2007-05-30

KET050 Feasibility Studies on Industrial Plants Dept of Chemical Engineering, Lund Institute of Technology

A Feasibility Study on Integrated Hydrogen Production

Presented to Norsk Hydro ASA

Norway

May 11, 2007

Principal investigators: Alexander Dingizian, Jens Hansson, Tony Persson, Henrik Svensson Ekberg, Per Tunå

Tutors:

Hans Eklund and Klaus Schöffel Norsk Hydro ASA

Christian Hulteberg and Hans T. Karlsson Department of Chemical Engineering

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Abstract A number of feasibility studies have been conducted on hydrogen production based on biomass as feedstock. One of the contributing participants is Norsk Hydro, who has assigned this project to investigate the possibilities of combining hydrogen production through water electrolysis and biomass gasification. As a scale indicator, the annual production of the projected plant is specified to 14 000 tonnes of hydrogen. In the first part of the project, the plant was designed, mainly by the use of literature references. Many of the current alternatives were evaluated and contrasted and a final design was set. The process layout contains an electrolysis part and a gasifier part. It was decided that the production rate of the electrolyser should stoichiometrically cover the oxygen input to the gasifier and tar cracker, and hence the ratio of production between the units is locked. The electrolyser of choice is an alkaline atmospheric Norsk Hydro unit, with a maximum capacity of 100 Nm3 H2/h. With respect to its large capacity and the low demands on the raw material size, a bubbling fluid bed gasifier is utilised. The electrolysis part contain a few pre-treatment and purification steps, including water purification, deionisation, deoxidation and drying, whereas the gasification requires more steps: The biomass is ground, dried and fed to the gasifier and the output gas purified through thermal tar cracking (tar removal), particle removal, desulphurisation, high and low temperature shift (CO reduction), scrubbing (removal of water soluble impurities) and PSA (pressure swing adsorption). The streams are subsequently combined and compressed according to the final product specifications. In order to improve the energy economy, the PSA exhaust gas is used in the drying step. Moreover, some of the electricity required in the electrolyser is generated by excess high pressure steam in turbines. The flow sheeting calculations were primarily conducted using Aspen Plus software. As this program does not contain models for all unit operations of the process, other methods were used as well. An economical analysis was carried out and a final production cost of the hydrogen was calculated, using the Ulrich and other factor estimation methods. Additional costs such as staff, maintenance etc. are estimated through rules of thumb. The final production cost estimate was US$ 4.97 /kg, which can be compared to US$ 1.91 /kg for hydrogen which derivates from coal gasification and US$ 6.58 /kg for hydrogen produced by on-site electrolysis. The dominant costs of the hydrogen plant include electricity, financial (loan interest) and the purchase of the electrolyser system. The sensitivity analysis conducted on these key factors, as well as on the biomass cost, showed that the hydrogen production cost varies primarily with the electricity price and the interest rate. Since the electricity cost influences the production price to such extent, power saving alternatives and a steady, low electricity price is important. Alternatively, the plant could be built without the electrolyser units.

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Table of contents

1 Introduction _______________________________________________________ 4

2 Screening and selection of equipment___________________________________ 6

2.1 Electrolysis__________________________________________________________ 7 2.1.1 Construction _____________________________________________________________ 7 2.1.2 Pre-treatment ____________________________________________________________ 8 2.1.3 Electrolysis gas purification _________________________________________________ 9 2.1.4 Choice of electrolyser______________________________________________________ 9

2.2 Pre-treatment ______________________________________________________ 10 2.2.1 Grinding _______________________________________________________________ 10 2.2.2 Drying_________________________________________________________________ 11 2.2.3 Feeding systems _________________________________________________________ 12 2.2.4 Choice of pre-treatment equipment __________________________________________ 12

2.3 The gasifier ________________________________________________________ 13 2.3.1 Gasifier types ___________________________________________________________ 13 2.3.2 Choice of gasifier ________________________________________________________ 15

2.4 Purification ________________________________________________________ 16 2.4.1 Tar removal ____________________________________________________________ 16 2.4.2 Reforming of hydrocarbons ________________________________________________ 18 2.4.3 Reformer design _________________________________________________________ 19 2.4.4 Water Gas Shift Reaction __________________________________________________ 21 2.4.5 Desulphurisation_________________________________________________________ 22 2.4.6 Reactor design __________________________________________________________ 22 2.4.7 Removal of particles______________________________________________________ 23 2.4.8 Scrubbing ______________________________________________________________ 23 2.4.9 Gas purification _________________________________________________________ 23 2.4.10 Choice of purification equipment____________________________________________ 24

2.5 Heat recovery_______________________________________________________ 24

2.6 Process Flow sheet___________________________________________________ 25

3 Process Design ____________________________________________________ 26

3.1 Electrolysis_________________________________________________________ 27

3.2 Oxygen utilisation ___________________________________________________ 27

3.3 Size reduction ______________________________________________________ 27

3.4 Drying_____________________________________________________________ 28

3.5 Feeding system______________________________________________________ 28

3.6 The gasifier ________________________________________________________ 29

3.7 Tar Cracker Heat balance and mass balance_____________________________ 30

3.8 Desulphurisation ____________________________________________________ 30

3.9 Water Gas Shift_____________________________________________________ 31 3.9.1 High Temperature Shift ___________________________________________________ 31 3.9.2 Low Temperature Shift____________________________________________________ 31

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3.10 Heat exchangers and power generation _________________________________ 32

3.11 Alternative power generation _________________________________________ 33

3.12 Summary __________________________________________________________ 34

4 Economic evaluation _______________________________________________ 37

4.1 Methods ___________________________________________________________ 37 4.1.1 Ulrich method___________________________________________________________ 37 4.1.2 The scale method ________________________________________________________ 38

4.2 Capital cost ________________________________________________________ 39

4.3 Operating cost ______________________________________________________ 41

4.4 Investment analysis__________________________________________________ 42

4.5 Sensitivity analysis __________________________________________________ 42 4.5.1 Electricity price _________________________________________________________ 42 4.5.2 Interest rate_____________________________________________________________ 43 4.5.3 Electrolyser unit cost _____________________________________________________ 44 4.5.4 Biomass cost____________________________________________________________ 44 4.5.5 Summary of the sensitivity analysis __________________________________________ 45

Discussion and conclusion ______________________________________________ 46

5 References _______________________________________________________ 48

Appendix A ___________________________________________________________ 50

Appendix B ___________________________________________________________ 51

Appendix C ___________________________________________________________ 52

Appendix D: Drying____________________________________________________ 53

Appendix E: Alternative power generation__________________________________ 54

Appendix F: Cost estimations ____________________________________________ 56

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1 Introduction With increasing demand for efficient and renewable energy to power our society, biomass and other renewable energy are constantly being revised as a feasible alternative to fossil fuels. Especially hydrogen with the next generation of fuel cells may prove to be a successful and sustainable alternative to oil. The most commonly used processes for hydrogen production today is oil cracking and steam reforming of natural gas, both based on fossil fuel. In this study an alternative solution to produce hydrogen from biomass and electrolysis has been conducted in cooperation with Norsk Hydro. The gasification technique has been studied in this assignment. Norsk Hydro is also developing a technique for electrolysis of water, therefore it is interesting to study if there will be any synergy effects by combining the two techniques, gasification and electrolysis. Combining the two processes will enable oxygen supply for the gasification from the electrolysis step. The electrolysis step produces hydrogen which can be added to the hydrogen from the gasified biomass. This process does furthermore fulfill the demand of reduced CO2-emissions since it is CO2-neutral, assuming that the electricity used in the electrolysis is based on renewable production. The electricity required for the electrolysis would partially be generated by the excess steam production from the gasification process and combustion of biomass or external power sources. The annual production of the plant is set to 14 000 tonnes of pressurised hydrogen. The location of the plant would be in Norway. The purpose of the feasibility study currently carried out is to perform a design of the hydrogen plant with emphasis on the integration and optimisation between the two process types. Choices of process units, energy optimisation and economical analysis are included.

Figure 1. Simple flow sheet of the process

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In figure 1 a simple schematic over the process is presented to illustrate the most important process units. The electrolyser is fed with water and electricity and producing hydrogen, it also supplies the gasifier with oxygen. In order to have an efficient gasification the biomass feedstock must be pre-treated. That is achieved by grinding followed by drying. In the gasifier, biomass, oxygen and steam are converted to syngas and tars. The gas needs to be purified prior to delivery and that is done in several steps including water-gas shift and pressure swing adsorption. Heat recovery is a necessary process step in order to produce on site electricity and to optimise the heating and cooling of the different process units.

The economical analysis indicates the feasibility of this alternative of hydrogen production. It includes capital costs, operating costs and through an investment analysis, the production cost of hydrogen is calculated. The effect from the variation of some of the dominating costs is studied in a sensitivity analysis.

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2 Screening and selection of equipment The methods for hydrogen production studied are gasification of biomass and electrolysis of water. Combining the two processes will enable oxygen supply for the gasification from the electrolysis step. Before the biomass enters the gasifier it must pre-treated. This includes operations like drying and size reduction. After the gasifier the produced gas must be purified before it can be stored. The gas contains organic impurities including tar and benzene but also other impurities such as carbon monoxide, carbon dioxide and soot particles. The purification involves unit operations like cracking, scrubbing and pressure swing adsorption. The electricity required for the electrolysis would partially be generated by the excess steam production from the gasification process and combustion of biomass or external power sources.

Figure 2. Simple flow sheet of the process

In order to define a system for the entire process, screening and selection of several equipments must be made, as almost every stage of the process can use different technological solutions. After the pre-treatment of the biomass the gasification of the biomass starts to produce syngas. The oxygen supply for the gasification comes from the electrolysis step. The produced syngas must undergo several cleaning steps in order to meet the specification for this plant. The heat recovered from the gasifier and the tar cracker is heat exchanged with water to produce superheated steam. The superheated steam is utilized to produce electricity in the turbine for the electrolysis step.

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2.1 Electrolysis The electrolysis process has been used for over 100 years in the production of hydrogen and the technology is now well-established. The first commercial system was built in 1902 and the first large plant was built by Hydro in 1927. The economic feasibility is chiefly determined by the cost of electricity. Electrolysis plants at a larger scale therefore demand a supply of cheap electric power, e.g. hydroplants. Smaller plants (50-500 Nm3/h) are often used in industries because the plants are easy to operate [1] and are usually cheaper than transported hydrogen.

2.1.1 Construction The classical, monopolar electrolyser is constructed as shown in figure 3.

Figure 3. The monopolar water electrolyser construction, with potassium hydroxide added to improve the performance

The electrolyser consists of a power source which transports electrons from the electrolyte at the anode to the electrolyte at the cathode. The electrolyte sections are separated by a diaphragm or a membrane which permits OH- ions and the electric current to pass through, whereas the actual electrolyte is stationary. This results in the following reactions at the electrodes: Cathode: )(2)(2)(2 22 aqOHgHelOH −− +→+ (VII)

Anode: −− ++→ elOHgOaqOH 2)()(½)(2 22 (VIII)

Cell reaction: )(½)()(2 222 qOgHlOH +→ (IX) Pure water has low conductivity and is thus not suitable. Therefore, a water solution usually of KOH, NaOH or HCl is used to increase the conductivity. With few exceptions, the electrolyte in conventional electrolysers has a 25-35 wt% content of KOH. [2]

H2 O2

cathode

anode

OH- OH-

KOH

H2O

e- e-

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Because of the harsh environment in the electrolyser, corrosive resistant material is required in all parts which are in contact with the electrolyte. Carbon steel is common, as are coatings with plastics, ceramics and nickel in especially exposed areas. The diaphragm is normally made by asbestos but in recent times, research has been done to replace the material with synthetic polymers. Another current development is to replace the diaphragm with a fluorinated cation exchange membrane [2]. The temperature of conventional electrolysis has to be kept below the boiling point of the electrolyte, i.e. at 70-90 oC at ambient pressure (some exceptions, with gaseous electrolyte exist). An increase in pressure will elevate the boiling point and hence, a higher temperature is possible, something which enhances the rate of production. A parallel plate cell has two separate electrodes (as is shown in figure 1), divided or not divided by a membrane/diaphragm, and is run in a monopolar mode. In the bipolar cell, several parallel plate cells are arranged in series and each electrode serves as anode on one side and as cathode on the other side. The bipolar cell is generally simpler to construct but corrosion and material problems can occur. However, the majority of the problems have been overcome, and most electrolysers today are constructed to run in the bipolar mode. [1]

Figure 4. a) Cell walls; b) Electrolyte; c) Cathode; d) Anode; e) Hydrogen outlet; f) Oxygen outlet; g) Gas collector; h) Diaphragm; i) Outer cathode; j) Outer anode; k) Bipolar electrode; l) Insulation [1]

2.1.2 Pre-treatment The electrolyser has a water consumption of about 0.8 l/Nm3 H2 plus losses for evaporation. There are high requirements of the feed water purity. It is necessary to avoid build-up of substances in the electrolyte solution which may poison the electrode or promote corrosion. Ion exchangers are used to obtain the acquired conductivity of 1 µS/cm or less before mixing the water with lye [3].

