - April New Gas Scrubber Removes H2S (Sulfolin)

1987 Marsh


Michigan Plant Opens up Nearby Sour-gas Field

1987 Chowdhury

Chemical Engineering - September

Membranes Set to Tackle Larger Separation Tasks

1987 KresseOil Gas Journal - October

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CO2 Removal Reduces Pipeline Corrosion at Two Storage Sites

1987 Goar

Chemical Engineering - December

Claus Oxygen - Based -Process Expansion (COPE)

1987 Youn-Poe-Sattler-Inlon

Oil Gas Journal - November

CO2 Recovery EOR SAGA - Ryan Holmes

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