NREL is a national laboratory of the U.S. Department of Energy, Office of Energy Efficiency & Renewable Energy, operated by the Alliance for Sustainable Energy, LLC.



### Contract No. DE

- AC36

- 08 GO28308 

# Techno

- Economic Analysis of
Biofuels Production Based on
Gasification

Ryan M. Swanson, Justinus A. Satrio, and
Robert C. Brown Iowa State University



# A lexandru Platon ConocoPhillips Company



# David D. Hsu National Renewable Energy Laboratory

Technical Report
NREL/TP

- 6A20

- 46587

November 2010NREL is a national laboratory of the U.S. Department of Energy, Office of Energy Efficiency & Renewable Energy, operated by the Alliance for Sustainable Energy, LLC.

National Renewable Energy Laboratory
1617 Cole Boulevard Golden, Colorado 80401
303

- 275

- 3000 • www.nrel.gov



### Contract No. DE

- AC36

- 08GO28308 

# Techno

- Economic Analysis of
Biofuels Production Based on
Gasification

Ryan M. Swanson, Justinus A. Satrio, and
Robert C. Brown Iowa State University



# A lexandru Platon ConocoPhillips Company



# David D. Hsu National Renewable Energy Laboratory

Prepared under Task No. BB07.7510

Technical Report
NREL/TP -6A20- 46587

November 2010NOTICE

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### iii 

# Foreword

The purpose of this techno-economic analysis is to compare a set of biofuel conversion technologies selected for their promise and near-ter m technical viability. Every effort is made to make this comparison on an equivalent basi s using common assumptions. The process design and parameter value choices underlying this analys is are based on public do main literature only. For these reasons, these results are not indicative of potential pe rformance, but are meant to represent the most likely pe rformance given the current state of public knowledge.

### iv 

# List of Acronyms

AGR acid gas removal ASU air separation unit BTL biomass to liquids CFB circulating fluidized bed

DCFROR discounted cash flow rate of return DME dimethyl-ether FCI fixed capital investment FT Fischer-Tropsch GGE gallon of gasoline equivalent HRSG heat recovery steam generator HT high temperature

IC indirect costs IGCC integrated gasification combined cycle

IRR internal rate of return ISU Iowa State University LHV lower heating value LT low temperature

MEA monoethanolamine MJ megajoule

MM million MTG methanol to gasoline MW megawatt

Nm 3 

### normal cubic meter

NREL National Rene wable Energy Laboratory PSA pressure swing adsorption PV product value Sasol South African Coal, Oil, and Gas Corporation SPR slurry phase reactor SMR steam methane reforming SWGS sour water-gas-shift TCI total capital investment

TDIC total direct and indirect cost

TIC total installed cost tpd tons per day TPEC total purchased equipment cost WGS water-gas-shift

### v 

# Executive Summary

This study compares capital and production costs of two biomass-to-liquid production plants based on gasification. The goal is to produce li quid transportation fuels via Fischer-Tropsch synthesis with electricity as a co-product. The biorefineries ar e fed by 2,000 metric tons per day

of corn stover. The first bior efinery scenario is an oxygen- fed, low-temperature (870°C), nonslagging, fluidized bed gasifier . The second scenario is an oxygen-fed, high-temperature (1,300°C), slagging, entrained flow gasifier. Both are followed by catalytic Fischer-Tropsch synthesis and hydroprocessing to naphtha-range (gasoline blend stock) and distillate-range (diesel blend stock) liquid fracti ons. (Hydroprocessing is a set of re finery processes that removes impurities and breaks down large molecules to fractions suitable for use in commercial formulations.)

Process modeling software (Aspen Plus) is utilized to organize the mass and energy streams and cost estimation software is used to generate e quipment costs. Economic analysis is performed to estimate the capital investment and operating co sts. A 20-year discounted cash flow rate of return analysis is developed to estimate a fuel product value (PV) at a ne t present value of zero
with 10% internal rate of return. All costs ar e adjusted to the year 2007. The technology is limited to commercial technology available for imple mentation in the next 5–8 years, and as a result, the process design is restricted to available rather than projected data.

Results show that the total capital investment required for n th 

### plant scenarios is $610 million and $500 million for high-temperature and low-temperature scenarios, respectively. PV for the hightemperature and low-temperature scenarios is estimated to be $4.30 and $4.80 per gallon of gasoline equivalent (GGE), respectively, based on a feedstock cost of $75 per dry short ton. The

main reason for a difference in PV between the scen arios is because of a higher carbon efficiency and subsequent higher fuel yield for the high-temp erature scenario. Sensitivity analysis is also performed on process and economic parameters. This analysis shows that total capital investment and feedstock cost are among the most influent ial parameters affecting the PV, while least influential parameters include per-pass Fisc her-Tropsch-reaction-conve rsion extent, inlet feedstock moisture, and catalyst cost.

In order to estimate the cost of a pioneer plant (first of its kind), an analysis is performed that inflates total capital investment and deflates the plant output for the first several years of operation. Base case results of this analysis estim ate a pioneer plant investment to be $1.4 billion and $1.1 billion for high-temperature and low-temp erature scenarios, respectively. Resulting PVs are estimated to be $7.60/GGE and $8.10/GGE for high-temperature and low-temperature pioneer plants, respectively.

### vi 

# Table of Contents

Introduction .................................................................................................................. ............................... 1 Background .................................................................................................................... ............................. 2 

### Biorenewable Resources ............................................................................................................2

Gasification .................................................................................................................. ..............2
Reactions ..................................................................................................................... .........3 Gasifier Types ......................................................................................................................4 Biomass Preprocessing ......................................................................................................... .....7 Syngas Cleaning................................................................................................................ .........9
End-Use Product ......................................................................................................................10

Power Generation ...............................................................................................................10 Synthetic Fuels and Chemicals ..........................................................................................11

Methanol to Gasoline .........................................................................................................1 1 Fischer-Tropsch ............................................................................................................... ..12 Techno-Economic Analysis .....................................................................................................1 3

Me t hodology ...................................................................................................................... ........................ 16 

### Down-Selection Process ..........................................................................................................16 Preliminary Criteria .......................................................................................................... .17
Scenario Selection ..............................................................................................................17

Scenarios Not Selected ......................................................................................................18 Project Assumptions ..........................................................................................................1 8 Process Description ..................................................................................................................19 High-Temperature Scenario Overview ..............................................................................19 Low-Temperature Scenario Overview ...............................................................................20 Area 100 Preprocessing .....................................................................................................22

Area 200 Gasification ........................................................................................................23 Area 300 Syngas Cleaning .................................................................................................25 Area 400 Fuel Synthesis ....................................................................................................27 Area 500 Hydroprocessing ................................................................................................29

Area 600 Power Generation ...............................................................................................29 Area 700 Air Separation ....................................................................................................30 Methodology for Economic Analysis ......................................................................................30 Methodology for Major Equipment Costs .........................................................................34 Methodology for Sensitivity Analysis ...............................................................................35 Methodology for Pioneer Plant Analysis ...........................................................................36

Results and Discussion ............................................................................................................................ 39 

### Process Results................................................................................................................ .........39 Cost Estimating Results ...........................................................................................................41 Capital and Operating Costs for n th 

### Plant ..........................................................................41 Sensitivity Results for n th 

### Plant ..........................................................................................43 Pioneer Plant Analysis Results ..........................................................................................45 Comparison with Previous Techno-Economic Studies ............................................................46 Summary of n th 

### Plant Scenarios ...............................................................................................48

Conclusions ................................................................................................................... ............................ 49
References .................................................................................................................... ............................. 50

### vii Appendix A. Techno -Economic Model Assumpti ons ........................................................................... 55 Appendix B. De tailed Costs .................................................................................................... ................. 59 

### Cost Summaries .......................................................................................................................59 Detailed Equipment Lists ...................................................................................................... ...61
Discounted Cash Flow .......................................................................................................... ...67 Appendix C. Scenario Modeling Details ......................................................................................... ........ 71 

### Property Method ......................................................................................................................71

Stream/Block Nomenclature ....................................................................................................7 1 Aspen Plus Calculator Block Descriptions ..............................................................................73 Aspen Plus Model Design Specifications ................................................................................84

Detailed Calculations ...............................................................................................................86 Appendix D. Process Flow Diagrams ............................................................................................. ...... 116 

### High-Temperature Scenario ...................................................................................................117 Low-Temperature Scenario ...................................................................................................127 Appendix E. St ream Data ....................................................................................................... ................ 138 

### High-Temperature Scenario ...................................................................................................139 Low-Temperature Scenario ...................................................................................................146

### viii 

# List of Figures

Figure 1. Overall process flow diagram for both scenarios ............................................................ 1 Figure 2. Typical energy content of the products of gasification of wood using air varied by equivalence ratio [12] ........................................................................................................ . 4 Figure 3. Design of fixed-bed (a) updraft and (b ) downdraft gasifiers sh owing reaction zones[13]

............................................................................................................................... .............. 5 Figure 4. Fluidized bed gasifier designs of (a) and (b) directly heated type and (c) and (d) indirectly heated type [16] .................................................................................................. 6 Figure 5. Entrained-flow gasifier [18] ........................................................................................ .... 7 Figure 6. Schematic of a biomass pretreatment vi a fast pyrolysis followed by an entrained-flow gasifier [17] ................................................................................................................. ........ 8 Figure 7. Main syngas conversion pathways [32] ........................................................................ 11 Figure 8. Fischer-Tropsch reactor types (a) multi-t ubular fixed bed and (b) slurry bed [31] ....... 13 Figure 9. Overall process flow diagram for HT scenario ............................................................. 20 Figure 10. Overall process flow diagram for LT scenario ............................................................ 22 Figure 11. Fischer-Tropsch product distribution as a function of chain growth factor ( ࢻ (using equation 11 [48] .............................................................................................................. .. 29 Figure 12. Sensitivity results for HT n th 

### plant scenario ................................................................ 44 Figure 13. Sensitivity results for LT n th 

### plant scenario ................................................................. 44 Figure 14. The effect of plant size on product va lue (per gallon of gaso line equivalent) for n th


### plant scenarios ............................................................................................................... .... 45 Figure 15. The effect of plant size on total capital investment for n th 

### plant scenarios ................. 45 Figure B-1. Economic analysis su mmary for HT scenario ........................................................... 59 Figure B-2. Economic analysis summary for LT scenario ........................................................... 60 Figure C-1. Stream nomencla ture used in model .......................................................................... 71 Figure C-2. Block nomenclat ure used in model ........................................................................... 71 Figure C-3. Heat and work stream nomenclature used in model .................................................. 72 Figure C-4. Decision diagra m for carbon balance ........................................................................ 80 Figure C-5. Decision diagra m for hydrogen balance .................................................................... 81 Figure C-6. Decision diagra m for oxygen balance ....................................................................... 82 Figure D-1. Overall plant area process flow diagram for HT scenario ...................................... 117 Figure D-2. Preprocessing area process flow diagram for HT scenario ..................................... 118 Figure D-3. Gasification ar ea process flow diagram for HT scenario ........................................ 119 Figure D-4. Syngas cleaning area process flow diagram for HT scenario ................................. 120 Figure D-5. Acid gas removal area proce ss flow diagram for HT scenario ............................... 121 Figure D-6. Sulfur recovery area proce ss flow diagram for HT scenario .................................. 122 Figure D-7. Fuel synthesis area process flow diagra m for HT scenario ..................................... 123 Figure D-8. Hydroprocessing area process flow diagram for HT scenario ................................ 124 Figure D-9. Power generation area process flow diagram for HT scenario ............................... 125 Figure D-10. Air separation unit process flow diagram for HT scenario ................................... 126 Figure D-11. Overall plant area proce ss flow diagram for LT scenario ..................................... 127 Figure D-12. Preprocessing area process flow diagram for LT scenario ................................... 128 Figure D-13. Gasification area process flow diagram for LT scenario ...................................... 129 Figure D-14. Syngas cleaning area proce ss flow diagram for LT scenario ................................ 130 Figure D-15. Acid gas removal area pro cess flow diagram for LT scenario .............................. 131

### ix Figure D-16. Sulfur recovery proces s flow diagram for LT scenario ......................................... 132 Figure D-17. Fuel synthesis area process flow diagram for LT scenario ................................... 133 Figure D-18. Syngas conditioning area process flow diagram for LT scenario ......................... 134 Figure D-19. Hydroprocessing area pro cess diagram for LT scenario ....................................... 135 Figure D-20. Power generation area proce ss flow diagram for LT scenario .............................. 136 Figure D-21. Air separation unit proce ss flow diagram for LT scenario .................................... 137



# List of Tables

Table 1. Reactions Occurring within th e Reduction Stage of Gasification .................................... 3

Table 2. Previous Techno-Economic Studies of Biomass-Gasification Biofuel Production

Plants ........................................................................................................................ ......... 15 Table 3. Process Configurations Consid ered in Down Selection Process .................................... 16 Table 4. Main Assumptions Used in n th 

### Plant Scenarios .............................................................. 18

Table 5. Stover and Char Elem ental Composition (wt %) ............................................................ 23 Table 6. Syngas Composition (M ole Basis) Leaving Gasifier for Gasification Scenarios

Evaluated...................................................................................................................... ..... 25 Table 7. Fischer-Tropsch Gas Cleanliness Requirements [31] ..................................................... 27 Table 8. Hydroprocessing Product Distribution [49] .................................................................... 29 Table 9. Main Economic Assumptions for n th 

### Plant Scenarios .................................................... 30 Table 10. Methodology for Capital Cost Estimation for n th 

### Plant Scenarios ............................... 32 Table 11. Variable Operating Cost Parameters Adjusted to $2007 .............................................. 33 Table 12. Sensitivity Parameters for n th 

### Plant Scenarios .............................................................. 35 Table 13. Pioneer Plant Analysis Parameters and Factors ............................................................ 38 Table 14. Power Generation and Usage ........................................................................................ 39 Table 15. Overall Energy Balance on LHV Basis ........................................................................ 40

Table 16. Overall Carbon Balance .............................................................................................. .. 41 Table 17. Capital Investment Breakdown for n th 

### Plant Scenarios ................................................ 42 Table 18. Annual Operating Cost Breakdown for n th 

### Plant Scenarios ......................................... 43 Table 19. Catalyst Replacement Costs for Both Scenarios (3-Year Replacement Period) .......... 43

Table 20. Pioneer Plant Analysis Results ..................................................................................... 46 Table 21. Comparison of n th 

### Plant LT Scenario to Tijmensen et al. Study IGT-R Scenario ........ 47 Table 22. Comparison of n th 

### Plant LT Scenario to Larson et al. Study FT-OT-VENT Scenario . 48

Table 23. Main Scenario n th 

### Plant Results .................................................................................... 48 Table B-1. Detailed Equipment List fo r Areas 100 and 200 of HT Scenario ............................... 61 Table B-2. Detailed Equipment List for Areas 300, 400, and 500 of HT Scenario ...................... 62 Table B-3. Detailed Equipment List fo r Areas 600 and 700 of HT Scenario ............................... 63 Table B-4. Detailed Equipment List fo r Areas 100 and 200 of LT Scenario ............................... 64 Table B-5. Detailed Equipment List fo r Areas 300, 400, and 500 of LT Scenario ...................... 65 Table B-6. Detailed Equipment List fo r Areas 600 and 700 of LT Scenario ............................... 66

Table B-7. Discounted Cash Fl ow Sheet for Construction Period and Years 1-8 of HT

Scenario....................................................................................................................... ...... 67
Table B-8. Discounted Cash Flow Sheet for Years 9-20 of HT Scenario .................................... 68

Table B-9. Discounted Cash Flow Sheet for Construction Period and Years 1-8 of LT Scenario69

Table B-10. Discounted Cash Flow Sheet for Years 9-20 of LT Scenario ................................... 70 Table C-1. Detailed Description of Stream and Block Nomenclature .......................................... 72

### x Table C-2. Reaction Extent Equations for Each Alkane Hydrocarbon ........................................ 75

Table C-3. Combustion Reactions to Determine Required Oxygen ............................................. 77 Table C-4. Hydroprocessi ng Product Blend ................................................................................. 84

Table E-1. Overall Plant Stream Data for HT Scenario .............................................................. 139 Table E-2. Preprocessing Area Stream Data for HT Scenario .................................................... 140

Table E-3. Gasification Area Stream Data for HT Scenario ...................................................... 141 Table E-4. Syngas Cleaning Area Str eam Data for HT Scenario ............................................... 142

Table E-5. Acid Gas Removal and Sulfur Recove ry Areas Stream Data for HT Scenario ........ 143 Table E-6. Fuel Synthesis Area Str eam Data for HT Scenario ................................................... 144 Table E-7. Hydroprocessing, Powe r Generation, and Air Separa tion Areas Stream Data for

HT Scenario ................................................................................................................... . 145
Table E-8. Overall Plant Stream Data for LT Scenario .............................................................. 146 Table E-9. Preprocessing Area Stream Data for LT Scenario .................................................... 147

Table E-10. Gasification Area Str eam Data for LT Scenario ..................................................... 148 Table E-11. Syngas Cleaning Area Str eam Data for LT Scenario ............................................. 149

Table E-12. Acid Gas Removal and Sulfur Recove ry Areas Stream Data for LT Scenario ...... 150 Table E-13. Fuel Synthesis Area St ream Data for LT Scenario ................................................. 151 Table E-14. Syngas Conditioning Area St ream Data for LT Scenario ....................................... 152 Table E-15. Hydroprocessing, Power Generation, and Air Separati on Areas Stream Data for

LT Scenario ................................................................................................................... .. 153

### 1 

# Introduction

This study investigates the economic feasibility of the thermochemical pathway of gasification of biomass to renewable transportation fuels. The ob jective is to compare capital investment costs and production costs for n th 

### plant biorefinery scenarios ba sed on gasification. The selected scenarios are high-temperature (slagging) gasification and low-temperature (dry-ash) gasification, each followed by Fischer-Trops ch synthesis and hydroprocessing. They are designed to produce liquid hydroc arbon fuels from 2,000 dry metric tons (2,205 dry short tons) per day of agricultural residue, namely corn stover . Corn stover is chosen as a feedstock in order to facilitate comparisons with biochemical and pyrolysis biofuels scenarios [1, 2].

The two scenarios were chosen from many op tions according to the following criteria:

1. The technology under consideration should be comm ercially ready in th e next 5–8 years.

2. The size of the biorefinery should be feasible with current agricultural productivity and

within a realistic feed stock collection area.

3. In addition, the desired end product should be compatible with the present fuel infrastructure, i.e., ga soline and/or diesel.

The high-temperature gasification scenario is based on a steam/oxygen-fed entrained flow, slagging gasifier similar to that described in Frey and Akunuri [3]. The low-temperature gasification scenario is based on a pressuri zed, steam/oxygen-fed fluidized bed gasifier developed by the Gas Technology Institute and repor ted by Bain [4]. The main areas of operation are feedstock preprocessing, gasification, s yngas cleaning, syngas condi tioning/upgrading, fuel synthesis, power generation, and air separati on (for oxygen production), as shown in Figure 1. Process modeling software is utilized to or ganize the mass and energy streams, and cost estimation software is used to generate equipm ent costs. Economic analysis is performed to estimate the capital investment and operating co sts. A 20-year discounted cash flow rate of return (DCFROR) analysis is de veloped to estimate a fuel produc t value (PV) at a net present
value of zero with 10% internal rate of return. All costs ar e adjusted to the year 2007.

Figure 1. Overall process flow diagram for both scenarios

### 2 

# Background

Biorenewable Resources The world population has long utilized materials that were in close proximity. The nearest

resource available to the huma n population is the organic matter in the environment around it. This organic matter is present for a limited am ount of time due to its decomposable nature. Brown [5] defines this material, or biorenewable resources, as organic material of recent biological origin. It is a rene wable resource if the rate of consumption is equal to the regeneration or growth and ther efore must be used only if pres erving biodiversity [6]. As a result, these resources have been important contributors to the world economy, providing foodstuffs, transportation, energy, and constructi on materials, as well as serving many other

functions.

Biorenewable resources for generating energy can be classified as woody biomass, energy crops, residues, and municipal waste [6 ]. The first two are primary resources while the remaining are secondary resources, meaning that their prim ary use has already occurred. Woody biomass includes logging products and energy crops includ e short-rotation trees su ch as poplar and fastgrowing grasses such as switchgr ass. Residues can come from logging processing or agricultural processing (e.g., corn stover). According to Perlac k et al. [7], the energy crop and agricultural residue potential in th e United States is 1.4 billion annua l tons. According to the U.S. Department of Energy’s “Roadmap for Agricult ure Biomass Feedstock Supply in the U.S.,” there is potential for 2 billion annual tons, including municipal waste and biosolids such as

manure.