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2.1.3 Electrolysis gas purification The product contains small amounts of impurities. Lye residuals from the electrolyte solution are removed with a gas scrubber. The small amount of oxygen (about 1 %) present in the product is removed with a catalytic deoxidiser. The deoxo reaction is:

)(2)()(2 222 lOHgOgH →+ (I) This is a strongly exothermic reaction why heat exchangers are often installed after the deoxo reactor. The catalysts are often platinum or palladium on carriers such as α-Al2O3, active charcoal, or aluminosilicates. After the gas scrubber and the deoxidization the gas holds a large amount of water which can be removed by absorption or condensation [3].

Figure 5. Process diagram of an electrolysis unit with the most important components a) Rectifier unit; b) Process water demineraliser ion exchange unit; c) Electrolytic cell's water electrolyser; d) Gas separator and cooler; e) Gas scrubber; f) Electrolyte tank and transfer pump; g) Gas holder; h) Filter; i) Compressor; j) Gas purifier; k) Drying; l) High-pressure storage tank and cylinder filling station [3]

2.1.4 Choice of electrolyser In the selection of the actual electrolyser, one of the main requirements is to acquire a pressurised hydrogen product, since this is required in the specifications. Given that the produced oxygen will be used in the gasification process, an increased pressure will reduce the extent of compression prior to its use in the gasifier. Moreover, the power consumption for pressure electrolysis does not increase significantly and hence, these units save both equipment and energy for the hydrogen compression [2]. There are a number of electrolysers industrially available, of which only a few are pressurised. The latter are currently manufactured by for example ELT and GTec (Lurgi type) [4], MTU, Teledyne [2] and Hydrogenics [1]. See table 1.

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Table 1. Conventional electrolysis systems [3], [1], [2], [4], [5], [6]

Lurgi system (ELT/GTec)

MTU Teledyne Hydrogenics Norsk Hydro

ABB &Cie

Cell type bipolar bipolar bipolar bipolar bipolar bipolar Operating pressure (bar)

30 30 7 10/25* Atmospheric ambient

Operating temperature (oC)

90 130 80 - 80 80

Electrolyte 25 % KOH 30 % KOH KOH 30 % KOH 25 % KOH 25 % KOH Current density (A/m2)

2000 7000-10000

3000 - - 2000

Cell voltage (V) 1.86 1.65-1.8 - - - 2.04 Current efficiency (%)

98.75 >99.5 - - 99.9 99.9

Power cons. (kWh/Nm 3)

4.3-4.65 4-4.4 5.6 4.2 4.3 4.9

Maximum prod. rate (Nm3/h)

760 - 140 60 100 -

* Hydrogen output pressure The optimal choice of electrolyser is the Lurgi type. This unit combines a high output pressure of both oxygen and hydrogen (30 bars) and operates at a high temperature. Moreover, the market position of this equipment is excellent, with more than 100 installed systems worldwide. To supply the demanded production rate of hydrogen, a number of parallel units will be needed. The maximum production rate for a single Lurgi electrolyser is 760 Nm3/h [4]. The pressurised oxygen product is used in the gasifier. The oxygen residuals in the hydrogen stream are removed in a catalytic deoxidiser. However, the Lurgi electrolyser is a very complex and expensive choice and therefore not a suitable choice for this project. Instead the atmospheric electrolyser from Norsk Hydro is chosen. When using an atmospheric electrolyser two compressors have to be installed, one for the hydrogen and one for the oxygen. The maximum production rate for a single electrolyser is 100 Nm3/h. To supply the demanded production of hydrogen several electrolysers will be needed.

2.2 Pre-treatment Prior to gasification the biomass has to be pre-treated to have the correct size and moisture content.

2.2.1 Grinding The biomass has to be ground prior to gasification. Different gasifiers require different particle size, the Bubbling Fluidised Bed, BFB, gasifier needs a particle size of 5 mm while the Entrained Flow, EF, gasifier requires a particle size less than 1 mm [7].

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There are two methods that are commonly used to grind wood to the required sizes. They are knife chippers and hammer mills. A knife chipper utilises blades that rotate at high speed (1800 rpm) and together with stationary blades they cut the wood to small size. However if the wood feedstock contains metal the blades will be damaged [7]. In a hammer mill the wood is crushed by large and heavy hammers that rotate against a breaker plate. Smaller pieces are pulverised between the hammers and a mesh at the bottom of the mill. The hammer mill requires an inlet particle size that is less than 4 cm, if so the hammer mill can reduce the biomass to sizes less than 0.6 mm [7].

2.2.2 Drying Drying is an important pre-treatment of the biomass as the efficiency of the gasifier increases with lower moisture content but drying costs increase rapidly when moisture contents go below 10 wt-% [8]. The biomass typically has a moisture content of 35 wt-% [8] when it is fed to the dryer. There are many different types of dryers that can be utilised and two of them are rotary cascade dryer and the steam dryer. The rotating dryer is made of a large cylindrical shell that slowly rotates and hot gas flow through the dryer. The rotary dryer has a drawback and that is that it requires large gas volumes since the heat transfer is low. Since the plant has plenty of excess steam, a steam dryer can be an alternative. Two examples of steam dryers are conveying steam dryer [9] and pressurised steam fluid bed dryer [10]. The pneumatic conveying steam dryer is fed by a pressure tight rotary valve or a plug screw. The drying steam is superheated indirectly in a heat exchanger and if lower moisture content is required a secondary heat exchanger can be utilised. The residence time is normally 5-60 seconds. As the material dries, steam is produced that must be ejected from the loop. Mechanical Vapour Compression can be used to produce 10-20 bars steam to be used as heating media for the heat exchangers. The pressurised steam fluid bed dryer is

Figure 6: Pneumatic conveying steam dryer [9]

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made of a central heat exchanger and sixteen cells where the biomass is held. Superheated steam is produced in the heat exchanger and is used to dry the feedstock. In order to remove entrained biomass particles in the used steam the steam passes an internal cyclone.

2.2.3 Feeding systems A pressurised gasifier requires a pressurised feeding system to avoid backflow of the gases produced from the gasifier. Often a locked hopper system is used to feed pressurised gasifiers as it can work at pressures up to 35 bars [8].

Figure 7. Lock hopper feeder [7].

Other feeders include piston and screw feeders that can also be used with pressurised systems. Screw feeders operate at a pressure of 5-25 bars and piston feeders at 45-150 bars [8].

2.2.4 Choice of pre-treatment equipment Pre-processing of the biomass begins with grinding the material into desirable particle size. The grinding is conducted by a hammer mill using large hammers to crush the biomass. The hammer mill has a great durability for handling feedstocks with contamination. Following the grinding it is important to dry the raw material. Low moisture content increases the efficiency of the gasifier. Heated air from the system is used to dry the raw material in a rotating dryer. The rotating dryer is chosen because the steam produced in the system is used for producing electricity which can be used in the electrolyser. Feeding the bio mass to the gasifier is performed with a piston feeder, chosen due to lack of information on the other systems.

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2.3 The gasifier When biomass is gasified, an insufficient amount of oxygen for stoichiometric combustion is fed and the gasification takes place at high temperature. Because of the deficancy of oxygen, components like H2 and CO are formed. Other unwanted components such as H2O, CO2, benzene, toluene, xylene, HCl, NH3, H2S, HCN, COS, alkaline metals, soot and ashes, are also formed. Therefore, the gas has to be cleaned. It is also feasible to use some of these byproducts because of their heat content [8], [11]. The gasification process of biomass is divided into three parts: pyrolysis, combustion and reduction. An overview of the gasification streams is displayed in figure 8. Pyrolysis will occur when the biomass is heated to 450-600 °C with no oxygen present. The result is pyrolysis gas and char coal. The combustion step supplies heat to the entire gasification process. In the reduction step synthesis gas is produced. The synthesis gas is a mixture of H2 and CO [8], [11].

Figure 8. Gasification process [11]

2.3.1 Gasifier types Fixed bed, fluid bed and entrained bed are the three main gasifiers. These different types will be described below.

Fixed bed Fixed bed gasifiers are the oldest ones and have been used for a long time. The fixed bed is simple and robust, which is an advantage. A disadvantage is that, the capacity has a limitation of 10-15 ton dry biomass per hour and the produced gas contains too much tar.

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The fixed bed can therefore not be used industrially to gasify biomass because of these disadvantages. There are at least three types of fixed bed gasifiers: updraft, downdraft and crossdraft [7], [8].

Figure 9. Fixed bed gasifier [12]

Fluid bed A fluid bed gasifier contains small particles, such as sand or alumina, to achieve an effective heat transfer and mixing of biomass, steam and oxygen. There are two types of fluid bed gasifiers: dense and circulating. In the circulating gasifier the small particles are fluidised (floating) and leaves the gasifiers together with the product gas. These particles are later separated from the product gas in a cyclone and brought back to the gasifier. The bubbling fluidised bed works almost like the circulating bed gasifier. The only difference is that the fluidised medium does not have enough velocity to make the whole bed float. It only bubbles through it [10]. Advantages with these gasifiers are that the biomass does not need much pre-treatment before entering the gasifier. The biomass still have to be dried but the size of the biomass particles entering the gasifier does not have to be smaller then 5 mm. Only a small amount of tar is produced and these kinds of gasifiers have a large capacity.

Figure 10. Bubbling fluidised bed [13]

Figure 11. Circulating fluidised bed [13]

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Entrained bed In an entrained bed gasifier a gas and a finely ground biomass is used as the feed. Later, the feed is mixed with an oxygen/steam flow and are gasified in a flame at high temperature. The gasification process takes place at 1200 °C and at pressures higher than 20 bars. An advantage is that the product gas does not contain any tar, since the gasification takes place at such high temperatures. Another advantage is that the gasifier has a large capacity. However, the biomass has to be ground to powder before entering the gasifier which results in high costs. A large amount of heat is generated and has to be taken care of, it can be used for production of electricity and steam. This brings a more complex design of the gasifier system [10].

Figure 12. Entrained bed gasifier [13]

2.3.2 Choice of gasifier It has been shown that the bubbling fluidised bed is the best choice for this process, this based on capacity and raw material size. There are however three factors which have to be decided.

1. Pure oxygen or air. 2. Pressurised or atmospheric pressure. 3. Direct or indirect heating of the gasifier.

1. Regarding the choice of air or pure oxygen, the choice is pure oxygen since the

nitrogen in the air takes up space. That means using pure oxygen from the electrolysis will reduce the size of the gasifier. The nitrogen also reduces the energy density in the gas and the problem with nitrogen removal is also avoided.

2. Pressurised gasification is chosen since it reduces the cost of the gasifier. If the

gasification occurs at 30 bars, the size of the gasifier can be reduced and the material costs will decrease. Additionally, a higher pressure reduces the tar production. A pressurised gasifier simplifies the gas purification and the cost for compression work is saved. On the contrary, the feeding process will become more difficult and will therefore be more expensive.

3. Direct heating of the gasifier is chosen since less tar is produced in the syngas

[10], [8].

