MASTER’S THESIS

Department of Applied Physics and Mechanical EngineeringDivision of Energy Engineering

M.Sc. in Sustainable Energy SystemsCONTINUATION COURSES

2006:54 PB • ISSN: 1653 - 0187 • ISRN: LTU - PB - EX - - 06/54 - - SE

2006:54 PB

Diana Carolina Cardenas Barrañon

Methanol andHydrogen Production

Energy and Cost Analysis

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MASTER THESIS

METHANOL AND HYDROGEN PRODUCTION: ENERGY AND COST ANALYSIS.

Diana Carolina Cárdenas Barrañón

Division of Energy Engineering Depertment of Applied Physics, Mechanical and Material Engineering

Luleå University of Technology S-971 87 Luleå- Sweden

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Preface This work has been carried out as a Master’s Thesis for the Master of Science Program in Mechanical Engineering with focus towards Energy Engineering. The research has been performed at the Division of Energy, Luleå University of Technology, during August 2005 through May 2006. I would like to thank to my supervisors Ph.D. student Sylvain Leduc and Doc. Jan Dahl for their great collaboration and support during my research. I would also like to thank all my teachers and classmates for all their help, feedbacks and friendly atmosphere. I also would like to thank my Father who has been my guide since I have memory and had been my deepest inspiration all this years of studies and work. Finally I will like to thank my dear boyfriend Per for his enormous patience, encouragement and for always being there for me. Diana Cardenas Luleå May 2006

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Abstract Methanol and hydrogen produced from biomass are promising carbon neutral fuels. Both are well suited for use in Fuel Cell Vehicles (FCVs) which are expected to reach high efficiencies, about a factor 2-3 better than current International Combustion Engine Vehicles (ICEVs). In addition they are quiet and clean, emitting none of the air pollutants SOx, NOx, VOS or dust. When methanol and hydrogen are derived from sustainable grown biomass, the overall energy chain can be greenhouse gas neutral. Technical and economic prospects of the future production of methanol and hydrogen from biomass have been evaluated. A technology review, including promising future components, was made resulting in a set of promising conversions concepts. Flowsheeting models were made to analyse the technical performance in ASPEN PLUS. Results were used for economic evaluations. Overall energy efficiencies are around 55 % HHV for methanol and around 60 % for hydrogen production. 400 MWth input systems produce biofuels at 9- 12 US$/GJ, this is above the current gasoline production price of 4-6 US$/GJ. This cost price is largely dictated by the capital investments. The outcomes for the various systems types are rather comparable, although concepts focussing on optimised fuel production with a little or no electricity co-production perform somewhat better. Long term cost reductions reside in cheaper biomass, technological learning, and application of large scale up to 2000 MWth. This could bring the production costs of biofuels in the 5-7 US$/GJ range. Biomass derived methanol and hydrogen are likely to become competitive fuels tomorrow.

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Table of Contents

Page 1. Introduction………………………………………………………………………… 5 1.1Background……………………………………………………………………… 5 1.1.1Biomass…………………………………………………………………... 6 1.2 Rationale…………………………………………………….…………………… 8 1.3 Objectives……………..………………………………………………………… 9 1.4 Methodology…………………………………………………………………… 9

2. System Description……………………………………………………………..9 2.1 Production of Biofuels…………………………………………………..9 2.1.1Pre-treatment of feedstock………………………………….9 2.1.2Gasification……………………………………………………………11 2.1.3Gas cleaning………………………………………………………..15 2.2 Syngas Processing…………………………………………………………17 2.2.1 Steam Reforming…………………………………………………..17 2.2.2 The water-gas shift reaction………………………………18 2.2.3 CO2 removal…………………………………………………………18 2.3 Methanol production…………………………………………………….19 2.4 Hydrogen production…………………………………………………….22 2.5 Power generation…………………………………………………………..24 2.5.1 Gas turbine…………………………………………………………….25 2.5.2 Heat Recovery Steam Generation (HRSG) ………….25 3. System Calculations…………………………………………………………25 3.1Modelling principles………………………………………………………..25 3.2 Results………………………………………………………………………………29 4. Economics………………………………………………………………………………..30 4.1Method……………………………………………………………………………….30 4.2Results……………………………………………………………………………….31 5. Discussion and Conclusions…………………………………………...34 6. References………………………………………………………………………………35 7. Annexes……………………………………………………………………………………..40

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1. INTRODUCTION 1.1 Background The interest risk for climate change is a growing concern for the global society. The third assessment report of the intergovernmental panel on climate change (IPCC) provides the strongest evidence so far that the global warming of the last 50 years is due largely to human activity, especially the carbon dioxide (CO2) emissions that arise when fossil fuels are burned (IPCC, 2001a). According to what is known as the Kyoto Protocol, the Conference of the Parties (COP) has agreed that by committed period 2008-2012, developed countries shall be legally committed to reduce their collective greenhouse gas (GHG) emissions by at least 5% compared to 1990 levels. Using renewable energy resources to substitute fossil fuels is one of the technological options to mitigating GHG emissions. The use of Biomass as a main energy source has increased since the last decades, (example of Sweden in Figure 1). The advantages of biomass is that it uses carbon dioxide from the atmosphere to grow, and releases as much carbon dioxide when consumed, it therefore does not contribute to climate change by emissions to the atmosphere of carbon dioxide or other greenhouse gases. Biomass is then considered as a renewable energy resource if the whole cycle of biomass conversion is well managed.

Figure 1. Energy supply in Sweden, excluding net electricity imports, 1970–2003 (Source: Statistics Sweden, Swedish Energy Agency processing). Carbon dioxide from public transportation has been increasing a lot since the past decade. One way to decrease the emissions is to innovate new energy conversion technologies in which the fuels for the transportation are carbon dioxide neutral for the atmosphere. How to reduce the increased GHG emissions related to the transportation is a challenging issue. Fuel cell vehicles are a promising technology for meeting future goals for zero emissions. Other zero emission energy is coming out and starts to be competitive to the fossil fuels: ethanol. In Brazil, ethanol from sugar cane is spreading very quickly. Using biofuels, the emissions of greenhouse gases from the transportation will decrease. Using on the other hand carbon sequestration when producing these biofuels, one may even decrease the amount of carbon dioxide in the atmosphere.

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This report focuses on different fuels, from the woody biomass, for public and individual transportation, methanol and hydrogen. Focus is pointed on the different technologies of production and the cost of production of these bio fuels. 1.1.1 Biomass Biomass is a carbonaceous material of biological origin derived from plants including wood, bark, straw, reeds, and agricultural crops, etc. Biomass energy is obtained directly from various agriculture crops or from raw material from the forest, or is obtained as by-products from the forest industry. Biomass is also recycled from household waste or derived from recycled paper and demolition wood. If properly grown and managed, biomass is a renewable energy resource. It does not contribute to climatic change through emissions to the atmosphere of carbon dioxide or other “greenhouse gases” because it absorbs the same amount of carbon in growing as it releases when consumed as fuel. The content of heavy metals in biomass is low, and if the ashes from biomass are recycled to forest and agriculture land, the use of biomass does not imply any major difference to the natural cycle of growth and decomposition (Yan, 1998). In this report, focus is mainly stressed on woody biomass. Woody biomass or lignocellulosic (FIGURE 2) is composed of carbohydrate polymers (cellulose and hemicellulose), lignin and a remaining smaller part (extractives, acids and salts and minerals) (Hamelinck et al. 2001).

FIGURE 2. Plant Cell Wall Composition (Shlesser, 1994).

Cellulose Cellulose has the same chemical structure in all types of biomass: a linear carbohydrate polymer with a high degree of polymerization (the average molecular weigh is 100,000). Cellulose chains aggregate into a crystalline structure that gives its mechanical strength. (Katofsky, 1993).

Hemicellulose Hemicelluloses are mixtures of polysaccharides comprised mainly of glucose, mannose, galactose, xylose, arabinose, methylglucuronic acid and galacturonic acid, and have a lower degree of polymerization than cellulose (the molecular weigh is<30,000). Unlike cellulose, hemicellulose does not form a crystalline structure, but is amorphous (Deglise et al., 1987, Katofsky, 1993).

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Lignin Lignin is a complex, amorphous, randomly linked, high molecular weight (2,100) (Okuda et al., 2004), ringed structure that helps bind the cellulosic fibers together. Dry wood is typically 45-50 % cellulose by weight, 15-25% hemicellulose, and 20-30% lignin (Sudo, et al. 1989). In addition, biomass contains 0.1-2% ash, and even smaller quantities of elements such as chlorine and sulphur (0.01-0.1%) and alkali metals (mainly potassium and sodium) (Katofsky, 1993).

Characteristics of Biomass (compared with coal) On a dry, ash free basis, biomass has a heating value that is roughly 60-70 % that of coal. As well, biomass has a lower mass density, higher initial moisture content (40-60% fresh woody biomass, compared to 2-12 % for most bituminous coals), and is dispersed over a wide area. Because of this initial moisture content, the receiving heating value of biomass is only about 30-40 % of the heating value of coal. Even some of the chemical properties of biomass make it superior to coal in many ways. As seen in TABLE 1, compared to coal, biomass contains far less inert material (ash) and significantly less sulphur. Biomass ash is usually free of toxic metals (such as arsenic) and other potentially hazardous materials that can make the disposal of coal ash difficult and complicate syngas processing. Biomass ash can actually be used as fertilizer, to restore nutrients to the land where it was grown (Katofsky, 1993). The most important difference between biomass and coal is the much higher volatile fraction of biomass. For this reason, biomass is considerably more reactive than coal. Cellulose, the main component of wood, pyrolyses (decomposes) more completely and at lower temperatures than coal (FIGURE 3). Biomass char, the carbon that remains after pyrolysis, is 10-30 times more reactive than coal char (Graboski, 1981) (FIGURE 3). Biomass chars gasify more rapidly and at lower temperatures than coal chars. Biomass is easier to gasify than coal in the sense that lower operating temperatures are required to achieve the same gasification rates and degree of conversion from solid to gas (Katofsky, 1993) TABLE 1.Compositional data and heating values for selected types of biomass, coal and natural gas (Katofsky, 1993). Feedstock

type HHV

(MJ/kg) Volatile matter

Fixed carbon Ash C H O N S Cl Ash

Douglas fir 20.37 87.3 12.6 0.1 50.64 6.18 43 0.06 0.02 0.00 0.10 Maple 18.86 87.9 11.5 0.6 49.89 6.09 43.27 0.14 0.03 0.00 0.58

Ponderosa pine 20.02 82.54 17.17 0.29 49.25 5.99 44.36 0.06 0.03 0.01 0.30

White oak 19.42 81.28 17.2 1.52 49.48 5.38 43.13 0.35 0.01 0.04 0.61 Poplar 19.38 82.32 16.35 1.33 48.45 5.85 43.69 0.47 0.01 0.10 1.43

Eucalyptus grandis 19.35 82.55 16.93 0.52 48.33 5.89 45.13 0.15 0.01 0.08 0.41

Sugarcane bagasse 17.33 73.78. 14.95 11.27 44.8 5.35 39.55 0.38 0.01 0.12 9.79

Wyoming subbitumou

s coal 26.78 44.68 46.12 9.2 68.75 4.89 15.55 0.89 0.69 0.00 9.24

Illinois no6 bituminous

coal 26.67 37.5 43.4 18.18 65.34 4.2 6.59 1.02 4.55 0.00 18.3

Natural gas 52.94 71.99 23.78 0.38 3.84 0.00 0.00

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FIGURE 3.Comparison of coal and cellulose pyrolysis (the main component of biomass) (scanned from Katofsky, 1993). 1.2 Rationale Methanol and hydrogen can be produced from biomass via gasification. Several routes involving conventional, commercial, or advanced technologies, which are under development, are possible. A scheme of the main process steps to convert biomass to methanol and hydrogen are shown in FIGURE 4.

Shift to adjustCO/H2 ratio

Pre-treatment:-drying-chipping

Gasifier Gas cleaningsection

Reformerfor higherhydrocarbons

Methanolsynthesis

H2separation

Methanol

Hydrogen

Gas turbine

Steamturbine

Electricity

Auxiliaries

Purge gas

steam

Shift to adjustCO/H2 ratio

Pre-treatment:-drying-chipping

Gasifier Gas cleaningsection

Reformerfor higherhydrocarbons

Methanolsynthesis

H2separation

Methanol

Hydrogen

Gas turbine

Steamturbine

Electricity

Auxiliaries

Purge gas

steam

FIGURE 4.Key components in biomass to methanol/hydrogen production concepts. (Hamelinck et al, 2001)

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1.3 Objectives The key objective of this work is to identify biomass to methanol and hydrogen conversion concepts that may lead to higher overall energy efficiencies and lower costs on longer term. Improved performance may be obtained by:

• Combined fuel and power production may lead to lower cost and possibly higher

overall thermal efficiencies because of cheaper reactor capacity and reduction of internal energy consumption of the total plant.

• Economies of scale; various system analyses have shown that the higher conversion

efficiencies and lower unit capital costs that accompany increased scale generally outweigh increased energy use and costs for transporting larger quantities of biomass. Furthermore, it should be noted that paper & pulp mills, sugar mills, and other facilities operate around the world with equivalent thermal inputs in the range of 1000-2000 MWth. Such a scale could therefore be considered for production of energy/fuel from (imported) biomass as well.