Many end products can be produced from these resources. Aside from the conventional use of biomass for human food consumption, livestock feed, and building materials, there are many new pathways to provide renewable alternatives to our transportation, infrastructure, and energy. Combustion of biomass offers a way to provide h eat and power and displaces coal and fuel oil. Fermentation of carbohydrates and liquefaction of biomass through fast pyrolysis yield liquid products with the potential to displace petroc hemicals. Gasification of biomass allows for chemical and liquid fuel synthesis, which is the focus of this study.

Developing an economy that invol ves biorenewable resources, especially biofuels, has many benefits. According to Greene et al. [8], biof uel production has the potential to provide a new source of revenue for farmers by generating $5 billion per year. Addi tionally, toxic and greenhouse gas emissions can be reduced by the use of biofuels. In the same study, Greene et al. report that 22% of the United St ates’s total greenhouse gas em issions could be reduced if biofuels were developed to replace half of the petroleum consumption. Arguably, the most important benefit of biofuel production is the potential for closing the carbon cycle.

Gasification Gasification is a high-temperature and catalytic pathway for produci ng biofuels. It is defined as the partial oxidation of solid, ca rbonaceous material with air, steam, or oxygen into a flammable gas mixture called producer gas or synthesis gas [5]. The synthesis gas contains mostly carbon monoxide and hydrogen with various amounts of carbon dioxide, water vapor, and methane. Typical volumetric energy conten t of synthesis gas is 4–18 MJ/Nm 3 

### [9]. Comparatively, natural gas (composed of mostly methan e) energy content is 36 MJ/Nm 3 

### [9]. Much of the energy content

### 3 of the biomass is retained in the gas mixture by partial oxidation rather than full oxidation of the biomass, which would result in th e release of mostly thermal ener gy. Historically, gasification of coal and wood produced “town gas,” which was subsequently burned in street lamps [10]. Additionally, during the Worl d Wars, vehicles were adapted to operate with gasification reactors [10]. During this same time period, Germany deve loped the catalytic synt hesis of transportation fuels from synthesis gas [11]. The same concept is still in use today by the South African Coal, Oil, and Gas Corporation (Sasol) to produce moto r fuels and liquid by-products using coal [11].

Reactions Four stages occur during gasification of carbonaceous material: drying, devolatilization, combustion, and reduction [9]. First, the moisture within the material is heated and removed through a drying process. Second, continued heat ing devolatilizes the ma terial where volatile matter exits the particle and comes into cont act with the oxygen. Third, combustion occurs,

where carbon dioxide and carbon monoxide are formed from carbon and oxygen. The combustion stage is very exothermic and provide s enough heat for the last stage, the reduction reactions, to occur. The last stage includes water gas reacti on, Boudouard reaction, water-gasshift reaction, and methanation reaction (Table 1) . As all these stages progress, solid fixed carbon remains present. The amount of fixed car bon varies depending on th e equivalence ratio.

Table 1. Reactions Occurring within the Reduction Stage of Gasification

Name Reaction

Water Gas 

### ܥ൅ܪ ଶ 

### ܱ ՜ ܥܱ ൅ ܪ ଶ

Boudouard 

### ܱܥ ൅ ܥ ଶ 

### ܱܥ2 ՜

Water-Gas-Shift 

### ܥܱ ൅ ܪ ଶ 

### ܱ ՜ ܥܱ ଶ 

### ൅ ܪ ଶ

Methanation 

### ܥܱ ൅ 3ܪ ଶ 

### ՜ ܪܥ ସ 

### ൅ ܪ ଶ 

### ܱ



### When the equivalence ratio (defined as the actual air/fuel ratio all divided by the stoichiometric air/fuel ratio) increases, solid fixed carbon (i.e., char) decreases until enou gh oxidizer is available for complete conversion (Figure 2) . This point of complete convers ion occurs at an equivalence ratio of approximately 0.25. At nearly the same point, the maximum synthe sis gas energy content (without accounting for sensib le energy) is reached.

### 4 Figure 2. Typical energy content of the products of gasification of wood using air varied by equivalence ratio [12]



### Gasifier Types Three types of reactors are used for gasification: fixed bed, fluidized bed, and entrained flow [13]. Fixed beds are fed with biomass from the to p of the reactor and form a bed, which gasifies as air moves through the bed (Figure 3). As the ma terial releases volatil e components, the char and ash exit through a grate at the bottom. T ypical operating temperature range is 750°–900°C. The two main types of fixed-bed gasifiers are upd raft and downdraft. The advantage of fixed-bed technology is its simplicity, but the technology is limited in scale-up and has low heat mixing due to high channeling potential within the reactor [14]. 0 5 10 15 20 25

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1Enthalpy (MJ/kg biomass, moisture&ash-free)

Equivalence ratio (mol/mol) total

g as , total

g as , sensible heat

g as , heatin g value

char

, total

char, sensible heat char, heating

value

### 5 (a) (b)

Figure 3. Design of fixed-bed (a) updraft and (b) downdraft gasifiers showing reaction zones [13]



### When the volumetric gas flow is increased through the grate, the fixed bed becomes a fluidized bed. Fluidized-bed gasifiers are so named because of the inert bed material that is fluidized by oxidizing gas creating turbulence through the bed material (Figure 4). Biomass enters just above the top of the bed and mixes with hot, inert material, creating very high heat and mass transfer. Operating temperature range is th e same as for the fixed bed. A dvantages of the fluidized bed include flexible feeds, uniform temperature distri bution across the bed, and large volumetric flow capability [15]. The main types of fluidized-bed gasifiers are circulating fluidized bed (CFB) and bubbling fluidized bed. Bubbling-bed gasifiers are directly heated from the combustion reactions occurring in the bed. They produce gas and the as h and char fall out the bottom or the side. The CFB recycles the char through a cyclone while the product leaves out th e top of the cyclone. There are also indirectly heated fluidized beds that use a hot material such as sand to provide the heat needed for gasification, as shown in Figur e 4. Fluidized beds have high carbon conversion efficiencies and can scale up easily [14].

### 6 (a) Bubbling fluidized-bed gasifier (b) Circulating fluidized-bed gasifier

(c) Indirectly heated gasifier via
combustor (d) Indirectly heated gasifier via heat exchange tubes

Figure 4. Fluidized bed gasifier desi gns of (a) and (b) directly heated type and (c) and (d) indirectly heated type [16]



### Another type of gasifier is the entrained-flow gasifier (Figure 5) . Normally operated at elevated pressures (up to 50 bar), it requires very fine fuel pa rticles gasified at high temperatures to ensure complete gasification during the short residen ce times in the reactor. The Energy Research Centre of the Netherlands has investigated this type of gasification a nd reported promise with biomass, as long as the biomass is pretreated to certain requirements [17]. To keep the residence time at approximately the time for a particle to fa ll the length of the reaction zone, fuel particles smaller than 1 mm and high temperatures (1100°

– 1500°C) are necessary for successful operation.

### 7 Figure 5. Entrained-flow gasifier [18]



### Entrained-flow gasification mixes the fuel with a steam/oxygen stream to form a turbulent flow within the gasifier. Ash-forming components melt in the gasifier and form a liquid slag on the inside wall of the gasifier, effectively protec ting the wall itself. The liquid flows down and is collected at the bottom. To form the slag, lime stone can be added as a fluxing material. For herbaceous biomass, such as switchgrass or corn stover, which is high in alkali content, there may be sufficient inherent fl uxing material present [18]. Ad vantages of entrained-flow gasification are that tar and me thane content are negligible a nd high carbon conversion occurs due to more complete gasification of the char . Syngas cleanup is simplified because slag is removed at the bottom of the gasifier, negating the need for cyclones and tar removal [19]. The disadvantages are that very high temperatures need to be maintained and the design and operation is more complex. An entrained-flow gasi fier co-firing up to 25% biomass with coal has been developed by Shell in Buggenum, Netherla nds. Another gasifier developed by Future Energy in Freiburg, Germany, uses waste oil and sl udges. Both are operating at commercial scale [17].

Biomass Preprocessing A degree of biomass processing is required before gasification can occur. Most gasifiers require smaller size feedstock than is ty pically collected during harvest. Therefore, a significant degree of size reduction is necessary. A typical setup fo r size reduction is using a two-step process in which a chipper accomplishes the primary reduction and a hammer mill is used for the secondary reduction [20]. In addition, the maximum moisture content for biomass gasification is 20%–30%

### 8 (wet basis), and for normal operation it is less th an 15% (wet basis) [9]. Therefore, a drying process is required to prepare th e feedstock for gasification.

The main benefit of drying biomass is to avoid us ing energy within the gasifier to heat and dry the feedstock [21]. Drier bioma ss also makes for more stable temperature control within the gasifier. Rotary dryers typically operate utilizi ng hot flue gas from a downstream process as the drying medium. They have high capacity but requir e long residence times. In addition, rotary dryers have a high fire hazard when using flue ga s [21]. To avoid using flue gas, rotary dryers can use superheated steam, essentially an inert gas, when a combined-cycle heat and power system is used downstream. That system has si gnificant steam available for use because of the steam produced in the steam cycle. An advantage of using steam for drying is better heat transfer

and therefore shorter residence time.

Pretreatment options for entrai ned-flow gasification include to rrefaction followed by grinding to 0.1-mm particles, grinding to 1-mm particles, pyrolysis to produce bio-oil/char slurry (bioslurry), and initial fluidized-bed gasifica tion of larger particle s coupled to an entrained-flow gasifier. Torrefaction, essentially an oxyge n-free roasting process, causes the biomass particles to be brittle, which makes for easy grinding but releases up to 15% of the energy in the biomass via volatile compounds [17]. The fluidized/entrained coupled option is attractive because of an overall energy efficiency of 80%–85%, but it is expensive because two gasifiers are used in

series.

The bioslurry option is illustrate d in Figure 6. Basically, a flash- pyrolysis process yields bio-oil and char and is followed by a slagging, entrained -flow gasifier. Because this process utilizes an entrained-flow gasifier, the feed must be pressu rized. Fortunately, the pyrol ysis slurry, already in an emulsified liquid state, can be pressurized easil y. Technology for slurry feeding is state of the art as a result of experience with coal slurries [17]. The bioslurry still c ontains 90% of the energy contained in the original biomass [22]. Another a dvantage is that no inert gas is needed for solids pressurization, avoiding dilution of the feed, wh ich would dilute the syngas. In the search for cost-effective methods for production of syngas, this option has poten tial, but it isn’t as developed as other technologies such as fluidized-bed gasifica tion. The biggest challenge is constructing and operating a large-scale pyrolysis process, because large-scale systems have not been demonstrated [17].

Figure 6. Schematic of a biomass pretreatment via fast pyrolysis followed by an entrained-flow gasifier [17]

### 9 Syngas Cleaning Because the raw syngas leaving the gasifier cont ains particulates, tars, alkali compounds, sulfur compounds, nitrogen compounds, and other contaminants, those components must be removed or reduced significantly. Partic ulates and tars have the pot ential for clogging downstream processes. Sulfur and nitrogen ha ve the potential to poison down stream processes, especially catalysts used in fuel synthesi s applications. Moreover, another motivation for cleaning syngas is meeting environmental emissions limits.

The syngas must be cooled before conventi onal gas cleanup is utilized. This can be accomplished in two ways: direct quench by injec tion of water and indirect quench via a heat exchanger. Direct quench is le ss expensive but dilute s the syngas. Direct quenching can also be used to clean up the gas by removing alkali species, particulates , and tars [23].

Particulates are defined as i norganic mineral material, ash, and unconverted biomass or char [24]. In addition, bed material from the gasifi er is included in the particulates. Switchgrass feedstock typically has 10% inorga nic material in the form of minerals. Many gasifiers operate with a 98%–99% carbon conversion efficiency wher e 1%–2% of the solid carbon is in the form of char [24].

Particulates are primarily removed by physical systems such as cyclones, in which the heavy particles fall down the center whil e the gases rise up and out of the cyclone. The initial step for particulate removal is usually a cyclone. Importa ntly, particulates should be removed before the gas is cooled down for cold gas cleaning. If re moved after gas cooling, tars can condense onto the particulates and potentially plug equipment. Alternatively, barrier filters, which operate above tar condensation temperatures, use metal or ceramic screens or filters to remove particulates while allowing the gas to remain hot. Barrier filters, however, have presented problems in sintering and breaking [24].

Even more critical to downstream syngas applica tions is tar removal. Tars are defined as higherweight organics and oxygenated aromatics heavie r than benzene (78 g/mol). They are produced from volatized material after pol ymerization [24]. A review by Miln e et al. [25] of tars produced during gasification covers differe nt removal methods. One method us es physical removal via wet gas scrubbing of tars in a scrubbing tower for the “heavy tars” followed by a venturi scrubber for lighter tars. This setup is similar to the direct quench cooling mentioned previously, as cooling

occurs as well. Tar concentrati on is reported to be lower than 10 ppm by volume at the exit of this setup [23]. The disadvantage of this setup is that wastewater treatment is required to dispose of the tar and can be expensive.

The other method for tar removal is catalytic or thermal conversion to non-condensable gas. This is also known as hot gas cleaning, as it occurs at temperatures at or above gasification temperatures. Catalytic conversi on can occur at temperatures as low as 800°C, and thermal conversion can occur at temperatures up to 1200° C. The energy required for thermal tar cracking may not be cost-competitive because of the te mperature rise required from the gasification temperature to crack the hi gh refractory tars [24].

### 10 Alkali compounds such as calcium oxide and pota ssium oxide are present in biomass. When gasified, both either vaporize or concentrate in the ash. Condensation of these compounds begins at 650°C, and they can deposit on cool surfaces causing equipment clogging, equipment corrosion, and catalyst deactivation [26]. Accord ing to Stevens [26], research on alkali adsorption filters using bauxite has been prom ising but not demonstrated on a large scale.

Stevens concludes that the best current method for alkali removal is using proven syngas cooling followed by wet scrubbing, where the addition of water cools the syngas and physically removes small particles and liquid droplets.

Wet scrubbing also removes ammonia that forms during gasification from the nitrogen in the biomass. Without proper removal, ammonia can deactivate catalysts as well. Complete ammonia removal can be accomplished through wet scrubbing [27]. For gasifiers coupled to a catalytic or thermal tar reformer, most of the ammonia can be reformed to hydrogen and nitrogen [27]. Sulfur in the biomass mostly forms hydrogen sulfide (H 2 

### S) with small amounts of carbonyl sulfide (COS). Hydrogen sulfide is removed by three main ways: chemical solvents, physical solvents, and catalytic sorbents. For chemical removal, amine-based solvents are typically utilized to chemically bond with H 2 

### S. Physical removal takes advant age of the high solubility of

H 2 

### S using an organic solvent. Typical setups of both chemical and physical removal involve an absorber unit followed by a solvent regenerato r unit, known as a stripper. Operation usually occurs at temperatures lower than 100°C a nd medium to high pressures (150–500 psi) [27]. Sulfur leaving these two systems is around 1–4 ppm and can require further removal, especially for fuel synthesis. In that case, a syngas polis hing step using a fixed-be d zinc oxide-activated carbon catalyst removes H 2 

### S and COS to achieve the parts per billion levels necessary for fuel synthesis. Halides, present in trace amounts in the biomass, can also be removed with the zinc oxide catalyst [27]. The water streams with the scrubbed-out im purities are sent to wastewater

treatment.

End-Use Product After syngas is cleaned of particulates, impurities, and contaminants, there is sufficient energy content for producing a higher valued product. There are three main large-scale biomass gasification pathways that have been researched and suggest ed for producing higher valued products: power generation, liquid-fu el synthesis, and chemical s ynthesis. According to Wender [28], the three largest comme rcial uses for syngas are ammonia production from hydrogen, methanol synthesis, and hydrocarbon synthe sis via the Fischer-Tropsch process.

Power Generation Power is most effectively generated using gasi fication by combusting the syngas in a gas turbine to provide mechanical work for a generator. Steam can then be generated by recovering heat from the hot syngas, and the steam in turn is used for mechanical work via a steam turbine. This gasifier plus gas turbine and st eam turbine setup is known as integrated gasification combined cycle (IGCC) power generation. The level of particulates, alkali me tals, and tar can decrease the performance of the gas turbine. Consonni and La rson [29] found that particulates can cause turbine blade erosion; 99% of 10-micron or sm aller particles should be removed. In addition, they also report that alkali metals corrode th e turbine blades and tars condense on the turbine blades, both hindering operation and escalating turbine failure. Fortunately, nearly all alkali

metals and tars can be removed us ing proven wet scrubbing techniques.

### 11 Using the IGCC approach to generate power, Bri dgwater et al. [30] a nd Craig and Mann [23] expect biomass to produce power with efficiencies in the range of 35%–40% with large scale systems (greater than 100 MW net output) at the high end of the range. Moreover, Craig and Mann suggest that future advanced turbine systems could reach 50% biomass-to-power efficiency.

Synthetic Fuels and Chemicals Instead of converting the energy content of th e syngas to power, the energy content can be condensed into a liquid energy carrier, or fuel. The conversion of syngas to fuels can only occur in the presence of proper catalysts [31]. The cat alytic reactions basically build up the small molecules in the syngas (i.e., carbon monoxide and hydrogen) into larger compounds that are more easily stored and transported. A summary of many catalytic pathways to fuels and chemicals is shown in Figure 7. In most catalytic synthesis reactions, syngas cleanliness requirements are very high. Most impurities a nd contaminants are removed to low parts-permillion and even parts-per-billion concentrations. This means that significant cost must be directed toward syngas cleaning.

Figure 7. Main syngas conversion pathways [32]



### Methanol to Gasoline Methanol is one of the top chemicals produced in the world [32]. Most commercially produced methanol is synthesized via steam methane refo rming and autothermal reforming. The synthesis of methanol from syngas is highly exotherm ic (equation 1). The reaction occurs over a

### 12 Cu/ZnO/Al 2 

### O 3 

### catalyst at temperatures of 220°–275°C and pressures of 50–100 bar with a catalyst lifetime of 2–5 years [31]. Wender [28] reports s yngas-to-methanol conversion efficiency can reach 99% with recycle, but per-pass efficiency is only about 25%.

ܪ2 ൅ ܱܥ ଶ 

### ܪܥ ՜ ଷ 

### ܪܱ (1 .eqn (

Although methanol can be used di rectly as a liquid fuel, it ca n also be converted into the conventional transportation fuel range. This process is known as the methanol-to-gasoline (MTG) process and was developed by the Mobil Oil Corporation [31] . In that process, methanol is heated to 300°C and dehydrated over alumina cat alyst at 27 atm, yiel ding methanol, dimethyl ether (DME), and water. The ex iting mixture reacts with a zeo lite catalyst at 350°C and 20 atm to produce 56% water and 44% hydrocarbons by we ight. Of the hydrocarbon product, 85% is in the gasoline range, with 40% of the gasoline-range product being aromatic. However, limitations on the aromatic content of ga soline have been proposed in legislation [31]. The thermal efficiency of the MTG process is 70% [11]. The overall MTG process usually includes multiple MTG reactors in parallel to perform periodic ca talyst regeneration by burning off coke deposits [11]. Mobil operated a commercial plant pr oducing 14,500 barrels per day in New Zealand during the 1980s [32]. A lternatively, the reaction process could be stopped directly after the methanol synthesis to focus on producing DME, because DME can be used as a diesel fuel because it has a high cetane number. DME is form ed from the dehydration reaction of methanol over an acid catalyst γ -alumina. Per-pass efficiency can be as high as 50%. Overall, syngas-toDME production efficiency is higher than that for syngas to metha nol [31]. However, DME is in gaseous form at atmospheric condi tions and needs to be pressuri zed for use in diesel engines [33]. Therefore, engine modifica tion is required; this is the main disadvantage for DME use as transportation fuel.

Fischer-Tropsch Fischer-Tropsch (FT) catalytic synthesis is a highly exothermic r eaction producing a wide variety of alkanes (equation 2).

ܥܱ ൅ 2.1ܪ ଶ 

### ՜ െ ሺ ܪܥ ଶ 

### ሻ െ ൅ܪ ଶ 

### ܱ (eqn. 2)

For gasoline-range products, high er temperatures (300°–350°C) and iron catalysts are typically used. For diesel-range and wax products, lower temperatures (200°–240°C) and cobalt catalysts are typically used [34]. Opera ting pressures are in the range of 10–40 bar. Product distribution can be estimated using the Anderson-Schulz-Flory chain-growth-probability model, in which longer hydrocarbon chains form as the temperatur e decreases. At high temp eratures, selectivity favors methane and light gases. This is a disadva ntage if liquid fuel production is the focus. At low temperatures, selectivity favors long-car bon-chain wax products requiring further hydrocracking to the diesel range in a separate unit, which adds more construction cost but is necessary for liquid fuel production.