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2.4 Purification

2.4.1 Tar removal The gas produced in the gasifier contains impurities. There are both inorganic impurities, like ammonia and hydrosulphurs, and organic impurities like tars and benzene. Tar can be defined as all organic contaminants with a molecular weight larger than benzene. When condensing (350-400°C) they foul downstream equipment, coat surfaces and enter pores in filters and sorbents [7], [14]. The organic impurities contain a lot of potential hydrogen; therefore they should be used to produce more hydrogen. The tars must first be cracked into smaller hydrocarbons. Fluidised beds produce tar at about 1-5 wt-% of the biomass feed. Roughly 10 % of the heat value is found in the tar [7], [14]. Tar is a complex mixture of different condensable organic compounds; they can be single ring and multiple ring aromatic compounds, oxygen containing hydrocarbons and complex polycyclic aromatic hydrocarbons. Due to the large mixture of organic compounds that are defined as tars they have been divided into four different classes depending on where in the reaction chain the tars are produced (se figure 8) [7], [14]. The four classes are: primary products, secondary products, alkyl tertiary products and condensed tertiary products.

The organic compounds derived from hemi cellulose, cellulose and lignin are classed as primary products. Phenolics and olefins are classed as secondary products. Alkyl tertiary products are mainly characterized as methyl derivatives of aromatic compounds. Condensed tertiary products comprise benzene, naphthalene, acenaphtylene, anthracene and pyrene. The different tar classes are present at various temperatures. The primary products are present at 500-800 °C the secondary products are present at 500-1000 °C, the alkyl tertiary products at 650-1000 °C and the condensed tertiary products at temperatures above 750 °C. There are three general methods for tar cracking or removal, thermal cracking, catalytic cracking and scrubbing [7], [14].

Figure 13. Tar classes [7]

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Figure 14. Temperature where the tar classes are present.[7]

Thermal tar cracking At temperatures above 1000-1200 °C tars are decomposed or reformed to hydrogen, carbon monoxide and other light gases without a catalyst, usually by addition of steam and/or oxygen which acts as selective oxidant. If enough amounts of oxygen are added the tar will be combusted. In order to achieve proper removal of tars the condensed tertiary products must be cracked, which requires temperatures of 1300°C. At this temperature soot is formed. The efficiency of the thermal treatment can increase if longer residence time is used. Residence time over 10 s at NTP is enough of time. Drawbacks are the loss of energy, the soot production and the low thermal efficiency [7], [3]. A thermal cracker can crack all hydrocarbons into carbon oxides and hydrogen. If the cracker works as the study in table 2 no reformer is necessary. [8]

Table 2. Mass flows in and out of the tar cracker [8].

Component In flow (kg/s) 850 °C Out flow (kg/s) 1300 °C CH4 1.29255 0 C2H4 0. 09093 0 C6H6 0.10145 0 Tar (C14H10) 0.127 0

Catalytic tar cracking Catalytic cracking is best applied in a secondary bed and it avoids the mentioned problems of thermal cracking. The catalyst is often dolomite or nickel based. Dolomite is

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a very effective catalyst for tar cracking when calcined. Experiments have been done and the result show that the tar content was reduced in product gas from 12 wt% to 0 wt%. But the use of dolomite catalyst in pressurised systems is however less efficient because of decalcination at high pressures. The technology is not yet fully proven and the catalyst consumption and costs are matters of concern [7], [14].

Oil scrubbing The third method is scrubbing, where the tar is removed at low temperature by advanced scrubbing with an oil based medium. The tar gets stripped from the oil and is re-burned in the gasifier [7], [14].

2.4.2 Reforming of hydrocarbons The hydrocarbons which are formed in the tar cracker and the organic impurities contain a lot of potential hydrogen. In the aspect of producing hydrogen it is therefore efficient to reform the hydrocarbons to hydrogen and carbon dioxide. The reformation of hydrocarbons to hydrogen and carbon monoxide occur in a general reaction, reaction II. The carbon monoxide produced in reaction II can react with water to produce carbon dioxide and hydrogen. This reaction is called the shift reaction, reaction III . The shift reaction is exothermic.

( ) 22 2 HmnnCOOnHHC mn ++↔+ (II)

222 HCOOHCO +↔+ (III) In reaction I, hydrocarbons react with steam in an endothermic reaction. The most common hydrocarbon is methane but they can also be propane, butane and naphtha fractions etc. The reaction is an equilibrium reaction and to achieve maximum hydrogen yield, some conditions must be met: High temperature at the end of the reactor, high excess of steam and low pressure. Normal operation conditions are at a temperature of 750 – 850 °C and a pressure of 15 – 30 bars. The amount of steam added to the reactor varies between a ratio of 2.5-6 molar steam to carbon [7],[3].

Catalyst In commercial plants, steam reforming are catalysed by supported nickel catalysts. The catalyst contains 15 – 25 wt % nickel oxide on a mineral carrier. Carrier materials are α-alumina, aluminosilicates, cement, and magnesia. Before start-up, nickel oxide must be reduced to metallic nickel. To prevent catalyst poisoning by sulphur components, uranium oxide and chromium oxide can used as a promoter. This is also reported to have a lower tendency to form carbon deposits [3], [15].

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The main catalyst poison is sulphur, which is present in most feed stocks. Concentrations as low as 0.1 ppm can form a deactivating layer on the catalyst. To compensate the activity loss of a poisoned catalyst the reaction temperature can be raised. This helps to reconvert the inactive nickel sulphide to active nickel sites [3], [15]. Worse than small constant sulphur concentrations are sulphur peaks in the feedstock. They lead to permanent deactivation of the catalyst and to mechanical problems because of carbon deposits. Other catalyst poisons are chlorine and other halogen compounds as well as heavy metals, especially lead, arsenic and vanadium [3], [15]. In the wood feedstock, alkaline and alkaline earth metals are also frequent.

Carbon deposits Soot formation occurs in the reformer because of the decomposition of methane or carbon monoxide, reaction IV or V.

24 2HCCH +→ (IV)

CCO2CO 2 +→ (V) Although hydrogen is produced in reaction IV, it is not desirable because of the soot production. High space velocities in critical regions also help to avoid soot formation [3], [15].

2.4.3 Reformer design The target of a reformer design for the production of hydrogen is to obtain a maximum temperature at the exit of the reformer using a minimum amount of fuel. A commonly used reformer is the tube reformer. The use of fuel is optimised if the heat delivered to each section of the reformer corresponds to the heat requirements of this section in the reaction. For small reformer plants with a hydrogen capacity of up to 1000 m3/h cylindrical furnaces are often used. These furnaces require the least expenditure with regard to steel construction and insulation materials and can be prefabricated in the workshop. The plant size is then limited by the maximum transport dimensions [3],[15]. A general design of a reformer plant is shown in figure 15.

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Figure 15. Hydrogen production by steam reforming, gas purification by pressure swing adsorption (Design: German Linde AG) a) Desulfurization; b) Feed preheater/superheater; c) Reformer; d) Waste-heat boiler; e) CO shift reactor (HT shift); f) Cooling of raw gas; g) Pressure swing adsorption; h) Off-gas puffer for fuel; i) Convection zone with steam production, steam superheating, and air preheating [3]

Partial oxidation of hydrocarbons Partial oxidation is the reaction of hydrocarbons with a sub-stoichiometric combustion at temperatures between 1350-1600°C and pressures up to 150 bars. Two basic reactions (VI, VII) occur, as well as reaction II described earlier. The hydrocarbons react with oxygen to form carbon monoxide or carbon dioxide and hydrogen [3],[15].

22 22 HmnCOOnHC mn +↔+ (VI)

222 2HmnCOnOHC mn +↔+ (VII)

Autothermal reactors Instead of a tube reformer where the energy comes from combustion in a furnace an autothermal reactor can be used. In an autothermal reactor the energy for the production of carbon monoxide and hydrogen is produced by catalytical partial oxidation of parts of the feedstock. The reactor is operated with pure oxygen and steam. Figure 16 shows three different types of autothermal reactors. [3]

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Figure 16. A) Water-cooled jacket reactor with burner for methane-rich gases; B) Secondary reformer with toric oxygen or air supply device; C) Secondary reformer with water-cooled tip without water jacket a) Raw gas inlet; b) Reformed gas outlet; c) Air, oxygen, and steam inlet; d) Catalyst; e) High-temperature catalyst; f) Inert material; g) Internal insulation; h) Multiple layers insulation; i) Burner; j) Water jacket [3]

2.4.4 Water Gas Shift Reaction The water gas shift reaction is usually divided into two industrial segments, high-temperature shift and low-temperature shift. The reaction that takes place is reaction II described earlier.

High-temperature shift The HT-shift reactor operates at temperatures ranging from 300 to 530 °C and pressures up to 70 bars. The reaction almost achieve equilibrium at space velocities between 3000 and 10000 m3 (STP) h–1 m–3. With sulphur free gases, an iron oxide-chromium oxide catalyst can be used. A cobolt-molybden catalyst can also be used but it requires sulphur in the feed gas [3].

Low-temperature shift LT-shift reaction occurs at 190-260 °C and is catalysed by a copper oxide-zinc oxide catalyst. The catalyst can contain stabilisers in the form of Cr2O3, Fe2O3 and Al2O3. Al2O3 acts as a carrier [3]. These catalysts are extremely sensitive to sulphur, contents of 0.1 ppm leading to their deactivation, so that a suitably large reactor volume must be used to ensure that any slow deactivation of the catalyst will not cause any premature shut-down of the unit [3].

22

Due to the equilibrium of the reaction, efficiency of the catalyst and the reaction temperature the shift reaction can remove carbon monoxide only down to a remaining concentration of ca. 0.2-0.5 vol % in the gas [3].

2.4.5 Desulphurisation The gas from the gasifier contains sulphuric compounds. The catalyst in the reformer and the low-temperature shift reactor is highly sensitive to sulphur and therefore a desulphurisation must be performed [3]. Impurities include hydrogen sulphide and other sulphur compounds can be removed by a zinc oxide bed. The adsorbent is active up to 100 bars and at temperatures ranging from 250 to 550 °C [3]. The inlet gas should have the temperature 350 – 400 °C. The gas leaves the reactor at the same temperature with sulphur content of less than 1 ppm [3].

2.4.6 Reactor design There are two different designs that can be used in this project depending on the amount of sulphur present in the outgoing gas from the tar cracker.

Figure 17. The two different designs of the water gas shift system

In design I only a HT-shift reactor is used with a cobolt-molybden catalyst. This design requires a gas with a large amount of sulphur impurities. Design II uses desulphurisation device and both LT- and HT-shift reactors.

23

2.4.7 Removal of particles The product gas contains particles in the shape of ash and fluidising medium. In order to remove larger particles, particles with a size larger then 5 µm, cyclones can be used. In order to remove smaller particles different types of particle filters can be used. [7]

2.4.8 Scrubbing A scrubber is utilised to remove water soluble impurities like ammonia, water and sulphuric components.

2.4.9 Gas purification There are two types of gas purification techniques that can be utilised to produce a high purity hydrogen stream, Pressure Swing Adsorption (PSA) and membrane separation. Membrane separation is still in the development phase. Pressure Swing Adsorption uses a number of adsorption columns in a cycle. The columns are packed with an adsorbent. There are different adsorbents to choose from, depending on the impurities in the feed stream (see table 4).

Table 4. The different adsorbents in the PSA.

Type Main adsorption duty Aluminum oxides water Silica gel water, CO2, C4+ Activated carbon CO2, CH4, C2+, N2 Molecular sieve (zeolite) CH4, CO, N2 Carbon molecular sieve O2 First the feed is fed to the first column (A) at high pressure to adsorb all gas molecules except hydrogen. Hydrogen passes the column and a clean hydrogen stream is produced. When the column becomes saturated it has to be desorbed. Desorption is achieved by reducing pressure which will purge the column. By combining three or more columns a continuous production of hydrogen can be achieved. Typical cycle times are in the range 3 to 10 minutes. A cycle can be described by the following [3]:

1. Adsorption at the pressure of the raw gas (highest system pressure), hydrogen production (adsorber A);

2. Co-current depressurisation, provision of gas for pressure equalisation and purging (adsorbers F and E);

3. Counter-current depressurisation, production of part of the tail gas (adsorber C); 4. Counter-current purging (lowest system pressure), production of another part of

the tail gas (adsorber D); 5. Counter-current repressurisation (adsorber B).

24

Figure 18. Pressure Swing Adsorption with six adsorbers.

The crude gas should have a hydrogen purity of 50 % for the plant to be economically operational. The temperature of the feed gas should not exceed 40 °C as that will result in a lower capacity in the plant [3].