1.4 Methodology The work consists of several steps. First of all a technology assessment on gasifiers, gas cleaning, syngas processing and combined cycles will be made to make inventory of possible configurations. Second, promising system configurations were selected for further performance modelling with help of the flowsheeting program Aspen plus. Aspen plus is used to calculate energy and mass balances. Third and economic evaluation is performed including economies of scale of the units. Finally the configurations are compared, conclusions drawn and recommendations are formulated. 2. SYSTEM DESCRIPTION 2.1 Production of Biofuels Syngas, a mixture of CO and H2, is needed to produce methanol or hydrogen. A train of processes to convert biomass to required gas specifications precedes the methanol reactor or hydrogen separation as was shown in Figure 4. 2.1.1 Pre-treatment of feedstock Biosyngas can be produced from biomass, but it needs a pre-treatment, which may be one of the following techniques:

• Chipping • Drying • Torrefaction • Powder

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Chipping

Chipping is generally the first step in biomass preparation. The fuel size necessary for fluidised bed gasification is between 0 and 50 mm (Pierik et al., 1995). The total primary energy requirements for chipping woody biomass are approximately 100 kJ/kg of wet biomass or 0.5 % of the higher heating value (Katofsky 1993).

Drying

The fuel is dried to 15 % or 10% depending on the gasifier applied in order to reduce the heat required for the gasification and to reduce the moisture content of the product gas. Drying consumes roughly 10% of the energy content of the feedstock. Although the heat of vaporisation of water is 2,250 kJ/kg, in practice approximately 3,500 kJ is required to evaporate one kg of water (Katofsky 1993). There are two ways for drying, either with steam or with flue gas. Heat in the flue gas that exist the heat recovery steam generator is used to dry the incoming biomass. Furthermore flue gas drying holds the risk of spontaneous combustion and corrosion (Consonni et al., 1994). It is not clear whether flue gas or steam drying is a better option in biofuel production. The specificities for steam or flue gas drying are shown on TABLE 2. TABLE 2. Requirements for steam or flue gas drying for a biomass feedstock from 50% moisture to 15-10% (Pierik and Curvers 1995).

Unit Steam Flue gas Temperature ˚C 200 Pressure bar 12 Energy use MJ/twe 2.8 2.4 / 3.0 Electricity consumption kWh/twe 40 40 - 100

To improve the combustion and transportation properties of the biomass, the forest fuel can be converted into several other forms, such as charcoal, torrefied wood, pellets, briquettes and wood powder

Torrefaction

Torrefaction is a thermal treatment at a temperature of 200 to 300 ˚C, at near atmospheric pressure and in the absence of oxygen. This mild thermal treatment not only destructs the fibrous structure and tenacity of the biomass (wood), but is also known to increase the calorific value and to invert the hydrophilic nature. During the process, the biomass partly devolatilises which leads to a decrease in mass, but the initial energy content of the biomass is mainly preserved in the solid product (fuel). The latter is of great importance to the overall energy efficiency of the biomass-to-biosyngas conversion chain. Due to the low moisture content of torrefied wood the transport cost is lower and the quality as a fuel better (Bergman et al., 2004).

Powder Biomass can be pulverized to particles of 100 mm or less. This however consumes huge amounts of electric energy.

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2.1.2 Gasification After pre-treatment, is the conversion into biosyngas by gasification and subsequently synthesize the required products. There are different kinds of gasification, such as:

• Low temperature gasification, 800 - 1000˚C: The biosyngas components H2 and CO typically contain only ~50% of the energy in the gas, while the remainder is contained in CH4 and higher (aromatic) hydrocarbons.

• High temperature gasification, >1200˚C: All the biomass is completely converted into

biosyngas (Boerrigter et al. 2004). FIGURE 5 below describes these processes.

FIGURE 5. Two biomass-derived gases via gasification at different temperature levels: biosyngas and product gas and their typical applications (Boerrigter et al. 2004). Thermo chemical gasification is the conversion by partial oxidation at elevated temperature of a carbonaceous feedstock such as biomass or coal into a gaseous energy carrier. In a gasifier, biomass is converted into gases (H2, CO, CO2, H2O, CH4, light hydrocarbons) and condensable tars at 800-1,200˚C. The gas also contains impurities originating from the fuel, like sulphur, nitrogen and chlorine compounds, and alkali metals. The final product distribution in a gasification gas largely depends on gas-feedstock contact type and process conditions. Gasification occurs in sequential steps:

• Drying, to evaporate moisture. • Pyrolysis, to give gas, vaporized tars or oils and a solid char residue. • Gasification or partial oxidation of the solid char, pyrolysis tars and pyrolysis

gases. The different types of gasifiers are the following:

• Fixed bed gasifiers • Fluidised bed gasifiers

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Fixed bed gasifiers

Downdraft gasifier

FIGURE 6. Downdraft gasifier (Ohlström et al.,

2001).

The downdraft gasifier features concurrent flow of gases and solids through a descending packed bed which is supported across a constriction known as a throat, where most of the gasification reactions occur. The reaction products are intimately mixed in the turbulent high-temperature region around the throat, which aids tar cracking. Some tar cracking also occurs below the throat on a residual charcoal bed, where the gasification process is completed. This configuration results in a high conversion of pyrolysis intermediates and hence a relatively clean gas (Bridgwater et al. 1995).

Updraft gasifier

FIGURE 7. Updraft

gasifier (Ohlström et al., 2001) .

In the updraft gasifier, the downward-moving biomass is first dried by the up flowing hot product gas. After drying, the solid fuel is pyrolysed, giving char which continues to move down to be gasified. Pyrolysis vapour are carried upward by the up flowing hot product gas. An advantage, especially of updraft (counter-current) gasifiers, is the high cold gas efficiency connected with the low syngas exit temperature. But this advantage is obtained at the expense of high tar content in the syngas.

Updraft or downdraft gasifiers require mechanically stable fuel pieces of one or few cm size, to guarantee an unblocked passage of gas through the bed. Straw, hay or paper must be pelleted or briquetted: This is an expensive procedure, but if it is performed directly with a harvester, it could simplify all transport and storage operations later on. To summarise: Fixed bed gasifiers are well suited for stable wood pieces. Fixed bed gasifiers are well suited for stable wood pieces (Henrich et al., 2004).

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Fluidised bed gasifiers

Figure 8. Fluidised bed gasifier (Ohlström et al., 2001).

Fluidised bed gasifiers are operated either in the bubbling or the circulating mode, usually with quartz sand as bed material and fuel particles of few mm up to 1 or few cm size. For biomass feedstock at least 750°C is required, to achieve a reasonably high gasification rate. The bed temperature must be kept below the ash softening temperature, since a sticky ash would glue together the bed particles and this agglomeration causes a breakdown of fluidisation (Henrich et al., 2004).

Circulating bed gasifiers The fluidising velocity in the circulating fluid bed is high enough to entrain large amounts of solids with the product gas. This configuration has been extensively developed for wood waste conversion in pulp and paper mills for firing lime and cement kilns and steam-raising for electricity generation (Bridgwater et al. 1995).

Twin fluid bed gasifiers These are used to give a gas of higher heating value from reaction with air than is obtained from a single air-blown gasifier. The gasifier is a pyrolyser, heated with hot sand from the second fluid bed, which is heated by burning the product char in air before recirculation to the first reactor. Steam is usually added to encourage the shift reaction to generate hydrogen and to encourage carbon-steam reactions (Bridgwater et al. 1995). Entrained flow gasifiers The entrained flow gasifier has been developed mainly for coal feedstock. The entrained flow gasifier operates at high temperature (1,300 ˚C) which is not necessary for biomass with high reactivity. The feedstock must also be crushed to fine-sized particles which is energy and intensive with biomass (Ohlström et al., 2001). Fluidised beds are attractive for biomass gasification. They are able to process a wide variety of fuels including those of high moisture and small size. They are easily scaled to large sizes suitable for electric power production. Disadvantage of fluidized beds include relatively high power consumption to move gas through the bed; high exit gas temperatures, which complicate efficient energy recovery; and relatively high particulate burdens in the gas due to the abrasive forces acting within the fluidized bed (Brown, 2003). Fluidised sand beds are suited for woods, which have a high ash melting point, usually > 1000°C. The ash melting points of cereal straw however can drop to even below 700°C (Henrich et al., 2004). TABLE 3 shows the different characteristics of some gasifiers, and TABLE 4 the gas composition form various kind of gasifiers.

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TABLE 3. Typical gasifier characteristics (Bridgewater et al. 1994)

Reaction (˚C)

Exit gas (˚C)

Tars Particulates Turn-down Scale-up ability

Current capacity max (t/h)

Min MWe

Max MWe

Fixed bed Downdraft 1000 800 v.low moderate good poor 0.5 0.1 1 Updraft 1000 250 v.high Moderate good good 10 1. 10 Cross current 900 900 v.high high fair poor 1 0.1 2 Fluid bed Single reactor 850 800 Fair High good good 10 1 20 Fast fluid bed 850 850 Low v.high good v.good 20 2 50 Circulating bed 850 850 Low v.high good v.good 20 2 100 Entrained bed 1000 1000 Low v.high poor good 20 5 100 Twin reactor 800 700 High High fair good 10 2 50 Moving bed Multiple hearth 700 600 High low poor good 5 1 10 Horizontal moving bed 700 600 High Low fair fair 5 1 10

Sloping hearth 800 700 Low Low poor fair 2 0.5 4 Screw /auger kiln 800 700 High low fair fair 2 0.5 4 Other Rotary kiln 800 800 High high poor fair 10 2 30 Cyclone reactor 900 900 low v.high poor fair 5 1 10 TABLE 4. Producer gas composition from various kinds of gasifiers (Brown, 2003).

Gaseous Constituents (vol.% dry) Energy content Gas quality

Gasifier Type H2 CO CO2 CH4 N2 HHV (MJ/m3) Tars Dust Air-blown updraft 11 24 9 3 53 5.5 High

(~10g/m3) Low

Air-blown downdraft 17 21 13 1 48 5.7 Low

(~1 g/m3) Medium

Air-blown fluidized bed 9 14 20 7 50 5.4 Medium

(~10g/m3) High

Oxygen-blown downdraft 32 48 15 2 3 10.4 Low

(~1 g/m3) Low

Indirectly heated fluidized bed 31 48 0 21 0 17.4 Medium

(~10g/m3) High

Processes based on indirect gasification are of interest with regard to methanol synthesis, as they offer a possibility to produce non-nitrogenous synthesis gas without any investment in a relatively expensive oxygen plant. In the indirect process, the gasification reactor is heated by hot bed material. The bed material is heated in a separate combustion unit operated in fluidised-bed principle by burning the mixture of residual char and bed material separated form the product gas (Ohlström et al., 2001).

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2.1.3 Gas Cleaning The syngas produced from the gasifier contains different kinds of contaminants: particles, tars, alkali, sulphuric, chloride and nitrogen compounds (Tijmensen 2000). These contaminants can lower catalyst activity in reformer, shift or methanol reactor, and cause corrosion in gas turbine or heat exchangers. Ash particles cause wear and corrosion throughout the plant. Particulate concentrations in raw gas from most fluidised bed gasifiers can be as high as 5000 ppmw. Severe gas cleaning is then required. The particulate concentration needs to be below 1 ppmw at the turbine inlet, with 99% of the particles smaller than 10 micron. This corresponds to a particulate concentration in the fuel gas before the combustor of about 3-5 ppmw (Consonni and Larson 1994b). Gas cleaning has to remove all components that may be harmful to the catalysts or other parts of the plant by corrosion, erosion or fouling. The gas cleaning is preferably operated at the same temperature of the downstream gas application to minimise efficiency loss by cooling. In general two routes can be distinguished:

• Wet low temperature gas cleaning, • Dry high temperature gas cleaning.

Wet low temperature gas cleaning

The subsequent cleaning steps are depicted in FIGURE .

FIGURE 9. Low temperature wet gas cleaning (Hamelinck et al., 2001).

• A cyclone separator removes most of the solid impurities, down to sizes of approximately 5 μm (Katofsky 1993).

• Before the bag filter the syngas is cooled to just above the water dew point. • New generation bag filters made from glass and synthetic fibres have an upper

temperature limit of 260 °C (Perry et al. 1987). At this temperature particulates and alkali, which condense on particulates, can successfully be removed (Consonni et al., 1994; Tijmensen 2000).

• The syngas is then scrubbed down to 40 °C below the water dew point, by means of water. Residual particulates, vapour phase chemical species (unreacted tars, organic gas condensates, trace elements), reduced halogen gases and reduced nitrogen compounds are removed to a large extend. Alkali removal in a scrubber is essentially complete (Consonni et al., 1994). With less than 30 ppm H2S in the syngas bulk removal of sulphur compounds is not necessary.

• A ZnO bed is sufficient to lower the sulphur concentration below 0.1 ppm. ZnO beds can be operated between 50 and 400 °C, the high-end temperature favours efficient utilisation. At low temperatures and pressures less sulphur is absorbed, therefore multiple beds will be used in series. The ZnO serves one year and is not regenerated (Katofsky 1993).

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• If CO2 removal is demanded as well, a solvent absorption process will be used like Rectisol or Sulfinol, this unit can also be placed downstream (Tijmensen 2000). H2S and COS are reduced to less than 0.1 ppm and all or part of the CO2 is separated. The sulphur in the acid gas output is concentrated to sulphuric acid or reclaimed as elemental sulphur in a Claus unit.

This method will have some energy penalty and requires additional waste water treatment, but on the short term it is more effective than hot dry gas cleaning.

High temperature dry gas cleaning

The subsequent cleaning steps are depicted in FIGURE .

FIGURE 10. High temperature dry gas cleaning (Hamelinck et al., 2001). As with wet gas cleaning the tar cracker is optional. Possibly present tars and oils are not removed during the downstream hot gas cleaning units since they do not condense at high temperatures. It is not clear to what extent tars are removed (Tijmensen 2000).

• For particle removal at temperatures above 400 °C sliding granular bed filters are used instead of cyclones.

• Final dust cleaning is done using ceramic candle filters (Hamelinck et al., 2001) or sintered-metal barriers operating at temperatures up to 720 °C; collection efficiencies greater that 99.8 % for 2 – 7 μm particles have been reported (Katofsky 1993). Still better ceramic filters for simultaneous SOx, NOx and particulate removal are under development (White et al. 1992).