Because of the highly exothermic reaction, the heat must be removed or the catalyst can be deactivated. Two main types of reactors have be en designed: a fixed-bed tubular reactor and a

### 13 slurry-phase reactor (Figure 8). He at removal is crucial to the pr ocess and has been the focus of reactor design development [31]. The fixed-be d reactor has many catalyst tubes with heat removal achieved by steam generation on the outside of the tubes [35]. The fixed-bed reactor is simple to operate and is well suited for wax production due to simple liquid/wax removal. However, it is more expensive to build because of the many tubes and has a high pressure drop across the reactor [36]. The sl urry-phase reactor (SPR) opera tes by suspending catalyst in a liquid and bubbling the syngas th rough from the bottom. Disadvantages of the SPR are more complex operation and difficult wax removal. Ho wever, SPRs cost approximately 40% less to build than fixed-bed reactors [36].

FT diesel is very low in sulfur , low in aromatic content, and has a high cetane number, making it very attractive as a conventi onal fuel alternative. Emissions across the board decrease when using FT diesel. A South Africa-based company, Sasol, has been producing transportation fuel since 1955 using the FT process and supplies 41% of South Africa’s transportation fuel requirements [31].

(a) (b)

Figure 8. Fischer-Tropsch reactor types (a) mult i-tubular fixed bed and (b) slurry bed [31]



### Techno-Economic Analysis For biofuels technologies to be utilized in commercial appli cations, the economic feasibility must be determined. A feasibility analysis is al so called a techno-economic analysis, in which the technical aspects of a project ar e coupled to the economic aspects. First, the basic theoretical configuration is developed and a mass and energy balance is performed. Second, cost estimation

### 14 allows the investment and production cost of a bi orefinery to be determin ed. With rising interest in biorenewable resources, many techno-econom ic studies have been performed on power generation and biofuel scenarios. These studies assist in understanding how the physical process relates to the cost of producing renewable alternatives. Accuracy of results from these studies is usually ±30% of the actual cost [5].

Previous studies of gasification-based biomass-to-liquid production plants estimate the cost of transportation fuels to range from $12/GJ to $16/GJ ($1.60–$2.00 per gallon of gasoline equivalent) [16, 37-40]. The same studies estimate total capital investment in the range of $191 million for 2,000 dry metric ton per day (tpd) in put [39] to $541 million for 4,500 dry metric tpd input [38].

A 1,650 dry metric tpd biomass-to-methanol plant based on gasification, w ith a production cost of $15/GJ ($0.90/gal of methanol), is reporte d by Williams et al. [16] in 1991$ for $45/dry

metric ton of biomass. Williams et al. also show s the production cost of methanol-derived natural gas to be $10/GJ ($0.60/gal of methanol). However, that study conclude s that if a carbon tax system was developed for lifecycle carbon emissi ons, then renewable methanol could become competitive with natural-gas-derived methanol at a tax of approximately $90 per metric ton of carbon. A more recent study by Larson et al. of switchgrass-to-hydrocarbons production in 2009 reports a production cost of $15.3/GJ ($1.90/gal of gasoline) in 2003$ for a 4,540 dry metric tpd (5,000 dry short tpd) plant ba sed on gasification [38].

Table 2 compares four biofuel production studie s based on gasification. A range of cost year, plant size, and feedstock cost show the diversity of characteristics and assumptions that technoeconomic studies use. In addition, resulting capital investment costs of the studies have a large range. For example, the capital investment costs of the Phillips et al. [39] and Tijmensen et al. [40] studies are $191 million and $387 million, respect ively, at similar plant sizes. Reasons for such a significant difference ar e choice of technologies and le vel of technology development. The Phillips et al. study is a target stu dy, meaning that it assu mes future technology improvement and therefore lower costs. Direct comparison is difficult because of the varying assumptions used by each study.

### 15 Table 2. Previous Techno-Economic Studies of Biomass-Gasification Biofuel Production Plants

Williams et al. [16] Phillips et al. [39] Tijmensen et al. [40] Larson et al. [38]

Cost Year 1991 2005 2000 2003

Plant Size (dry metric ton per day) 1,650 2,000 1,741 4,540

Feedstock Generic
biomass Poplar Poplar Switchgrass

Fuel Output Methanol Ethanol FT liquids Diesel, gasoline

Feedstock Cost ($/dry short ton) 41 35 33 46

Capital Investment ($MM) N/A 191 387 541

Product Value ($/GJ) 15 12 16 15

Product Value ($/GGE) 1.90 1.60 2.00 1.85

### 16 

# Methodology

The following steps were undertaken to pe rform the analysis in this study:

• Collect performance information on relevant technologies for systems under evaluation

• Perform down-selection process with develope d criteria to identify most appropriate

scenarios

• Design process models using Aspen Pl us process engineering software

• Size and cost equipment using Aspen Icarus Process Evaluator software, literature references, and experimental data

• Determine capital investments and perf orm discounted cash flow analysis

• Perform sensitivity analysis on process and economic parameters

• Perform pioneer plant cost growth and performance analysis.

Down-Selection Process A number of process configurations for the gasi fication-based, biomass-to-liquids (BTL) route were initially considered. Thes e configurations are listed in Table 3 and discussed in the following sections.

Table 3. Process Configurations Considered in Down Selection Process

Gasifier Block Entrained flow, slagging gasifier

Fluid bed, dry ash gasifier

Transport gasifier, dry ash (e .g., Kellog, Brown, and Root)

Indirect gasifier, dry ash (e.g., Battelle-Columbus Labs)

Syngas Cleaning Water scrubbing

Catalytic tar conversion/reduction

Thermal tar conversion/reduction

Amine-based acid gas removal

Physical sorbent-based acid gas removal (e.g., Selexol, Rectisol)

Fuel Synthesis Fischer-Tropsch

Mixed alcohols

Methanol to gasoline (MTG)

Dimethyl ether

Syngas fermentation

### 17 Preliminary Criteria The initial technology confi guration options are reviewed and sc reened in accord ance with three

criteria.

First, the technology under consider ation should be commercially ready in the next 5–8 years and preferably with high technology developmen t. High technology development increases the likelihood that a configuration will perform at the scale defined in this study. For example, coal gasification has been demonstrat ed commercially at large scal es [11]. While similar scale biomass gasifiers have not been proven commer cially, the technology development of coal is assumed to apply for biomass in 5–8 years.

Second, the size of biorefinery s hould be feasible with typical agricultural pr oductivity and

within a realistic coll ection area. For example, if one-third of total land use surrounding the biorefinery is for stover collecti on and each acre provides conservati vely 1 short dry ton per year, then the required collection radius is 35 mile s and the amount of biomass transported to the biorefinery is approximately 2,300 short tons (2,090 metric tons) per day. The collection area

with a 35-mile radius is assumed to be realistic . In addition, previous st udies by Tijmensen et al., Phillips et al., and Lau et al. have used similar plant sizes [39-41].

Third, the desired product should be compatib le with the present transportation fuel infrastructure, i.e., gasoline and diesel-range hydrocarbons.

Scenario Selection For the gasification area, two ga sifiers are selected for modeli ng. First, an entrained-flow, slagging gasifier is chosen because of its commerci al application with coal (GE, Siemens, Shell, and ConocoPhillips) and its potential for use with biomass. Moreover, process modeling of this gasifier is simple because it can be closely approximated at thermodynamic equilibrium [3]. Second, a fluidized-bed, dry-ash gasifier is chos en because of experience at the Gas Technology Institute and because of data availability. Also , a report by Bain [4] at the National Renewable Energy Laboratory contains collected and analyz ed data for fluidized-bed gasification. In addition, Iowa State University is currently operating an atmos pheric-pressure fluidized-bed gasifier as either ai r or oxygen/steam fed.

The syngas cleaning area is chos en to include configurations that have less technological complexity than previous studies. Phillips et al. [39] and Larson et al. [38] both employ an external catalytic tar-r eforming process for dry-ash gasifi cation. Because of low technological development in tar conversion and its inherent complexity, a direct-con tact syngas quenching and scrubbing system is instead chosen for this study. In the case of the slagging gasifier, high temperatures inhibit tar formation, yet quenching and particulate and a mmonia removal are still required. An amine-based chemical absorber/strip per configuration is chosen for removal of hydrogen sulfide and carbon dioxide. This configurat ion is chosen due to data availability, as compared to proprietary physical gas cleani ng process such as Rectisol and Selexol.

Two fuel synthesis configurat ions under consideration pr oduce liquid hydrocarbons: FischerTropsch (FT) synthesis and MTG. FT synthesis has been proven in operation at commercial scale for many years by Sasol [11]. Because of more accessible data and long industrial experience, FT synthesis is the only fuel synt hesis option chosen. Because of this selection, a

### 18 post-synthesis fuel upgrading area is also necessary, as FT produc ts need to be separated and hydroprocessed.

Scenarios Not Selected The indirect, dry-ash gasifier a nd the mixed alcohol synthesis conf igurations are not considered due to previous work by Phillip s et al. [39]. The transport ga sifier design, t hough a promising technology, is not considered because of r eactor complexity, unproven commercial-scale operation, and lack of public domain data. Tar c onversion via external thermal or catalytic cracking is not considered due to lack of public domain data and commercial scale experience. Acid gas removal using proprie tary technology (e.g., Rectisol or Selexol) is not considered because of a lack of public operational data. MTG, including methanol synthesis, is not considered because of time constraints and limited operational data. DME and syngas

fermentation are not considered because of limited commercial scale experience and incompatibility with present fuel infrastructure.

Project Assumptions The main project assumptions for process a nd economic analysis are listed in Table 4

. A more extensive list can be found in Appendix A. The process design is assumed to incorporate an “n th


### plant” level of implementation experience. This is a significant assumption for a process concept that has yet to undergo detailed engineeri ng development and see its first commercial application. The design a nd operation of an n th 

### plant is likely to diverge from the design and operation considered in this study. However, the n th 

### plant evaluation is chosen to provide an analysis similar to studies completed by othe r groups, including NREL and Pacific Northwest National Laboratory.

Table 4. Main Assumptions Used in n th Plant Scenarios

Main Assumptions

The plant is modeled as n th plant
Plant capacity is 2,000 dry metric ton/day

Feedstock is corn stover at 25% moisture

Feedstock ash content at 6%

Feedstock is purchased at pl ant gate for $75/dry short ton

All financial values are adjusted to 2007 cost year
Plant is 100% equity financed

Fuel PV is evaluated at 10% internal rate of return
Plant initiates operation in 5–8 year time frame
Plant life is 20 years
Plant availability is 310 days per year (85%)

### 19 Process Description

High-Temperature Scenario Overview The high-temperature (HT) or sl agging scenario is a 2,000 dry me tric ton (2,205 dry short ton) per day corn stover-fed gasifica tion biorefinery that produces na phtha-range and distillate-range liquid fractions to be used as blend stock for gasoline and diesel, respectively, as well as electricity for export. It is based on pressurized, oxygen-blown, entrained-flow gasification. The

HT scenario is an n th 

### plant design, meaning significant design, engineering, and operating experience has been achieved.

Figure 9 shows seven main areas of operation for the HT scenario. Feedstock preprocessing (Area 100) is where the stover is chopped, dr ied, and ground to 1-mm size and 10% moisture. Gasification (Area 200) c ontains the stover pressurization fo r solids feeding, gasification, and slag removal. Synthesis gas cleaning (Area 300) contains cold gas cl eaning technologies where the syngas is quenched and scrubbed of partic ulates, ammonia, hydrogen sulfide, and carbon

dioxide. Area 300 also contains the water-gas-shift reaction, wh ich occurs before the hydrogen sulfide and carbon dioxide removal. This adjust s the ratio of hydrogen to carbon monoxide for optimal fuel synthesis. Fuel synthesis (A rea 400) contains syngas boost pressurization, contaminant polishing via zinc oxide guard beds, Fischer-Tr opsch reactor, and hydrocarbon gas/liquid separation. Hydroprocessing (Area 500) produces the final fuel blend and is treated as a black box utilizing published data. Power ge neration (Area 600) cont ains gas and steam turbines along with a heat rec overy steam generator. Area 700 contains the air separation unit (ASU) where oxygen is separated from air and pre ssurized for use in the gasifier. For cost analysis uses only, a balance-of-plant area includ es the cooling tower area, cooling water system, waste solids and liquids handli ng area, and feed water system. Detailed process flow diagrams can be found in Appendix D and detailed st ream data can be found in Appendix E.

### 20 Dryer Acid Gas

Removal Water,

Solids Fischer Tropsch Hydroprocessing

Naphtha, Distillate to storage Fuel Gas
to A600

Slag Grinder Entrained

Flow

Gasifier Lockhopper

Water

Sour Water

Gas Shift LO-CAT Sulfur

Cake

Compressor Pressure Swing

Adsorption

Hydrogen

Power Combustion Flue Gas

Steam

Syngas 

### A700

A100 A200 A300 A400 A500

A600 Chopper

Power Generation Steam Stover Air

Condensate Nitrogen Oxygen

Dry Stover Syngas Carbon Dioxide

1 2 3 4 5 6 7

8

9 10 11 12

13

14 15

16

17 18 19

20 21 22 24

23 25

26 27

28

29 33

Oil/Water Separation

Water

34

31

32

30 Direct Quench

Stream 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 T (°C) 25 25 90 90 16 32 149 120 200 120 220 249 1300 50 203 203 190 P (bar) 1.01 1.01 1.01 1.01 1.1 1.01 28 1.98 1.98 1.98 1.03 28 26.62 26.62 26.62 25.93 10 m (Mg/day) 2667 2667 2222 2222 2189 2903 744 444 4000 4000 271 180 3823 114 4000 3867 550

Stream 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 T (°C) 240 40 40 50 50 62 76 30 200 42 45 45 45 35 35 35 37 P (bar) 24.8 24.82 24.82 3.45 1.93 22.75 26 1.01 24.96 23.6 23.58 23.58 23.58 22.2 22.2 22.2 1.03 m (Mg/day) 4417 1502 2965 1755 3.1 3570 3570 4.38 1404 1068 2308 156 33 427 641 52 378

Figure 9. Overall process flow diagram for HT scenario (parallelograms enclosing numbers in the

diagram designate individual process streams, which are detailed in the accompanying table)



### Recycle streams are utilized to provide better syngas-to-FT-products conversion. Unconverted syngas in the fuel synthesis area is recycled to the syngas cleani ng area to remove carbon dioxide

and allow for further conversion in the Fischer- Tropsch reactor. A small portion of unconverted syngas is sent to a steam boiler to raise steam required for drying the biomass. The balance of unconverted syngas is combusted in a gas turbine and waste h eat is recovered in a steam generator for steam turbine power. Power genera ted is used throughout the plant and excess is

sold.

Some of the largest consumers of power are the ASU and hydropr ocessing area at 11.6 MW and 2.2 MW, respectively. Another consumer of power is the hammer mill for grinding the dried biomass in Area 100; it requires 3.0 MW. The amin e/water solution recirculation pump in Area 300 requires approximately 0.9 MW. Syngas co mpressors throughout the plant require a significant amount of power as well. Gross plant power pr oduction is 48.6 MW and net electricity for export is 13.8 MW.

Low-Temperature Scenario Overview The low-temperature (LT) or dry-ash scenario is a 2,000 dry metric ton (2,205 dry short ton) per day corn stover-fed gasification biorefinery th at produces naphtha-range and distillate-range

### 21 liquid fractions to be used as blend stock for gasoline and diesel, respectively, as well as electricity for export. It is based on a pressurized, oxygen/steam- blown fluidized- bed gasifier developed by Gas Technology Institu te. The LT scenario is an n th 

### plant design, meaning significant design, engineer ing, and operating experi ence has been achieved.

Figure 10 shows the seven main areas of operation for the LT scenario. Feedstock preprocessing (Area 100) is where the stover is chopped, dr ied, and ground to 6-mm size and 10% moisture. Gasification (Area 200) c ontains the stover pressurization fo r solids feeding, gasification, and char and ash removal. Synthesis gas cleaning (A rea 300) contains cold gas cleaning technologies where the syngas is quenched and scrubbed of particulates, ammonia, hydrogen sulfide, and carbon dioxide. Fuel synthesis (Area 400) contains syngas boost pressurization, contaminant polishing via zinc oxide beds, Fischer-Tropsch reactor, and h ydrocarbon gas/liquid separation.

Also included within Area 400 is the steam me thane reformer (SMR) to reduce methane content and water-gas-shift (WGS) to adjust the ratio of hydrogen and carbon monoxide.

Hydroprocessing (Area 500) produces the final fuel blend and is treated as a black box utilizing published data. Power generation (Area 600) contai ns gas and steam turbines along with a heat recovery steam generator. Area 700 contains the ASU that separates oxygen from air and pressurizes it for use in the ga sifier. Detailed process flow di agrams can be found in Appendix D and detailed stream data can be found in Appendix E.

### 22 Figure 10. Overall process flow diagram for LT scenario (parallelograms enclosing numbers in the diagram designate individual process streams, which are detailed in the accompanying table)



### Recycle streams are utilized to provide better FT products conversion. Unconverted syngas in the fuel synthesis area is recycled to the syngas cleaning area to remove carbon dioxide and allow for further conversion in the Fischer-Tropsch r eactor. The balance of unconverted syngas is combusted in a gas turbine and waste heat is r ecovered in a steam generator for steam turbine power. Power generated is used throughout the plant and excess is sold. Unconverted carbon within the gasifier in the form of char is co llected and combusted in a furnace to produce heat, thereby generating steam for the drying of the biomass.

Some of the largest consumers of power are the ASU and hydropr ocessing area at 9.1 MW and 1.7 MW, respectively. Another consumer of power is the hammer mill for grinding the dried biomass in Area 100; it requires 1.1 MW. The amin e/water solution recirculation pump in Area 300 requires approximately 0.7 MW. Syngas co mpressors throughout the plant require a significant amount of power as well. Gross plant power pr oduction is 40.7 MW and net electricity for export is 16.3 MW.

Area 100 Preprocessing The preprocessing area contains all the unit opera tions required for preparing the biomass for feeding into the gasifier. Biomass enters the plan t gate at 25 wt% moisture on wet basis in bales. The corn stover composition is shown below in Table 5. Ash content is assumed to be 6% by Stream 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 T (°C) 25 25 90 90 16 32 149 120 200 120 200 243 870 870 1500 40 40 P (bar) 1.01 1.01 1.01 1.01 1.10 1.01 28.00 1.98 1.98 1.98 1.00 28.00 27.55 27.55 1.00 26.17 26.17 m (Mg/day) 2667 2667 2222 2222 1744 2313 569 444 4000 4000 1471 180 2930 215 119 6000 2198

Stream 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 T (°C) 40 50 50 62 78 870 304 60 200 32 32 32 35 35 50 50 P (bar) 26.17 3.45 2.10 22.89 26.50 25.81 25.12 1.01 25.00 22.89 22.89 22.89 22.89 22.89 22.89 22.89 m (Mg/day) 1380 1593 2.3 2709 2709 3707 3708 3.4 560 1463 2073 168 331 1132 41 293

### 23 weight. The composition of char formed in the gasifier is also shown in Table 5. Forklifts transport the bales to conveyors where the stove r is separated from any metal in a magnetic separator. The first modeled opera tional area is a primary biomass chopper to complete the initial size reduction step and prepare stover for drying.

Table 5. Stover and Char Elemental Composition (wt %)



### The next area of operation is the direct-contact steam drying, whic h is modeled as a rotary steam dryer with exiting biomass moisture of 10% on we t basis. For steam dryers, Amos [21] suggests a 9:1 ratio of steam to evaporated moisture. Th erefore, 4,000 metric tpd steam is utilized in a loop and heated to 200°C from the hot combustion flue gases exiting the syngas- or char-fired

combustor in Area 200. Steam mixes with 25°C bioma ss and enters the dryer. At the exit, steam at 120°C returns to the combustor for reheating, and dried biomass exits at 90°C and is conveyed to the grinding area.

The grinding area has the same configuration as the chopping area ex cept that the grinder requires significantly more power because of th e larger size reduction. The grinder reduces the size of the biomass to 1-mm and 6-mm particles for the HT and LT scenarios, respectively. The power requirements of the grinder for the HT and LT scenarios are 3,000 kW and 1,100 kW, respectively. Energy requirements for grinding are determined using the correlations for specific energy (kWh per short ton), which is adapted from Mani et al. [42].

Area 200 Gasification The gasification area of the plant produces synthesi s gas using pressurized gasifiers. Also in this area, slag, char, and ash are removed. This area also includes lockhoppers for biomass pressurization and a fired combustor that provid es heat to raise steam for drying the stover.