2.4.10 Choice of purification equipment The post-treatment of the produced gas in the gasifier includes a thermal tar cracker. This technology is chosen since it is the most developed and robust method. Thereafter, two different water gas shift system designs are available. The design with only one shift reactor, the one with a cobalt-molybdenum catalyst requires a large amount of sulphur impurities. The amount of sulphur impurities is not known and therefore the other design with two shift reactors and a desulphurisation reactor is chosen. Depending on the fraction of hydrocarbons in the outgoing gas from the tar cracker, a reformer unit might be necessary. The gas stream is led to a scrubber to remove water-soluble by-products. Then the water is removed by absorption in a molecular sieve unit. A PSA is subsequently utilised for the removal of the remaining impurities and a clean hydrogen gas is obtained.

2.5 Heat recovery The hot gases from the thermal tar-cracker and the reformer are heat-exchanged with water to produce super-heated steam used in the generation of electricity. The steam expands over a turbine to produce electricity. The produced electricity is used in the electrolysers to produce oxygen and hydrogen. Some of the electric power needed in the electrolyser is generated by exchanging heat from the gasifier production stream into a separate steam circuit, which is expanded in a power generator. The remaining electricity required is purchased or produced in a power plant.

25

Generator

The

rmal

tar

crac

ker

1300

°C

300 °C

Particles

Z

nO

346 °C

FeC

rO

HT-shiftDesulfurization

162 °C

CuZ

nO

LT-shift

Filter

Steam 500 °C85 bar

Generator

StorageBiomass

Grinding

Cyc

lone

Dry

ing

250

°C

FeedingG

asifi

er B

FB

850

°C

Water Deionisation Electrolysis

KOHH2O

Deoxidastion

O2

H2 + H2O

35 wt%

15 wt%

AshesAir

Dry

er

Scrubber

HCLNH3mm

PSA

H2

Biproductsunreacted

4,8 kW/kg H2

Hydrogen Storage

450 °C

30 bar

Combustion

Air

H2O 85 bar

Steam mix

Steam splitter

H2O H2O

H2O

H2O

Crack splitter

Crack Mixer

Boiler

2.6

Pro

cess

Flo

w s

hee

t T

he

revi

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f th

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26

3 Process Design Gasification of the biomass to produce syngas is the first part of the production of hydrogen gas. The produced syngas must undergo several cleaning steps in order to meet the specification for this plant. The biomass has to be pre-treated before entering the gasifier. The first pre-treatment step is the size reduction. In this investigation the right particle size is set to 5 mm. After the feedstock is transported from the storage room the size reduction starts and it is done by a hammer mill. The biomass feedstock has a moisture content of 35 % and must be dried to a moisture content of 15 % to increase the efficiency of the gasifier. The gas from the PSA is used to dry the biomass. Air is mixed with the gases from the PSA in order to cool them. An air temperature around 250 °C is needed in the dryer. The biomass can be pyrolised and unusable if the temperature exceeds 250 °C. After the dryer the feedstock is transported to the gasifier. The Gasification is performed at 30 bars, which complicates the feeding problem. This problem is solved by using a piston feeder. When the gasification process is finished, a tar cracker removes the tar and other hydrocarbons like methane and benzene. Mass and energy balances have been made over the different pre-treatment steps so the electricity use can be calculated. Then the electricity usage is used to calculate the economy costs. In order to calculate the mass and energy balances Aspen Plus is a useful tool. Aspen Plus is a powerful flow sheeting program that allows the simulation of the entire plant. This software includes many predefined unit operations such as heat exchangers, reactors and turbines. However the program has difficulties with unrealistic operations and therefore the physical properties for every unit must be known. As a first iteration a process design that matched the design in figure 17 was used. But due to the lack of unit operations such as the gasifier and the electrolysis the schedule had to be changed. See appendix B for the complete process schedule in Aspen plus. One of the problems with Aspen Plus is that the user has to have an approximation of the result otherwise Aspen Plus does not converge. For the unit operations which could not be calculated in Aspen plus the calculations have been made by hand. The following conditions have to be met in order to accomplish the predefined target:

• An annual hydrogen production of 14 000 tonnes. • The electrolysis should only produce the amount of oxygen which is used in the

gasifier and the tar cracker, thus minimizing the use of electricity.

27

3.1 Electrolysis In Aspen Plus there is no pre-programmed electrolysis, therefore the electrolysis is simulated as separator that separates hydrogen from oxygen in water. The energy consumption is calculated outside of Aspen Plus. The energy consumption is 4.3 kWh/Nm3 H2. The total power consumption is 44 MW to produce the oxygen stream required.

3.2 Oxygen utilisation The oxygen produced in the electrolysis shall support the needs in both the gasifier and the tar cracker. The need for the total system is 0.324 kg oxygen for every kg product gas from the gasifier, see figure 20 [8].

3.3 Size reduction The biomass has to be ground prior to gasification. The Bubbling Fluidised Bed, BFB, gasifier needs a particle size of 5 mm. The size reduction is done by a hammer mill. The following function describes how much electricity is needed to reduce the biomass to a certain particle size for any kind of wood [8].

75.0*06.16 −= xy y = electricity needed in kWe/MWth

1

x = the size of the reduced particles

Gasifier Tar Cracker

1 kg

Oxygen 0.324 kg

56.5 % of oxygen stream

43.5 % of oxygen stream

Figure 20. Oxygen utilisation

28

To grind the biomass to a particle size of 5 mm, an electricity use of 0.0048 MWe/MWth is needed. Since the energy content in the biomass is 59.1 MW the electric output needed to reduce the size to 5 mm will be 0.284 MWe/kg biomass.

Figure 21. Mass and energy balance over the grinder 1 MWth: Megawatt thermal. Overall power in megawatt MWe: Megawatt electric. Electric output in megawatt

3.4 Drying The biomass feedstock has a moisture content of 35 % and must be dried to a moisture content of 15 %. The exhaust gas from the PSA combustion is used for drying the biomass. See appendix D.

3.5 Feeding system A lock-hopper system was chosen at first, but since the lock-hopper consumes a lot of inert gas an alternative for the lock-hopper was chosen, the piston feeder. The piston feeder would require 8 kWe/MWth. [16] Energy use for the piston feeder: 8 kWe/MWth*59.1 MWth = 472.8 kWe

Figure 22. Mass and energy balance over the feeding system

29

3.6 The gasifier A bubbling fluidised bed gasifier is used to gasify the biomass. The pressure is 30 bars and pure oxygen is used in the gasifier. The table below shows the composition of the mass flows in and out from the gasifier. The data is scaled from [14]. Table 5. Mass flows and mass fractions for the bubbling fluidised bed gasifier at 30 bars with pure oxygen as an oxidative medium. Component In (kg/s) Out (kg/s) Mass fraction (%) Biomass 3.73 Pure oxygen 0.796 Steam 1.19 H2 0.0737 1.29 H2O 1.634 28.6 CO 0.75 13.17 CO2 2.54 44.5 C2H4 0.0708 12.4 O2 0 0 C10H8 0.048 0.84 H2S 0.0267 0.468 N2 0.0377 0.66 CH4 0.49 8.6 C6H6 0.038 0.67 Total 5.714 5.714

Figure 23. Mass and energy balance over the gasifier The efficiency of the gasifier is calculated as the energy content in the product gas divided by the energy content in the biomass based on the lower heating value of willow wood. See appendix C.

Energy content biomass input: MWs

kg

kg

MJ1,5973,383,15 =⋅

Energy content in productgas: MW 48,1

30

The efficiency of the gasifier: 81.01.59

1.48 ==η

The rest of the energy is present as sensible and latent heat in the product gas. [8]

3.7 Tar Cracker Heat balance and mass balance In order to raise the temperature inside the tar cracker, methane must be combusted with oxygen. The mass flow in and out is assumed to be equal. There are two general combustion reactions taking place in the cracker:

OHCOOCH 2224 2 +→+ (VIII)

OHCOOtar 222 +→+ (IX) To implement the calculations over the cracker a Gibbs reactor is simulated in Aspen Plus. The reactor reaches an equilibrium state at a temperature of 1290 °C. Table 6. Mass flows and energy in and out of the tar cracker Stream In (from gasifier) Oxygen stream Out Mass Flow (kg/s) 5.75 1.036 6.790 Temperature (°C) 850 90 1300 Pressure (bar) 30 30 30 H2 (kg/s) 0.0742 0 0.1488 H2O (kg/s) 1.645 0 2.235 CO (kg/s) 0,7576 0 2.250 CO2 (kg/s) 2.5597 0 2.091 C2H4 (kg/s) 0.07132 0 0 O2 (kg/s) 0 1.036 0 Tar (kg/s) 0.0483 0 0 H2S (kg/s) 0.0269 0 0.0269 N2 (kg/s) 0.0380 0 0.0380 CH4 (kg/s) 0.4947 0 0 C6H6 (kg/s) 0.0385 0 0 Enthalpy (kJ/kg) -7198.0 60.0 -6151.6 Enthalpy (MW) 41.42 0.062 41.77 The tar cracker fulfils its general assignment, removing tar from the product stream. The tar cracker also removes other hydrocarbons like methane and benzene. This is a big advantage because it eliminates the need of a reformer. According to the volume of the stream and a residence time of 10 seconds the tar cracker must at least be 14.6 m3.

3.8 Desulphurisation In the desulphurisation H2S and COS gets adsorbed by a zinc-oxide bed. In Aspen Plus there is no good unit operation alternative for simulation of an adsorption bed, therefore the desulphurisation have been simulated as a separator. The separator is simulated to

31

separate the sulphuric components from the feed. No reaction occurs and therefore no heat is needed to be added or removed. The total volume of the zinc-oxide bed is 69.44 m3 and is based on a H2S concentration of 100 ppm.

3.9 Water Gas Shift The water gas shift reactors assignment is to reduce the amount of carbon monoxide with water (steam) to produce hydrogen and carbon dioxide. Steam is not added to the reactors because there is steam in the product gas from the tar cracker. The reaction which occurs is exothermal and therefore no heat is added to the reactors.

3.9.1 High Temperature Shift To implement the calculations over the high temperature shift reactor a Gibbs reactor is simulated in Aspen Plus. The reactor reaches an equilibrium state at a temperature of 490 °C. Because of the generated heat in the reactor, the feed doesn’t have to have the reaction temperature. To determine the feed temperature a specification have been set in Aspen Plus so the net duty in the high temperature shift should be 0 W. Table 7. Mass flows and energy in and out of the high temperature shift reactor. Stream In Out Mass Flow (kg/s) 6.76 6.76 Temperature (°C) 346 500 Pressure (bar) 30 30 H2 (kg/s) 0.1488 0.246 H2O (kg/s) 2.235 1.367 CO (kg/s) 2.250 0.901 CO2 (kg/s) 2.091 4.211 C2H4 (kg/s) 0 0 O2 (kg/s) 0 0 Tar (kg/s) 0 0 H2S (kg/s) 0 0 N2 (kg/s) 0.0380 0.03796 CH4 (kg/s) 0 0 C6H6 (kg/s) 0 0 Enthalpy (kJ/kg) -7744.177 -7990.88 Enthalpy (MW) 52.376 54.044 Because of the thermal reaction the feed only needs to be heated to 346 °C before entering the high temperature shift reactor. The high temperature shift reactor reduces the amount of carbon monoxide and water significantly but there are still large amounts of the two components left in the product gas. Using a GHSV of 2000 gives a catalyst bed volume of 12.96 m3.

3.9.2 Low Temperature Shift For the low temperature shift reactor the same assumptions have been made as for the high temperature shift reactor, but in this case the reactor reaches an equilibrium at

32

215 °C. As before the reaction that occurs is exothermal and therefore the feed does not have to be heated to the reaction temperature. To calculate the temperature of the feed a specification have been set in Aspen Plus so the net duty in the reactor is 0 W. Table 8. Mass flows and energy in and out of the high temperature shift reactor. Stream In Out Mass Flow (kg/s) 6.790 6.790 Temperature (°C) 158.5 225 Pressure (bar) 30 30 H2 (kg/s) 0.246 0.3040 H2O (kg/s) 1.367 0.8483 CO (kg/s) 0.901 0.0946 CO2 (kg/s) 4.211 5.478 C2H4 (kg/s) 0 0 O2 (kg/s) 0 0 Tar (kg/s) 0 0 H2S (kg/s) 0 0 N2 (kg/s) 0.03796 0.03796 CH4 (kg/s) 0 0 C6H6 (kg/s) 0 0 Enthalpy (kJ/kg) -8317.094 -8647.458 Enthalpy (MW) 56.250 58.485 The feed can have the temperature 159 °C and still reach the reaction temperature because of the exothermal reaction that occurs. After the low temperature shift the amount of carbon monoxide is almost down to 1 wt%. . Using a GHSV of 2000 gives a catalyst bed volume of 8.28 m3.