• Processes for alkali removal in the 750 – 900 °C range are under development and expected to be commercialised within few years. Lead and zinc are not removed at this temperature. High temperature alkali removal by passing the gas stream through a fixed bed of sorbent or getter material that preferentially adsorbs alkali via physical adsorption or chemisorption was discussed by Turn et al. (1998). Below 600 °C alkali metals condense onto particulates and can more easily be removed with filters (Katofsky 1993). Nickel based catalysts have proved to be very efficiency in decomposing tar, ammonia and methane in biomass gasification gas mixtures at about 900 °C. However sulphur can poison these catalysts (Tijmensen 2000). It is unclear if the nitrogenous component HCN is removed. It will probably form NOx in a gas turbine (Verschoor et al., 1991).

• Halogens are removed by Na and Ca based powdered absorbents. These are injected in the gas stream and removed in the dedusting stage (Verschoor et al., 1991).

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2.2 Syngas Processing The syngas produced by the gasification of biomass, consists mainly of H2, CO, CO2 and CH4. Syngas processing consists of three steps, which are all optional.

2.2.1Steam reforming In the presence of a suitable catalyst (usually nickel based), methane, tars and other light hydrocarbons are reformed into CO and H2 at high temperatures (Katofsky 1993). Steam reforming is the most common method of producing a synthesis gas from natural gas or gasifier gas. The highly endothermic process takes place over a nickel-based catalyst. Reactions are:

CH4 + H2O→CO + 3H2 Δh˚=206 MJ/kmol (R1) C2H4 + 2H2O→2CO + 4H2 Δh˚=210 MJ/kmol (R2) C2H6 + 2H2O→2CO + 5H2 Δh˚=347 MJ/kmol (R3)

The water gas shift reaction (Reaction R4) takes place as well, and brings the reformer product to chemical equilibrium (Katofsky 1993). Reforming is favoured at lower pressures, but elevated pressures benefit economically (smaller equipment). Reformers typically operate at 1 – 3.5 MPa. Steam methane reformer (SMR) uses steam as the conversion reactant and to prevent carbon formation during operation. Tube damage or even rupture can occur when the steam to carbon ratio is allowed to drop below acceptable limits. The specific type of reforming catalyst used, and the operating temperature and pressure are factors that determine the proper steam to carbon ratio for a safe, reliable operation. Typical steam to hydrocarbon-carbon ratios range from 2:1 for natural gas feeds with CO2 recycle, to 3:1 for natural gas feeds without CO2 recycle, propane, naphtha and butane feeds (King et al., 2000). Usually full conversion of higher hydrocarbons in the feedstock takes place in an adiabatic pre-reformer. This makes it possible to operate the tubular reformer at a steam to carbon ratio of 2.5. When higher hydrocarbons are still present, the steam to carbon ratio should be higher: 3.5. In older plants, where there is only one steam reformer, the steam to carbon ratio was typically 5.5. A higher steam: carbon ratio favours a higher H2:CO ratio and thus higher methanol production. However more steam must be raised and heated to the reaction temperature; this decreases the process efficiency. Neither is additional steam necessary to prevent coking (Katofsky 1993).

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2.2.2 The water-gas shift reaction In a shift reactor the ratio CO:H2 is changed via the water-gas shift reaction:

CO + H2O ↔ CO2 + H2 (R4) This reaction is exothermic and proceeds nearly to completion at low temperatures. Modern catalysts are active as low as 200 °C (Katofsky 1993) or 400 °C (Maiya et al. 2000). Due to high catalyst selectivity all gases except those involved in the water-gas shift reaction are inert. The reaction is independent of pressure. Conventionally the shift is realised in a successive high temperature (360 °C) and low temperature (190 °C) reactor. Nowadays, the shift section is often simplified by installing only one CO-shift converter operating at medium temperature (210 °C). To shift as much as possible CO to H2, and profit from the kinetics of high temperatures, the dual shift reactor is applied in the hydrogen production concepts in the present study. For methanol synthesis, the gas can be shifted partially to a suitable H2:CO ratio, therefore ‘less than one’ reactor is applied. The temperature may be higher because the reaction needs not to be complete and this way less process heat is lost. Theoretically the steam: carbon monoxide ratio could be 2:1. On lab scale good results are achieved with this ratio (Maiya et al. 2000). In practice extra steam is added to prevent coking (Tijmensen 2000); the ratio is set 3:1.

2.2.3 CO2 removal To get the ratio (H2-CO2)/(CO +CO2) to the value desired for methanol synthesis, part of the carbon oxides could be removed. This can be done by partially scrubbing out carbon oxides, most effectively carbon dioxides. For this purpose different physical and chemical processes are available. The various technologies for CO2 removal from gas streams have been described by many authors. A technology overview was made in a previous STS report (Hamelinck et al. 2000 Annex A). Generally a division can be made into:

• Chemical absorption • Physical absorption • Membranes • Distillation

The two absorption options are widely applied, and at present the most suitable for application to a broad range of CO2 containing streams.

• Chemical absorption using amines is the most conventional and commercially best-proven option.

• Physical absorption, using Selexol, has been developed since the seventies and is an economically more attractive technology for gas streams containing higher concentrations of CO2.

19

2.3 Methanol Production Conventionally, methanol is produced in two-phase systems: the reactants and products forming the gas phase and the catalyst being the solid phase. Synthetic methanol production first began in 1923 at BASF’s Leuna, Germany plant. The process required a high pressure (250 – 350 atm) and catalyst selectivity was poor. Processes under development at present focus on shifting the equilibrium to the product side to achieve higher conversion per pass. Examples are the gas/solid/solid trickle flow reactor, with a fine adsorbent powder flowing down a catalyst bed and picking up the produced methanol; and liquid phase methanol processes where reactants, product, and catalyst are suspended in a liquid. (Hamelinck et al. 2001).

Synthesis Gas Conditioning

For optimal production of methanol, three parameters are of particular importance: a) The ratio of CO2 to CO should be optimised for methanol production, similar to the ratio in steam reformed natural gas. b) The H2:CO ratio of the synthesis gas. The synthesis of methanol is most efficient when the feed gas contains the correct ratio of components. The ideal gas composition is given by:

22

22 =+−

=COCOCOHR Equation 2-1

c) The concentration of inert materials (e.g. N2, CH4) should be minimised. The CO2/CO ratio and the stoichiometry number can be adjusted by the water-gas-shift reaction and then CO2 removal. If the raw gas from the gasifier contains significant quantities of CH4, steam reforming may also be used, as follows:

CH4 + H2O → 3H2 + CO (R5) The water-gas-shift reaction is a catalytic process operating at 200-475 ˚C to convert CO and steam to H2 and CO2, via the reaction:

CO + H2O → H2 + CO2 (R6) For the production of methanol from biomass only partial conversion is required. Excess CO2 may then be removed by one of several commercially available processes. Methanol Synthesis Once the economic optimum synthesis gas is available the methanol synthesis takes place. This typically uses a copper-zinc catalyst at temperatures of 200-300 ˚C and pressures of 50-100 bar. Only a portion of the CO in the feed gas is converted to methanol in one pass through the reactor, due to the low temperature at which the catalyst operates. The unreacted gas is recycled at a ratio typically between 2.3 and 6 (Hamelinck et al., 2001).

20

Methanol is produced from syngas via the following two reactions:

CO+2H2 ↔ CH3OH Δh˚ = -90.7 MJ/kmol (R7) CO2+3H2 ↔ CH3OH+H2O Δh˚ = -49.5 MJ/kmol (R8)

Methanol synthesis is exothermic, and gives a net decrease in molar volume. A crude methanol production is condensed out by cooling the product gas of the methanol synthesis reactor, and is then sent to a distillation column (Hamelinck et al, 2001, Ohlström et al., 2001). (R7) is the primary methanol synthesis reaction. Since (R8) results in the loss of some of the hydrogen as water, under ideal circumstances there would be no CO2 in the feed. However, it is well established that a small amount of CO2 in the feed (1-2%) acts as a promoter of the primary methanol synthesis reaction and helps maintain catalyst activity (Hamelinck et al, 2001). Some side reactions are also possible:

2CH3OH↔CH3OCH3+H2O (dimethyl-ether) (R9) CO + H2↔ CH2O (formaldehyde) (R10) 2nH2+nCO↔CnH2n+1OH + (n-1) H2O (higher alcohols) (R11)

Methanol Purification

The crude methanol from the synthesis loop contains water produced during synthesis as well as other minor by-products. Purification is achieved in multistage distillation, with the complexity of distillation dictated by the final methanol purity required (Schuck, 2002). The following figure presents a combined production of methanol.

FIGURE 4. Process flow diagram of methanol production.

Optional are a gas turbine or boiler to employ the unconverted gas, and a steam turbine; resulting in electricity co-production.

Pre- treatment

Gasifier

Gas cleaning section

Reformer for higher hydrocarbons

Shift to adjust CO/H2 ratio

Dried Wood

Raw Gas

Clean Gas

CO, H2

Methanol production

21

Wood based methanol process integrated to CHP production The evaluated methanol process is based on oxygen-driven gasification. After the gasification the gas is reformed with a catalyst to achieve a maximum conversion of hydrocarbons into gases (CO, H2). Water scrubbing of the gas is needed for removing the remaining condensable tars, solids, and ammonia. In addition, the composition of the gas must be converted by CO conversion (shift) units to adjust the stoichiometric ratio in the range required by methanol synthesis. Sulphur compounds and CO2 must be removed with special scrubbing processes. In principal, all process units presented are commercial technology, although no wood-based methanol plants are under operation (Ohlström et al., 2001).

FIGURE 5. Methanol production combined to CHP production at a pulp and paper mill. Steam production and utilisation is combined to a steam cycle of the pulp mill (Ohlström et al., 2001).

22

2.4 Hydrogen Production Hydrogen can be combusted or combined with oxygen inside a fuel cell to produce energy. However, hydrogen gas does not occur naturally in large quantities on earth; it must be separated from other compounds such as water or fuels that contain carbon such as fossil fuels or biomass. Hydrogen can be produced by two processes from biomass: gasification and pyrolysis.

Hydrogen from gasification During gasification, the following reaction occurs:

CaHb + a/2O2 → b/2H2 + aCO (R12) A variety of secondary reactions such as hydro cracking, steam gasification, hydrocarbon reforming, and water-gas shift reactions also take place. The feedstock react with oxygen under severity operating conditions (1,150°C -1,425°C at 400-1,200 psig). In hydrogen production plant, there is an air separation unit (ASU) upstream of the gasifier. Using oxygen rather than air avoids downstream nitrogen removal steps. In some designs, the gasifiers are injected with steam to moderate operating temperatures and to suppress carbon formation (Simbeck et al. 2002). The hot syngas could be cooled directly with a water quench at the bottom of the gasifier or indirectly in a waste heat exchanger (often referred to as a syngas cooler) or a combination of the two. The characteristics of the biomass gasification process are also similar to those of coal gasification systems, but the biomass gasifier operates at lower temperatures and has different clean-up requirements (Ogden, 1999).

Steam methane reforming

The gasification step is followed by a similar steam reforming process; a reforming step which converts methane and the higher hydrocarbons in the syngas to hydrogen. Steam methane reforming (SMR) is an endothermic reaction conducted under 30 atm and temperatures exceeding 870°C. Conventional SMR is a fired heater filled with multiple tubes to ensure uniform heat transfer.

CH4 + H2O ↔ 3H2 + CO (R13) Commercially, the steam to carbon ratio is between 2 and 3. Higher stoichiometric amounts of steam promote higher conversion rates and minimize thermal cracking and coke formation. Because of the high operating temperatures, a considerable amount of heat is available for recovery from both the reformer exit gas and from the furnace flue gas. A portion of this heat is used to preheat the feed to the reformer and to generate the steam for the reformer. Additional heat is available to produce steam for export or to preheat the combustion air. Methane reforming produces a synthesis gas (syngas) with a 3:1 H2/CO ratio. The H2/CO ratio decreases to 2:1 for less hydrogen-rich feedstock such as light naphtha. The addition of a CO shift reactor could further increase hydrogen yield from SMR according to Reaction (R14).

CO + H2O → H2 + CO2 (R14)

23

Shift conversion

The shift conversion may be conducted in either one or two stages operating at three temperature levels. High temperature (350°C) shift utilizes an iron-based catalyst, whereas medium and low (205°C) temperature shifts use a copper based catalyst. Assuming 76% SMR efficiency coupled with CO shift, the hydrogen yield from methane on a volume is 2.4:1.

Hydrogen purification

There are two options for purifying crude hydrogen. Most of the modern plants use multi-bed pressure swing adsorption (PSA) to remove water, methane, CO2, N2, and CO from the shift reactor to produce a high purity product (99.99%). Alternatively, CO2 could be removed by chemical absorption followed by methanation to convert residual CO2 in the syngas. FIGURE 6 below shows a block diagram for the production of hydrogen:

FIGURE 6. Block diagram of the process configurations for the production of hydrogen from biomass. PSA= Pressure Swing Absorption (Katofsky et al, 1993, Komiyama et al., 2001).