Dried and ground stover enters the area and is im mediately conveyed to a lockhopper system for pressurized feeding. Carbon dioxide from the syngas cleaning area is used as pressurization gas. According to Lau et al. [41], a lockhopper system is the best setup for pressurized feeding of solids, despite higher operating costs due to hi gh inert gas usage. A proven track record with

biomass is the main reason for their recomm endation. The power requirement of a lockhopper system using biomass is 0.082 kW/metric tpd, resulting in a 180-kW system. Higman and van der Burgt [43] report inert gas usage as 0.09 kg/kg for 25-bar app lications. This results in a 180 metric ton per day carbon dioxide addition into the hopper. It is assumed that only 5% of the inert gas leaks into the gasifier whil e the rest is vented by the lockhopper. Element Stover Char

Ash 6.00 0

Carbon 47.28 68.05
Hydrogen 5.06 3.16
Nitrogen 0.80 0.29

Chlorine 0 0

Sulfur 0.22 0.15
Oxygen 40.63 28.34

### 24 Pressurized biomass is then conveyed into the gasi fier. Oxygen at 95% purity is produced fromthe air separation unit. A fixed 0.35 mass ratio of oxygen to biomass is used for the entrainedflow gasifier as reported by Henr ich [18]. Steam addition to the ga sifier is set at 0.48 mass ratio

of steam to biomass in accordance with Probste in and Hicks [11] and explained further in Appendix C. This gasifier operates at a temperat ure of 1,300°C, meaning that equilibrium can be modeled according to Frey and Akunuri [3]. The reactions shown in equations 3–9 are modeled using equilibrium constants.

ܥ ൅ 2ܪ ଶ 

### ՞ ܪܥ ସ 

### (eqn. 3)

2ܥ ൅ 1.5ܱ ଶ 

### ՞ ܥܱ ൅ ܥܱ ଶ 

### (eqn. 4)

ܥܱ ൅ ܪ ଶ 

### ܱ ՞ ܥܱ ଶ 

### ൅ ܪ ଶ 

### (eqn. 5)

2ܥܱ ൅ ܱ ଶ 

### ՞ 2ܥܱ ଶ 

### (eqn. 6)

ܵ൅ܪ ଶ 

### ՞ ܪ ଶ 

### ܵ (eqn. 7)

0.5ܰ ଶ 

### ൅ 1.5ܪ ଶ 

### ՞ ܰܪ ଷ 

### (eqn. 8)

ܪ ൅ ܱܥ ଶ 

### ܪ ൅ ܱܵܥ ՞ ܵ ଶ 

### (9 .eqn (

The LT scenario gasifier uses a 0.26 mass ra tio of oxygen to biomass at a gasification temperature of 870°C. This ratio is developed from the data found in an IGT gasifier study by Bain [4]. In that study, Bain develops mass bala nces for an IGT gasifi er operating with woody biomass. Steam addition to the gasifier is ca lculated using a 40/60 steam to oxygen mass ratio, consistent with experiments performed at Iowa State University using corn stover feedstock and a steam/oxygen-blown fluidized-bed gasifier. Lo w-temperature gasification cannot be modeled at equilibrium with or without approach temperatures for reactions. Instead, an elemental mass balance calculation and adjustment is performe d to ensure all inlet and outlet streams are accounted for across the gasifier. For details on the LT gasifier mass balance calculation, see Appendix C.

Yield from each gasifier is di fferent. As Table 6 shows, hydroc arbons and tars are not produced in the high-temperature gasifier because of near equilibrium conditions. Also, more hydrogen formation occurs in the high-temp erature gasifier as a result of the water-gas-shift reaction (equation 5) and because thermodynamically, ne arly no methane, ethane, and ethylene are produced. The low-temperature gasifier, on the other hand, produces a significant amount of methane, ethane, and ethylene in the syngas, requiring downstream reforming. In the HT scenario, slag forms from the ash when the ash me lts and flows on the inside walls. It is then collected at the bottom and removed for storage and subsequent waste removal. In accordance with Frey and Akunuri [3], it is assumed that 95% of the ash in the stover becomes slag while the rest becomes fly ash.

### 25 Table 6. Syngas Composition (Mole Basis) Leaving Gasifier for Gasification Scenarios Evaluated

Component High Temperature (mole fraction) Low Temperature (mole fraction)

Carbon Monoxide 0.264 0.240 Hydrogen 0.310 0.200

Carbon Dioxide 0.137 0.274

Water 0.280 0.194 Nitrogen 0.002 0

Methane 6×10 -6 0.055

Ethane 0 0.0061 Ethylene 0 0.013

Ammonia 3.1×10 -5 0.0094 Hydrogen Sulfide 6.72×10 -4 0.00112 Carbonyl Sulfide 2.6×10 -5 0 Tar (Anthracene) 0 5.00×10 -4

Oxygen 0 0
Argon 0.006 0.006



### Directly after the low-temperat ure gasifier, initial syngas clea ning by cyclones captures char and ash. The cyclones are split into two trains be cause of high volumetric gas flow. Each train contains a medium efficiency cyclone, follo wed by high efficiency cyclones to capture particulates. Overall particul ate removal efficiency for the cyclone area is 99%. Nearly particulate-free syngas travels to the more ri gorous syngas cleaning in Area 300. Captured char

in the LT scenario is collected and combusted in a fluidized-bed combustor providing energy for heating low-pressure steam, which is used fo r drying the stover. Syngas produced in the HT scenario contains fly ash, which is subsequent ly removed in a direct-water-quench unit. The

combustion area in the HT scenario receives unconverted syngas from the fuel synthesis area, as char is not produced. For both scenarios, the co mbustor is assumed to operate adiabatically, resulting in an exit flue gas temperature of approximately 1,800°C. Hot flue gas heats 120°C steam to 200°C and loops to the stover drying area.

Area 300 Syngas Cleaning After the initial particulate removal accomplished by the cyclones, the syngas still contains some particulates and all of the ammonia, hydrogen su lfide, and other contaminants. Area 300 removes these species using a cold-gas-cleaning approach , which is presently proven in many commercial configurations. Hydrogen sulfid e and carbon dioxide, collective ly known as acid gas, are absorbed via amine scrubbing. The LO-CAT hydr ogen sulfide oxidation pr ocess is used to separate carbon dioxide from hydr ogen sulfide with subsequent r ecovery of solid sulfur. In addition, the HT scenario contains a sour water-gas-shift (SWGS) process (sour because of the presence of sulfur), whereas the LT scenario situ ates the water-gas-shift (WGS) process directly upstream from the Fischer-Tropsch reactor.

Because of a less than optimal hydrogen to carb on monoxide ratio from the gasifier, a WGS reaction is necessary at some point in the process to adjust to the optimum Fischer-Tropsch ratio of 2.1. Therefore, a significant WG S activity is required, so a sizable amount of carbon dioxide is produced. To keep that carbon dioxide from bu ilding up in downstream processes, the SWGS

### 26 reactor is located before the acid gas remova l area. This SWGS unit operation is the most significant difference between the HT and LT scenarios in this area.

In the HT scenario, the syngas arri ving from the gasifier is cooled by direct-contact water quench to the operating temperature of the SWGS unit. In addition to cooling, the direct water quench removes all of the fly ash, sludge, and black water to prevent downs tream plugging. At this point, a portion of the syngas is diverted to the SWGS unit, which is modeled at equilibrium conditions and has an exit gas te mperature of 300°C. A ratio of 3:1 water to carbon monoxide is reached by addition of steam to the SWGS reactor. After the syngas is combined, the gas is further cooled to prepare for the acid gas remo val. In the LT scenari o, the direct quench unit condenses the syngas, removing approximately 90% of ammonia and 99% of solids. Tar is condensed in this unit and can be recycled back into the gasifier using a slurry pump, but this configuration is not modeled. A water treatment f acility for the direct-quench effluent is not

modeled but is accounted for in a balance-of-plant cost.

The next step for cleanup is the removal of acid gas (carbon dioxide and hydrogen sulfide) through the use of an amine-based solvent in a chemical gas absorp tion system. At this point in the cleaning process, hydrogen su lfide and carbon dioxide conten t is approximately 900 ppm and 30% on a molar basis, respectively. Sulfur must be removed to at least 0.2 ppm for FischerTropsch synthesis [31]. According to th e Gas Processors Suppliers Association Engineering
Data Book [44], amine-based systems are only capa ble of removing sulfur down to 4 ppm. Therefore, a zinc oxide guard bed is required to remove the difference. In this study, 20% concentrated monoethanolamine (MEA), capable of absorbing 0.4 mole acid gas per mole amine, is used as the absorbent. The process setup is based on a report by Nexant Inc. [27]. Hydrogen sulfide leaves the top of the absorber at 4 ppm and CO 2 

### leaves at 2%, which is 99% and 90% removal, respectively. The clean syngas is now ready for polishing to final cleanliness requirements. A stripper is utilized to desorb the acid gas and regene rate the amine solution. Before the acid gas and amine solution enters the stripper, a heat exchanger raises the temperature to 90°C.

Acid gas is brought to the LO-CAT sulfur reco very system to isolate hydrogen sulfide and

convert it to solid sulfur. The LO-CAT system sold and owned by Gas Technology Products uses oxygen and a liquid solution of ferric iron to oxidi ze hydrogen sulfide to el emental solid sulfur [45]. This system is suitable for a range of 150 lb to 20 tons per day sulfur recovery and 100 ppm

to 10% H 2 

### S concentration in sour gas, as reported by Nexant Inc. [27]. The sulfur production in this model is approximately 3 metric tpd with H 2 

### S concentration of approximately 150 ppm, which is within the reported ranges. First, the H 2 

### S is absorbed/oxidized, forming solid sulfur and water, while the ferric iron converts to ferrous iron. The second vessel oxidizes the ferrous iron

back to ferric iron and the sulfur cake is removed while the iron solution is recycled back into the absorber [46]. The carbon dioxide gas stream fr om the absorber is sp lit, and a portion is compressed and used in biomass pressurization wh ile the rest is vented to the atmosphere.

### 27 Area 400 Fuel Synthesis Conversion from syngas to liquid fuel occurs in Area 400, the fuel synthesis area. The major operations in this area are zinc oxide/activat ed carbon gas polishing, steam methane reforming (SMR) (only in the LT scenario), water-gas-shift (only in the LT scenario), Fischer-Tropsch (FT) synthesis, hydrogen separation via pressure swin g adsorption (PSA), FT products separation, and unconverted syngas distribution. Another major di fference between the LT and HT scenarios is

in this area. Area 400 in the LT s cenario contains the water-gas- shift reaction and steam methane reformer, as recycle streams contain a high enough content of methane and ethylene to significantly accumulate and cause dilution.

A compressor is the first operation in Area 400, boosti ng the pressure to 25 bar for FT synthesis. Then the syngas is heated to 200°C and passes through zinc oxide/activated carbon fixed-bed sorbent. This polishing guard bed acts as a barrier to any upstream non-normal contaminant concentrations and provides sulfur removal down to synthesis requirements. To limit downstream catalyst poisoning, the syngas stream must be cleaned of these components. Removal to 50 ppb sulfur is possible with zinc oxide sorbent [27]. To comply with reported requirements, the sorbent removes sulfur to approximately 200 ppb. In addition to sulfur, halides are removed by the sorbent. Syngas contamin ant level requirements for Fischer-Tropsch synthesis are shown in Table 7.

Table 7. Fischer-Tropsch Gas Cleanliness Requirements [31]

Contaminant Tolerance Level

Sulfur 0.2 ppm (200 ppb)

Ammonia 10 ppm

HCN 10 ppb

Halides 10 ppb



### Methane, nitrogen, and carbon dioxide act as inert ga ses in the FT synthesis. At this point in the LT scenario, an SMR step is utilized. Syngas is heated to 870°C through a fired heater and passed through a reformer nickel-based catalyst to reduce methane, ethylene, and ethane content. It is assumed that the SMR can be modeled to ope rate at equilibrium. Steam is added to bring the steam to methane ratio to approximately 6.0, wh ich at 870°C and 26 bar results in about 1.5% equilibrium methane content in the exit stream [4 7]. For the HT scenario, the SMR step is not necessary. The WGS reaction is now employe d for the LT scenario to increase the H 2 

### :CO ratio. A portion of the gas is diverted through the fixed catalyst bed while the rest bypasses the reactor,

similar to the SWGS unit in the HT scenario.

The exiting H 2 

### :CO ratio after WGS is slig htly above 2.1 in order for the excess hydrogen to be separated and used in the hydropr ocessing area. A PSA process is employed to isolate a stream of hydrogen. Because only a small amount of hydrogen needs to be separated from the syngas stream for downstream use, a small percentage of the syngas is directed to the PSA unit. Hydrogen removal efficiency within the PSA uni t is assumed to be 85%, and the unit produces pure hydrogen [41]. After the PSA , the syngas rejoins the main gas line and enters the FT

reactor.

### 28 The Fischer-Tropsch synthesis reactor operate s at 200°C and 25 bar using a cobalt catalyst according to equation 10. Per-pass carbon monoxide conversion in the reacto r is set at 40%. The product distribution follows the Anderson-Schulz- Flory alpha distributi on where chain growth factor, α , depends on partial pressures of H 2 

### and CO and the temperature of the reactor, as reported by Song et al. [48] for cobalt ca talyst and shown in equation 11, where ݕ is the molar fraction of carbon monoxide or hydrogen and ܶ݁݉݌ is the reactor operating temperature in kelvin. The reactor is based on a fixed-bed type reactor and that choice is reflected by the low per-pass CO conversion.

ܥܱ ൅ 2.1ܪ ଶ 

### ՜ െ ሺ ܪܥ ଶ 

### ሻ െ ൅ܪ ଶ 

### ܱ (eqn. 10)

· 0.2332൤ ൌ ߙ ݕ ை஼



### ݕ ை஼ 

### ݕ ൅ ுଶ 

### ൅ 0.6330൨ · ሾ 1 െ 0.0039 ሺ ܶ݁݉݌ െ 533 ሻሿ (eqn. 11)

To ensure that the hydrocarbon pr oduct distribution leans toward the production of diesel fuel, the value of alpha should be at least 0.85 and preferably greate r than 0.9, as shown in Figure 11. The reactor operating temperat ure needed to achieve a ch ain growth value of 0.9 is approximately 200°C. This produces 30 wt% wa x in the FT products, requiring hydrocracking

before addition to the final fuel blend. All exiti ng effluent is cooled to 35°C and the liquid water and hydrocarbons are separated in a gas/liquid knock-out separator. Unconverted syngas is split into four streams: direct recycl e to the FT reactor, recycle to the acid gas removal area, purge to the combustor in Area 200, and to the gas turbine in the power generation area. The LT scenario does not include sending a syngas stream to the combustor in Area 200 because char is used. Overall CO conversion is 66% due to recycling syngas. The recycle rati o is approximately 1.95

for both scenarios.

### 29 Figure 11. Fischer-Tropsch product distribution as a function of chain growth factor ( ࢻ (using equation 11 [48]



### Area 500 Hydroprocessing FT products from the fuel synthesis area cont ain significant amounts of high-molecular-weight wax. Hydrogen is required to crack these high- molecular-weight parrafins to low-molecularweight hydrocarbons. A product distri bution is specified in Table 8, as detailed in Shah et al. [49]. It is assumed that the hydroprocessing area contains a hydrocracker for converting the wax

fraction and a distillation sect ion for separating naphtha, diesel , and lighter-molecular-weight hydrocarbons. Also, hydrogen is assumed to be recycled within this area as needed. Methane and propane are separated and used to fuel the gas turbine in the power generation area. The hydroprocessing area is mode led as a “black box.”

Table 8. Hydroprocessing Product Distribution [49]

Component Mass Fraction

Methane 0.0346
LPG (propane) 0.0877
Gasoline (octane) 0.2610
Diesel (hexadecane) 0.6167



### Area 600 Power Generation A gas turbine and steam turbin e provide the means to produce the power that is required throughout the plant and the excess power that is generated for export. Unconverted syngas from Fisher-Tropsch synthesis and fuel gas from hydroprocessing are combusted in a gas turbine, producing hot flue gas and shaft work. The flue ga s exchanges heat with wa ter in a heat-recovery steam generator to produce steam for the steam tu rbines, which subsequen tly produce more shaft work. Electric generators attached to both the gas turbine and the steam turbine produce electricity from the shaft work. 0.00 0.20 0.40 0.60 0.80 1.00

0.00 0.20 0.40 0.60 0.80 1.00Weight Fraction

Chain Growth Factor, α Weight Fraction of Alkanes across Chain Growth Factor Range C1 = methane, C2 = ethane, C3 = propane, etc.

C1-C4

C5-C11

C12-C19

C20-C120

### 30 Area 700 Air Separation Because 95% purity oxygen is used for both scenar ios, a cryogenic ASU is employed rather than purchasing oxygen. A two-column cryogenic oxygen/nitrogen separation system is employed with subsequent oxygen compression and nitrogen vent. Air pre-cooling is accomplished by exchanging heat with exiting nitrogen. This ar ea requires a significan t amount of power, as explained in the results s ection, which is provided by the power generation area.

Methodology for Economic Analysis Capital investment and PV of each scenario ar e determined by finding all equipment costs and operating costs for the construc tion and operation of a plant for 20 years. Total capital investment is based on the total equipment cost plus additional installati on costs and indirect costs (such as engineering, construction, and c ontingency costs). Annual operating costs are

determined and a discounted cash flow rate of re turn analysis is developed. PV per unit volume of fuel is determined at a net present value of zero and a 10% internal ra te of return. The major economic assumptions used in this analysis are listed in Table 9. A detailed list of assumptions can be found in Appendix A.

Table 9. Main Economic Assumptions for n th Plant Scenarios

Parameter Assumption

Financing 100% equity

Internal rate of return (after taxes) 10%

General plant depreciation period 7 years (all areas except Area 600)

Steam plant depreciation period 20 years (Area 600 only)

Construction period 2.5 years with total capital investment spent at 8%, 60%, and 32% per year during years before operation

Startup time 0.5 years, where during that time revenues, variable operating costs, and fixed operating costs are 50%, 75%, and 100% of normal, respectively.

Income tax rate 39%

Contingency 20% of fixed capital investment

Electricity cost 5.4 cents/kWh

Working capital 15% of fixed capital investment

Land purchase 6% of total purchased equipment cost

Plant availability 310 days per year (85%)



### Unit operations from the scenario s are sized and costs estimated using Aspen Icarus Process Evaluator software based on the Aspen Plus simu lation data. Unique equipment costs for such equipment as the gasifier and Fischer-Tropsch sy nthesis reactor are estimated externally using literature references. Additionally, some equipmen t, such as the biomass dryer and lockhoppers, require literature references to determine the sizing, whereby their costs are subsequently

### 31 estimated using Aspen Icarus. The hydroprocessing plant area is modeled as a “black box” and

therefore its costs are estimated as an overall scaled area cost from literature.

The costs of each equipment item and process ar ea are scaled based on a scaling stream and scaling size factor ( ݊ ) according to equation 12, where the size factor is between 0.6 and 1.0, depending on the equipment type.

ݐݏ݋ܥ ௪௡௘ 

### ݐݏ݋ܥ ൌ ଴ 

### ൤ כ ݁ݖܵ݅ ݁ܽ݉ݎݐܵ ௪௡௘ 

### ݁ݖܵ݅ ݁ܽ݉ݎݐܵ ଴ 

### ൨ ௡ 

### (eqn. 12)

All purchased equipment costs determined via Aspe n Icarus contain an in stallation factor that accounts for piping, electrical, and other costs requir ed for installation. However, this installation factor tends to be sign ificantly lower than metrics suggested by Peters et al. [50]. Therefore, rather than using the software-derived installation factors, an overall installation factor is applied to most equipment. A 3.02 overall installation fact or is used, as suggested by Peters et al. for solid-liquid plants. Basically, the purchased equi pment cost of a piece of equipment is multiplied by the installation cost to determine its installe d cost. For the gasification unit, a 2.35 installation factor is used according to a National Energy Te chnology Laboratory study by Reed et al. [51]. It is assumed that all gas compressors receive a 1.2 installation factor, whic h is consistent with Aspen Icarus. The Chemical Engineering Plant Cost Index is used to bring the cost to 2007 U.S.

dollars whenever a source for an estimated cost is from a previous year [52]. For multiple unit operations that operate in parallel or in trains, a train cost factor is applied. The reason for the factor, as reported by Larson et al . [38], is because those units share some piping, electrical, and other installation costs. It is ap plied as shown in equation 13, where ݊ is the number of units in

the train and ݉ is the train factor with a value of 0.9.