3.10 Heat exchangers and power generation Heat exchangers are utilised in order to optimise the energy flow in the plant. The most important exchangers are presented in table 8. Table 9. Heat exchangers Exchanger Duty (MW) Area (m2) After tar cracker 12.3 358.2 After desulphurisation 0.53 12.26 After HT-shift 4.5 964.6 After LT-shift 3.9 390.5 The heat exchangers after the tar cracker and after HT-shift are used to produce high-pressure (HP) steam at 85 bars. The steam is led to a high-pressure turbine where it is expanded to 30 bars to produce electricity. A part of the steam is led to the gasifier and the remaining part is expanded to 50 mbar in order to produce even more electricity. To simulate the heat exchangers a shortcut-model is used in Aspen. All heat exchangers

33

operate counter currently. The heat transfer coefficient is approximated to 0.85 kW/m2°C when heat transfer is between liquid-liquid, 0.05 kW/m2°C for gas-liquid and 1.1 kW/m2°C for boiling liquid-gas.. The turbines are simulated isentropic with a efficiency of 0.72 and a mechanical efficiency of 1. The HP-turbine generates 1.76 MW work and the LP-turbine generates 5.19 MW.

3.11 Alternative power generation The price of electricity is very variable and is often growing. Because of this alternative ways of power generating has been investigated. Biomass can be used in a power plant. The biomass gets combusted and heats water to steam at a pressure of 200 bars. The superheated steam then expands over a turbine and power is generated. The turbine works isocratic with an efficiency of 72 %. [17] To generate 1 MW electric power the needs are 0.358 kg willow wood per second. The electricity price with a biomass plant would be US$ 0.080 /kWh and with the current electricity price (US$ 0.074 /kWh), it is cheaper to purchase electricity. See appendix E for the calculations.

34

3.12 Summary

The results of the calculations regarding the major parts in the plant are summarized in table 10. Table 10. Equipment list.

The energy balances for the process were summarised and an energy flow diagram was created, see figure 24.

Item Design Materials – Vessel/Catalyst/Adsorbent

Condition (ºC/bar) (Hot-Cold)

Gasifier 74 MW -/-/- 850 / 30 Cracker 14.6 m3 -/-/- 1300 / 30

Electrolysis 41.99 MW -/-/- 80 / 1 Desulphurisation 43.65 m3 stainless steel/-/ZnO-Co/Mo 300 / 30

HT-shift 12.96 m3 stainless steel/Fe-CrO/- 500 / 30 LT-shift 8.28 m3 stainless steel/Cu-ZnO/- 225 / 30 Scrubber 1.383 m3 stainless steel/-/- 50 / 30

PSA 19.24 m3 stainless steel/-/Alumina-Active carbon-Zeolit

50 / 30

Steam boiler 8.48 MW stainless steel/-/- 850 / 1 Heat exchanger

HEX1 29.6 m2 / 12.1 MW

-/-/- 1300 / 30 - 300 / 30 25 / 85 - 500 / 85

Heat exchanger HEX2

0.7 m2 / 0.52 MW

-/-/- 1300 / 30 - 1070 / 30 300 / 30 - 346 / 30

Heat exchanger HEX3

82.9 m2 / 4.33 MW

stainless steel /-/- 500 / 30 - 162 / 30 25 / 85 - 300 / 85

Heat exchanger HEX4

313.7 m2 / 4.03 MW

stainless steel /-/- 225 / 30 - 50 / 30 25 / 85 - 144 / 85

Hydrogen compressor 2.89 MW -/-/- 80 / 1 – 897 / 30 Oxygen compressor 1.38 MW -/-/- 80 / 1 – 797 / 30 Product compressor 3.76 MW -/-/- 50 / 30 - 100 / 850

High pressure turbine 1.76 MW -/-/- 462 / 85 - 352 / 30 Low pressure turbine 5.19 MW -/-/- 352 / 30 - 33 / 0.05

35

Process

Biomass 59,1 MW

Water from heat

exchangers 0,9 MWElectricity 47,3 MW

Hydrogen 57 MW

Water 1,24 MW

Losses 48,93 MW

Exhaust gas 0,81 MW

Figure 24. Energy flow diagram over the process The process has a total efficiency of 54.5%. There are major losses in the process. The largest energy consumption comes from the electrolyser and the efficiency of the gasifier is only 81 %. It would be beneficial for the plant economy to use the outgoing heat exchanger water for district heating purposes, but unfortunately most of these streams are of low temperature, making this operation unprofitable. There are five major units which produce hydrogen, the gasifier, the tar cracker, the high temperature shift, the low temperature shift and the electrolyser. The pressure swing adsorption unit has a negative influence on the production because 20 % of the hydrogen is lost during purification. The distribution of the hydrogen production is displayed in figure 25.

36

-0,200

-0,100

0,000

0,100

0,200

0,300

0,400

0,500

0,600

Fra

ctio

n o

f th

e to

tal

amo

un

t o

f p

rod

uce

d h

ydro

gen

Gasifier Tar cracker HT-shift LT-shift PSA TotalGasification

Electrolyser

Process unit

Distribution of hydrogen production

Figure 25. Hydrogen production distribution The electrolyser is the single largest hydrogen producing unit, but the total hydrogen produced in the gasifier process is 44.8% of the total. There are several electricity consuming units. They are all summarized in figure 26.

-10

-5

0

5

10

15

20

25

30

35

40

45

Eff

ect (

MW

)

Electrolysis O2 comp H2 comp Grinder Feeder Pumps Productcomp

Electricitygenerated

Process unit

Electricity usage

Figure 26. Distribution of the electricity usage

37

4 Economic evaluation In order to create a wider picture of the hydrogen production plant, an economic evaluation is carried out. The vast majority of the costs of importance are estimated through a few methods, including the Ulrich method and a scale method based on a literature reference [14]. The module cost (including equipment, apparatus, piping, installation, transportation), the contracting and contingency and the cost for auxiliary facilities are considered capital costs. The operating costs break down into fixed capital (raw material, product and spare parts in storage), direct operating cost (electricity, staff, raw material, maintenance and waste treatment) and indirect operating cost (overhead, administration and R&D). Adding the above mentioned costs and using an economic life span of 15 years and an interest rate of 10 %, the total cost of the plant is estimated. To make this number comparable, the hydrogen production price per kg is calculated.

4.1 Methods The method used to estimate the direct and indirect cost for each process unit is primarily the Ulrich method. When the Ulrich method does not provide the required information, the scale method is used.

4.1.1 Ulrich method The Ulrich method consists of databases for equipment cost (CP). Adjustments for the real conditions are made by different adjustment factors; FBM for installation etc., FP for pressure, FM for material. The purpose is to calculate a module cost (CBM, se equation X, XI) for each unit. This value contains the direct and indirect construction cost for each unit. CBM = CP*FBM (X) CBM = CP*FP *FM (XI) CP is the purchased equipment cost, expressed in US$(1982), for standard conditions regarding material (carbon steel), temperature and pressure. The module costs are updated with the Swedish net price index for respective year, NPI , and the exchange rates

(SEK/US$) for respective year , KKV , to adjust for the change in currency value over the years. All prices are expressed in US$ (2007) (see equation XII, XIII ).

fCC BMUS *)2007$( = (XII)

1982

2007

1982

2007

)(

)(

)(

)(

KK

KK

NP

NP

V

V

I

If = (XIII)

38

The module cost includes cost contribution from piping and instrumentation, materials and labour for installation, buildings, site development, freight, insurance, construction overhead and engineering [18]. Table 11. Price updating data [19] Net price index 1982 (1980=100)

1982)( NPI 121.67

Net price index 2007 (1980=100)

2007)( NPI 243

Net price index 2002 (1980=100)

2002)( NPI 234.38

Exchange rate 1982 (SEK/US$)

1982)( KKV 7

Exchange rate 2007 (SEK/US$)

2007)( KKV 6

Exchange rate 2002 (€ /US$)

( )2002KKV 0.88

4.1.2 The scale method The scale method is based upon known cost for major equipment given by experts or found in literature [14]. The components investment costs depend on its size and are calculated by scaling from a known scale and cost (see equation XIV). The prices are given in M€(2002).

IS

SCC

R

**0

10€(2002)

= (XIV)

were C0 and C1 are the cost at known scale and desired scale respectively, S0 and S1 are the known scale and desired scale respectively, R is the power scaling factor or scale exponent and I is the overall installation factor. This method is used when Ulrich method cannot provide the desired information. Similar to the Ulrich method the costs are updated with net price index and change rates into US$ (2007) (se equation XV, XVI).

fCCUS *€(2002))2007$( = (XV)

( )20022002

2007 *)(

)(KK

NP

NP VI

If = (XVI)

where ( )2002KKV is the exchange rate for € to US$.

39

It is assumed that the factored estimation method takes in consideration the same cost contributions as Ulrich method and therefore the total costs from the different methods are comparable.

4.2 Capital cost The capital costs are mainly estimated through previously mentioned methods. The detailed calculations are shown in Appendix F. The results for the bare module costs of each item are summarised in table 12. In table 13, contracting and contingency and auxiliary facilities are added on.

40

Table 12. Bare module cost for each item in 2007 US$ ITEM ITEM COST (US$) (US$) % OF PLANT COSTS Biomass pretreatment 7 565 263 Grinding 108 867 0.1 Drying 7 448 693 3.9 Feeder 7 703 0.0 Gasifier units 16 068 464 Gasifier 16 068 464 8.4 Heat exchangers 4 427 879 HEX1 3 154 475 1.6 HEX2 515 782 0.3 HEX3 757 621 0.4 HEX4 517 843 0.3 After-treatment from gasif. 11 019 291 Cracker 601 595 0.3 Cyclone 1 37 662 0.0 Particle filter 581 741 0.3 Cyclone 2 10 271 0.0 Desulphurisation 1 006 592 0.5 HT shift 1 372 906 0.7 LT shift 769 156 0.4 Scrubber 254 216 0.1 PSA 6 977 166 3.6 Electrolysis 80 373 282 Hydro electrolysis system 73 923 948 38.5 Oxygen compressor 2 278 344 1.2 Hydrogen compressor 4 170 991 2.2 Power generation 8 257 652 ST1 718 994 0.4 ST2 4 534 290 2.4 Steam boiler 3 004 368 1.6 Pumps 235 385 Pump 1 141 231 0.1 Pump 2 94 154 0.0 Pump 3 17 975 0.0 Pump 4 103 569 0.1 Storage vessels 7 803 373 Biomass 21 570 0.0 H2 256 784 0.1 Product compression 7 525 019 3.9 TOTAL MODULE COST 136 981 988

41

Table 13. Summary of the capital cost Costs (US$) Per cent of capital cost Module cost 136 981 988 71.4 Contracting & contingency 20 638 965 10.7

Auxiliary facilities 34 398 275 17.9 TOTAL CAPITAL COST 191 774 784

4.3 Operating cost The operating cost are mainly estimated through rules of thumb and at times from current prices. Table 14 summarises the costs. See appendix F for detailed calculations. Table 14. Summary of the operating costs ITEM ANNUAL COST (US$) % OF ANNUAL OPERATING COST Material storage Biomass 8 132 0.0 H2 13 824 0.0 Spare parts 687 966 1.4 FIXED CAPITAL 709 922 1.4 Raw material Biomass 6 802 403 13.7 Water 82 451 0.2 Desulphurisation adsorbent 6 111 121 12.4 Electricity 25 271 228 51.3 Operating staff 1 028 571 2.1 Management 154 286 0.3 Lab work 154 286 0.3 Maintenance 6 879 655 13.9 Waste treatment 127 047 0.3 DIRECT OPERATING COST 46 531 689 94.4 Overhead for staff 874 286 1.8 Administration 218 571 0.4 R&D 848 832 2.0 INDIRECT OPERATING COST 1 941 689 4.2 TOTAL OPERATING COST 49 295 834

42

4.4 Investment analysis By the use of the annuity method, the Depreciation of Capital Cost is calculated (see Appendix F) and the sum of this and the annual operating cost make the Total annual production cost. Using the annual production volume, the production cost for the hydrogen product is calculated (table 15). In table 16, the base case is used (see table 15). Table 15. Investment analysis – annual costs Depreciation of Capital cost (US$) 25 213 355

Operating cost (US$) 49 295 834

Annual H2 production (kg) 13 884 902

Production cost (US$/kg H2) 4.97

Table 16. Base case Electricity price (US$/kWh) [19] 0.074 Vkk,current 6 Interest rate 10 % Economic life span 15 years Hydro electrolyser system cost (MUS$/unit) 0.75

4.5 Sensitivity analysis A few key costs in the economical analysis are believed to have a significant impact on the financial feasibility of the hydrogen plant. By varying the magnitude of some of these, a picture of the sensitivity of the plant is created.