Pressure swing absorption (PSA) Pressure swing absorption (PSA) systems have reached a level of performance such that they can produce hydrogen with extremely high purity (99.999%) at recovery rates of 90 % or higher. In Pressure Swing Adsorption (PSA), molecules are physically bound to a surface at high pressure, and released at low pressure (Katofsky 1993). This technology can be applied for various purification purposes, like in hydrogen or oxygen production. Hydrogen purification

Pre- treatment

Gasifier

Steam methane reforming

CO Shift reaction

PSA H2 separation CO2 removal

Dried Wood

Raw Gas

Syngas

Hydrogen

24

by adsorption was first performed commercially by Union Carbide Corp. in 1966. Since then over 400 H2 PSA plans have been installed around the world. Most H2 PSA plants use activated carbon or zeolite adsorbents or both, sometimes in layers with alumina or other adsorbents for impurity removal (LaCava et al. 1998). The adsorption surfaces have to be large and can be selective to particular molecules. Two basic categories are carbonaceous and zeolitic adsorbents, as extensively described by Katofsky (1993). Zeolites are both naturally occurring and manmade, and are also called molecular sieves. Broadly defined, they are silicates of aluminium with alkali metals. The ability of a substance to adsorb a particular gas depends on several factors including pore size, pore size distribution, void fraction and surface activity. Some zeolites contain metal cations, which can attract certain gas molecules. There are literally hundreds of different types of zeolites, with pore sizes ranging from 3 to 10 Å. The size of the gas molecule to be adsorbed is therefore important when selecting which zeolite to use. Macroscopic properties are also important. Sufficient macroporosity is required to permit rapid diffusion of gases from the surface of the adsorbent into the microscopic structure. Greater macroporosity also reduces pressure losses and allows for rapid desorption during bed regeneration. For hydrogen purification from synthesis gas, two sets of PSA beds are placed in series ( FIGURE 14). The gas is cooled down to a temperature of about 40°C before entering the PSA unit. PSA-A removes all the CO2 and H2O, PSA-B removes all residual gasses but 84 % of the hydrogen. By recycling 80 % of the liberated gas from PSA-B to PSA-A, the overall hydrogen recovery is above 90 %. The produced hydrogen is extremely pure (99.999%) and is liberated almost at feed pressure. Besides pure hydrogen, also a highly pure CO2 stream and a combustible purge gas stream, undiluted by inert compounds, are produced (Katofsky 1993).

FIGURE 14. PSA set up for hydrogen purification (Katofsky 1993)

2.5 Power Generation The off gas after the methanol or hydrogen production section can still contain a significant amount of chemical energy. This gas stream may be combusted in a gas turbine, to generate electricity.

25

2.5.1Gas turbine The gas turbine consists of two distinct parts. First the off gas, at high pressure, is put through a combustor. Adding pressurised air generates a large hot stream, which is expanded so electricity is generated. Part of the work is used to drive the air compressor. In some cases the syngas has a caloric value too low for the ensure stable combustion in the combustor. The mass flow will become too large if the standard combustion temperature is to be maintained. Therefore, a lower combustion temperature is necessary when syngas with a low caloric value is put through a gas turbine, also adjustments to the gas turbine have to be made (Ree, Oudhuis et al. 1995). This reduction of the gas turbine combustion temperature is called de-rating. Another form of de-rating is raising the pressure of the air put through the turbine. Heavy de-rating has strong impacts on overall performance and economics foe combined cycles. Roughly spoken, a syngas with LHV higher than 6 MJ/Nm3 can be burned without de-rating, at 1200°C and 14 bar. The burner may need modifications if the syngas is lower than 10 MJ/Nm3 (Walter et al. 1998). Syngas with a caloric value below 5 MJ/Nm3 will probably not be suited for the gas turbine, but this depends on the gas turbine used (LHV’s of 2.5 have been used in gas turbines). When a gas turbine is used the exhaust gas carries a large amount of heat. Steam from cooling the exhaust gas is fed to the Heat Recovery Steam Generation (HRSG). The exhaust gas of the gas turbine cannot be cooled below 170°C due to environmental restrains (Faaij 1997). A gas turbine using HRSG is called combined cycle. 2.5.2 Heat recovery steam generation (hrsg) At different places in the whole process cooling down is necessary, e.g. the exhaust gas from the gas turbine, the cooling down of the syngas after gasification, etc. Water is pressurised in advance of the heat exchanger, so high pressure steam is generated. The superheated steam is expanded in a partly condensing gas turbine to produce electricity. Also, steam required in the process (drying, gasification, shifting and reforming) can be taken from the HRSG at different pressures. 3. SYSTEM CALCULATIONS 3.1 Modelling principles Modelling of the concepts has been performed with the flowsheet program Aspen Plus, on basis of the flowsheet presented in FIGURE 15 for methanol and FIGURE 16 for hydrogen.

26

W

W

FIGURE 15. Aspen flowsheet used for the calculation of the mass and energy balances for the methanol production.

Reforming Section

Shift Section

Gasification and cleaning section

CO2 Removal

MeOH Section

Gas Turbine

Steam Turbine

Gasifier

Reformer

Shift reactor

Scrubber

MeOH Reactor

Separator

Gasifier

Compresor Expander

Heat Recovery

Power

Power

27

W

W

FIGURE 16. Aspen flowsheet used for the calculation of the mass and energy balances for hydrogen production.

Reforming Section

Shift Section

Gasification and cleaning section

CO2 Removal H2

Production

Gas Turbine

Steam Turbine

Gasifier

Reformer

Scrubber

Heat Recovery

28

The selected systems are modelled in Aspen Plus, a widely used process simulation program. In this flowsheeting program, chemical reactors, pumps, turbines, heat exchanging apparatus, etc are virtually connected by pipes. Every component can be specified in detail: reactions taking place, efficiencies, dimensions of heating surfaces and so on. For given inputs, product streams can be calculated, or one can evaluate the influence of apparatus adjustments on electrical output. The plant efficiency can be optimised by integrating the heat supply and demand. The resulting dimensions of streams and units and the energy balances can subsequently be used for economic analyses. The pre-treatment section is not modelled, their energy use and conversion efficiencies are included in the energy balances, though. The models start with the gasification section with a composition of biomass given in TABLE 5. Only the base scale of 80 dry tonne/hour (430 MWth) biomass is modelled.

TABLE 5. Biomass composition (Consonni, Larson 1994).

The heat supply and demand within the plant is carefully matched, aimed at maximising the production of superheated steam for the steam turbine. A summation of all heat inputs and outputs in a heat bin is too simple, since it does not take the quality of heat into account. Heat integration of heat demand and supply within the considered plants here is done by hand. The intention is to keep the integration simple by placing few heat exchangers per gas/water/steam stream. First, an inventory of heat supply and demand is made. Streams matching in temperature range and heat demand/supply are combined: e.g. heating before the reformer by using the cooling after the reformer. When the heat demand is met, steam can be raised for power generation. Depending on the amount and ratio of high and low heat, process steam is raised in heat exchangers, or drawn from the steam turbine. Steam for gasification and drying is almost

Feed wood Equivalent fuel Moisture content 50 LHV, MJ/Kg 8.11 HHV, MJ/Kg 10 CH4 6.950 CO 10.780 Dried wood CO2 19.720 Moisture content 15 C2H2 0.016 LHV, MJ/Kg 15.5 C2H4 3.157 HHV, MJ/Kg 17 C2H6 2.229 C3H8 0.042 Bone dry wood H2 0.344 C 49.98 N2 0.016 H 6.12 O2 0.721 O 42.49 C 5.250 N 0.55 H2O 50.000 S 0.06 NH3 0.317 Ash 0.8 SO2 0.048 LHV, MJ/Kg 18.66 H2S 0.006 HHV, MJ/Kg 20

29

always drawn from the steam cycle, unless a perfect match is possible with a heat-supplying stream. The steam entering the steam turbine is set at 86 bar and 501 °C. 3.2 Results TABLE 6 summarises the outcomes of the flowsheet models. TABLE 6. Results of the Aspen performance calculations, for 430MWth input HHV systems of the methanol and hydrogen production concepts considered.

Methanol Production HHV Fuel Output (MW) HHV Efficiency (%) Material (MJ/Kmol) Net electricity 1) Fuel 2) (ton/hr) IGT Gasifier, Scrubber, Steam Reforming, 187 43 54 28.2 Shift Reactor, Gas Phase Methanol Reactor, Combine Cycle.

Hydrogen Production HHV Fuel Output (MW) HHV Efficiency (%) Material (MJ/Kmol) Net electricity 1) Fuel 2) (ton/hr) IGT Gasifier, Scrubber, Steam Reforming, 225 26 60 5.7 Shift Reactor, Pressure Swing Adsorption, Combine Cycle. 1) Net electrical output is gross output minus internal use. Gross electricity is produced by gas turbine and/or steam turbine. The internal electricity use stems from pumps, compressors, oxygen separator, etc. 2) The electricity part is assumed to be produced from biomass at ηe = 45 % HHV efficiency (Faaij et al. 1998). The Fuel only efficiency is calculated by η = Fuel/(MWth,in – Electricity/ηe). Based on experiences with low calorific combustion elsewhere (van Ree et al. 1995; Consonni and Larson 1994a) the streams in this study, which were projected to be combusted in a gas turbine, will give stable combustion. The overall plant efficiency for the methanol concept lie in a close range: methanol 54% and hydrogen 60%.

30

4. ECONOMICS 4.1 Method The methanol and hydrogen production costs are calculated by dividing the total annual costs of a system by the produced amount of fuel. The total annual costs consist of:

• Annual investments • Operating and Maintenance • Biomass feedstock • Electricity supply / demand (fixed power price)

The total annual investment is calculated by a factored estimation (Peters and Timmerhaus 1980), based on knowledge of major items of equipment as found in literature. The uncertainty range of such estimates is up to ± 30 %. The installed investment costs for the separate units are added up. The unit investments depend on the size of the components (which follow from the Aspen+ modelling), by scaling from known scales in literature (see TABLE 21 in Annex H), using Equation 4-1:

Cos ta / Costb = (Sizea / Sizeb ) R Equation 4-1

with R = scaling factor Various system components have a maximum size, above which multiple units will be placed in parallel. Hence the influence of economies of scale on the total system costs decreases. This aspect is dealt with by assuming that the base investment costs of multiple units are proportional to the cost of the maximum size: the base investment cost per size becomes constant. The total investment costs include auxiliary equipment and installation labour, engineering and contingencies. If only equipment costs, excluding installation, are available, those costs are increased by applying an overall installation factor of 1.86. This value is based on 33% added investment to hardware costs (instrumentation and control 5%, buildings 1.5% grid connections 5%, site preparation 0.5%, civil works 10%, electronics 7%, and piping 4%) and 40 % added installation costs to investment (engineering 5%, building interest 10%, project contingency 10%, fees/overheads/profits 10%, start-up costs 5%) (Faaij et al. 1998). The annual investment cost follows from Equation 4-2 which takes the technical and economic lifetime of the installation into account. The interest rate is 10 %.

( )( ) ⎟⎟

⎠

⎞⎜⎜⎝

⎛ −⋅

+−⋅×

+−

=t

ettt

t

annual ttt

IRI

IR

IRIe

e

111

111

Equation 4-2

with Iannual = Annual investment costs IR = Interest rate = 10% It = Total investment (sum of the unit investments) te = Economical lifetime= 15 years tt = Technical lifetime = 25 years

31

Operational costs (maintenance, labour, consumables, residual streams disposal) are taken as a single overall percentage (4 %) of the total installed investment (Faaij et al. 1998; Larson et al. 1998). It is assumed in this is study that the biomass price is 110 SEK/MWh and electricity supplied to or demand from the grid costs is 360 SEK/ MWh. The annual load is 8000 hours. 4.2 Results Results of the economic analysis are given in TABLE 7. The 400 MWth conversion facilities deliver methanol at 93.9 SEK/GJ and hydrogen at 75.75 SEK/GJ. It can be seen that the costs for the gasification system, syngas processing and power generation generally make up the larger part of the investment. Developments in the gasification and reforming technology are important to decrease the investments. On a longer term capital costs may reduce due to technological learning: a combination of lower specific component costs and overall learning. A third plant may be 15% cheaper leading to 8 – 15% fuel cost reduction (Hamelinck et al, 2001). The last rows of TABLE 7 show potential fuel production costs in smaller or bigger installations. Going to 1000 and 2000 MWth scales the fuel production costs reach costs levels as low as 78 SEK/GJ for methanol and 63.71 SEK/GJ for hydrogen. On the long term different cost reductions are possible concurrently (Tijmensen 2000). Capital costs for a third plant built are 15% lower, and a large (2000 MWth) plants profit from economies of scale. FIGURE 17. Optimistic view scenario. Different cost reductions are predictable: (1) technological learning reduces capital investment by 15% and (2) application of a large scale (2000 MWth) reduces unit investment cost.

1

2

1

2

32

TABLE.7 Economic analysis for the concepts considered, Costs in MSEK Jan 2006

UNIT METHANOL HYDROGEN Total Pre-treatment MSEK 384.272 384.272Gasification System - IGT - 719.843 719.843 Oxigen plant - 278.010 278.010Gas Cleaning - Tar Cracker - Cyclones - 9.716 9.716 HTHX - 49.345 49.345 Bag house filter - 7.383 7.383 Condensing Scrubber - 10.514 10.514Syngas Procesing - Compresor - 17.602 17.602 Steam Reformer - 493.267 493.267 Shift reactor (installed) - 128.366 128.366 Selexol CO2 removal (ins) - 138.445 0.000Methanol Production - Make up Compressor - 114.678 0.000 Gas Phase Methanol - 75.373 0.000 Refining - 144.273 0.000Hydrogen Production - PSA units A+B - 0.000 371.442Power Isle - Gas Turbine + HRSG - 299.515 154.492

Steam Turbine + steam system - 92.862 108.457

Total Installed Invesment MSEK 2963.463 2732.708Total Invesment corrected MSEK 901.613 831.408 Biomass input dry tonne/hr 80.000 80.000 Biomass input MWth 428.400 428.400 Load hours 8000.000 8000.000 Biomass input GJ/year 12.338 12.338 Annual Cost Capital MSEK 118.539 109.308 OM MSEK 39.851 36.748 Biomass MSEK 376.992 376.992 Cost/Income Power MSEK -79.978 -58.821Total annual costs 455.403 464.227 Producction Fuel output MW HHV 186.809 225.001 Power output Mwe 27.800 20.424 Efficiency fuel % 53.6 59.3 Cost of fuel produced 80 MWth SEK/GJ 137.293 110.751 400MWth SEK/GJ 93.901 75.748 1000MWth SEK/GJ 84.468 68.138 2000MWth SEK/GJ 78.980 63.711

33

Diesel and gasoline production costs vary strongly depending on crude oil process, but for an indication: current gasoline market prices lie in the range of: 12.98 US $/GJ (EIA, 2006). FIGURE 18 shows the production costs for methanol and hydrogen in this study plant. A comparison was done between the biofuels and gasoline production costs, and it can be seen that methanol and hydrogen production costs are slightly under the gasoline price.