ݐݏ݋ܥ ௧௥௔௜௡ 

### ൌ ݐݏ݋ܥ ௨௡௜௧ 

### כ ݊ ௠ 

### ) eqn. 13)

Table 10 explains the methodology used to estimate capital inve stment. After total purchased equipment cost (TPEC) and total installed cost (TIC) are determined, indirect costs are applied. Indirect costs (IC) include engineering and supervision, construction expenses, and legal and contractors’ fees at 32%, 34%, and 23% of TPEC, respectively [50]. Project contingency is
added as 20% of total direct and indirect cost (TDIC). TDIC is set as the sum of IC and total installed costs (TIC). Adding pr oject contingency to TDIC give s the fixed capital investment. Total capital investment (TCI) is determined by adding working capital to fixed capital investment and thereby repres ents the overall investment re quired for each scenario.

### 32 Table 10. Methodology for Capital Cost Estimation for n th Plant Scenarios

Parameter Method

Total Purchased Equipment Cost (TPEC) Aspen Icarus Process Evaluator,
references

Total Installed Cost (TIC) TPEC Installation Factor

Indirect Cost (IC) 89% of TPEC a

Total Direct and Indirect Costs (TDIC) TIC + IC

Contingency 20% of TDIC

Fixed Capital Investment (F CI) TDIC + Contingency

Working Capital (WC) 15% of FCI

Total Capital Investment FCI + WC

a Indirect costs are broken down into engineering and supervision, construction expenses, and legal and contractors’ fees at 32%, 34%, and 23%, respectively, for a total of 89% of TPEC.



### Raw material costs are inflated to 2007 U.S. dollars using th e Industrial Inorganic Chemical Index also used by Phillips et al. [39]. Annual variable operating costs are determined from

material stream flows. Variable operating costs and their respective cost methods are shown in Table 11. Natural gas for use in the gas turb ine to produce power during startup and backup periods is assumed to be employed 5% of the an nual operating time. Solids disposal costs are for the handling and removal of ash in the LT scenar io and slag in the HT scenario. Wastewater disposal cost is applied to the sludge and black water produced during direct syngas quench. Catalyst costs are not calculated on an annual basis because the catalysts for all reactors are assumed to be replaced every 3 years. Instead th ey are accounted for in the discounted cash flow analysis.

### 33 Table 11. Variable Operating Cost Parameters Adjusted to $2007

Variable Operating Costs Cost information

Feedstock $75/dry short ton

LO-CAT Chemicals $176/metric ton of sulfur produced as reported in Phillips et al. [39]

Amine Makeup $1.09/lb as reported in Phillips et al. and set as 0.01% of the circulating rate [39]

Process Steam $8.20/ton (Peters et al.) [50]

Cooling Water $0.31/ton (Peters et al.)

Hydroprocessing $4.00/barrel produced as reported by Robinson and Dolbear [53]

Natural Gas (for backup) $6.40/thousand standard cubic feet as the average wellhead price for 2007 [54]

Ash/Char Disposal $23.52/ton [39]

Wastewater Disposal $3.30/ hundred cubic feet [39]

Electricity $0.054/kWh a

Sulfur $40.00/ton [39]

Fischer-Tropsch Catalyst (cobalt) $15/lb and 64 lb/ft 3 density; applied on first operation year and then every three years a

Water-Gas-Shift Catalyst (copper-zinc) $8/lb and 900 kg/m 3 ; applied on first operation year and then every three years. Sour shift and normal WGS are assumed to operate with same catalyst a

Steam Methane Reforming Catalyst (nickel-aluminum) $15/lb and 70 lb/ft 3 ; applied on first operation year and then every three years a

Pressure Swing Adsorption $2/lb a

a Assumed.



### Fixed operating costs include employee salaries, overhead and maintenan ce, and insurance and taxes. Salaries are calculated similarly to Phil lips et al. [39], where employees include a plant manager, shift supervisors, lab technicians, maintenance technicians, shift operators, yard workers, and office clerks. The labor index devel oped by the Bureau of Labor Statistics [55] is used to adjust the labor cost to $2007. Overhead is calculat ed as 60% of total salaries;

maintenance cost and taxes/insurance cost are both 2% of total instal led equipment cost in accordance with Aden et al. [56].

For the DCFROR analysis, the capital investment is spent over a 2.5 year construction period, with 8% in the first half-year, followed by 60% and 32% for the second a nd third years. Working capital is applied in the year before operation and recovered at the e nd of the plant life. A

standard modified accelerated co st recovery system (MACRS) is used, with the steam plant depreciating over 20 years and th e rest of the plant over a 7-y ear period consistent with IRS

### 34 allowances. The project life is 20 years. Plant availability of 310 days per year (85%) is assumed and affects raw materials purch ase as well as fuel production. The PV per gallon of gasoline equivalent is calculated for a set net present value of zero including a 10% inte rnal rate of return.

Methodology for Major Equipment Costs The software used for determining equipment costs is not capable of estima ting every unit in this study. Some units such as the gasifiers and Fi scher-Tropsch reactors ar e unique pieces of equipment that are underestimated if estimated as a simple vertical pressure vessel. Therefore, literature sources are used to help estimate si zes and costs of many units. The following section details a few of the more important units.

The biomass dryer costs are estimated by determ ining the drying contact area. According to Couper [57], typical rotary dryers have a diamet er of 6 feet and solids holdup of 8%. Assuming a bulk density of 100 kg/m 3 

### for ground stover and 1,000 kg/m 3 

### for moisture in the stover, the resulting total surface area re quired for drying is 1,880 m 2 

### . The surface area provides enough information for estimating the costs because ro tary dryer costs are estimated based on surface area in Aspen Icarus. Details on dr yer sizing can be found in the Deta iled Calculations section of Appendix C.

The lockhopper system sizes are estimated by refe rring to a U.S. Department of Energy report completed by Combustion Engineering, Inc. [ 58], in which residence times and operating pressures are given. The biomass receiving bin, loc khopper, and feed bin costs are then estimated with Aspen Icarus. Details on lockhopper sizing can be found in the Detailed Calculations section of Appendix C.

The high-temperature gasifier cost is estimated from Reed et al. [ 51]. The total bare erected cost (installed cost) of a train of eight high-temperature E-Gas gasifiers (2,500 metric tpd coal) including syngas cooling costs is $638 million ($ 2006). It is assumed that the syngas cooling accounts for 20% of that cost, and therefore the estimated installed cost in millions of $2006 for a 2,000 metric tpd high-temperature gasifier follows the formula in equation 14, resulting in $57

million installed cost.

ݐݏ݋ܥ ு்௚௔௦௜௙௜௘௥ 

### ൌ ሺ638 · 80%ሻ

8 ൤ · ܶܯ2000

ܶܯ2500 ൨ ଴.଻ 

### (eqn. 14)

Installed cost for a fluidized-bed ga sifier is described in Larson et al. [38] and is calculated as shown in equation 15, where ݐݏ݋ܥ ௔௚_଴௦௜௙௜௘௥ 

### is $6.41 million ($2003), ܵݎݐ݁ܽ݉ ܵ݅ݖ݁ ଴ 

### is 41.7 metric ton per hour, and ݊ is 0.7. The gasifier is evaluated at 300 short tpd because that appears to be the highest proven capacity for a GTI gasifi er. Therefore, seven fluidized-bed gasifiers are used in parallel. It is assumed that the gasifier train follows the train cost formula (equation 13), resulting in $19 million installed cost.

ݐݏ݋ܥ ௜௙௜௘௥ ௦௚௔்௅ 

### ݐݏ݋ܥ ൌ ௥௘௜௙௜௦଴_௚௔ 

### ൤ כ ݁ݖܵ݅ ݁ܽ݉ݎݐܵ

݁ݖܵ݅ ݁ܽ݉ݎݐܵ ଴ 

### ൨ ௡ 

### (eqn. 15)

### 35 In a similar manner, the FT reactor is estimated as described in Larson et al. [38], where base installed cost is $10.5 million ($2003), base sizing value is 2.52 million scf/h of synthesis gas flow, and sizing exponent is 0.72. An installation f actor of 3.6 is assumed for the FT reactor as found in Peters et al. [50] for liquid production plants. This allows the purchased cost of the unit

to be back-calculated.

The acid gas removal (AGR) area cost is evaluated using information from Phillips et al. [39], following equation 12 where the base stream si ze is 4,000 short tpd and the base cost is $5.45

million. The stream size is the mass flow of the synthesis gas entering the AGR as the sum of fresh syngas from gas scrubbing and unconverted syngas from the fuel synthesis area.

Capital investment for the hydroprocessing area is found in Robinson et al. [53]. That study reports a volumetric unit cost of $4,000 per barr el per standard day. Assuming the typical hydroprocessing refinery produces 25,000 barrels per day, the base cost, C 0 

### , is $100 million. Assuming a scaling exponent of 0.65, the cost of Area 500 is found using equation 12. The cost details for both of the gasifiers, the AGR area, the FT reactor, and the hydroprocessing area can

be found in the Detailed Calcul ations section of Appendix C.

Methodology for Sensitivity Analysis Sensitivity parameters are chosen to reflect the change in PV. The parameters are either economic or process parameters. The sensitivity bounds are chosen based on what is expected to be observed in the construction and operation of a biomass-to-liquids production plant. The chosen favorable, baseline, and unfavorable sensitivity variables are shown in Table 12.

Table 12. Sensitivity Parameters for n th Plant Scenarios

Parameter Favorable Baseline Unfavorable

Availability (hours/year) 8,000 7,446 7,000

Balance of Plant (% of TPEC) a 8 12 16

Catalyst Cost (%) b 50 100 200

Catalyst Lifetime (years) 5 3 1

CO Conversion in FT Reactor (%) 30 40 50

Compressor Install Factor 1.0 1.2 3.0

Contingency (% of TDIC) c 10 20 30

Feedstock Cost ($/dry short ton) 50 75 100

Feedstock Moisture (% wet) 20 25 30

Price of Electricity (¢/kWh) 7.0 5.4 3.0

Total Capital Investment (% of baseline) 70 100 130

a TPEC=total purchased equipment cost.

b All catalyst costs are varied over this range.

c TDIC=total direct and indirect cost.

### 36 Methodology for Pioneer Plant Analysis Economic analysis is based on an n th 

### plant design, and before a pr oject is undertaken the pioneer (1 st 

### ) plant cost is important to estimate. We use a method beg un by the RAND Corporation that estimates pioneer plant costs a nd plant performance. Using this methodology, two main areas of

the n th 

### plant economic analysis are adjusted: capital investment and plant performance. Through a series of parameters, a cost in flation factor is generated to in flate the capital investment. In addition, a plant performance factor is calculated that reduces the fuel sales, feedstock purchase, and variable operating costs for the first several y ears that the plant is in operation. Each year, the plant performance factor is increased until fu ll performance is attain ed. For the purpose of determining a range of pioneer plant costs, baseline, optimistic, and pessimistic values are chosen. The details of the RAND methodology ca n be found in Merrow et al. [59]. The following section explains the reasoning behind the parameters chosen for the scenarios.

Cost growth and plant performance factors are calculated as shown in equations 16 and 17 in accordance with Merrow et al. [59].

ܵܧܫܶܫܷܴܲܯܫ כ 0.02125 െ ܹܧܶܰܥܲ כ 0.00297 െ 1.1219 ൌ ݄ݐݓ݋ݎܩ ݐݏ݋ܥ

ܵܵܧܰܧܸܫܷܵܮܥܰܫ כ 0.00111 ൅ ܻܶܫܺܧܮܲܯܱܥ כ 0.01137 െ ܱܰܫܶܫܰܫܨܧܦ ܶܥܧܬܴܱܲ כ 0.06351 െ (eqn. 16)

4.12 െ ܳܵܧܮܣܤ כ 0.33 ൅ ܲܵܧܹܵܶܧܰ כ 9.69 െ 85.77 ൌ ݂.ݎܲ݁ ݐ݈ܲܽ݊

ܵܦܫܮܱܵ כ 17.91 െ ܧܵܶܣܹ כ (17 .eqn (

The variables used in determining pioneer pl ant performance (equation 17) are defined as

follows:

NEWSTEPS (0+): The feedstock handling area is chosen as a new step because of the large scale, which has not been demonstrat ed with biomass. The gasifier and solids feeding are also included as a new step becau se a pressurized biomass feeding system has

not been demonstrated at a commercial scale except for in limited campaigns.

BALEQS (0 to 100): The mass and energy balances cannot be valid ated with current plant data, so a value of zero is chosen.

WASTE (0 to 5): Waste streams for gasifi cation include scrubber sludge, black water, gasifier slag, fly ash, and sulfur. The scrubbe r sludge and black wa ter require chemical treatment and the sulfur requires special ha ndling. A mid-range value of 2.5 is chosen.

SOLIDS (0 or 1): Solids are presen t; therefore a value of 1 is used.

Variables used in determining pi oneer plant cost grow th (equation 16) are defined as follows:

PCTNEW (0 to 100%): The percentage cost of the gasifier, solids pressurizing, and solids feeding out of the total purchased equipment cost.

### 37 IMPURITIES (0 to 5): There are two major r ecycle streams in the gasification process, and there is the possibility of inert component buildup. There is also a potential for equipment corrosion due to sulfur compone nts, hydrogen chloride, and hydrogen, so a value of 4 is assigned.

COMPLEXITY (0+): There are nine conti nuously linked steps in the gasification process. These include feedstock handling, so lids feeding, gasificat ion, amine scrubbing, sour water-gas-shift, pressure swing adsorption, Fischer-Tropsch synthesis, hydroprocessing, and air separation.

INCLUSIVENESS (0 to 100): Land costs and st artup costs are consid ered in the TCI; however, they have not been rigorously investigated. A value of 33% is used.

PROJECT DEFINITION (2 to 8): The gasificati on platform is considered to be in the study design stage so a value of 7 is assigned.

The factors are applied to the capital investment and plant performance as shown in equations 18 and 19. Expenses and revenues affected by the plant performance factor are fuel sales, feedstock purchase, co-product credits, and variable operating costs.

ܫܥܶ ௣௜௢௡௘௘௥ 

### ൌ ܫܥܶ ௛௧௡ 

### ݄ݐݓ݋ݎܩ ݐݏ݋ܥ (18 .eqn (

The ݐݏ݋ܥ ݐݓ݋ݎܩ݄ factor causes the TCI of the pioneer plant to increase from the n th 

### plant cost.

ݐݏ݋ܥ ௥௘௘௡௢௜௣ 

### ሺݐሻ ൌ ݐݏ݋ܥ ௡௧௛ 

### ሺݐሻ כ ሺ݈ܲܽ݊ݐ ܲ݁ݎ݂. ൅ 20 כ ሺ ݐെ1 ሻ ሻ

100 (eqn. 19)

ݐݏ݋ܥ ௡௧௛ 

### ሺ ݐ ሻ is the n th 

### plant expense or revenue at year ݐ . The plant performanc e factor is applied at year 1 and increases by 20% each year unt il 100% performance is reached. The chosen parameters and calculated factors for baseline, op timistic, and pessimistic scenarios are shown in

Table 13.

### 38 Table 13. Pioneer Plant Analysis Parameters and Factors

Parameter Baseline Optimistic Pessimistic Range

NEWSTEPS 2 1 3 0+

BALEQS 0 0 0 0–100

WASTE 4 3 5 0–5

SOLIDS 1 1 1 0 or 1

Plant Performance 38.18 49.93 22.31 0–100

PCTNEW 19 (9) a 10 (5) a 25 (20) a 0–100

IMPURITIES 4 3 5 0–5

COMPLEX 9 6 12 0+

INCLUSIV. 33 50 0 0–100

PROJ. DEF. 7 6 8 2–8

Cost Growth (HT) 0.47 0.63 0.30 0–1

Cost Growth (LT) 0.50 0.65 0.31 0–1

a Value in parentheses is value chosen for LT scenario.

### 39 

# Results and Discussion

Process Results With a higher carbon conversion efficiency (see mass and energy balance di scussions below), the HT scenario produces 61.0 GGE of fuels per dr y metric ton of feed stock compared to 47.2 GGE/metric ton for the LT scenario. Resulting a nnual fuel outputs are 41.7 million GGE/yr for the HT scenario and 32.3 million GGE/yr for the LT scenario.

Along with lower fuel yield, the LT scenario cons umes less power (Table 14). The LT scenario and HT scenario total power usage are 15 and 22 MW, respectively. Major contributions to this result are a lower grinder power due to less strict biomass size requirement, lower pressurized oxygen consumption in the gasifier, and genera lly lower downstream mass flow rates throughout the plant for the LT scenario. Lower syngas yiel d also means that there is less unconverted syngas and fuel gas from the hydroprocessing area available for the gas tu rbine. Therefore, the LT scenario generates 31 MW compared to 36 MW generated by the HT scenario. Due to unoptimized flow rates of the recycle streams, the LT scenario actually generates a net 16 MW of power, which is more than the 14 MW produced in the HT scenario. Reducing the net power generation is achievable by increasing the recycl e ratio and thereby increasing conversion, but a consequence is higher flow rates and therefore larger and more expensive equipment. The focus of this study is to produce liquid fuels. Ho wever, procedures to optimize recycle ratios, equipment sizes, and fuel production rates are no t within the scope of this study and are not

undertaken.

Table 14. Power Generation and Usage

Power (MW) HT Scenario LT Scenario

USAGE Chopper 0.50 0.50

Grinder 2.96 1.10 Lockhopper System 0.18 0.18 Lean Amine Solution Pump 0.86 0.69 Syngas Booster Compressor 1.25 0.96

PSA Compressor 0.15 0.11 Recycle Compressor 0.39 0.29 Hydroprocessing Area 2.24 1.73 Oxygen Compressor (ASU) 3.61 2.80 Air Compressor (ASU) 7.94 6.31

Sour Gas Shift Steam Compressor 1.59 0

CO 2 Compressor 0.39 0.39 Total Usage 22.06 15.06

GENERATION

Gas Turbine 26.25 21.02

Steam Turbine 9.63 10.40

Total Generated 35.88 31.42 Net Export 13.82 16.36

### 40 An energy balance of the scenarios shows that the biomass-to-fuels efficiency for the LT and HT
scenarios is 39% and 50% on a lo wer-heating-value (LHV) basis, re spectively (Table 15). When the net electricity is added, the efficiencies are 43% and 53% on an LHV basis, respectively. The LT scenario is expected to be lower because mass and energy loss occurs in the production and removal of char and tar. Char and tar energy loss sums to 7.5% of the en ergy in the biomass. In this scenario, char is combuste d in a fluidized-bed combustor to provide heat for biomass drying. Biomass drying in the HT scenario is accomplished by a syngas purge. However, the most significant energy loss in the LT scenario, about 25 %, occurs across the gasifier. One reason for higher energy loss in the LT scenario is because conversion efficiency in creases with increasing operating temperature. A second reason is the lo ss of energy during the cooling of the syngas after the gasifier. More effective capture of the energy in the hot syngas would increase the overall energy efficiency.

High exothermicity of the FT reaction causes a si gnificant portion of the chemical energy in the syngas to leave as thermal energy in both scen arios. A higher loss across the FT reactor is

observed in the HT scenario due to higher flow rates. Energy cl osure, as shown in Table 15, is approximately 90% for both scenarios. It is assume d that the last 10% is due mostly to heat loss from the cooling of the syngas by direct quench rath er than capturing the h eat and raising steam.

Table 15. Overall Energy Balance on LHV Basis

HT Scenario LT Scenario

IN

Biomass 100.0 100.0

OUT

Fuel 49.7 38.5 Net Electricity 3.5 4.2

Power Gen Losses 4.2 3.1

FT Reactor Losses 16.2 12.5

Gasifier Losses 12.1 24.9

Char - 6.3

Tar - 1.2 Syngas Purge 1.8 - Total a 87.5 90.7

a The balance of energy is assumed to come from various heating and cooling losses.



### A carbon balance analysis shows that 26% and 34% of the carbon in the biomass is passed on to

the fuels for the LT and HT scenar ios, respectively (Table 16). Approximately 99% of the carbon is accounted for. Major carbon losses include carbon dioxide flue gases, LO-CAT venting, and lockhopper venting. Char leaving the LT scenario is accounted for in the A200 flue gas because the char is combusted for process heat. Also, because the LT scenario produces low-molecularweight hydrocarbons in the gasifi cation process, a small fracti on is dissolved in the liquid

effluent of the wet scrubber. Ca rbon dioxide also dissolves in th e wet scrubber effluent stream.

Another carbon loss comes from the hydrocarbons that dissolve in the acid gas removal area.