4.5.1 Electricity price The plant is a heavy user of electricity due to the electrolysis – despite energy saving arrangements such as steam turbines; the net electricity use is about 40 MW. Electricity represents 58.6 % of the operating costs. Looking at figures of the recent increases in electricity price [22], it would appear that it is an aspect to take into account. See diagram 1.

43

Sensitivity analysis - electricity price

3,50

4,00

4,50

5,00

5,50

6,00

6,50

7,00

7,50

0,02 0,04 0,06 0,08 0,10 0,12 0,14 0,16 0,18

Electricity price (US$/kWh)

Pro

du

ct c

ost

(U

S$/

kg)

Diagram 1. Sensitivity analysis – electricity price. The circle indicates the hydrogen production cost in the base case.

4.5.2 Interest rate The interest cost for the loan is the dominating operating cost, with almost half of the annual cost. A change of the interest rate would probably have effects on the product price and hence, it is varied between 5 and 25 %. See diagram 2.

Sensitivity analysis - interest rate

3,50

4,00

4,50

5,00

5,50

6,00

6,50

7,00

7,50

0,00 0,05 0,10 0,15 0,20 0,25 0,30

Interest rate

Pro

du

ct c

ost

(U

S$/

kg)

Diagram 2. Sensitivity analysis – interest rate. The circle indicates the hydrogen production cost in the base case.

44

4.5.3 Electrolyser unit cost Since the electrolyser system cost alone is 38.5 % of the direct plants costs, it is necessary to analyse a variation of the electrolyser unit price. Depending on the customer demands, a cleaner or more impure product might be of demand, and thus more or less advanced (expensive) equipment is needed. For example, an electrolyser which produces hydrogen with a purity of 99.998 % costs 0.75 MUS$, whereas the corresponding price for 99.8 % purity is 0.59 MUS$. In the projected plant 98 units are needed, making the balance even more sensitive to a price variation. See diagram 3.

Sensitivity analysis - electrolyser unit cost

3,50

4,00

4,50

5,00

5,50

6,00

6,50

7,00

7,50

0,4 0,5 0,6 0,7 0,8 0,9 1,0 1,1

Electrolyser price (MUS$/unit)

Pro

du

ct c

ost

(U

S$/

kg)

Diagram 3. Sensitivity analysis – electrolyser unit cost. The circle indicates the hydrogen production cost in the base case.

4.5.4 Biomass cost Since the raw material cost is often important in chemical plants, a sensitivity analysis of the biomass cost was carried out. See diagram 4.

45

Sensitivity analysis - biomass cost

3,50

4,00

4,50

5,00

5,50

6,00

6,50

7,00

7,50

40 60 80 100 120

Biomass cost (US$/tonne)

Pro

du

ct c

ost

(U

S$/

kg)

Diagram 4. Sensitivity analysis – biomass cost. The circle indicates the hydrogen production cost in the base case.

4.5.5 Summary of the sensitivity analysis The sensitivity analysis over the factors above are summarised in diagram 5.

3,50

4,00

4,50

5,00

5,50

6,00

6,50

7,00

7,50

-80 -40 0 40 80 120

Increase/decrease (%)

Pro

du

ct c

ost

(U

S$/

kg)

Biomass

Electricity

Interest rate

Electrolyser

Diagram 5. Summary of the sensitivity analysis

46

Discussion and conclusion As the need for renewable energy is increasing, the research for alternative more environment friendly and efficient solutions for energy generation and fuel production is of high importance. In this project a feasibility study on a hydrogen plant has been conducted. The hydrogen plant combines gasification of biomass and electrolysis, eliminating the need of an oxygen plant, which is otherwise needed for oxygen gasification. The electrolyser system however requires a great amount of electricity. Parts of the electricity can be produced from the excess heat generated in the gasification system but the majority has to be purchased. This raises the issue of where this electricity is derived from – if it has been generated from fossil fuels, the environmental advantages of this process are reduced to a great extent. An on-site biomass power plant could be a CO2 neutral solution. The estimation of the biomass cost for electricity produced this way however indicates that it is more expensive per kWh than the purchased electricity. The capital investment and operating costs other than the raw material must also be taken into account, and it is therefore not feasible with this arrangement with the current electricity price. According to the project specifications, the product should be pressurised to 850 bars – this leads to large costs in the compressors, both in the module and operating costs. There is also the question of purity; if the product for example is to be utilised in Fischer-Tropsch processes, the demands might be different. If the purity demands were lower, some purification steps could be removed. A different electrolyser could also be chosen. Since the cost of the electrolyser system is the dominating module cost (roughly 50 %) it is worth studying the variation of this in a sensitivity analysis. The analysis shows that the hydrogen price is reduced about 10 % if the system with lower purity is used instead of the one with higher purity. Another discussion point is the choice of atmospheric electrolysers instead of pressurised. The pressured system would be more expensive than the atmospheric but the latter requires additional compressors, which increases the module plant cost further. The combined power consumption of the atmospheric system with the extra compressors is higher than the pressurised, which makes the second alternative more advantageous in the long run. It is probably worth looking into the further development of pressurised electrolysers. The energy flow diagram shows that there are large energy losses – the gasifier is the major contributor, losing 19 % of its heat. A more efficient gasifier could be the solution to this. There is also a lot of hot heat exchanger water, which indicates a possibility of utilisation in district heating, and hence improving the process economy. Calculations however show that most of the water is of low exergy content and it is not worthwhile to exploit this energy as district heating.

47

The estimated production price is $ 4.97 /kg H2. This can be compared to the production cost in the literature which ranges from $ 1.91 /kg H2 for hydrogen made from gasification of coal and shipped by pipeline, to $ 6.58 /kg H2 for on-site electrolysis. The electricity price is the largest operating cost, making up for over 50 % of the total operating costs. The sensitivity analysis shows that the production price is very sensitive to fluctuations in the electricity price, varying 20¢/kg per 1¢/kWh difference, i.e. the production cost is US$ 3.74 /kg instead of US$ 4.97 /kg (25 % lower) at an electricity cost of 2.1¢/kWh instead of the base case cost of 7.4¢/kWh. It can therefore be concluded that the plant economy is very much dependent on how the electricity is obtained – if it for example were situated in a country with access to cheap hydroelectric power such as Norway, its feasibility is increased. Alternatively, the plant could be built without the electrolysers and instead rely on gasification alone. In that case the electricity consumption would decrease considerably. But as the gasifier requires the oxygen produced in the electrolyser, a separate unit for the oxygen production would be necessary. Despite that, this might be a more profitable solution with the current electricity prices. The sensitive analysis of the interest rate indicates that this is somewhat important; the production price of the hydrogen increases to US$ 5.43 /kg from US$ 4.97 /kg (+9 %) if the interest rate would increase by 50 %. A sensitivity analysis of the biomass cost was also carried out. It shows that the hydrogen production cost is just as insensitive to changes in the biomass cost as it is for variation of the interest rate; a doubling of the base case cost of biomass only increases the production cost to US$ 5.42 /kg compared to US$ 4.97 /kg (+9 %). To summarize, the projected plant shows that there is a potential of biomass gasification as an alternative method of producing hydrogen. The use of electrolysis to replace the oxygen plant and to cover roughly half of the hydrogen production is less feasible since the electricity prices are high and still on the rise. If this process design is to be used, a low and stable electricity price is necessary to accomplish a steady production price. In all cases, a more energy efficient plant in terms of heat losses and energy recycling is of importance.

48

5 References 1. www.hydrogenics.com, 2007-01-15

2. Ullmann’s Encyclopedia of Industrial Chemistry (Electrochemistry:) Article Online Posting Date: January 15, 2003

3. Ullmann’s Encyclopedia of Industrial Chemistry (Hydrogen:) Article Online Posting Date: June 15, 2000

4. www.elektrolyse.de, 2007-02-21

5. http://www.directindustry.com/industrial-manufacturer/hydrogen-generator-73995.html, 2007-02-21

6. www.hydro.com/electrolysers/en/ 2007-04-19

7. A Critical Analysis of Biomass to Fuel Concepts Glans, Martin October 2005

8. Förstudie gällande en förgasningsanläggning för produktion av fordonsbränsle från biomassa Vlassiouk, Igor 2005

9. http://www.barr-rosin.com/english/products/super-heated-steam-drying.htm, 2007-02-18

10. http://www.niro.co.uk/ndk_website/NUKF/CMSDoc2.nsf/WebDoc/ndkk5hvee2, 2007-02-21

11. A Feasibility Study of Production of Diesel/Petrol from Biomass Glans, Martin; Hoffstedt, Christian; Pallin, Axel; Ström, Lise-Lott; Suki, Anna; Vlassiouk, Igor 2005-05-25

12. http://www.auri.org/news/ainapr06/images/chart.jpg, 2007-02-18

13. www.gtp-merichem.com/images/art/syngas_fig1_2.gif, 2007-02-18

14. Production of FT transportation fuels from biomass; technical options, process analysis and optimisation, and development potential Hamelinck, Carlo N; Faaij, André P.C; den Uil, Herman; Boerringter, Harald March 2003

15. Ullmann’s Encyclopedia of Industrial Chemistry (Gas Production:) Article Online Posting Date: December 15, 2006

49

16. Entrained flow gasification of biomass; Ash behaviour, feeding issues, and system analyses Boerrigter, H.; Coda, B.; Cieplik, M.K.; Hemmes, K.; van der Drift, A. 27 April 2004

17. Teknik för elproduktion Stenström, Stig; Institutionen för Kemiteknik Lunds Universitet, 2004

18. ProjekteringsHandboken 2007 Karlsson, Hans T., Institutionen för Kemiteknik, 2007

19. Statistiska centralbyrån, www.scb.se

20. http://bioenergiportalen.agriprim.com/?p=1590

21. http://www.svo.se/minskog/templates/svo_se_vanlig.asp?id=8401

22. http://www.chem.uu.nl/nws/www/publica/e2003-08.pdf

23. Prospects for Bioenergy in Europe - Supply, Demand and Trade Ericsson, Karin, Lund University 2006

24. Hydrogen Project – Production of Hydrogen from Renewable Resources Hulteberg, Christian; Jönsson, Christopher; Lindblom, Henrik; Olsson, Henrik; Åkesson, Jessica Lund Institute of Technology, May 2002

25. A Feasibility Study on Microbial Hydrogen Production Andersson, E.; Andersson M.; Burström, H.; Goitom, A.; Malmborg, A.; Svärd, T.