Methanol Hydrogen

9.5 US$

75.8 SEK

11.8 US$

93.9 SEK

0

10

20

30

40

50

60

70

80

90

100

Pric

e/G

J

Gasoline Production Cost 12.98 US$/GJ(EIA 2006)

FIGURE 18. Comparison between biofuels and gasoline production costs in US$/GJ.

34

5. DISCUSSION AND CONCLUSIONS The methanol and hydrogen production plants were modelled using the Aspen Plus flowsheet program. Mass balances and internal heat demand and supply were carried out. Both models gave directly the dimensions of streams and units for the economic calculation. The biofuel production costs are calculated by dividing the total annual costs of a system by the produced amount of fuel. Unit sizes, resulting from the plant modelling, are used to calculate the total installed capital of biofuel plants, larger units benefit from cost advantages. Conventional production costs of gasoline are about 13 US $/GJ according to Energy Institution of America. Production costs of methanol and hydrogen from biomass are competitive with conventional prices of gasoline according to this study. Nevertheless considering the 30% uncertainty rage in the estimates of the economic method applied one should be careful in ranking the concepts. Long term cost reductions mainly reside in slightly lower biomass costs, technological learning, and application of large scales (2000 MWth). This could bring the methanol and hydrogen production costs in the range of gasoline/diesel and even lower for sure. Hydrogen as the ultimate fuel for fuel cell vehicles, has a high fuel economy and low costs per km driven, and will compete with gasoline. However, hydrogen requires new distribution infrastructure which is the main bottleneck.

A methanol distribution system is probably easier to realise and FCVs fuelled by on-board reformed methanol will initially have a greater range.

35

6. REFERENCES Bergman, P.C.A., Boersma, A.R., Kiel, J.H.A., Prins, M.J., Ptasinski, K.J., Janssen, F.J.J.G. “Torrefaction for Entrained-Flow Gasification of Biomass.” Presented at The 2nd World Conference and Technology Exhibition on Biomass for Energy, Industry and Climate Protection, Rome, Italy, 10-14 May 2004. Brown, L.F. “A Comparative Study of Fuels for On-Board Hydrogen Production for Fuel–Cell-Powered Automobiles.” International Journal of Hydrogen Energy 26 (2001) 381-397. Brown, R.C. “Biorenwable resources, Engineering New Products from Agriculture.” ISBN 0-8138-2263-7, 2003. Boerrigter, H., Van Der Drift, A. ““BIOSYNGAS” Description of R&D trajectory necessary to reach large-scale implementation of renewable syngas from biomass.” ECN-C--04-112, December 2004. Bridgwater, A.V. “The Technical and Economic Feasibility of Biomass Gasification for Power Generation.” Fuel Vol 74, No5, pp. 631-653, 1995. Chmielniak, T., Sciazko, M. “Co-gasification of biomass and coal for methanol synthesis”. Applied Energy 74, 393–403, 2003. Consonni S, Larson E. “Biomass-gasifier/aeroderivative gas turbine combined cycles. Part A: Technologies and Performance modelling. Part B: Performance Calculations and Economic Assessment.” in Proceeding of Mechanical Engineers’ 8th congress on gas turbines in Cogeneration and utility, industrial and independent power generation, Portland, Oregon, 1994. De Boer, A.J., Den Uil, H. “An Evaluation of tree Routes for the Production of Liquid Fuels from Biomass.” ECN-R-97-001, January 1997. Deglise , X., Magne, P. “Pyrolysis and Industrial Charcoal.” Biomass: Regenerable Energy. Ed. D.O. Hall and R.P. Overend. Chichester: John Wiley & sons, 1987. EIA, 2006, Energy Information Administration, http://www.eia.doe.gov. Faaij A. “Long term perspectives of biomass integrated gasification with combined cycle technology: Cost and efficiency and a comparison with combustion.” Utrecht University, 1998. Faaij A. “Production of methanol and hydrogen from biomass via advanced conversion concepts - preliminary results.” 1st World Conf. & Exhibition on Biomass for Energy and Industry, Sevilla, Spain, June 2000. Gabroski, M.S. “Kinitics of Char Gasification.” Biomass Gasification: Principles and Technology. Ed. T.B. reed. Park Ridge New Jersey: Noyes Data Corporation, 1981.

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Grégoire Padró, C., E., Putsche, V. “Survey of the Economics of hydrogen Technologies.” Technical Report, National Renewable Energy Laboratory, September 1999. Hamelinck, C.N. “Potential for CO2 sequestration and enhanced coal bed methane production in the Netherlands.” Netherlands agency for energy and the environment NOVEM, Utrecht, the Netherlands, 51 pp + annexes, 2000. Hamelinck, Faaij. “Future prospects for production of methanol and hydrogen from biomass.” September 2001 Hamelinck, C.N., van Hooijdonk, G., Faaij, A. “Prospects for Ethanol from Lignocellulosic Biomass: Techno-economic Performance as Development Progresses.” ISBN 90-393-2583-4, November 2003. Henrich, E, Weirich, F. “Pressurized Entrained Flow Gasifiers for Biomass.” Environmental Engineering Science 21 (1) 53-64, 2004. Hon, D.N.S., Shiraishi, N. “Wood and Cellulosic Chemistry. Second edition, revised and expanded”. ISBN 0-8247-0024-4 (2001). Intergovernmental Panel on Climate Change (IPCC). “Climate Change 2001: the scientific basis.” Cambridge University Press, Cambridge, 2001a. Iwasaki, W. “A Consideration of the Economic Efficiency of Hydrogen Production from Biomass.” International Journal of Hydrogen Energy, 28, (2003), 939-944. Johansson, T.B. “Renewable Energy - Sources for Fuels and Electricity.” Earthscan Publications Ltd, London, 1993. ISBN 1-85383-155-7. Katofsky, Ryan E. “The Production of fluid fuels from biomass.” Princeton, N.J. 1993. King, D.L., Bochow Jr, C.E. “What should an owner/operator know when choosing an SMR/PSA plant?” Hydrocarbon Processing(Issue), pp 39-48, 2000. KPI, Konsument Pris Index, http://www.scb.se/templates/Product____33769.asp. Komiyama, H., Mitsumori, T., Yamaji, K., Yamada, K. “Assessment of energy systems by using biomass plantation.” Fuel 80 (2001) 707-715. LaCava, A.I., Shirley, A.I., Ramachandran, R. “How to specify pressure-swing adsorption units - key components of PSA units.” Chemical Engineering, 105(Issue), pp 110-118, 1998. Lide, D.R., “Handbook of Chemestry and Pysics.” 73rd edition, 1992-1993, ISBN: 0-8493-0566-7. Lynd, L.R. “Overview and Evaluation of Fuel Ethanol from Cellulosic Biomass: Technology, Economics, the Environment, and Policy.” Annual Reviews, Energy Environment, 21, 403-465, 1996.

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Maiya PS, Anderson TJ, Mieville RL, Dusek JT, Picciolo JJ and Balachandran U. “Maximizing H2 production by combined partial oxidation of CH4 and water gas shift reaction.” Applied Catalysis A: General, 196: 65-72, 2000. Mann, M.K. “Technical and economic analyses of hydrogen production via indirectly heated gasification and pyrolysis,” in Proceedings of the 1995 US DOE Hydrogen Program Review, Vol. 1, NREL/CP-430-20036-Vol. 1, pp. 205-236, 1995. Milne, T.A:, Elam, C.C., Evans, R.J. “Hydrogen from biomass, state of the art and Research challenges.“ Report for the international Energy Agency, IEA/H2/TR-02/001. Ogden, J. “Prospects for building a hydrogen energy infrastructure.” Annual Review Energy Environment 24:227-279, 1999. Ogier, J.C., Ballerini, D., Leygue, J.P., Rigal, L., Pouriquié, J. “Production d’éthanol à partir de biomasse lignocellulosique.” Oil & Gas Science and Technology – Revue de l’IFP, Vol. 54, No1, pp 67-94, 1999. Ohlström, M., Laurikko, J., Mäkinen, T., Pipatti, R. “New Concepts for Biofuels in Transportation, Biomass-based Methanol Production and Reduced Emissions in Advanced Vehicles.” ISBN 951-38-5780-8, 2001. Okuda, K., Man, X., Umetsu, M., Takami, S., Adschiri, T. “Efficient conversion of lignin into single chemical species by solvothermal reaction in water–p-cresol solvent.” Journal of Physics: Condensed Matter 16 S1325–S1330, 2004. Olsson, L. Ethanol Production from lignocellulosic Materials: Fermentation and On Line Analysis. Doctoral Thesis, April 1994. Perry, R.H., Green, D.W., Maloney, J.O. “Perry’s chemical handbook.” Sixth edition, Singapore: McGraw-Hill Book Co., 1987. Pierik, J., Curvers, A. “Logistics and pretreatment of biomass fuels for gasification and combustion.” Netherlands Energy Research Foundation ECN, Petten, 1995. Pridmore, A., Bristow, A. “The role of hydrogen in powering road transport.” Tyndall Centre for Climate Change Research Working Paper 19, April 2002. Prins, M.J., Ptasinski, J.K., Janssen, J.J.G. “Exergetic Optimisation of a Production Process of Fischer-Tropsch Fuels from Biomass.” Fuel Processing Technology 86, 375-389 (2004). Reith, J.H., Veenkamp, J.M., van Ree, R. “Co-Production of Bio-Ethanol, Electricity and Heat from Biomass wastes, Potential and R&D Issues.” Netherlands Energy Research Foundation ECN. ECN, Nedalco, ATO contribution to The First European Conference on Agriculture & Renewable Energy 6-8 May 2001, RAI, Amsterdam, The Netherlands ECN-RX--01-011 Shleser, R. “Ethanol Production in Hawaii. Prepared for the State of Hawaii.” Department of Business, Economic Development and Tourism. Honolulu: Energy Division, Dept. of Business, Economic Development and Tourism, State of Hawaii, 1994.

38

Simbeck, D.R., Chang, E. “Hydrogen Supply: Cost Estimate for Hydrogen Pathways – Scoping Analysis”. National Renewable Energy Laboratory, SR-540-32525, July 2002. Schuck, S. “Wood for Alcohol Fuels Status of Technology and Cost/Benefit Analysis of Farm Forestry for Bioenergy.” Report for the RIRDC/Land & Water Australia/FWPRDC/MDBC, November 2002. Sørensen, Å.L. Economies of Scale in Biomass Gasification Systems. IIASA, IR-05-030, March 2003. Sudo, S., Takahashi, F., Tekeuchi, M. “Chemical Properties of Biomass.” Biomass Handbook. Ed. Osamu Kitani and Carl W. Hall. New York: Gordon and Breach Science Publishers, 1989. Thomas.C.E., James, B.D., Lomax Jr, F.D., Kuhn Jr, I.F. “Fuel Option for the Fuel Cell Vehicle: Hydrogen, Methanol, or Gasoline?” International Journal of Hydrogen Energy 25 551-567, 2000. Tijmensen, M. “The Production of Fischer Tropsch Liquids and Power Through Biomass Gasification.” Utrecht University, NWS-E-2000-29, ISBN 90-73958-62-8, November 2000. Tijmensen, M., Faaij; A., Hamelinck, C.N., van Hardeveld, M. “Exploration of the possibilities for production of Fischer Tropsch liquids and power via biomass gasifcation.” Biomass and Bioenergy 23 129 – 152, 2002. Tillmann, D. A. “Biomass cofiring: the technology, the experience, the combustion consequences.” Biomass & Bioenergy 19, 365, 2000. Turn S.Q., Kinoshita, C.M., Ishimura, D.M., Zhou, J., Hiraki, T.T., Masutani, S.M., “A review of sorbent materials for fixed bed alkali getter systems in biomass gasfier combined cycle power generation applications.” Journal of the Institute of Energy, 71: pp 163-177, 1998. Van der Drift, B., van ree, R., Boeringter, H., Hemmes, K. “Bio-syngas: key intermediate for Large Scale Production of Green Fuels and Chemicals.” Presented at The 2nd World Conference and Technology Exhibition on Biomass for Energy, Industry and Climate Protection in Rome, Italy, 10-14 May 2004. Van Ree, R. “Advanced biofuels for transportation.” ECN-RX--04-055, June 2004. Verschoor, M., Melman, A.G. “System Study high-temperature Gas Cleaning at IGCC Systems.” Summary of a system study by ECN, Kema, Stork Boilers and TNO as part of the Netherlands Clean Coal Programme. Novem brochure SO2.28 91.06, 1991. Wadmark, G. Batelle-projektet/slutrapport. Statens Energiverk/Swedish Energy Administration, Report number BF-86/4, 1984. Wenzl, H. F. J. “The chemical Technology of Wood.” New York and London, Academic Press. 1970.

39

White, L.R., Tompkins, T.L., Hsieh, K.C., Johnson, D.D. “Ceramic filters for hot gas cleanup.” International gas turbine and aeroengine congress and exposition, Cologne, Germany, 8, 1992. Wyman, C.E., Hinman, N.D. “Ethanol: Fundamentals of Production from Renewable Feedstocks and Use as a Transportation Fuel.” Applied Biochemistry and Biotechnology, Vol. 24/25, 1990. Yan, J. “Biomass Gasification, Power Generation Technologies, The State-of-the-Art of Research, Development and Demonstration.” Technical Report, Energy Processes, Department of Chemical Engineering and Technology, Royal Institute of Technology, ISSN 1104 – 3466, April 1998.