### 41 Table 16. Overall Carbon Balance

HT Scenario LT Scenario

kmol/h % kmol/h %

IN

Biomass 3,281 100 3,281 100

OUT

Fuel 1,111 34 862 26

A300 CO 2 Vent 1,458 45 1,294 39

A600 Flue Gas 334 10 302 9

A200 Flue Gas 39 1 227 7

Lockhopper Vent 159 5 162 5

Wet Scrubber
Effluent 154 5 318 10

Tar 0

- 35 1

Dissolved Hydrocarbons 0 - 46 1

Total 3,257 99 3,245 99



### Throughout both scenarios, steam and cooling water are required as utilit ies. Because a pinch analysis (a method to optimize heat exchange) is not undertaken for this study, integration of the heat streams is not optimized. Therefore, it is assumed that the resul ting heating and cooling requirements within the model represent steam and cooling water utilities, whereby they are recycled at a ratio of 9:1. In other words, fr esh steam and cooling water utility input to the
scenarios are calculated at 10% of the required circulating rate.

Cost Estimating Results

Capital and Operating Costs for n th 

### Plant The breakdown of costs by area and resulting tota l capital investment is shown in Table 17. Total capital investment amounts for the HT a nd LT scenarios are $606 million and $498 million, respectively. Major areas of inve stment are the gasification area in the HT scenario and the fuel synthesis area in the LT scenario. Moreover, thes e two areas contain significant differences in capital investment between the scenarios. The inst alled cost of the entrained-flow gasifier is significantly higher than the cost of the fluidized-bed gasifiers, even when seven of the latter are used in parallel. Area 400 costs for the LT scenario are higher than for the HT scenario due to the

steam methane reformer and additional heat -exchange equipment required for the high operational temperature. A signifi cant portion of the capital cost is due to gas compression, such as the air compressor in the air separation un it and the syngas booster compressor. Due to high purchase costs, compressors make up approxima tely 18% of the TPEC for each scenario. Detailed accounting of equipment found in each process area can be found in Appendix B.

### 42 Table 17. Capital Investment Breakdown for n th Plant Scenarios

HT Scenario LT Scenario

Area Installed Cost Installed Cost

$MM % $MM %

A100: Preprocessing 23 7 23 9

A200: Gasification 68 22 28 11

A300: Syngas Cleaning 34 11 29 12

A400: Fuel Synthesis 49 16 59 23

A500: Hydroprocessing 33 11 30 12

A600: Power Generation 46 15 39 15

A700: Air Separation Unit 24 8 20 8

Balance of Plant 33 11 27 11

Total Installed Cost 309 254

Indirect Cost 130 107

Total Direct and Indirect
Cost 439 361

Contingency 88 72

Fixed Capital Investment 527 433

Working Capital 79 65

Total Capital Investment 606 498



### Annualized costs for operation of the plant are sh own in Table 18. The percentage displayed also represents percentage of PV. Th e largest annual incurred costs fo r both scenarios are the average return on investment and feedstock purchase. Utilities such as steam and cooling water are higher for the LT scenario due to heating and coo ling of the syngas before and after the SMR and steam input to the SMR. Waste disposal costs ar e equal because equal amount s of ash or slag are by-products of both plants. Annua l hydroprocessing area costs and income taxes are higher for the HT scenario because of its higher fuel pr oduction rate. Fixed costs and capital depreciation are also higher due to higher TCI.

Catalyst costs are not determined on an annua l basis because catalysts are assumed to be replaced only every 3 years. Table 19 contains ca talyst replacement costs. The catalyst cost for the ZnO guard bed and PSA unit are equal across the scenarios because the volumes of the units are assumed to be the same. The FT catalyst for th e HT scenario is significantly more expensive because of a higher gas flow rate and hence, high er catalyst use. Using a DCFROR analysis, the PVs at a net present value of zero for the LT and HT scenarios are $4.83/GGE and $4.27/GGE, respectively. Further de tail of the yearly cash flow of th e life of the plant can be found in Appendix B.

### 43 Table 18. Annual Operating Cost Breakdown for n th Plant Scenarios

HT Scenario LT Scenario
$MM % $MM %
Average Return on
Investment 58.2 32.7 48.3 31.0

Feedstock 51.3 28.9 51.3 32.9
Capital Depreciation 26.3 14.8 21.7 13.9
Average Income Tax 21.9 12.3 18.0 11.6

Fixed Costs 14.3 8.1 12.4 8.0

Hydroprocessing 4.4 2.5 3.0 2.0

Steam 2.7 1.5 3.5 2.2
Cooling Water 2.3 1.3 3.5 1.6

Waste Disposal 1.5 0.3 1.5 0.3

Other Raw Material Costs 1.4 0.8 1.3 0.8
Co-product Credits (5.6) -3.1 (6.6) -4.2

Table 19. Catalyst Replacement Costs for Bo th Scenarios (3-Year Replacement Period)

Catalyst HT
scenario LT
scenario

Water-Gas-Shift (copper-zinc) $115,000 $105,000

Steam Reforming (nickel-aluminum) N/A $103,000

ZnO Guard Bed $424,000 $424,000
PSA Packing $497,000 $497,000 Fischer-Tropsch (cobalt) $7,687,000 $6,128,000



### Sensitivity Results for n th 

### Plant The results of sensitivity analysis are summarized in Figure 12 and Figure 13 for the HT and LT scenarios, respectively. Total cap ital investment and feedstock purchase cost have the highest effect on the PV at approximately ±$0.80 and ±$0.40 per GGE, respectively, for both scenarios. Due to the high percentage of equipment cost fo r compressors, the compressor installation factor has a very high effect on PV as well. When the compressor installation factor is increased to 3.0,

which is the usual installation f actor for most of the equipmen t, the PV increases by $0.71 and $0.78 per GGE for the LT and HT scenarios, respec tively. Parameters with a lesser, but still significant, effect are the continge ncy factor (as percenta ge of total direct and indirect costs) and plant availability, both with an effect of approximately ±$0.20 per GGE. Parameters with the least effect are generally characte ristic of the process rather than of the economics. For example, catalyst life, feedstock moisture, and carbon monoxide conversion in the FT reactor affect the PV less than ±$0.15 per GGE.

### 44 Figure 12. Sensitivity results for HT n th plant scenario

Figure 13. Sensitivity results for LT n th plant scenario



### Additionally, the size of th e plants can be varied by feedstock i nput rate. The effects of plant size on PV and TCI are shown in Figure 14 and Figur e 15, respectively. When the plant size is reduced to 500 metric ton per da y, the two scenarios approach equa l PV. Also, as the plant size is reduced from the baseline, the difference in cap ital investment decreases. As the plant size increases past the baseline, the slope of the PV levels out, suggesting that the benefits of lower PV may not be worth the significant in crease in capital cost (Figure 14). 4.23 4.23 4.23 4.23 4.24 4.18 4.08 4.07 4.19 3.85 3.47

4.32 4.33 4.34 4.35

4.40 4.36 4.45 4.46 4.98 4.69 5.05

$3.25 $3.50 $3.75 $4.00 $4.25 $4.50 $4.75 $5.00 $5.25 Per pass CO Conversion in FT 50:40:30 (%) Electricity Price 7.0:5.4:3.0 (¢/kWh) Feedstock Moisture 20:25:30 (%wet) Catalyst Cost 50:100:200 (%) Catalyst Life 5:3:1 (year) BOP 8:12:16 (% of TPEC) Availability 8000:7446:7000 (hr/yr) Contingency 10:20:30 (% of TDIC) Compressor Install Factor 1.1:1.2:3.0 Feedstock Cost 50:75:100 ($/ton) Total Capital 70:100:130 (%)

Product Value (2007$/GGE)

4.79 4.79 4.78 4.77 4.80 4.73 4.63 4.62 4.74 4.29 3.99

4.88 4.91 4.90 4.92

4.98 4.93 5.02 5.04 5.61 5.37 5.68

$3.75 $4.00 $4.25 $4.50 $4.75 $5.00 $5.25 $5.50 $5.75 Per pass CO Conversion in FT 50:40:30 (%) Catalyst Cost 50:100:200 (%) Feedstock Moisture 20:25:30 (%wet)

Electricity Price 7.0:5.4:3.0 (¢/kWh) Catalyst Life 5:3:1 (year) BOP 8:12:16 (% of TPEC) Availability 8000:7446:7000 (hr/yr) Contingency 10:20:30 (% of TDIC) Compressor Install Factor 1.0:1.2:3.0 Feedstock Cost 50:75:100 ($/ton) Total Capital 70:100:130 (%)

Product Value (2007$/GGE)

### 45 Figure 14. The effect of plant size on product value (per gallon of gasoline equivalent) for n th plant
scenarios

Figure 15. The effect of plant size on total capital investment for n th plant scenarios



### Pioneer Plant Analysis Results The total capital investment fo r a base-case pioneer plant is expected to double from the n th 

### plant scenario, as detailed in Table 20. PV for a base case pioneer plant is estimated to increase to $7.60 and $8.10 per GGE for the LT and HT scenarios, respectively. Table 20 also shows estimates for the optimistic and pessimistic cases . An important observation is that the pioneerplant PV for the LT scenario is actually lower than for the HT scenario. The reason behind this inverted result is because of the higher capital cost inflation (cost grow th factor) in the HT scenario due to higher gasifi cation area capital costs. 0.00 1.00 2.00 3.00 4.00 5.00 6.00 7.00

0 500 1000 1500 2000 2500 3000 3500 4000 4500 5000 5500Product value ($/GGE)

Plant Size (metric ton/day) HT Scenario

LT Scenario

0 200 400 600 800 1000 1200

0 500 1000 1500 2000 2500 3000 3500 4000 4500 5000 5500Total capital investment ($MM)

Plant Size (metric ton/day) HT Scenario

LT Scenario Base

Base

### 46 Table 20. Pioneer Plant Analysis Results

Analysis HT Scenario LT Scenario

TCI ($MM) PV ($/GGE) TCI ($MM) PV ($/GGE)

n th Plant 606 4.27 498 4.83

1 st Plant Base 1,400 8.10 1,080 7.60

1 st Plant Optimistic 1,030 6.30 830 6.30

1 st Plant
Pessimistic 2,200 12.60 1,740 11.50



### Comparison with Previous Techno-Economic Studies Two previous BTL studies that specifically use biomass feedstock, low-te mperature gasification, and Fischer-Tropsch synthesis technology are Tijm ensen et al. [40] and Larson et al. [38]. In order to compare with these stud ies, major economic and process parameters from the present n th


### plant LT scenario are adjusted to reflect similar va lues of the previous studies. First, the plant size of the present study is adjusted to increase equipment costs and raw materials purchases. As a result, the annual biomass input and TCI are affected. Second, av ailability in hours per year, rate of return, cost year, and f eedstock cost are adjusted. The co mbined effect of all adjusted parameters causes the present study’s produc t value to reflect the comparison study.

A comparison to the IGT-R scenario (which employs a low-temperature IGT gasifier and a steam methane reformer) in Tijmensen et al. shows that fuel product value is higher in the present study, as summarized in Table 21. Of all the scen arios developed by Tijmensen et al., the IGT-R scenario is most similar to the present study b ecause of the reformer. The IGT-R scenario has a TCI of $387 million, a feedstock cost of $33 per short ton, and a product value of $1.90/GGE. An important characteristic of the Tijmensen et al. study is that it does not include a hydroprocessing area. Therefore, it might be expected that th e TCI would be higher for the present study, which includes hydroprocessing. However, that is not the case, as the TCI of the present study using Tijmensen et al. parameters is $339 million, which is lower than the reported $387 million of the earlier study. Another importa nt observation is that the annual fuel production for the present study, with adjusted parameters, is 30.2 million gallons per year, compared to 39.8 million gallons per year of FT products reported by Tijmensen et al. One reason for lower annual fuel production in th e present study is becau se of a loss during hydroprocessing. Therefore, due to lower annual fuel production and hence lower fuel revenue, the present study has a higher product va lue compared to Tijmensen et al.

### 47 Table 21. Comparison of n th Plant LT Scenario to Tijmensen et al. Study IGT-R Scenario [40]

Parameter Tijmensen et al. Study (IGT-R Scenario) Present Study n th

Plant LT
Scenario Present Study w/ Tijmensen et al.
Parameters

Plant Size (dry tons/day) 1,920 2,205 1,920

Annual Biomass Input (ton s) 640,000 684,100 640,000
Total Capital Investment ($MM) 387 498 339

Availability (hour/year) 8,000 7,446 8,000

Rate of Return (%) 10 10 10

Cost Year 2000 2007 2000
Feedstock Cost ($/sho rt ton) 33.00 75.00 33.00

Efficiency (%, LHV, incl. elec.) 50.1 42.7 42.7

Fuel Yield (MMGGE/yr) 39.8 32.3 30.2 Product Value ($/GJ) 16.50 39.80 25.17 Product Value ($/GGE) 2.00 4.83 3.05



### A comparison to the FT-OT-VENT scenario (whi ch is low-temperature gasification with carbon dioxide vent and once through FT synthesis) report ed by Larson et al. is summarized in Table 22. In a similar fashion to the prev ious comparison, the parameters ar e adjusted to approximate the comparison study. Some important observations are ma de from this comparison. First, the TCI of the present study with adjusted parameters is sign ificantly higher. Second, the net electricity is significantly lower for the presen t study. Third, the PV is signifi cantly higher for the present study. Essentially, the Larson et al. study generate s more revenue from selling electricity and recovers the capital investment in less time. In addition, annual operating costs for the Larson et al. study are lower than for the present study. Therefore, because of lower revenue from electricity and higher annual opera ting costs, the present study ha s a higher fuel product value when compared on a similar basis to Larson et al.

### 48 Table 22. Comparison of n th Plant LT Scenario to Larson et al. Study FT-OT-VENT Scenario [38]

Parameter Larson et al.Study (FT-OTVENT Scenario) Present Study n th

Plant LT
Scenario Present Study w/
Larson et al.
Parameters

Plant Size (dry tons/day) 5,000 2,205 5,001

Annual Biomass Input (tons) 1,458,000 684,000 1,459,000
Total Capital Investment ($MM) 541 498 678

Availability (hour/year) 7,000 7,446 7,000

Debt/Equity (% Equity) 60 100 60

Rate of Return (%) 12 10 12

Cost Year 2003 2007 2003

Electricity Price (cents/kWh) 4.0 5.4 4.0

Net Electricity (MW) 207 16.3 37.1
Feedstock Cost ($/sho rt ton) 46.00 75.00 46.00

Plant Yield (MMGGE/yr) 63.3 32.3 68.9
Product Value ($/GJ) 15.25 39.80 26.80
Product Value ($/GGE) 1.85 4.83 3.25



### Summary of n th 

### Plant Scenarios The HT scenario requires more power and capita l investment, yields more fuel per ton of feedstock, and subsequently produce s more fuel per year compared to the LT scenario. The total capital investment requirements for the LT and HT scenarios are $498 million and $606 million, respectively. Despite higher cap ital investment for the HT scen ario, the product value (PV) is

lower. PVs for the LT and HT scenarios are $4.83 and $4.27 per gallon of gasoline equivalent, respectively. The main reason for a lower PV is because of increased fuel revenue. The main n th


### plant scenario results are shown in Table 23. A detailed summary of costs can be found in Appendix B.

Table 23. Main Scenario n th Plant Results

Scenario TCI ($MM) TPEC ($MM) Fuel Yield (GGE/metric ton) Annual Fuel Output (MMGGE/yr) Net Electricity Export (MW) PV ($/GGE)

High Temperature 605.9 145.7 61.0 41.7 13.8 4.27

Low Temperature 498.3 120.4 47.2 32.3 16.4 4.83

.

### 49 

# Conclusions

This analysis compares capital and operating cost for two biomass-to-liquids scenarios: hightemperature (HT) gasification and low-temperat ure (LT) gasification. Th e selection of these scenarios allows for direct comparison betweentwo modes of gasification: slagging and nonslagging. The slagging, entrained-fl ow gasifier employed for the HT scenario results in higher plant costs (about 20%) than the LT scenario, which employs a fl uidized bed gasifier. On the other hand, the higher carbon conversi ons for the HT gasifier result in a lower PV compared to the LT scenario. A biomass-to-liquids plant is ex pected to produce fuels co sting in the range of $4–$5 per gallon gasoline equivalent using presen t gasification and Fischer-Tropsch synthesis technology. The factors chiefly resp onsible for this relatively hi gh PV are feedstock costs and investment return on the capital to build a $500 million to $650 million plant to process 2,000 metric tons of biomass per day. A pioneer pl ant analysis estimates that the total capital investment for a pioneer plan t would double and PV would incr ease by approximately 60% from

the n th 

### plant scale. This uncertain ty suggests that economics are still a major challenge for biomass-to-liquids production plants.

Parameters with the most sensitive effect on PV are total capital cost, feedstock purchase cost, and compressor installation factor, affecting the PV ±$0.40–$0.80 per gallon. Less expensive

biomass feedstock that is lower in ash content th an that used in the pr esent study will have higher fuel yield and have the potential to signifi cantly decrease PV. Gas compression is a major portion of capital investment, and sensitivity analysis shows that installation costs of compressors have a high effect on PV. Factors with little effect on the PV are mostly related to the process, such as carbon monoxi de conversion in the FT reactor, feedstock inlet moisture, and catalyst lifetime.

Because of time and resource constraints, the techno-economic study presented includes a few shortcomings. The process configuration is not fu lly optimized for heat integration. While some recycle streams are included, a complete heat exchange network for heat recovery is not conceptualized. In addition, some areas such as FT product separation and hydroprocessing are not modeled rigorously and can be improved wi th detailed mass and energy flows. Further studies can benefit by optimizing he at exchange networks and mode ling in detail areas such as FT product separation and hydroprocessing. More de tailed modeling will likel y result in changes to economic components such as equipment cost, in stallation factors, and contingency. However, because increasing model detail may result in tradeoffs between the different economic components, the net effect of increased m odel detail on PV is difficult to predict.

This analysis tracks PV based on commercial te chnology for which sufficient public domain data

existed in 2007. In order to evaluate the econom ics of biofuels produced from gasification based on future scenarios, particular attention will be needed on the most sensitive parameters— feedstock cost and capital cost. Attention should also focus on commercial biomass-to-fuels gasification plants that come online. As these plants do come online, cost growth and plant performance factors will improve, ther eby decreasing the pioneer plant PV.

### 50 

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### 54 50. Peters, M.S.; Timmerhaus, K.D.; West, R.E. Plant Design and Economics for Chemical Engineers . New York, NY: McGraw-Hill, 2003.

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### 55 

# Appendix A. Techno-Economic Model Assumptions

Financial Assumptions

• Capital investment

o Equity: 100%

o Working capital (% of FCI): 15%

• Depreciation model

o Zero Salvage Value for both gene ral plant and steam/ power plant

o Type of depreciation: Double-Declini ng-Balance Depreciation Method (DDB) as per IRS Modified Accelerated Cost Recovery System (MACRS) guidelines

 Depreciation period (years):

• General plant: 7

• Steam/power system: 20

• Construction & startup:

o Construction period (years): 2.5

 % spent in year “-3”: 8%

 % spent in year “-2”: 60%

 % spent in year “-1”: 32%

o Startup time (years): 0.5

 Revenues (% of normal): 50%

 Variable costs (% of normal): 75%

 Fixed cost (% of normal): 100%

• Other

o Internal rate of return: 10%

o Income tax rate: 39%

o Operating hours per year: 8,406

Capital Costs

• Cost year for analysis: 2007; cost escalation is applied us ing the Chemical Engineering

Plant Cost Index

• The plant is designed based on the state of the technology, at the n th 

### plant level of experience

• Most equipment installation factors are applied using Peters et al. fo r solid-fluid plants (i.e., 3.02 installation factor)

### 56 • Materials of construction are carbon steel, stainless steel, alloys, and refractory where

necessary

• Sensitivity parameters involving changes in equipment size or capacity use scaling exponents available in literature

Operating Costs

• Working capital is assumed to be 15% of the total capital investment

• Annual maintenance materials are 2% of the total installed equipment cost

• Boiler feedwater and wastewater treatment costs are derived from prior NREL work

• Fresh cooling water and steam costs are calculat ed at 10% of the required circulation rate, meaning a 9:1 ratio of water recycling

• Employee salary estimation is same as that chosen by Phillips et al.

• Employee salaries are indexed to the year of 2007 following the data of the Bureau of

Labor Statistics

Feedstock, Products, and By-Products

• Feedstock is corn stover (comprisi ng stalks, leaves, cobs, and husks)

o Moisture content in the feedstock is 25%

• Feed rate is 2,000 dry metric ton per day

o The feedstock delivery logist ics are not considered

o The feedstock is delivered to the feed handling area of the plant

• Feed cost is assumed to be $75/dry short ton at the gate

• Gasoline and diesel products ar e sold for over-the-fence prices

• Gasoline energy content is 115,000 BTU/gallon

• Fly ash and slag incur a solids waste disposal cost

• Solid sulfur and electricit y are sold as by-products

Process Assumptions For both scenarios, most of the process was mode led with the aid of Aspen Plus software. The process was divided by l ogical process areas, which are named below.