26. Rasten, Egil (Norsk Hydro), E-mail conversation, April 17, 2007

27. Data och diagram – Energi- och kemitekniska tabeller Mörtstedt, Sten-Erik; Hellsten, Gunnar

28. Projekteringsmetodik Karlsson, Hans T., Institutionen för Kemiteknik, 1992

29. Public perception related to a hydrogen hybrid internal combustion engine transit bus demonstration and hydrogen fuel Energy Policy 35 (2007) 2249–2255 Hickson, Allister; Phillips, Al; Morales, Gene

30. The characteristics of inorganic elements in ashes from a 1 MW CFB biomass gasification power generation plant Fuel Processing Technology 88 (2007) 149–156 Liao, Cuiping; Wu, Chuangzhi; Yan, Yongjie:

31. Thellin, Gunnar; Personal communication, April 26, 2007

50

Ap

pen

dix

A

Generator

The

rmal

tar

crac

ker

1300

°C

300 °C

Particles

Z

nO

346 °C

FeC

rO

HT-shiftDesulfurization

162 °C

CuZ

nO

LT-shift

Filter

Steam 500 °C85 bar

Generator

StorageBiomass

Grinding

Cyc

lone

Dry

ing

250

°C

Feeding

Gas

ifier

BF

B 8

50 °

C

Water Deionisation Electrolysis

KOHH2O

Deoxidastion

O2

H2 + H2O

35 wt%

15 wt%

AshesAir

Dry

er

Scrubber

HCLNH3mm

PSA

H2

Biproductsunreacted

4,8 kW/kg H2

Hydrogen Storage

450 °C

30 bar

Combustion

Air

H2O 85 bar

Steam mix

Steam splitter

H2O H2O

H2O

H2O

Crack splitter

Crack Mixer

Boiler

51

Appendix B

Q Duty (kW)

W P ower(Wat t)

CRACKER

Q=-83

VVXCRACK

Q=12300

HPTURB

W=- 1760240

DESULPH

HT-SHIFT

Q=9

LT-SHIFT

Q=-0

SCRUBBER PSA

ELEKTRO

H2MIX

VVXCRAPU

W=66 978

VVXHTPU

W=39196

PSACOMBU

Q=7314

STEAMMIX

STEAMSPL

VVXHTSHI

Q=4460

LPTURB

W=-5191610

CRACKSPL

CRACKMIX

VVXDESUL

Q=525

H2COMP

W=3710030

VVXLTSHI

Q=3919

VVXLTPU

W=1380

VVXPSA

Q=8642

VVXPSAPU

W=53022

O2ELSPLI

H2ELCOMP

W=1814940

O2ELCOMP

W=921513

52

Appendix C Table I. Heat value for different biomass Biomass HHV (MJ/kg) Heating value

(MWh/tonne) Moisture (%)

Branches and tops of trees [20]

9.5 2.6 45

Branches and tops of trees [21]

11.52 3.2 35

Dry chips [20] 12 4.6 12 Bark chips [20] 7.3 2 55 Saw dust [20] 8.4 2.3 50 Salix chip [20] 7.9 2.2 50 Milled peat [20] 9.5 2.6 50 Wood pellets [20] 16.8 4.7 11 Fuelwood [20] 13.8 3.8 25 Straw [20] 14.4 4 15 Reed phalaris [20] 14.3 4 14 Grain [20] 15 4.2 11 Hemp [20] * * * Oil [20] (Eo1) 42.7 11.9 <0.01 Willow wood [22] 19.88 5.52 15

53

Appendix D: Drying The heat needed for drying the biomass is taken from exhausted gas from the PSA combustion. Massflow biomass: 3.73kg/s (15% moisture)

Massflow biomass: skg /88.465.0

85.0*73.3 = (35% moisture)

Water removed OHm2

: 4.88-3.73 = 1.15kg/s

x1 = 0.35 x2 = 0.15 T1 = 10°C H1 = 41.8 kJ/kg H2 = 2675.25 kJ/kg The gas stream: 0.5 kg H2O in 17 kg total mluft = 16.5kg/s

xin = kgkg /0303.05.16

5.0 =

T = 424°C Hair,in = Cpair*T+x in*∆Hvap+x*Cpsteam*T

Cpair = 1 Ckg

kJ

°*

∆Hvap = 2501 kJ/kg

Cpsteam = 1.855 Ckg

kJ

°*

⇒ Hair,in = 523.6 kJ/kg 1.15=16.5*(x3-0.303) ⇒ x3 = 0.099 Tluft,ut = 255°C 523.6 = 1*T+0.1*2501+0.1*1.855*T ⇒T = 230.7 °C

Energy needed in the dryer: =−

=1000

* 1222

HHmE OHOH 3.03 MW

54

Appendix E: Alternative power generation

H1=3395 kJ/kg Enthalpy for superheated steam at 200 bars H2,S = 2533 kJ/kg Isentropic enthalpy for saturated steam at 0.02 bars η = 0.72 Isentropic efficiency H2 = 2774 kJ/kg Enthalpy for saturated steam at 0.02 bars h3 = 73.5 kJ/kg Enthalpy for water at 0.02 bars

•−

⋅=1000

)(1 21

2

HHmsMJ OH ()

kg/s61.12

=⇒ OHm

1000

)(15.085.0 01

,

HHHH biomasscalbiomass

−⋅−⋅=

Hcal,biomass = 19.9 MJ/kg Calometric heating value for willow wood H0 = 111.86 kJ/kg Enthalpy for water at 25°C

MJ/kg42.16=⇒ biomassH

Losses in furnace = 9 %

)(91.0 212HHmmH OHbiomassbiomass −⋅=⋅⋅

Combustion Biomass

Generator

Turbine

Condensor

Water 0.02 bars

2

3

Superheated steam 200 bars, 550 °C

Saturated steam, 0.02 bars

1

55

kg/MW358.0kg/s358.0 ==⇒ biomassm

Table II. Feasibility calculation for biomass power plant Biomass cost (US$/kg) [23] 0.063 Biomass need (kg/kWh) 1.29 Cost US$/kWh 0.081

56

Appendix F: Cost estimations All design data are taken from Aspen Plus calculations unless otherwise indicated.

Pre-treatment

Grinding The power consumption for the grinder is about 179 kWh/tonne [24]. Apparitional cost is estimated using Ulrich method for hammer mills [table 16.32, 19]. See table III. Table III. Grinding Power consumption kWh/t 179.00 Flow (tonnes/h) 13.4 Cp 22712 Fbm 2.8 Power consumption kWh/year 2741257 App cost (US$) 1982 63595

App cost (US$) 2007 108867

Drying The apparitional cost for the dryer is estimated using the scale method [table 6, 10]. See table IV. Table IV. Drying Mass flow (tonnes/h) 13.4 Bas cost c0 (M€) 8.5 Base scale s0 33.5 Scale factor R 0.8 Overall installation factor 2.0 Cost (M€ 2002) 8.16766 Cost (US$) 2002 7187529

Cost (US$) 2007 7448693

Feeder The apparitional cost for the feeder is estimated using Ulrich method for a screw feeder [table 16.32, 19] due to lack of information about the piston feeder. The effect needed is 593.2 kW [16]. See table V.

57

Table V. Feeding Biomass flow rate (tonnes/h) 13.4 Biomass density (tonnes/m3) 0.146 Biomass flow rate (m3/h) 91.78 Biomass flow rate (ft3/h) 3241.21 Cp (US$ 1982) 4500

Cost (US$ 2007) 7703 Effect need (kW) 593.2 Power consumption (kWh/year) 4745600

Gasifier The cost for the gasifier is estimated using the scale method for a CFB gasifier [table 6, 10]. The constructial differences between a CFB gasifier and bubbling fluidised bed does not affect the cost significantly [18]. See table VI. Table VI. The gasifier Mass flow (tonnes/h) wet 13.4 Mass flow (tonnes/h) dry 8.71 Dry per cent (mass%) 65 Base cost c0 (M€) 44.3 Base scale s0 68.8 Scale factor R 0.7 Overall installation factor 1.7 Cost (M€ 2002) 17.6194 Cost (US$ 2002) 15505076

Cost (US$ 2007) 16068464

Heat exchangers The heat exchangers following the tar cracker and the desulphuriser operates at high temperatures (max temperature ca 1300 o C) and therefore factored estimation method for high temperature heat exchangers is used. See table VII. Table VII. Heat exchangers 1 and 2 Power duty (MW) 12.10 0.52 Base cost c0 (M€) 8.1 8.1 Base scale s0 138.1 138.1 Scale factor R 0.6 0.6 Overall installation factor 1.8 2.0 Cost (M€ 2002) 3.459 0.566 Cost (US$ 2002) 3043874 497698 Cost (US$ 2007) 3154475 515782

58

The heat exchangers following the shift reactors operate at lower temperatures and therefore Ulrich method for plate-and-frame heat exchangers is used. See table VIII. Table VIII. Heat exchangers 3 and 4 HEX3 HEX4 DeltaTL 61.4 15.1 A (m2) 965.0 390.0 A (ft2) 10387.5 4198.1 cp 340434 232690 Fp 30 bar 1.3 1.3 Cost (US$ 1982) 442564 302498 Cost (US$ 2007) 757621 517843

After-treatment

Tar cracker The cost for the tar cracker is estimated using the scale method [table 6, 10]. See table IX. Table IX. Tar cracker Temperature (oC) 1300 Pressure (bar) 30 Gas flow (m3/s) 1.46 Flow NTP (Nm3/s) 7.60 Residence time (s) NTP 10 Vreact (NTP) 76.0 Vreact @ 30 bar, 1300 oC 14.6 Temperature (oC) 1300 Pressure (bar) 30 Gas flow in (m3/s) 0.79 Gas flow out (m3/s) 1.46 Mean gas flow (m3/s) 1.13 c0 (M€) 3.6 Base scale s0 34.2 scale factor R 0.7 Cost (M€ 2002) 0.659661 Cost (US$) 2002 580502

Cost (US$) 2007 601595

Particle removal The particle removal step consists of two cyclones and a particle filter. The cost for the cyclones is estimated by Ulrich [figure 5.55, 19]. See table X.

59

Table X. Particle removal Cyclone 1: After gasifier Cyclone 2: Before scrubber Feed rate m3/s 0.777 0.247 Cp 5500 1500 Fbm (stainless steel) 4 4 Cost (US$) 1982 22000 6000

Cost (US$) 2007 37662 10271 The cost for the particle filter is estimated using the factored estimation method [table 6, 10]. See table XI. Table XI. Particle removal Feed rate (m3/s) 0.777 Base cost c0 (M€) 1.9 Base scale s0 12.1 Scale factor R 0.7 Overall installation factor 2.0 Cost (M€ 2002) 0.638 Cost (US$ 2002) 561344

Cost (US$) 2007 581741

Desulphuriser The cost of the desulphuriser consists of adsorbent cost and vessel cost. The adsorbent cost is considered an operating cost since the adsorbent bed is consumed and replaced annually. It is estimated by scaling values from a previous project [25]. The height and the gas velocity from the reference are kept and the adsorbent amount and the diameter are assumed to be proportional to the flow rate. The vessel cost is estimated using Ulrich method for vertically oriented process vessels [figure 5-44 b, 19]. See table XII.

60

Table XII. Desulphuriser Adsorbent Flow (m3/s) 0.53 G (Nm3/s) 7.58 Temp (oC) 300 Pressure (bar) 30 Source Total vol 12.57 Co/Mo vol ratio 0.032 ZnO vol ratio 0.159 Gsource (Nm3/s) 0.42 Scale factor (G/Gsource) 181.86 Adsorbent volumes Total vol (m3) 436.46 Co/Mo vol (m3) 13.89 ZnO vol (m3) 69.44 Adsorbent specs Price (SEK/kg) Density (kg/m3) Price (US$/m3) Cost (US$) Co/Mo 200 2400 68571 95238 ZnO 200 2600 74286 515874 Adsorbent cost (US$) 6111121

Vessel

Diameter (m) 6.09 Height (m) 15.00 Cp 30000.00 Fp 2.60 Fm (stainless steel) 4.00 Fbm 19.60 Cbm (US$) 1982 588000 Cost (US$) 2007 1006592

Total Desulphuriser (US$) 1617704

Shift reactors The cost for the shift reactors vessels is estimated by Ulrich method for vertically oriented process vessels [figure 5-44 b, 19]. The adsorbent and catalysts cost is estimated using values from a previous project [25]. See table XIII.