40

7. ANNEX A : Properties of MeOH and H2 TABLE 8.Properties of methanol and hydrogen (Lyde, 1993). Property Unit Methanol Hydrogen Chemical Formula CH3OH H2 Molecular Weight kg/kmol 32.042 2.02(x) Composition, Weight % Carbon 37.5 0 Hydrogen 12.6 100 Oxygen 49.9 0 Melting Point °C -97.7 -259.4 Relative Density 0.79 Density at 20 °C kg/m³ 791 Critical properties Critical temperature K 512.6 33 Critical pressure MPa 8.092 1.293 Critical density g/cm³ 0.272 0.031 Critical compression factor 0.224 0.307 Viscosity @25˚C mPas 0.544 8.81 .10-3 Auto ignition temperature °C 385 1,050–1,080 Flammability limits: Lower / Higher volume % 6 / 37 4.0 / 75.0 Heat of Formation Liquid kJ/mol -239.1 Gas kJ/mol -201.3 0.0 Gibbs Free Energy Liquid kJ/mol -166.6 Gas kJ/mol -162.62 Entropy Liquid J/mol.K 126.8 Gas J/mol.K 43.9 130.7 Heat capacity@ cst P @25 ˚C J/gK 2.53 J/mol.K 44.06 Boiling point (at atmospheric pressure) °C 64.6 -252.8 Heat of Vaporization kJ/mol 35.21 0.9 Vapour pressure @ 25 ˚C kPa 17 Octane no.(1) Research octane no. 107 130+ Motor octane no. 92 – (RON + MON)/2 100 –

41

ANNEX B: Abbreviations BIG/CC Biomass Integrated Gasification Combined Cycle COP Conference of the parties FCV Fuel Cell Vehicle GCH Gas converted Hydrogen P GCM Gas converted Methanol Process GHG Greenhouse gas HHV Higher Heating Value ICEV Internal Combustion Engine Vehicle IPCC Intergovernmental Panel on Climate Change PSA Pressure Swing Absorption. Syngas abbreviation for synthesis gas Synthesis gas A gas containing primarily hydrogen (H2) and carbon monoxide

(CO), or mixture of H2 and CO; intended for synthesis in a reactor to form methanol and /or other hydrocarbons (synthesis gas may also contain CO2, water, and other gases).

WGS Water Gas Shift

42

ANNEX C: Aspen Floowsheet for Methanol Production.

FUEL1

OXIGEN

STEAM-G

2SCRUBBER

3

W-WATER

CLE-H20

GASIFIER

C-WATER1

STEAM-1

HX1COMP1

5

REFORMER

6

HX3 7

C-WATER2

STEAM-2

HX2

S-8BAR

S-8BAR2

SHIFT-R

8

STEAM-SR

9

CO2

RE-CO2

COMP3

10HX4

11

C-WATER3

STEAM-3

RE-MEOH

12

HX5

C-WATER4

STEAM-4

13SE-MEOH

F-GASES

MEOH

COMP2

S-1BAR

CSYN

AIR COMP

AIR

CAIR

GAS-2

GT

LPCSYN

MIX

15

HR SG

COOLGAS

18

HST

SPLIT-1

MST

SPLIT-2

LSTST-W1 ST-W2 ST-WOUTW

GT-W1 GT-WOUTW

43

ANNEX D: Stream Results for Methanol Production. TABLE. 9 Stream Results for the Syngas System in Methanol Production. Substream: MIXED Mole Flow kmol/hr FUEL STEAM-G OXIGEN 1 2 3 4 5 CH4 611.179 0 0 175.088 175.088 175.083 175.083 175.083 CO 542.953 0 0 1793.638 1793.638 1793.633 1793.633 1793.633 CO2 632.150 0 0 966.627 966.627 966.318 966.318 966.318 C2H2 0.867 0 0 2.44E-05 2.44E-05 2.44E-05 2.44E-05 2.44E-05 C2H4 158.762 0 0 0.005 0.005 0.00459 0.00459 0.005 C2H6 104.579 0 0 0.010 0.010 0.00953 0.00953 0.010 C3H8 1.344 0 0 1.78E-06 1.78E-06 1.78E-06 1.78E-06 1.78E-06 H2 240.744 0 0 2154.057 2154.057 2154.046 2154.046 2154.046 N2 0.806 0 0 13.749 13.749 13.749 13.749 13.749 O2 31.788 0 744.028 5.21E-15 5.21E-15 0 0 0 C 616.655 0 0 1.75E-22 1.75E-22 0 0 0 H2O 521.249 1321.545 0 1476.899 1476.899 337.933 337.933 337.933 NH3 26.260 0 0 0.373 0.373 0.031 0.031 0.031 SO2 1.057 0 0 2.19E-07 2.19E-07 2.13E-07 2.13E-07 2.13E-07 H2S 0.248 0 0 1.305 1.305 1.293 1.293 1.293 CH30H 0 0 0 0.001 0.001 4.72E-05 4.72E-05 4.72E-05 Total Flow kmol/hr 3490.640 1321.545 744.0279 6581.752 6581.752 5442.101 5442.101 5442.101 Total Flow kg/hr 79360 23808 23808 126976 126976 106437.115 106437.115 106437.115 Total Flow cum/hr 10470.479 1456.704 905.986 21522.962 12329.785 139756.703 10912.765 13032.589 Temperature C 600 250 25 900 400 40 563.523 727 Pressure bar 20 34.5 20 30 30 1.013 35.013 35.013 Vapor Frac 0.823 1 1 1 1 1 1 1 Liquid Frac 0.177 0 0 0 0 0 0 0 Solid Frac 0 0 0 0 0 0 0 0 Enthalpy kcal/mol -22.274 -56.299 -0.0445 -26.970 -31.468 -29.441 -25.209 -23.765 Entropy cal/mol-K 2.014 -13.502 -6.0361 10.541 5.569 8.895 9.689 11.265

44

TABLE. 9 Continuation of the Stream Results for the Syngas System in Methanol Production. Substream: MIXED Mole Flow kmol/hr 6 7 STEAM-SR 8 9 CO2 10 11 12 13 F-GASES MeOH CH4 43.771 43.771 0 43.771 43.771 0 43.771 43.771 43.771 43.771 43.624 0.146 CO 1924.945 1924.945 0 1482.208 1482.208 0 1482.208 1482.208 592.883 592.883 592.588 0.295 CO2 966.318 966.318 0 1409.056 7.045 1402.01037 7.045 7.045 7.045 7.045 6.806 0.239 C2H2 2.44E-05 2.44E-05 0 2.44E-05 2.44E-05 0 2.44E-05 2.44E-05 2.44E-05 2.44E-05 2.33E-05 1.06E-06 C2H4 0.005 0.005 0 0.005 0.005 0 0.005 0.005 0.005 0.005 0.005 5.57E-05 C2H6 0.010 0.010 0 0.010 0.010 0 0.010 0.010 0.010 0.010 0.009 0.0001 C3H8 1.78E-06 1.78E-06 0 1.78E-06 1.78E-06 0 1.78E-06 1.78E-06 1.78E-06 1.78E-06 1.75E-06 2.90E-08 H2 2547.982 2547.982 0 2990.719 2990.719 0 2990.719 2990.719 1212.070 1212.070 1211.747 0.322 N2 13.749 13.749 0 13.749 13.749 0 13.749 13.749 13.749 13.749 13.735 0.014 O2 0 0 0 0 0 0 0 0 0 0 0 0 C 0 0 0 0 0 0 0 0 0 0 0 0 H2O 206.621 206.621 1030.844 794.728 794.728 0 794.728 794.728 794.728 794.728 1.607 793.122 NH3 0.031 0.031 0 0.031 0.031 0 0.031 0.031 0.031 0.031 0.005 0.025 SO2 2.13E-07 2.13E-07 0 2.13E-07 2.13E-07 0 2.13E-07 2.13E-07 2.13E-07 2.13E-07 9.01E-08 1.23E-07 H2S 1.29336047 1.29336047 0 1.29336047 0 1.293 0 0 0 0 0 0 CH30H 4.72E-05 4.72E-05 0.00E+00 4.72E-05 4.72E-05 0.00E+00 4.72E-05 4.72E-05 889.325 889.325 9.106 880.219 Total Flow kmol/hr 5704.725 5704.725 1030.844 6735.569 5332.266 1403.304 5332.266 5332.266 3553.616 3553.616 1879.233 1674.384 Total Flow kg/hr 106437.115 106437.115 18570.950 125008.065 63261.789 61746.276 63261.789 63261.789 63261.789 63261.789 20746.735 42515.054 Total Flow cum/hr 12931.120 6467.952 23.527 7585.590 6032.186 1522.416 4858.318 4509.230 2853.555 1013.695 1589.273 50.540 Temperature C 672.85 200 210 200 200 200 268.433 230 230 30 30 30 Pressure bar 35 35 35 35 35 35 50 50 50 50 30.12 30.12 Vapor Frac 1 1 0 1 1 1 1 1 1 0.528 1 0 Liquid Frac 0 0 1 0 0 0 0 0 0 0.472 0 1 Solid Frac 0 0 0 0 0 0 0 0 0 0 0 0 Enthalpy kcal/mol -21.896 -25.758 -64.962 -31.086 -14.978 -92.348 -14.475 -14.761 -28.162 -34.457 -9.335 -62.640 Entropy cal/mol-K 11.181 5.549 -30.381 2.532 2.461 -2.086 2.739 2.191 -7.492 -22.940 1.083 -48.662

45

TABLE. 10 Stream Results for the Steam System in Methanol Production.

Substream: MIXED

Mole Flow kmol/hr C-WATER1 STEAM-

1 CLE-H2OW-

WATER S-1BAR S-8BAR S-8BAR2 C-

WATER2 STEAM2 C-

WATER3 STEAM3C-

WATER4 STEAM4 CH4 0 0 0 0.005 0 0 0 0 0 0 0 0 0 CO 0 0 0 0.005 0 0 0 0 0 0 0 0 0 CO2 0 0 0 0.309 0 0 0 0 0 0 0 0 0 C2H2 0 0 0 1.18E-08 0 0 0 0 0 0 0 0 0 C2H4 0 0 0 2.19E-07 0 0 0 0 0 0 0 0 0 C2H6 0 0 0 3.13E-07 0 0 0 0 0 0 0 0 0 C3H8 0 0 0 8.07E-12 0 0 0 0 0 0 0 0 0 H2 0 0 0 0.012 0 0 0 0 0 0 0 0 0 N2 0 0 0 4.94E-05 0 0 0 0 0 0 0 0 0 O2 0 0 0 0 0 0 0 0 0 0 0 0 0 C 0 0 0 0 0 0 0 0 0 0 0 0 0 H2O 2470.125 1543.134 109802 110941 1301.673 1301.673 1301.673 1301.673 1301.673 499.576 118.000 2775.422 2775.422 NH3 0 0 0 0.342 0 0 0 0 0 0 0 0 0 SO2 0 0 0 6.65E-09 0 0 0 0 0 0 0 0 0 H2S 0 0 0 0.012026 0 0 0 0 0 0 0 0 0 CH30H 0 0 0 0.000757 0 0 0 0 0 0 0 0 0 Total Flow kmol/hr 2470.125 1543.134 109802 110942 1301.673 1301.673 1301.673 1301.673 1301.673 499.576 118 2775.422 2775.422 Total Flow kg/hr 44500 27800 1978117 1998656 23450 23450 23450 23450 23450 9000 2125.8 50000 50000 Total Flow cum/hr 44.153 150060 1963 2041.668 102139.3 21941.85 14385.729 23.267 102221 8.930 4920.576 49.610 53393 Temperature C 10 896.571 10 40 671 1347.981 792.081 10 671.758 10 230.265 10 101.537 Pressure bar 1 1 1.0133 1.0133 1 8 8 1 1 1 1 1 1 Vapor Frac 0 1 0 0 1 1 1 0 1 0 1 0 0.623 Liquid Frac 1 0 1 1 0 0 0 1 0 1 0 1 0.377 Solid Frac 0 0 0 0 0 0 0 0 0 0 0 0 0 Enthalpy kcal/mol -69.017 -49.833 -69.017 -68.433 -52.096 -44.883 -50.919 -69.017 -52.089 -69.017 -56.085 -69.017 -60.957 Entropy cal/mol-K -41.108 1.500 -41.108 -39.147 -0.647 0.939 -3.600 -41.108 -0.639 -41.108 -6.298 -41.108 -18.898

46

TABLE. 11 Stream results for the Gas Turbine in Methanol Production. Substream: MIXED Mole Flow kmol/hr F-GASES AIR CAIR CSYN LPCSYN CH4 43.624 0 0 3.23E-09 3.23E-09 CO 592.588 0 0 15.073 15.073 CO2 6.806 0 0 637.079 637.079 C2H2 0 0 0 1.70E-16 1.70E-16 C2H4 0.005 0 0 1.47E-18 1.47E-18 C2H6 0.009 0 0 2.56E-21 2.56E-21 C3H8 1.75E-06 0 0 0 0 H2 1211.747 0 0 12.518 12.518 N2 13.735 3762.028 3762.028 3775.765 3775.765 O2 0 989.330 989.330 3.83E-06 3.83E-06 C 0 0 0 1.85E-21 1.85E-21 H2O 1.607 0 0 1306 1306 NH3 0.005 0 0 0 0 SO2 9.01E-08 0 0 8.89E-08 8.89E-08 H2S 0 0 0 1.20E-09 1.20E-09 CH30H 9.106 0 0 2.86E-11 2.86E-11 Total Flow kmol/hr 1879 4751 4751 5747 5747 Total Flow kg/hr 20747 137045 137045 157792 157792 Total Flow cum/hr 1589 113758 10887 25456 378572 Temperature C 30 15 513 1200 519 Pressure bar 30.120 1 28.792 27.792 1 Vapor Frac 1 1 1 1 1 Liquid Frac 0 0 0 0 0 Solid Frac 0 0 0 0 0 Enthalpy kcal/mol -9.335 -0.072 3.535 -13.546 -19.725 Entropy cal/mol-K 1.083 0.799 1.321 6.107 7.132