Area 100 - Preprocessing

• Biomass is dried down to 10%

o Steam raised from hot flue gas is used to dry the feedstock

o Steam to moisture removal ratio is set at 9:1 in accordance with Amos

o Heat is provided by combus ting char and unreacted syngas

• Grinder reduces biomass to 6-mm or less

### 57 o The energy required for grinding is calc ulated separately using literature correlations by Mani et al.

Area 200 - Gasification

• Scenario 1: Entrained-flow gasifier is modeled using thermodynamic equilibrium

• Scenario 2: Fluidized-bed gasifier is modeled using a mass balance calculation

• 95% purity oxygen produced from air separation unit provides oxidizer

• Carbon dioxide is used as solids pressurization gas

• All char produced in LT scenario is combusted for process heat

Area 300 - Syngas Cleaning

• Particulates, tar, and partial ammonia removal via wet scrubbing

o Scrubbing water is recy cled at 90% rate

o Particulate handli ng (not modeled)

 High-temperature gasifier: particulate decant slurry is sent back into slagging gasifier

 Low-temperature gasifier: particulate decant slurry is piled and landfilled;

excess water is sent to aerobic water treatment (not modeled)

o Makeup water compensates for water lost via particulate slurry

 Process water condensate is used as makeup water

• Sour water-gas-shift occurs at equi librium and is modeled as such

• Carbon dioxide, hydrogen sulfide, and ex cess ammonia removal via amine scrubbing acid gas removal (AGR) at pressure

o 99% of sulfur is removed and 90% of carbon dioxide

o Monoethanolamine (MEA) is the scrubbing solvent

o Carbon dioxide is vented fo llowing LO-CAT removal of H 2 

### S

• Hydrogen sulfide is converted to solid sulf ur via LO-CAT oxidat ion (99% conversion)

• Ammonia can be disposed of by d ecomposition (not modeled) in

o Gasifier burner (slagging gasifier)

o Char and syngas combustor (fluidized-bed gasifier)

• Zinc oxide and activated carbon guard bed po lishing assumed (not modeled in detail)

Area 400 - Fuel Synthesis

### 58 • Water-gas-shift occurs at equilib rium and is modeled as such

• Pressure swing adsorption (PSA) is employed to remove excess H 2 

### at an efficiency of 85% and 99% purity

o The PSA system employs two trains with six reactors each to account for all stages of pressurization, de pressurization, purging, etc.

o PSA adsorbers are filled 2/3 with activa ted carbon and 1/3 with molecular sieve

• Syngas is catalytically converted to fuels by one-step Fisc her-Tropsch synthesis followed by wax hydrocracking and fuel separation

o FT synthesis employs cobalt catalyst

o 40% syngas conversion to fuels

o Part of the unconverted syngas is recycled

 A fraction of the recycle is sent to the AGR to prevent CO 2 

### buildup

 The overall recycle ratio is about 1.9

• A syngas purge is used as fuel in the combus tor side of the biomass dryer (only in HT scenario)

• Excess syngas is sent to a ga s turbine for power production

Areas 500, 600, 700

• Hydroprocessing and product distillation costs are estimated as a “black box” based on literature capital cost and operating cost information from Robinson et al.

o Literature yield data is used for estima ting the relative yields of gasoline and

diesel

Miscellaneous

• Combustion occurs with 20% excess oxygen

### 59 

# Appendix B. Detailed Costs

Cost Summaries

Figure B-1. Economic analysis summary for HT scenario Product Value ($/gal) 

### $4.26

Total Production at Operating Capacity (MM gal / year) 41.7 Product Yield (gal / Dry US Ton Feedstock) 61.0

Delivered Feedstock Cost $/Dry US Ton $75 Internal Rate of Return (After- Tax) 10% Equity Percent of Total Investment 100%

Capital Costs Operating Costs (cents/gal product)
Area 100: Pretreatment $22,700,000 7% Feedstock 123.0 28.9%
Area 200: Gasification $67,800,000 22% Steam 6.4 1.5% Area 300: Syngas Cleaning $33,500,000 11% Cooling Water 5.5 1.3% Area 400: Fuel Synthesis $49,400,000 16% Other Raw Materials 3.4 0.8% Area 500: Hydrocracking/Hydrotreating $33,000,000 11% Waste Disposal 1.3 0.3% Area 600: Power Generation $45,600,000 15% Hydroprocessing 10.6 2.5% Area 700: Air Separation $24,300,000 8% Fixed Costs 34.4 8.1%
Balance of Plant $33,100,000 11% Co- product credits - 13.3 - 3.1% Capital Depreciation 63.0 14.8% Total Installed Equipment Cost $309,400,000 Average Income Tax 52.4 12.3% Average Return on Investment 139.5 32.7%
Indirect Costs 129,700,000 (% of TPI) 21.4% Operating Costs ($/yr) Project Contingency 79,000,000 Feedstock $51,300,000
Steam $2,700,000 Total Project Investment (TPI) $605,900,000 Cooling Water $2,300,000
Other Raw Matl. Costs $1,400,000 Installed Equipment Cost per Annual Gallon $7.42 Waste Disposal $1,500,000 Total Project Investment per Annual Gallon $14.52 Hydroprocessing $4,400,000
Fixed Costs $14,300,000

Loan Rate N/A Co- product credits - $5,600,000 Term (years) N/A Capital Depreciation $26,300,000 Capital Charge Factor 0.176 Average Income Tax $21,900,000 Average Return on Investment $58,200,000 Gasifier Efficiency - HHV % 82.1 Gasifier Efficiency - LHV % 87.9 Total Plant Electricity Usage (KW) 22,065 Overall Plant Efficiency (incl. electricity) - HHV % 52.7 Electricity Produced Onsite (KW) 35,880 Overall Plant Efficiency - LHV % 53.0 Electricity Purchased from Grid (KW) 0 Electricity Sold to Grid (KW) 13,815 Availability (%) 85.0% Plant Hours per year 7446 Plant Electricity Use (KWh/gal product) 6.1 

### HT Biomass- to- Liquids Scenario Summary

2,000 Dry Metric Tonnes Biomass per Day

All Currency in 2007$ and Volume in Gallons Gasoline Equivalent (GGE) High Temperature Entrained Flow Gasifier, Sulfur Removal, Fischer- Tropsch Synthesis, Hydroprocessing, Combined Cycle Power

### 60 Figure B-2. Economic analysis summary for LT scenario Product Value ($/gal) 

### $4.83

Total Production at Operating Capacity (MM gal / year) 32.3 Product Yield (gal / Dry US Ton Feedstock) 47.2

Delivered Feedstock Cost $/Dry US Ton $75 Internal Rate of Return (After- Tax) 10% Equity Percent of Total Investment 100%

Capital Costs Operating Costs (cents/gal product)

Area 100: Pretreatment $22,700,000 9% Feedstock 158.9 32.9%

Area 200: Gasification $28,200,000 11% Steam 10.9 2.2% Area 300: Syngas Cleaning $29,300,000 12% Cooling Water 7.8 1.6% Area 400: Fuel Synthesis $58,700,000 23% Other Raw Materials 4.1 0.8% Area 500: Hydrocracking/Hydrotreating $29,500,000 12% Waste Disposal 1.5 0.3%

Area 600: Power Generation $38,900,000 15% Hydroprocessing 9.4 2.0% Area 700: Air Separation $19,500,000 8% Fixed Costs 38.4 8.0%

Balance of Plant $27,200,000 11% Co- product credits

- 20.4 -4.2% Capital Depreciation 67.2 13.9% Total Installed Equipment Cost $253,900,000 Average Income Tax 55.9 11.6% Average Return on Investment 149.5 31.0%

Indirect Costs 107,200,000 (% of TPI) 21.5% Operating Costs ($/yr) Project Contingency 65,000,000 Feedstock $51,300,000

Steam $3,500,000 Total Project Investment (TPI) $498,300,000 Cooling Water $3,500,000

Other Raw Matl. Costs $1,300,000 Installed Equipment Cost per Annual Gallon $7.86 Waste Disposal $1,500,000 Total Project Investment per Annual Gallon $15.43 Hydroprocessing $3,000,000

Fixed Costs $12,400,000

Loan Rate N/A Co- product credits

- $6,600,000 Term (years) N/A Capital Depreciation $21,700,000 Capital Charge Factor 0.177 Average Income Tax $18,000,000 Average Return on Investment $48,300,000 Gasifier Efficiency - HHV % 64.3 Gasifier Efficiency - LHV % 68.8 Total Plant Electricity Usage (KW) 15,044 Overall Plant Efficiency - HHV % 43.0 Electricity Produced Onsite (KW) 31,420 Overall Plant Efficiency - LHV % 43.3 Electricity Purchased from Grid (KW) 0 Electricity Sold to Grid (KW) 16,376 Availability (%) 85.0% Plant Hours per year 7446 Plant Electricity Use (KWh/gal product) 5.4 