61

Table XIII. Shift reactors HT shift LT shift Pressure (bar) 30 30

Temperature (oC) 500 225

Gas velocity (m/s) 0.3 0.3

Gas flow (m3/s) 0.72 0.46

Flow NTP (Nm3/s) 7.63 7.57 Vreakt 12.96 8.28

Area (m2) 2.4 1.53

d (m) 1.75 1.40

SV (h-1) at 30 bars 2000 2000

Retention time (s) 1.8 1.8

Height (m) 6.4 6.4

Inner surface area (m2) 34.46 26.77

Cp 11000 10000

Fp (30 bars) 2.60 2.60

Fm stainless clad 4.00 4.00

Fbm 9.00 9.00

Cbm 99000 90000

Cp (US$) 2007 169477 154070

HT Shift LT Shift

Catalyst price (Fe-CrO) (SEK/kg) 250.00 Catalyst price (Cu/ZnO) (SEK/kg) 200.00

Density Fe-CrO 2600 Density Cu/ZnO 2600

Catalyst price (Fe-CrO) (US$/m3) 92857.143 Catalyst price (Cu/ZnO) (US$/m3) 74286

Catalyst cost (US$) 1203429 Catalyst cost (US$) 615086

Scrubber The cost for the scrubber is estimated Ulrich method for vertically oriented process vessels [figure 5-44 b, 19]. See table XIV. Table XIV. Scrubber Flow (m3/s) 0.461 Speed (m/s) 0.5 Area (m2) 0.922 Residence time (s) 3 Height (m) 6 Fp 20-30 bar 1.6 Fm 4 Cp 11000 Fbm 13.5 Cost (US$ 1982) 148500

Cost (US$ 2007) 254216

62

PSA The cost for the PSA-unit is estimated by Ulrich method for vertically oriented process vessels. The dimensions and the catalysts cost are calculated by scaling the values from a literature reference. The height and the gas velocity from the reference are kept and the catalyst amount and the diameter are assumed to be proportional to the flow rate. Note that the linearity is only assumed in a narrow interval, i.e. the mass flows are similar in quantity. [figure 5-44 b, 19], [25]. See table XV.

63

Table XV. PSA Pressure (bar) 30 Temperature (oC) 50 Gas velocity (m/s) 0.3 Gas flow (m3/s) 0.108 Flow NTP (Nm3/s) 2.74 Gsource (Nm3/s) 1.1389 dsource (m) 2 G/dsource 0,5694 d (m) 4,8010 h (m) (source) 4 Cp 50000 Fp 2.60 Fm 4.00 Fbm 19.60 Cbm (US$ 1982) 980000 Cost (US$ 2007) 1677653

Adsorbents Source Total ads vol 8 Alumina vol ratio 0.25 Activated carbon vol ratio 0.375 Zeolit vol ratio 0.375 G source (Nm3/s) 1.14 G (Nm3/s) 2.74 Scale factor (G/Gsource) 2.40 Adsorbent volumes Total cat vol (m3) 19.24 Alumina vol (m3) 4.81 Activated carbon vol (m3) 7.21 Zeolit vol (m3) 7.21 Adsorbent specs Price (SEK/kg) Density (kg/m3) Price (US$/m3) Cost (US$) Alumina 20 2000 5714 27480 Activated carbon 20 700 2000 14427 Zeolit vol 30 800 3429 24732 Adsorbent cost (US$) 66639 Total cost (US$) per ads 1744292

Total cost (US$) 4 adsorbers 6977166

64

Electrolysis The cost of 4.5 MNOK for the Hydro electrolysis system includes pre-treatment of the feed water and post-treatment of the hydrogen gas [26]. The power consumption for the electrolyser is 4.3 kWh/Nm3 H2 [6]. See table XVI. Table XVI. Electrolysis

Power generation

Steam turbines Power is generated by two steam turbines. The costs for these are estimated by Ulrich method. The first steam turbine (ST 1) is non-condensing [figure 5-21, 19]. Since the second turbine (ST 2) expands the steam to a low pressure and therefore condensates a part of the steam, a condensing steam turbine is used for ST 2 [table 16.32, 19]. See table XVII. Table XVII. Power generation ST 1 ST 2 Shaft power (hp) 2392.24 7057.04 Shaft power (kW) 1759 5189 Cp 120000 756772 Fbm 3.5 3.5 Cost (US$) 1982 420000 2648703

Cost (US$) 2007 718994 4534290

Steam boiler The gas from the PSA is used to generate steam in a steam boiler. The cost for the steam boiler is estimated with Ulrich method for steam boilers. See table XVIII.

H2 flow (Nm3/h) 9765 Power needed (MW) 41.99 Capacity/unit (Nm3/h) 100 Number of units needed 98 Cost per unit (MNOK) 4.5 Cost per unit (US$) 754326

Cost (US$ 2007) 73923948

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Table XVIII. Steam boiler Heating duty (kW) 7521 Cp 600000 Fp 1.25 Temp 30 bars vapour (oC) [27] 234.00 Temp (oC) 500.00 Superheating (oC) 266.00 Ft 1.30 Fbm 1.80 Cost (US$) 1982 1755000

Cost (US$) 2007 3004368

Pumps The costs for the pumps used for the feed water to the heat exchangers are estimated by Ulrich method for centrifugal pumps [table 5.49, 19]. See table XIX. Table XIX. Pumps Pump 1 Pump 2 Pump 3 Pump 4 Power kW 66.96 39.1 1.39 52.38 Power used kW 66.96 39.10 1.39 52.38 Cp 15000 10000 3000 11000 Suction pressure 85 85 4 85 Fp 2.6 2.6 1 2.6 Fm (cast iron) 1 1 1 1 Fbm 5.5 5.5 3.5 5.5 Cost (US$ 1982) 82500 55000 10500 60500

Cost (US$ 2007) 141231 94154 17975 103569

Storage vessels The biomass is stored for 48 hours and the produced hydrogen is stored for 6 hours. The cost for these is calculated by Ulrich method. The biomass is contained in bins at atmospheric conditions and the hydrogen in a spherical vessel at 850 bar [figure 5-61, 19]. See table XX.

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Table XX. Storage vessels Biomass H2 Flow (tonnes/h) 13.4 1.712 Storage time (h) 48 6 Storage volume (tonnes) 643.2 10.27296 Biomass density (tonnes/m3) 0.243 [24] Storage volume (m3) 2646.9 Cp 6000 5000 Fp 8.5 Fm 2.5 FBM 2.1 30 CBM (US$ 1982) 12600 150000

Cost (US$ 2007) 21570 256784

Compression The oxygen and hydrogen produced in the electrolyser is compressed from 1 bar to 30 bars. The produced hydrogen is compressed from 30 to 850 bars before storage. The cost for the compressors is estimated by the factored estimation method [table 6, 10]. See table XXI. Table XXI. Compression Product O2 H2 Initial pressure (bar) 30 1 1 Storage pressure (bar) 850 30 30 Power needed (kW) 3760 922 1878 Required power 3760 922 1878 Base cost c0 (M€) 12.9 12.9 12.9 Base scale s0 13200 13200 13200 Scale factor R 0.9 0.9 0.9 Overall installation factor 1.9 1.9 1.9 Cost (M€ 2002) 8.25 2.50 4.57 Cost (US$ 2002) 7261179 2198461 4024748 Cost (US$ 2007) 7525019 2278344 4170991

Contracting and contingency A cost for contracting and contingency corresponding to 15 % of the capital cost is added. This includes costs for the use of external competence and all unforeseen costs which occur during the construction of the plant and until the start-up. The rate is chosen with reference to [28].

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Auxiliary facilities Facilities which are not covered above are referred to as auxiliary. These include:

• office buildings and accessories • chimneys • waste water cleaning (capital investment) • storage tanks • paving • workshop • laboratory

The added cost for auxiliary facilities is set to 25 % of the direct plants cost [28].

Operating costs

Fixed capital

Material in storage The cost caused by raw material and product kept in storage is specified as a cost under material in storage. Biomass corresponding to 48 hours of use and product H2 from a run time of 6 hours is stored. These requirements are specified by the project task. See table XXII. Table XXII. Biomass H2 Flow (tonnes/h) 13.4 1.71 Storage time (h) 48 6 Storage volume (tonnes) 643.2 10.3

Storage volume (m3 at 70 oC, 30 bars) 4838 Storage volume (Nm3) 115630

Storage volume at 25 oC 850 bars (m3) 6 Price (US$/tonne and US$/Nm3 resp) 63* 0.60 [29] Value of stored material (US$) 52705 69378 N 10 10 Interest rate 0.15 0.15 fA 0.1993 0.1993 Cost (US$/year) 10501 13824 * The biomass price is calculated as follows: Heat value biomass (MJ/kg) 19.88 [14] Biomass price (€/GJ) 3.5 [23] Biomass price (US$/tonne 2002) 61.23 Biomass price (US$/tonne 2007) 63 The cost of spare parts in storage is calculated through 10 % multiplied by the annual cost of maintenance (see table 11), as specified.

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Direct operating cost

Raw material The price of biomass is calculated as above. The water price used is a mean value of the interval mentioned in [28]. See table XXIII. Table XXIII. Raw material Biomass Water Flow (tonnes/h) 13.4 7.20 Runtime (h/year) 8000 8000 Tonnes/year 107200 57600 Cost (US$/tonne) 63 1.43 Cost (US$/year) 6753600 82451

Electricity The electricity price used includes the electricity market price, electricity certificate, net fees and tax. Using a mean value of some of the latest prices according to Statistiska Centralbyrån, the electricity price can be set to 51.85 öre/kWh or US$ 0.074 /kWh [19]. See table XXIV. Table XXIV. Electricity costs Power usage (kWh/year) Power generation (kWh/year) Effect (MW) Electrolysis 335 942 426 -41.99 Grinder 2 741 257 -0.34 Feeder 4 745 600 -0.59 Pumps 848 480 -0.11 Product compression 30 080 000 -3.76 O2 compression 7 376 000 -0.92 H2 compression 15 024 000 -1.88 ST 1 14 072 000 1.76 ST 2 41 512 000 5.19 Total 396 757 763 55 584 000 Diff -341 173 763 -42.65 Cost (US$) 25 271 228

Staff The amount of operating staff needed for the plant is estimated through Ulrich’s unit method. [table 7-6, page 133, 29]. The cost for management and laboratory work can both be set to 15 % each of the cost for the operating staff [28]. See table XXV.

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Table XXV. Staff

Unit No of units No of staff needed/shift/unit

No of staff needed/shift

Grinder 1 0.3 0.3 Dryer 1 0.3 0.3 Feeder 1 0.2 0.2 Gasifier 1 0.5 0.5 HEX 4 0.1 0.3 Cracker 1 0.2 0.2 Cyclone 1 0.05 0.05 Particle filter 1 0.05 0.05 Desulphuriser 1 0.2 0.2 High temperature shift 1 0.5 0.5 Low temperature shift 1 0.5 0.5 Scrubber 1 0.2 0.2 PSA 1 0.2 0.2 Electrolyser 1 0.5 0.5 Steam turbine 2 0.1 0.2 Pump 4 0.05 0.1 Storage vessel 2 0.2 0.4 Product compressor 1 0.1 0.1 Total operating staff 5 Mean salary (US$/month) 2857 Shifts 3 Cost operating staff (US$) 1028571 Cost management (US$) 154286 15 % of operating staff cost Cost lab work (US$) 154286 15 % of operating staff cost

Maintenance A factor of 5 % of the direct plant cost is added to cover maintenance costs [28]. See table 11 in the main document.

Waste treatment The cost for the disposal of the catalysts and adsorbents are assumed to be included in their respective price.

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As a by-product of the gasification, ash is produced. The choice of disposal includes landfill in road works and return of the material to the forest. See table XXVI. Table XXVI. Ash disposal Ash-% dry basis 2.44 [31, table 1] Mass flow (tonnes/h), 15% wc 13.4 Mass flow (tonnes/h), dry basis 11.39 Ash flow (tonnes/h) 0.278 Ash flow (tonnes/year) 2223 Cost (SEK/tonne) 400 [31] Cost (US$/tonne) 57.14 Ash disposal cost (US$) 127047

Indirect operating cost All of the estimated indirect operating costs are based on [chapter 7, 29].

Overhead for staff The overhead for staff is included in the indirect operating costs. As a rule of thumb, it estimates to 70 % of the cost for operating personnel and to 50 % of the management and laboratory staff. See table 11.

Administration The administration cost work out as 25 % of the overhead for staff. See table 11 in the main document.

Distribution and sales Sales and distribution costs are calculated as 10 % of the total operating costs. See table 11 in the main document.

Research and development (R&D) The cost for R&D for this plant assumes to correspond to 2 % of the total operating costs. See table 11 in the main document.

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Investment analysis To be able to finance the plant, a loan of the whole plant construction cost is assumed. The annuity method is utilised to estimate the annual interest cost [page 38, 29]. An economic life span (N) of 10 years is used in the calculations. The capital cost is multiplied with the annuity factor and using this method the depreciation of the capital cost is calculated, as seen in table XXVII. Table XXVII. N 15 Interest (%) 10 Annuity factor, fA 0.131 Total plant cost (US$) 191 774 784 Annual cost (US$) 25 213 355