47

TABLE. 12 Stream Results for the Steam Turbine in Methanol Production. Substream: MIXED Mole Flow kmol/hr 18 19 20 21 22 23 24 26

CH4 0 0 0 0 0 0 0 0 CO 0 0 0 0 0 0 0 0 CO2 0 0 0 0 0 0 0 0

C2H2 0 0 0 0 0 0 0 0 C2H4 0 0 0 0 0 0 0 0 C2H6 0 0 0 0 0 0 0 0 C3H8 0 0 0 0 0 0 0 0

H2 0 0 0 0 0 0 0 0 N2 0 0 0 0 0 0 0 0 O2 0 0 0 0 0 0 0 0 C 0 0 0 0 0 0 0 0

H2O 5738 5738 3443 2295 2295 1148 1148 1148 NH3 0 0 0 0 0 0 0 0 SO2 0 0 0 0 0 0 0 0 H2S 0 0 0 0 0 0 0 0

CH30H 0 0 0 0 0 0 0 0 Total Flow kmol/hr 5738 5738 3443 2295 2295 1148 1148 1148

Total Flow kg/hr 103376 103376 62025 41350 41350 20675 20675 20675 Total Flow cum/hr 3953 8223 4934 3289 4257 2129 2129 11636

Temperature C 501 363 363 363 319 319 319 102 Pressure bar 86 34.5 34.5 34.5 25 25 25 3 Vapor Frac 1 1 1 1 1 1 1 1 Liquid Frac 0 0 0 0 0 0 0 0 Solid Frac 0 0 0 0 0 0 0 0

Enthalpy kcal/mol -54.166 -55.204 -55.204 -55.204 -55.531 -55.531 -55.531 -57.183 Entropy cal/mol-K -11.809 -11.605 -11.605 -11.605 -11.537 -11.537 -11.537 -10.974

TABLE. 13 Stream Results for the work produced by the turbines in Methanol Production. GT-W1 GT-WOUT ST-W1 ST-W2 ST-WOUT TOTAL WORKPOWER MW 20.13 -20.74 -6.86 -7.72 -9.91 -30.65

48

TABLE 14. Stream Results for the Heat Recovery Steam Gas System for Methanol Production. Substream: MIXED Mole Flow kmol/hr 15 18 LPCSYN COOLGAS CH4 0 0 3.23E-09 3.23E-09 CO 0 0 15.073 15.073 CO2 0 0 637.079 637.079 C2H2 0 0 1.70E-16 1.70E-16 C2H4 0 0 1.47E-18 1.47E-18 C2H6 0 0 2.56E-21 2.56E-21 C3H8 0 0 0 0 H2 0 0 12.518 12.518 N2 0 0 3775.8 3775.8 O2 0 0 3.83E-06 3.83E-06 C 0 0 1.85E-21 1.85E-21 H2O 5738.2 5738.2 1306.3 1306.3 NH3 0 0 0.001 0.001 SO2 0 0 8.89E-08 8.89E-08 H2S 0 0 1.20E-09 1.20E-09 CH30H 0 0 2.86E-11 2.86E-11 Total Flow kmol/hr 5738.2 5738.2 5746.8 5746.8 Total Flow kg/hr 103375.8 103375.8 157791.6 157791.6 Total Flow cum/hr 2619.1 3952.6 378572.1 266641.5 Temperature C 328.291 501 519.029 284.972 Pressure bar 86 86 1 1 Vapor Frac 1 1 1 1 Liquid Frac 0 0 0 0 Solid Frac 0 0 0 0 Enthalpy kcal/mol -56.077 -54.166 -19.725 -21.633 Enthalpy kcal/kg -3112.724 -3006.656 -718.369 -787.859 Enthalpy MMkcal/hr -321.785 -310.820 -113.354 -124.319 Entropy cal/mol-K -14.612 -11.809 7.132 4.282 Entropy cal/gm-K -0.811 -0.656 0.260 0.156

49

ANNEX E: Aspen Flosheet for Hydrogen Production.

FUEL

OXIGEN

STEAM-G 1 2

CLE-H2O

W-WASTE

3

4

5

6

GASIFIER

7

ST EAM-SR

8

B8

B1

B6

9

B9

CO2

10 B10

H2

11 B14

FUELGAS

R-GASES

19

B2

C-WAT ER1

STEAM-1

B5

S-8BAR

S-8BAR2

B11

S-1BAR

B12

C-WAT ER2

STEAM-2

B13

C-WAT ER3

ST EAM-3

B15

CSYN

AIR-COMP

AIR

C-CAIR

GT

LPCSYN

B17

B19

S-86BAR

34

35B21

HST

36

MST LST

39 42

ST-W1 ST -W2 ST -WOUTW

GT-W1 GT-WOUTW

B29

48

B32 B33

51

52

53

54

50

ANNEX F: Stream Results for Hydrogen Production. TABLE 15. Stream Results for the Syngas System in Hydrogen Production. Substream: MIXED Mole Flow kmol/hr FUEL STEAM-G OXIGEN 1 2 3 4 5 6 7 STEAM-SR CH4 611.179 0 0 175.088 175.088 175.083 175.083 175.083 43.771 43.771 0 CO 542.953 0 0 1793.638 1793.638 1793.633 1793.633 1793.633 1924.945 1924.945 0 CO2 632.150 0 0 966.627 966.627 966.318 966.318 966.318 966.318 966.318 0 C2H2 0.867 0 0 2.44E-05 2.44E-05 2.44E-05 2.44E-05 2.44E-05 2.44E-05 2.44E-05 0 C2H4 158.762 0 0 0.0046 0.0046 0.0046 0.0046 0.0046 0.0046 0.0046 0 C2H6 104.579 0 0 0.0095 0.0095 0.0095 0.0095 0.0095 0.0095 0.0095 0 C3H8 1.344 0 0 1.78E-06 1.78E-06 1.78E-06 1.78E-06 1.78E-06 1.78E-06 1.78E-06 0 H2 240.744 0 0 2154.057 2154.057 2154.046 2154.046 2154.046 2547.982 2547.982 0 N2 0.806 0 0 13.749 13.749 13.749 13.749 13.749 13.749 13.749 0 O2 31.788 0 744.028 5.21E-15 5.21E-15 0 0 0 0 0 0 C 616.655 0 0 1.75E-22 1.75E-22 0 0 0 0 0 0 H2O 521.249 1321.545 0 1476.899 1476.899 337.932 337.932 337.932 206.620 206.620 1030.84437 NH3 26.260 0 0 0.373 0.373 0.031 0.031 0.031 0.031 0.031 0 SO2 1.057 0 0 2.19E-07 2.19E-07 2.13E-07 2.13E-07 2.13E-07 2.13E-07 2.13E-07 0 H2S 0.248 0 0 1.305 1.305 1.293 1.293 1.293 1.293 1.293 0 METHA-01 0 0 0 0.001 0.001 4.72E-05 4.72E-05 4.72E-05 4.72E-05 4.72E-05 0 Total Flow kmol/hr 3490.640 1321.54 744.028 6581.752 6581.752 5442.100 5442.100 5442.100 5704.724 5704.724 1030.844 Total Flow kg/hr 79360 23808 23808 126976 126976 106437.095 106437.095 106437.095 106437.095 106437.095 18570.9499 Total Flow cum/hr 10470.479 1456.704 905.986 21522.962 12329.785 139756.675 10912.763 13032.586 12931.117 6467.950 3868.820 Temperatura C 600 250 25 900 400 40 563.522867 727 672.85 200 200 Pressure bar 20 34.5 20 30 30 1.013 35.013 35.013 35 35 10 Vapor Frac 0.823 1 1 1 1 1 1 1 1 1 1 Liquid Frac 0.177 0 0 0 0 0 0 0 0 0 0 Solid Frac 0 0 0 0 0 0 0 0 0 0 0 Enthalpy kcal/mol -22.274 -56.299 -0.044 -26.970 -31.468 -29.441 -25.209 -23.765 -21.896 -25.758 -56.451 Entropy cal/mol-K 2.014 -13.502 -6.036 10.541 5.569 8.895 9.689 11.265 11.181 5.549 -11.548

51

TABLE 15. Continuation of the Stream Results for the Syngas System in Hydrogen Production. Substream: MIXED Mole Flow kmol/hr 8 9 CO2 10 19 H2 11 R-GASES FUEL-GAS 48 CH4 43.771 43.771 0 43.771 35.017 0 78.787 63.030 15.757 15.757 CO 1482.208 1482.208 0 1482.208 1185.766 0 2667.974 2134.379 533.595 533.595 CO2 1409.056 1409.056 1409.05562 0 0 0 0 0 0 0 C2H2 2.44E-05 2.44E-05 0 2.44E-05 1.95E-05 0 4.39E-05 3.51E-05 8.79E-06 8.79E-06 C2H4 0.005 0.005 0 0.005 0.004 0 0.008 0.007 0.002 0.002 C2H6 0.010 0.010 0 0.010 0.008 0 0.017 0.014 0.003 0.003 C3H8 1.78E-06 1.78E-06 0 1.78E-06 1.42E-06 0 3.20E-06 2.56E-06 6.41E-07 6.41E-07 H2 2990.719 2990.719 0 2990.719 382.812 2833.766 539.765 431.812 107.953 107.953 N2 13.749 13.749 0 13.749 10.999 0 24.749 19.799 4.950 4.950 O2 0 0 0 0 0 0 0 0 0 0 C 0 0 0 0 0 0 0 0 0 0 H2O 794.727385 794.727385 794.727385 0 0 0 0 0 0 0 NH3 0.031 0.031 0 0.031 0.025 0 0.055 0.044 0.011 0.011 SO2 2.13E-07 2.13E-07 0 2.13E-07 1.70E-07 0 3.83E-07 3.06E-07 7.65E-08 7.65E-08 H2S 1.293 1.293 0 1.293 1.035 0 2.328 1.862 0.466 0.466 METHA-01 4.72E-05 4.72E-05 0 4.72E-05 3.78E-05 0 8.50E-05 6.80E-05 1.70E-05 1.70E-05 Total Flow kmol/hr 6735.568 6735.568 2203.783 4531.785 1615.665 2833.766 3313.684 2650.947 662.737 662.737 Total Flow kg/hr 125008.045 125008.045 76329.493 48678.553 34891.400 5712.533 77857.420 62285.936 15571.484 15571.484Total Flow cum/hr 26507.635 15547.566 3529.803 11845.354 6358.372 11180.784 13027.880 10422.304 2605.576 1226.086 Temperatura C 200 40 40 40 40 39.884 39.884 39.884 39.884 168.907 Pressure bar 10 10 10 10 6.617 6.617 6.617 6.617 6.617 20 Vapor Frac 1 0.888 0.644 1 1 1 1 1 1 1 Liquid Frac 0 0.112 0.356 0 0 0 0 0 0 0 Solid Frac 0 0 0 0 0 0 0 0 0 0 Enthalpy kcal/mol -31.062 -33.513 -84.698 -8.706 -19.667 0.103 -21.587 -21.587 -21.587 -20.677 Entropy cal/mol-K 5.073 -1.516 -16.273 3.970 13.203 -3.394 14.499 14.499 14.499 14.729

52

TABLE 16. Stream Results for the Steam System in Hydrogen Production. Substream: MIXED Mole Flow kmol/hr C-WATER1 STEAM-1 CLE-H2O W-WASTE S-1BAR S-8BAR S-8BAR2 C-WATER2 STEAM-2 C-WATER3 STEAM-3 CH4 0 0 0 0.005 0 0 0 0 0 0 0 CO 0 0 0 0.005 0 0 0 0 0 0 0 CO2 0 0 0 0.309 0 0 0 0 0 0 0 C2H2 0 0 0 1.18E-08 0 0 0 0 0 0 0 C2H4 0 0 0 2.19E-07 0 0 0 0 0 0 0 C2H6 0 0 0 3.13E-07 0 0 0 0 0 0 0 C3H8 0 0 0 8.07E-12 0 0 0 0 0 0 0 H2 0 0 0 0.012 0 0 0 0 0 0 0 N2 0 0 0 4.94E-05 0 0 0 0 0 0 0 O2 0 0 0 0 0 0 0 0 0 0 0 C 0 0 0 0 0 0 0 0 0 0 0 H2O 1543.1 1543.1 109802.2 110941.1 1301.7 1301.7 1301.7 1301.7 1301.7 1500 1500 NH3 0 0 0 0.342 0 0 0 0 0 0 0 SO2 0 0 0 6.65E-09 0 0 0 0 0 0 0 H2S 0 0 0 0.012 0 0 0 0 0 0 0 METHA-01 0 0 0 0.001 0 0 0 0 0 0 0 Total Flow kmol/hr 1543.1 1543.1 109802.2 110941.8 1301.7 1301.7 1301.7 1301.7 1301.7 1500 1500 Total Flow kg/hr 27800 27800 1978116.8 1998655.7 23450 23450 23450 23450 23450 27022.92 27022.92 Total Flow cum/hr 27.583 150059.7 1962.7 2041.7 102139.3 19788.3 11875.5 23.267 102221.4 26.812 42399.7 Temperatura C 10 896.571112 10 40 671 1189.24379 608.576774 10 671.757923 10 101.537082 Pressure bar 1 1 1.013 1.013 1 8 8 1 1 1 1 Vapor Frac 0 1 0 0 1 1 1 0 1 0 0.91506754 Liquid Frac 1 0 1 1 0 0 0 1 0 1 0.08493246 Solid Frac 0 0 0 0 0 0 0 0 0 0 0 Enthalpy MMkcal/hr -106.504 -76.900 -7578.338 -7592.174 -67.813 -60.772 -68.630 -89.839 -67.803 -103.527 -87.018 Entropy cal/mol-K -41.108 1.500 -41.108 -39.147 -0.647 -0.232 -5.457 -41.108 -0.639 -41.108 -11.035

53

TABLE 17. Stream results for the Gas Turbine in Hydrogen Production. Substream: MIXED Mole Flow kmol/hr 48 AIR C-CAIR CSYN LPCSYN CH4 15.757 0 0 1.74E-25 1.74E-25 CO 533.595 0 0 0.00222373 0.00222373 CO2 0 0 0 549.360 549.360 C2H2 8.79E-06 0 0 0 0 C2H4 0.002 0 0 0 0 C2H6 0.003 0 0 0 0 C3H8 6.41E-07 0 0 0 0 H2 107.953 0 0 0 0 N2 4.950 1540.422 1540.422 1545.377 1545.377 O2 0 405.097 405.097 52.085 52.085 C 0 0 0 4.62E-29 4.62E-29 H2O 0 0 0 139.964 139.964 NH3 0.011 0 0 1.14E-10 1.14E-10 SO2 7.65E-08 0 0 0.466 0.466 H2S 0.46560983 0 0 8.45E-15 8.45E-15 METHA-01 1.70E-05 0 0 8.97E-24 8.97E-24 Total Flow kmol/hr 662.737 1945.518 1945.518 2287.254 2287.254 Total Flow kg/hr 15571.484 56115.181 56115.181 71686.665 71686.665 Total Flow cum/hr 1226.086 46579.938 4328.130 10122.573 153891.957 Temperature C 168.907 15 522.241 1200 535.872 Pressure bar 20 1 30 27.828 1 Vapor Frac 1 1 1 1 1 Liquid Frac 0 0 0 0 0 Solid Frac 0 0 0 0 0 Enthalpy kcal/mol -20.677 -0.072 3.604 -15.708 -21.958 Entropy cal/mol-K 14.729 0.799 1.325 8.288 9.299

54

TABLE 18. Stream Results for the Steam Turbine in Hydrogen Production. Substream: MIXED Mole Flow kmol/hr 35 36 51 52 39 53 54 42 CH4 0 0 0 0 0 0 0 0 CO 0 0 0 0 0 0 0 0 CO2 0 0 0 0 0 0 0 0 C2H2 0 0 0 0 0 0 0 0 C2H4 0 0 0 0 0 0 0 0 C2H6 0 0 0 0 0 0 0 0 C3H8 0 0 0 0 0 0 0 0 H2 0 0 0 0 0 0 0 0 N2 0 0 0 0 0 0 0 0 O2 0 0 0 0 0 0 0 0 C 0 0 0 0 0 0 0 0 H2O 4344.807 4344.807 2172.404 2172.404 2172.404 1086.202 1086.202 1086.202 NH3 0 0 0 0 0 0 0 0 SO2 0 0 0 0 0 0 0 0 H2S 0 0 0 0 0 0 0 0 METHA-01 0 0 0 0 0 0 0 0 Total Flow kmol/hr 4344.807 4344.807 2172.404 2172.404 2172.404 1086.202 1086.202 1086.202 Total Flow kg/hr 78272.920 78272.920 39136.460 39136.460 39136.460 19568.230 19568.230 19568.230Total Flow cum/hr 3243.357 15867.272 7933.631 7933.631 17328.603 8664.302 8664.302 16448.704Temperature C 550 273.493 273.493 273.493 169.452 169.452 169.452 97.645 Pressure bar 86 12 12 12 4.5 4.5 4.5 2 Vapor Frac 1 1 1 1 1 1 1 1 Liquid Frac 0 0 0 0 0 0 0 0 Solid Frac 0 0 0 0 0 0 0 0 Enthalpy kcal/mol -53.655 -55.825 -55.825 -55.825 -56.642 -56.642 -56.642 -57.205 Entropy cal/mol-K -11.169 -10.665 -10.665 -10.665 -10.433 -10.433 -10.433 -10.244

TABLE 19. Stream Results for the work produced by the turbines in Hydrogen Production. GT-W1 GT-WOUT ST-W1 ST-W2 ST-WOUT TOTAL WORK POWER MW 8.40 -8.06 -9.87 -11.73 -12.37 -20.42

55

TABLE 20. Stream Results for the Heat Recovery Steam Gas System for Hydrogen Production. Substream: MIXED Mole Flow kmol/hr S-86BAR 35 LPCSYN 34 CH4 0 0 1.74E-25 1.74E-25 CO 0 0 0.002 0.002 CO2 0 0 549.360 549.360 C2H2 0 0 0 0 C2H4 0 0 0 0 C2H6 0 0 0 0 C3H8 0 0 0 0 H2 0 0 0 0 N2 0 0 1545.377 1545.377 O2 0 0 52.085 52.085 C 0 0 4.62E-29 4.62E-29 H2O 4344.807 4344.807 139.964 139.964 NH3 0 0 1.14E-10 1.14E-10 SO2 0 0 0.466 0.466 H2S 0 0 8.45E-15 8.45E-15 METHA-01 0 0 8.97E-24 8.97E-24 Total Flow kmol/hr 4344.807 4344.807 2287.254 2287.254 Total Flow kg/hr 78272.920 78272.920 71686.665 71686.665 Total Flow cum/hr 3307.241 3243.357 153891.957 159322.528 Temperature C 562.754 550.000 535.872 564.419 Pressure bar 86 86 1 1 Vapor Frac 1 1 1 1 Liquid Frac 0 0 0 0 Solid Frac 0 0 0 0 Enthalpy kcal/mol -53.522 -53.655 -21.958 -21.706 Entropy cal/mol-K -11.010 -11.169 9.299 9.604

56

ANNEX G : Ecomonics Table N-2 Cost of system components in MSEKJan 2006 1) MSEK Scale Base scale Overal Maximun

UNIT KPI

JAN2006 Factor instalation Size 18) 279,59 Factor 17) Pre-treatment 2) Conveyers 3) 3,521 0,8 33,5 wet tonne/hr 1,86 110 Grinding 3) 4,124 0,6 33,5 wet tonne/hr 1,86 110 Storage 3) 10,059 0,65 33,5 wet tonne/hr 1,86 110 Dryer 3) 76,452 0,8 33,5 wet tonne/hr 1,86 110 iron removal 3) 3,722 0,7 33,5 wet tonne/hr 1,86 110 feeding system 3,4) 4,124 1 33,5 wet tonne/hr 1,86 110 384,272 Gasification System IGT 5) 383,266 0,7 68,8 dry tonne/hr 1,69 75 Oxigen plant 6) 444,628 0,85 41,7 tone O2/hr 1 - Gas Cleaning Tar Cracker 3) 31,184 0,7 34,2 m3 gas/s 1,86 52 Cyclones 3) 26,155 0,7 34,2 m3 gas/s 1,86 180 HTXH 7) 70,316 0,6 39,2 kg steam/s 1,86 - Bag house filter 3) 16,095 0,65 12,1 m3 gas/s 1,86 64 Condensing Scrubber 3) 26,155 0,7 12,1 m3 gas/s 1,86 64 Syngas Procesing Compresor 8) 111,660 0,85 13,2 Mwe 1,72 - Steam Reformer 9) 94,559 0,6 1390 Kmol total/hr 2,3 - Shift reactor (installed)

10) 371,194 0,85 15,6 Mmol CO+H2/hr 1 - Selexol CO2 removal (ins) 11) 544,218 0,7 9909 kmol CO2/hr 1 - Methanol Production Make up Compressor 8) 111,660 0,85 13,2 Mwe 1,72 Gas Phase Methanol 12) 70,416 0,6 87,5 tonne MeOH/hr 2,1 - Refining 13) 151,898 0,7 87,5 tonne MeOH/hr 2,1 - Hydrogen Production PSA units A+B 14) 281,665 0,7 9600 kmol feed/hour 1,69 - Power Isle Gas Turbine + HRSG 3,15) 190,124 0,7 26,3 Mwe 1,86 - Steam Turbine + steam system 3,16) 51,303 0,7 10,3 Mwe 1,86 -

1) The costs of the components were found in MUS$2001. For this study the purchasing costs of the components were calculated for January 2006 by using the KPI (“Consumers Price Index” for Sweden). 2) Total pre-treatment approximately sums up to a base cost of 8.15 MUS$2001 at a base scale of 33.5 tonne wet/hour with an R factor of 0.79. 3) Based on first generation BIG/CC installations. Faaij et al. (1995) evaluated a 29 MWe BIG/CC installation (input 9.30 kg dry wood/s, produces 10.55 Nm3 fuel gas/s) using vendor quotes. When a range is given, the higher values are used (Faaij et al. 1998). The scale factors stem from Faaij et al. (1998). 4) Two double screw feeders with rotary valves (Faaij et al. 1995).

57

5) 29.74 MUS$1991 (includes already added investment to hardware) for a 1650 dry tonne/day input IGT gasifier, R = 0.7 (Williams et al. 1995). Maximum input is 400 MWth HHV (Tijmensen 2000). 6) Air Separation Unit: Plant investment costs are given by Van Dijk (van Dijk et al. 1995): I = 0.1069·C0.8508 in MUS$1995 installed, C = Capacity in tonne O2/day. The relation is valid for 100 to 2000 tonne O2/day. Williams et al. (1995) assume higher costs for small installations, but with a stronger effect of scale: I = 0.260·C0.712 in MUS$1991 fob plus an overall installation factor of 1.75 (25 and 40%). Larson et al. (1998) assume lower costs than Van Dijk, but with an even stronger scaling factor than Williams: 27 MUS$1997 installed for an 1100 tonne O2 per day plant and R=0.6. The production of 99.5% pure O2 using an air separation unit requires 250 – 350 kWh per tonne O2 (van Dijk et al. 1995; van Ree 1992). 7) High temperature heat exchangers following the gasifier and at other locations are modelled as HRSG’s. A 39.2 kg steam/s unit costs 6.33 MUS$1997 fob, overall installation factor is 1.84 (Larson et al. 1998). 8) Katofsky (1993) assumes compressors to cost 700 US$1993 per required kWmech, with an installation factor of 2.1. The relation used in the present study stems from the compressor manufacturer Sulzer quoted by (2000). At the indicated base-scale, total installed costs are about 15 % higher than assumed by Katofsky. Multiple compressors, for fuel gas, recycle streams, or hydrogen, are considered as separate units. Overall installation factor is taken 1.72 because the base unit matches a 400 MWth plant rather than a 70 MWth plant. 9) Investments for steam reformer vary from 16.9 MUS$1993, for a throughput of 5800 kmol methane/hour with an overall installation factor of 2.1 (Katofsky 1993) to 7867 k$1995 for a 6.2 kg methane/s (1390 kmol/hour), overall installation factor is 2.3 (van Dijk et al. 1995). These values suggest a strong effect of scaling R=0.51, while Katofsky uses a modest R=0.7. In the present study the values of Van Dijk are used in combination with an R factor of 0.6. The total amount of moles determines the volume and thus the price of the reactor. 10) Investment for shift reactors vary from 9.02 MU$1995 for an 8819 kmol CO+H2/hr reactor, and an overall installation factor is 1.81 (Williams et al. 1995) to 30 MUS$1994 installed for a 350000 Nm3/hr CO+H2/hr (15625 kmol/hr) reactor (Hendriks 1994). Williams assumes an R=0.65, but comparison of the values suggest only a weak influence of scale (R=0.94), in the present study the values from Hendriks are used, R is set 0.85. A dual shift is costed as a shift of twice the capacity. 11) Costs for CO2 removal through Selexol amounts 14.3 MUS$1993 fob (overall installation factor is 1.87) for an 810 kmol CO2/hr unit, R = 0.7 (Katofsky 1993) up to 44 MUS$1994 installed for a 9909 kmol CO2/hour unit (Hendriks 1994). The value from Hendriks is assumed to be right, since his research into CO2 removal is comprehensive. 12) Van Dijk et al. (1995) estimate that a Methanol Reactor for a 2.1 ktonne methanol per day plant costs 4433 kUS$1995 (fob) or 9526 kUS$1995 installed (overall installation factor is 2.1). The total plant investment in their study is 138 MUS$1995, or 150 MUS$2001. Katofsky (1993) estimates the costs for a 1056 tonne methanol/day plant to be 50 MUS$1995 fob, this excludes the generation and altering of syngas, but includes make-up and recycle compression and refining tower. Correspondence with mr. Van Ooijen (2001) of Akzo Nobel and mr. De Lathouder (2001) of DSM Stamicarbon revealed that a 1000 tpd plant costs about 160 MUS$2001, and a 2000 tpd plant 200 MUS$2001 (this suggests a total plant scale factor of 0.3). These values come near the ones mentioned by Katofsky. This implies that the values given by Van Dijk are too optimistic and should be altered by a factor 1.33. It is therefore assumed that the base investment for the methanol reactor only is 7 MUS$2001, the installation factor is 2.1. The influence of scale on reactor price is assumed to be not as strong as for the complete plant: 0.6. 13) Cost number for methanol separation and refining is taken from Van Dijk, increased with 33 % as described in note 12. 14) PSA units (excluding the recycle compressor) cost 23 MUS$1993 for a 9600 kmol feed/hour throughput, R= 0.7 (Katofsky 1993). 15) Scaled on Gas Turbine size. 16) Steam system consists of water and steam system, steam turbine, condenser and cooling. Scaled on Steam Turbine size. 17) Overall installation factor. Includes auxiliary equipment and installation labour, engineering and contingencies. Unless other values are given by literature, the overall installation factor is set 1.86 for a 70 MWth scale (Faaij et al. 1998). This value is based on 33% added investment to hardware costs (instrumentation and control 5%, buildings 1.5% grid connections 5%, site preparation 0.5%, civil works 10%, electronics 7%, and piping 4%) and 40 % added installation costs to investment (engineering 5%, building interest 10%, project contingency 10%, fees/overheads/profits 10%, start-up costs 5%). For larger scales, the added investments to hardware decreases slightly. 18) Maximum sizes from Tijmensen (2000).