### LT Biomass- to- Liquids Process Engineering Analysis

2,000 Dry Metric Tonnes Biomass per Day

All Currency in 2007$ and Volume in Gallons Gasoline Equivalent (GGE) Low Temperature Fluidized Gasifier, Sulfur Removal, Fischer- Tropsch Synthesis, Hydroprocessing, Combined Cycle Power61Detailed Equipment ListsHigh-Temperature Scenario Equipment ListsTable B-1. Detailed Equipment List for Areas 100 and 200 of HT ScenarioEquipment NumberNumberRequiredNumberSpares Equipment NameOriginal Equip Cost(per unit) Base YearTotal Original Equip Cost(Req'd & Spare) in BaseYearScaled UninstalledCost in 2007$ Installed Cost Base Year Installed Cost in 2007$ Cost SourceA100.CONV1 2 Bale Transport Conveyor $400,000 2000$800,000 $1,066,531 $1,296,000$1,727,781Aden et al. 2002A100.CONV2 2 Bale Unwrapping Conveyor $150,000 2000$300,000 $399,949 $357,000$475,940Aden et al. 2002A100.CONV3 1 Belt Press Discharge Conveyor $50,000 2000$50,000 $66,658 $94,500$125,984Aden et al. 2002A100.SCALE 2 Truck Scales $34,000 2000$68,000 $90,655 $167,960$223,918Aden et al. 2002A100.FORK1 4 1 Truck Unloading Forklift $18,000 2000$90,000 $119,985 $90,000$119,985Aden et al. 2002A100.FORK2 4 Bale Moving Forklift $18,000 2000$72,000 $95,988 $72,000$95,988Aden et al. 2002A100.SLAB 1 Concrete Feedstock-Storage Slab $450,655 2000$450,655 $600,797 $991,441$1,321,754Aden et al. 2002A100.MAGSEP 1 Magnetic Separator $13,863 1998$13,863 $18,700 $18,022$24,310Aden et al. 2002A100.A100CHOP.CHGRIN01 4 Chopper $105,100 2007 $420,400 $420,400 $1,105,258$1,105,258Aspen IcarusA100.A100CHOP.CHMIX01 1 Chopper Conveyor $61,400 2007 $61,400 $61,400 $185,428$185,428Aspen IcarusA100.A100CHOP.CHSEP01 1 Chopper Screen with Recycle Conveyor $20,800 2007 $20,800 $20,800 $62,816$62,816Aspen IcarusA100.A100DRY.DRDRY01 10 Dryer $633,700 2007 $6,337,000 $6,337,000 $15,201,647$15,201,647Aspen IcarusA100.A100GRIN.GRGRIN01 4 Grinder $167,100 2007 $668,400 $668,400 $1,757,266$1,757,266Aspen IcarusA100.A100GRIN.GRMIX01 1 Grinder Conveyor $61,400 2007 $61,400 $61,400 $185,428$185,428Aspen IcarusA100.A100GRIN.GRSEP01 1 Grinder Screen with Recycle Conveyor $20,800 2007 $20,800 $20,800 $62,816$62,816Aspen IcarusA100Subtotal$9,434,718 $10,049,464 $21,647,582 $22,676,317A200.A200COMB.CBREAC01 1 Combustor - Steam Boiler $1,450,500 2007 $1,450,500 $1,450,500 $4,380,510$4,380,510Aspen IcarusA200.A200SLAG.SLREAC01 1 Entrained Flow, Slagging Gasifier $23,234,043 2006 $23,234,043 $24,433,879 $54,600,000$57,419,616Reed et al. 2007A200.A200SLAG.SLSEP01 1 Slag collector/separator $35,100 2007 $35,100 $35,100 $106,002$106,002Aspen IcarusA200.A200SLAG.SLSEP03 3 Direct Quench Syngas Cooler $396,200 2007 $1,188,600 $1,188,600 $3,589,572$3,589,572Aspen IcarusA200.GSHOP01 1 Biomass Receiving Hopper $151,400 2007 $297,900 $297,900 $899,658$899,658Aspen IcarusA200.GSTANK01 1 Lockhopper $229,100 2007 $229,100 $229,100 $691,882$691,882Aspen IcarusA200.GSTANK02 1 Biomass Feeding Bin $228,900 2007 $228,900 $228,900 $691,278$691,278Aspen IcarusA200Subtotal$26,664,143 $27,863,979 $64,958,902 $67,778,518Process Evaluatorsoftware62Table B-2. Detailed Equipment List for Areas 300, 400, and 500 of HT ScenarioEquipment NumberNumberRequiredNumberSpares Equipment NameOriginal Equip Cost(per unit) Base YearTotal Original Equip Cost(Req'd & Spare) in BaseYearScaled UninstalledCost in 2007$ Installed Cost Base Year Installed Cost in 2007$ Cost SourceA300.A300AGR.AGRarea 1 High Pressure Amine System $6,949,800 2005 $6,949,800 $7,798,857 $20,988,396$23,552,549Phillips et al. 2007A300.A300SGS.SGCOMP01 2 Sour Water Gas Shift Steam Compressor $1,381,900 2007 $2,763,800 $2,763,800 $3,316,560$3,316,560Aspen IcarusA300.A300SGS.SGREAC01 1 Sour Water Gas Shift Reactor $66,600 2007 $66,600 $66,600 $201,132$201,132Aspen IcarusA300.A300SUL.SUCOL01 1 LO-CAT Absorber $23,800 2007 $23,800 $23,800 $71,876$71,876Aspen IcarusA300.A300SUL.SUREAC01 1 LO-CAT Oxidizer Vessel $1,000,000 2007 $1,000,000 $1,000,000 $3,020,000$3,020,000Phillips et al. 2007A300.A300SUL.SUSEP01 1 Sulfur Separator $15,900 2007 $15,900 $15,900 $48,018$48,018Aspen IcarusA300.CLCOMP01 2 Carbon Dioxide Compressor $1,181,200 2007 $2,362,400 $2,362,400 $2,834,880$2,834,880Aspen IcarusA300.CLDRUM01 1 Liquid Collection Tank $29,600 2007 $29,600 $29,600 $89,392$89,392Aspen IcarusA300.CLHEAT03 1 Direct Quench Syngas Cooler $91,500 2007 $91,500 $91,500 $276,330$276,330Aspen IcarusA300.CLMIX01 1 Venturi Scrubber $27,100 2007 $27,100 $27,100 $81,842$81,842Aspen IcarusA300Subtotal$13,330,500 $14,179,557 $30,928,426 $33,492,579A400.FSCOMP01 2 Booster Syngas Compressor $1,007,100 2007 $2,014,200 $2,014,200 $2,417,040$2,417,040Asen IcarusA400.FSCOMP02 1 Recycle Syngas Booster Compressor $748,400 2007 $748,400 $748,400 $898,080$898,080Asen IcarusA400.FSCOMP03 1 PSA Booster Compressor $1,461,700 2007 $1,461,700 $1,461,700 $1,754,040$1,754,040Asen IcarusA400.FSHEAT01 1 Syngas Heater $73,400 2007 $73,400 $73,400 $221,668$221,668Asen IcarusA400.FSHEAT03 1 Syngas Cooler $137,400 2007 $137,400 $137,400 $414,948$414,948Asen IcarusA400.FSHEAT04 1 Recycle Syngas Pre-heater $21,500 2007 $21,500 $21,500 $64,930$64,930Asen IcarusA400.FSREAC01 1 Fischer-Tropsch Reactor $8,888,889 2003 $8,888,889 $11,617,468 $32,000,000$41,822,886Larson et al. 2005A400.FSSEP01 2 ZnO Sulfur Removal Beds $61,000 2007 $122,000 $122,000 $368,440$368,440Asen IcarusA400.FSSEP02 12 Pressure Swing Absorption Unit $33,300 2007 $399,600 $399,600 $1,206,792$1,206,792Asen IcarusA400.FSSEP03 1 FT knock-out Column $39,600 2007 $39,600 $39,600 $119,592$119,592Asen IcarusA400.FSSEP04 1 Water Separator $47,900 2007 $47,900 $47,900 $144,658$144,658Asen IcarusA400Subtotal$13,954,589 $16,683,168 $39,610,188 $49,433,074A500.HYREAC01 1 Hydroprocessing Unit $9,377,483 2007 $9,377,483 $9,377,483 $28,320,000$28,320,000Robinson & Dolbear 2007A500.HYTANK01 1 Diesel 30-day Storage Tank $1,167,600 2007 $1,167,600 $1,167,600 $3,526,152$3,526,152Aspen IcarusA500.HYTANK02 1 Gasoline 30-day Storage Tank $371,900 2007 $371,900 $371,900 $1,123,138$1,123,138Aspen IcarusA500Subtotal$10,916,983 $10,916,983 $32,969,290 $32,969,29063Table B-3. Detailed Equipment List for Areas 600 and 700 of HT ScenarioEquipment NumberNumberRequiredNumberSpares Equipment NameOriginal Equip Cost(per unit) Base YearTotal Original Equip Cost(Req'd & Spare) in BaseYearScaled UninstalledCost in 2007$ Installed Cost Base Year Installed Cost in 2007$ Cost SourceA600.COMBB 1 Combustion Turbine - Electric Generator $22,404,000 2007 $22,404,000 $22,404,000 $26,884,800$26,884,800Aspen IcarusA600.CWPUMP 1 1 Cooling Water Pump $5,900 2007 $11,800 $11,800 $35,636$35,636Aspen IcarusA600.ECON1_HRSG 1 Heat Recovery Steam Generator $202,200 2007 $202,200 $202,200 $610,644$610,644Aspen IcarusA600.HPPUMP 1 1 High Pressure Steam Pump $266,700 2007 $533,400 $533,400 $1,610,868$1,610,868Aspen IcarusA600.HPSEP 1 High Pressure Steam/Water Separation $107,400 2007 $107,400 $107,400 $324,348$324,348Aspen IcarusA600.LPEXP_ELECGEN 1 Combined Steam Turbine - Electric Gen. $4,709,600 2007 $4,709,600 $4,709,600 $5,651,520$5,651,520Aspen IcarusA600.LPSEP 1 Low Pressure Water/Steam Separation $108,800 2007 $108,800 $108,800 $328,576$328,576Aspen IcarusA600.O2COMP 1 Air Compressor $8,431,900 2007 $8,431,900 $8,431,900 $10,118,280$10,118,280Aspen IcarusA600Subtotal$36,509,100 $36,509,100 $45,564,672 $45,564,672A700.COMP1 2 Air Compressor $3,346,500 2007 $6,693,000 $6,693,000 $8,031,600$8,031,600Aspen IcarusA700.COOLER 1 Air Cooler $27,200 2007 $27,200 $27,200 $82,144$82,144Aspen IcarusA700.GOXCLR-1 1 Oxygen Compressor Cooler $23,300 2007 $23,300 $23,300 $70,366$70,366Aspen IcarusA700.GOXCLR-2 1 Oxygen Compressor Cooler $23,000 2007 $23,000 $23,000 $69,460$69,460Aspen IcarusA700.GOXCMP-1 2 Oxygen Compressor $1,489,600 2007 $2,979,200 $2,979,200 $3,575,040$3,575,040Aspen IcarusA700.HIGH-P.cond 1 High Pressure Column Condenser $20,300 2007 $20,300 $20,300 $61,306$61,306Aspen IcarusA700.HIGH-P.cond acc 1 High Pressure Column Condenser Accumulator $40,500 2007 $40,500 $40,500 $122,310$122,310Aspen IcarusA700.HIGH-P.reflux pump 1 1 High Pressure Column Reflux Pump $14,300 2007 $28,600 $28,600 $86,372$86,372Aspen IcarusA700.HIGH-P.tower 1 High Pressure Column Tower $314,300 2007 $314,300 $314,300 $949,186$949,186Aspen IcarusA700.INTRC1 1 Air Compressor Intercooler $338,300 2007 $338,300 $338,300 $1,021,666$1,021,666Aspen IcarusA700.INTRC2 1 Air Compressor Intercooler $304,500 2007 $304,500 $304,500 $919,590$919,590Aspen IcarusA700.INTRC3 1 Air Compressor Intercooler $222,500 2007 $222,500 $222,500 $671,950$671,950Aspen IcarusA700.LOW-P.reb 1 Low Pressure Column Reboiler $19,600 2007 $19,600 $19,600 $59,192$59,192Aspen IcarusA700.LOW-P.tower 1 Low Pressure Column Tower $2,581,600 2007 $2,581,600 $2,581,600 $7,796,432$7,796,432Aspen IcarusA700.TSA 1 Water Knock-out Drum $35,900 2007 $35,900 $35,900 $108,418$108,418Aspen IcarusA700.TURB-1 2 Gas Expander $86,100 2007 $172,200 $172,200 $520,044$520,044Aspen IcarusA700.WK01 1 Water Knock-out Drum $57,700 2007 $57,700 $57,700 $174,254$174,254Aspen IcarusA700Subtotal$13,881,700 $13,881,700 $24,319,330 $24,319,330$124,691,733 $130,083,951 $259,998,390 $276,233,779$139,654,741 $145,694,026 $291,198,196 $309,381,833TotalTotal (with BOP)64Low-Temperature Scenario Equipment ListsTable B-4. Detailed Equipment List for Areas 100 and 200 of LT ScenarioEquipment NumberNumberRequiredNumberSpares Equipment NameOriginal Equip Cost(per unit) in Base Year Base YearTotal Original Equip Cost(Req'd & Spare) in BaseYearScaled UninstalledCost in 2007$ Installed Cost Base Year Installed Cost in 2007$ Cost SourceA100.CONV1 2 Bale Transport Conveyor $400,000 2000$800,000 $1,066,531 $1,296,000$1,727,781Aden et al. 2002A100.CONV2 2 Bale Unwrapping Conveyor $150,000 2000$300,000 $399,949 $357,000$475,940Aden et al. 2002A100.CONV3 1 Belt Press Discharge Conveyor $50,000 2000$50,000 $66,658 $94,500$125,984Aden et al. 2002A100.SCALE 2 Truck Scales $34,000 2000$68,000 $90,655 $167,960$223,918Aden et al. 2002A100.FORK1 4 1 Truck Unloading Forklift $18,000 2000$90,000 $119,985 $90,000$119,985Aden et al. 2002A100.FORK2 4 Bale Moving Forklift $18,000 2000$72,000 $95,988 $72,000$95,988Aden et al. 2002A100.SLAB 1 Concrete Feedstock-Storage Slab $450,655 2000$450,655 $600,797 $991,441$1,321,754Aden et al. 2002A100.MAGSEP 1 Magnetic Separator $13,863 1998$13,863 $18,700 $18,022$24,310Aden et al. 2002A100.A100CHOP.CHGRIN01 4 Chopper $105,100 2007 $420,400 $420,400 $1,105,258$1,105,258Aspen IcarusA100.A100CHOP.CHMIX01 1 Chopper Conveyor $61,400 2007 $61,400 $61,400 $185,428$185,428Aspen IcarusA100.A100CHOP.CHSEP01 1 Chopper Screen with Recycle Conveyor $20,800 2007 $20,800 $20,800 $62,816$62,816Aspen IcarusA100.A100DRY.DRDRY01 10 Dryer $633,700 2007 $6,337,000 $6,337,000 $15,201,647$15,201,647Aspen IcarusA100.A100GRIN.GRGRIN01 4 Grinder $167,100 2007 $668,400 $668,400 $1,757,266$1,757,266Aspen IcarusA100.A100GRIN.GRMIX01 1 Grinder Conveyor $61,400 2007 $61,400 $61,400 $185,428$185,428Aspen IcarusA100.A100GRIN.GRSEP01 1 Grinder Screen with Recycle Conveyor $20,800 2007 $20,800 $20,800 $62,816$62,816Aspen IcarusA100Subtotal$9,434,718 $10,049,464 $21,647,582 $22,676,317A200.A200COMB.CBCYC01 3 Combustor Cyclone (medium efficiency) $35,400 2007 $106,200 $106,200 $320,724$320,724Aspen IcarusA200.A200COMB.CBCYC02 3 Combustor Cyclone (high efficiency) $6,700 2007 $20,100 $20,100 $60,702$60,702Aspen IcarusA200.A200COMB.CBMIX01 1 Ash Storage Vessel $142,800 2007 $142,800 $142,800 $431,256$431,256Aspen IcarusA200.A200COMB.CBREAC01 1 Combustor - Steam Boiler $1,450,500 2007 $1,450,500 $1,450,500 $4,380,510$4,380,510Aspen IcarusA200.A200CYC.CYCYC01 2 1st train, medium efficiency cyclone $20,300 2007 $40,600 $40,600 $122,612$122,612Aspen IcarusA200.A200CYC.CYCYC02 4 1st train, high efficiency cyclone $24,900 2007 $99,600 $99,600 $300,792$300,792Aspen IcarusA200.A200CYC.CYMIX02 1 Char Collector and conveyor $84,400 2007 $84,400 $84,400 $254,888$254,888Aspen IcarusA200.GSREAC01 7 Fluidized Bed Gasifier (Pressurized) $1,096,170 2003 $7,673,191 $10,028,594 $14,843,424$19,399,838Larson et al. 2005A200.GSTANK01 7 Biomass Receiving Hopper $71,700 2007 $501,900 $501,900 $1,247,712$1,247,712Aspen IcarusA200.GSTANK02 7 Lockhopper $47,700 2007 $333,900 $333,900 $830,068$830,068Aspen IcarusA200.GSTANK03 7 Biomass Feeding Bin $47,700 2007 $333,900 $333,900 $830,068$830,068Aspen IcarusA200Subtotal$10,787,091 $13,142,494$23,622,756$28,179,17065Table B-5. Detailed Equipment List for Areas 300, 400, and 500 of LT ScenarioEquipment NumberNumberRequiredNumberSpares Equipment NameOriginal Equip Cost(per unit) in Base Year Base YearTotal Original Equip Cost(Req'd & Spare) in BaseYearScaled UninstalledCost in 2007$ Installed Cost Base Year Installed Cost in 2007$ Cost SourceA300.A300AGR.AGRarea 1 High Pressure Amine System $6,050,000 2005 $6,050,000 $6,789,129 $18,271,000$20,503,168Phillips et al.2007A300.A300SUL.SUCOL011LO-CAT Absorber$16,200 2007$16,200$16,200$48,924$48,924Aspen IcarusA300.A300SUL.SUREAC011LO-CAT Oxidizer Vessel$1,000,000 2007$1,000,000 $1,000,000$3,020,000$3,020,000Phillips et al.2007A300.A300SUL.SUSEP011Sulfur Separator$16,200 2007$16,200$16,200$48,924$48,924Aspen IcarusA300.CLCMP012Carbon Dioxide Compressor$1,176,900 2007$2,353,800 $2,353,800$2,824,560$2,824,560Aspen IcarusA300.CLHEAT012Direct Quench Recycle Cooling$188,800 2007$377,600 $377,600$1,140,352$1,140,352Aspen IcarusA300.CLHEAT021Venturi Recycle Cooling$91,500 2007$91,500 $91,500$276,330$276,330Aspen IcarusA300.CLMIX011Venturi Scrubber$26,800 2007$26,800$26,800$80,936$80,936Aspen IcarusA300.CLSEP032Direct Quench Syngas Cooler$188,800 2007$377,600$377,600$1,140,352$1,140,352Aspen IcarusA300.CLSEP041Venturi Liquid Collection Tank$74,500 2007$74,500$74,500$224,990$224,990Aspen IcarusA300Subtotal$10,384,200 $11,123,329$27,076,368$29,308,536A400.A400COND.CDHEAT011Syngas Heater$60,500 2007$60,500$60,500$182,710$182,710Aspen IcarusA400.A400COND.CDHEAT021Syngas Pre-heater Furnace$1,949,500 2007$1,949,500 $1,949,500$5,887,490$5,887,490Aspen IcarusA400.A400COND.CDHEAT031Reformed Syngas Waste Heat Boiler$396,600 2007$396,600$396,600$1,197,732$1,197,732Aspen IcarusA400.A400COND.CDHEAT041Syngas Cooler #2$41,200 2007$41,200$41,200$124,424$124,424Aspen IcarusA400.A400COND.CDREAC011Steam Methane Reformer$1,650,800 2007$1,650,800 $1,650,800$4,985,416$4,985,416Aspen IcarusA400.A400COND.CDREAC021Water Gas Shift Reactor$136,600 2007$136,600$136,600$412,532$412,532Aspen IcarusA400.A400COND.CDSEP012ZnO Sulfur Removal Beds$46,400 2007$92,800$92,800$280,256$280,256Aspen IcarusA400.FSCOMP012Booster Syngas Compressor$921,600 2007$1,843,200 $1,843,200$2,211,840$2,211,840Aspen IcarusA400.FSCOMP021Recycle Syngas Booster Compressor$725,400 2007$725,400$725,400$870,480$870,480Aspen IcarusA400.FSCOMP031PSA Booster Compressor$1,482,100 2007$1,482,100 $1,482,100$1,778,520$1,778,520Aspen IcarusA400.FSDRUM011PSA Knock-out$1,482,100 2007$1,482,100 $1,482,100$4,475,942$4,475,942Aspen IcarusA400.FSHEAT031Syngas Cooler$165,200 2007$165,200$165,200$498,904$498,904Aspen IcarusA400.FSHEAT041Recycle Syngas Pre-heater$24,300 2007$24,300$24,300$73,386$73,386Aspen IcarusA400.FSREAC011Fischer-Tropsch Reactor$7,303,889 2003$7,303,889 $9,545,928$26,294,000$34,365,342Larson et al. 2005A400.FSSEP0212Pressure Swing Absorption Unit$30,500 2007$366,000$366,000$1,105,320$1,105,320Aspen IcarusA400.FSSEP031FT knock-out Column$72,100 2007$72,100$72,100$217,742$217,742Aspen IcarusA400.FSSEP041Water Separator$39,200 2007$39,200$39,200$118,384$118,384Aspen IcarusA400Subtotal$17,792,289 $20,034,328$50,596,694$58,668,036A500.HYREAC011Hydrocracking/Hydrotreating Unit$7,927,152 2007$7,927,152 $7,927,152$23,940,000$23,940,000Robinson & Dolbear 2007A500.HYTANK011Gasoline 30-day Storage Tank$646,300 2007$646,300$646,300$1,951,826$1,951,826Aspen IcarusA500.HYTANK021Diesel 30-day Storage Tank$1,200,700 2007$1,200,700 $1,200,700$3,626,114$3,626,114Aspen IcarusA500Subtotal$9,774,152 $9,774,152$29,517,940$29,517,94066Table B-6. Detailed Equipment List for Areas 600 and 700 of LT ScenarioEquipment NumberNumberRequiredNumberSpares Equipment NameOriginal Equip Cost(per unit) in Base Year Base YearTotal Original Equip Cost(Req'd & Spare) in BaseYearScaled UninstalledCost in 2007$ Installed Cost Base Year Installed Cost in 2007$ Cost SourceA600.COMBB 1 Combustion Turbine - Electric Generator $18,607,700 2007 $18,607,700 $18,607,700 $22,329,240$22,329,240Aspen IcarusA600.CWPUMP 1 1 Cooling Water Pump $5,900 2007 $11,800 $11,800 $35,636$35,636Aspen IcarusA600.ECON1_HRSG 1 Heat Recovery Steam Generator $202,200 2007 $202,200 $202,200 $610,644$610,644Aspen IcarusA600.HPPUMP 1 1 High Pressure Steam Pump $266,700 2007 $533,400 $533,400 $1,610,868$1,610,868Aspen IcarusA600.HPSEP 1 High Pressure Steam/Water Separation $107,400 2007 $107,400 $107,400 $324,348$324,348Aspen IcarusA600.LPEXP_ELECGEN 1 Combined Steam Turbine - Electric Gen. $5,056,300 2007 $5,056,300 $5,056,300 $6,067,560$6,067,560Aspen IcarusA600.LPSEP 1 Low Pressure Water/Steam Separation $108,800 2007 $108,800 $108,800 $328,576$328,576Aspen IcarusA600.O2COMP 1 Air Compressor $6,331,200 2007 $6,331,200 $6,331,200 $7,597,440$7,597,440Aspen IcarusA600Subtotal$30,958,800 $30,958,800 $38,904,312 $38,904,312A700.COMP1 2 Air Compressor $3,119,600 2007 $6,239,200 $6,239,200 $7,487,040$7,487,040Aspen IcarusA700.COOLER 1 Air Cooler $24,300 2007 $24,300 $24,300 $73,386$73,386Aspen IcarusA700.GOXCLR-1 1 Oxygen Compressor Cooler $23,300 2007 $23,300 $23,300 $70,366$70,366Aspen IcarusA700.GOXCLR-2 1 Oxygen Compressor Cooler $23,000 2007 $23,000 $23,000 $69,460$69,460Aspen IcarusA700.GOXCMP-1 2 Oxygen Compressor $1,514,700 2007 $3,029,400 $3,029,400 $3,635,280$3,635,280Aspen IcarusA700.HIGH-P.cond 1 High Pressure Column Condenser $20,300 2007 $20,300 $20,300 $61,306$61,306Aspen IcarusA700.HIGH-P.cond acc 1 High Pressure Column Condenser Accumulator $36,300 2007 $36,300 $36,300 $109,626$109,626Aspen IcarusA700.HIGH-P.reflux pump 1 1 High Pressure Column Reflux Pump $14,300 2007 $28,600 $28,600 $34,320$34,320Aspen IcarusA700.HIGH-P.tower 1 High Pressure Column Tower $279,900 2007 $279,900 $279,900 $335,880$335,880Aspen IcarusA700.INTRC1 1 Air Compressor Intercooler $338,300 2007 $338,300 $338,300 $405,960$405,960Aspen IcarusA700.INTRC2 1 Air Compressor Intercooler $304,500 2007 $304,500 $304,500 $919,590$919,590Aspen IcarusA700.INTRC3 1 Air Compressor Intercooler $222,500 2007 $222,500 $222,500 $671,950$671,950Aspen IcarusA700.LOW-P.reb 2 Low Pressure Column Reboiler $19,600 2007 $39,200 $39,200 $118,384$118,384Aspen IcarusA700.LOW-P.tower 1 Low Pressure Column Tower $1,538,900 2007 $1,538,900 $1,538,900 $4,647,478$4,647,478Aspen IcarusA700.TSA 1 Water Knock-out Drum $30,100 2007 $30,100 $30,100 $90,902$90,902Aspen IcarusA700.TURB-1 2 Gas Expander $89,200 2007 $178,400 $178,400 $538,768$538,768Aspen IcarusA700.WK01 1 Water Knock-out Drum $64,800 2007 $64,800 $64,800 $195,696$195,696Aspen IcarusA700Subtotal$12,421,000 $12,421,000 $19,465,392 $19,465,392Total $101,552,251 $107,503,567 $210,831,043 $226,719,704Total (with BOP) $113,738,521 $120,403,995 $236,130,768 $253,926,06867Discounted Cash FlowHigh-Temperature ScenarioTable B-7. Discounted Cash Flow Sheet for Construction Period and Years 1-8 of HT ScenarioDCFROR WorksheetYear-2 -10 1 2 3 4 5 6 7 8Fixed Capital Investment$50,890,395 $316,115,651 $168,595,014Working Capital$79,028,913Loan Payment$0$0$0$0$0$0$0$0LoanInterestPayment$0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0LoanPrincipal$0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0GGE (Gallon of Gasoline Equiv.) Sales$133,364,635 $177,819,513 $177,819,513 $177,819,513 $177,819,513 $177,819,513 $177,819,513 $177,819,513Diesel Sales$0$0$0$0$0$0$0$0By-Product Credit$4,173,208 $5,564,277 $5,564,277 $5,564,277 $5,564,277 $5,564,277 $5,564,277 $5,564,277Plant Performance1.001.001.001.001.001.001.001.00Total Annual Sales$137,537,843 $183,383,791 $183,383,791 $183,383,791 $183,383,791 $183,383,791 $183,383,791 $183,383,791Annual Manufacturing CostRaw Materials$44,894,145 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594SWGS catalysts$114,621$0$0 $114,621$0$0 $114,621$0Steam reforming catalysts$0$0$0$0$0$0$0$0ZnO$424,410$0$0 $424,410$0$0 $424,410$0Pressure Swing Adsorption Packing$497,135$0$0 $497,135$0$0 $497,135$0FT catalysts$7,686,720$0$0 $7,686,720$0$0 $7,686,720$0Other Variable Costs$11,727,856 $13,403,264 $13,403,264 $13,403,264 $13,403,264 $13,403,264 $13,403,264 $13,403,264Fixed Operating Costs$14,345,785 $14,345,785 $14,345,785 $14,345,785 $14,345,785 $14,345,785 $14,345,785 $14,345,785Total Product Cost$79,690,672 $79,056,643 $79,056,643 $87,779,529 $79,056,643 $79,056,643 $87,779,529 $79,056,643Annual DepreciationGeneral PlantDDB$128,361,546 $91,686,818 $65,490,585 $46,778,989 $33,413,564 $23,866,831 $17,047,737SL$64,180,773 $53,483,977 $45,843,409 $40,931,615 $38,982,491 $38,982,491 $38,982,491Remaining Value$320,903,864 $229,217,046 $163,726,461 $116,947,472 $83,533,909 $59,667,078 $42,619,341Actual$128,361,546 $91,686,818 $65,490,585 $46,778,989 $38,982,491 $38,982,491 $38,982,491Steam PlantDDB$5,819,551 $5,383,084 $4,979,353 $4,605,902 $4,260,459 $3,940,925 $3,645,355 $3,371,954SL$3,879,700 $3,777,603 $3,688,410 $3,612,472 $3,550,382 $3,503,044 $3,471,767 $3,458,414Remaining Value$71,774,458 $66,391,374 $61,412,021 $56,806,119 $52,545,660 $48,604,736 $44,959,380 $41,587,427Actual$5,819,551 $5,383,084 $4,979,353 $4,605,902 $4,260,459 $3,940,925 $3,645,355 $3,458,414Net Revenue($76,333,925) $7,257,245 $33,857,210 $44,219,371 $61,084,198 $61,403,732 $52,976,416 $100,868,733Losses Forward($76,333,925) ($69,076,681) ($35,219,471)$0$0$0$0Taxable Income($76,333,925) ($69,076,681) ($35,219,471) $8,999,900 $61,084,198 $61,403,732 $52,976,416 $100,868,733Income Tax$0$0$0 $3,509,961 $23,822,837 $23,947,455 $20,660,802 $39,338,806Annual Cash Income$57,847,171 $104,327,147 $104,327,147 $92,094,301 $80,504,310 $80,379,692 $74,943,459 $64,988,341Discount Factor1.211.11 0.909090909 0.826446281 0.751314801 0.683013455 0.620921323 0.56447393 0.513158118 0.46650738Annual Present Value$645,181,377$52,588,337 $86,220,783 $78,382,530 $62,901,646 $49,986,843 $45,372,241 $38,457,845 $30,317,541Total Capital Investment + Interest$61,577,378 $347,727,216 $247,623,927Net Present Worth$068Table B-8. Discounted Cash Flow Sheet for Years 9-20 of HT ScenarioDCFROR WorksheetYear 9 10 11 12 13 14 15 16 17 18 19 20Fixed Capital InvestmentWorking Capital($79,028,913)Loan Payment $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0LoanInterestPayment$0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0LoanPrincipal$0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0GGE (Gallon of Gasoline Equiv.) Sales $177,819,513 $177,819,513 $177,819,513 $177,819,513 $177,819,513 $177,819,513 $177,819,513 $177,819,513 $177,819,513 $177,819,513 $177,819,513 $177,819,513DieselSales$0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0By-Product Credit$5,564,277 $5,564,277 $5,564,277 $5,564,277 $5,564,277 $5,564,277 $5,564,277 $5,564,277 $5,564,277 $5,564,277 $5,564,277 $5,564,277Plant Performance1.001.001.001.001.001.001.001.001.001.001.001.00Total Annual Sales$183,383,791 $183,383,791 $183,383,791 $183,383,791 $183,383,791 $183,383,791 $183,383,791 $183,383,791 $183,383,791 $183,383,791 $183,383,791 $183,383,791Annual Manufacturing CostRaw Materials$51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594SWGS catalysts$0 $114,621$0$0 $114,621$0$0 $114,621$0$0 $114,621$0Steam reforming catalysts$0$0$0$0$0$0$0$0$0$0$0$0ZnO$0 $424,410$0$0 $424,410$0$0 $424,410$0$0 $424,410$0Pressure Swing Adsorption Packing$0 $497,135$0$0 $497,135$0$0 $497,135$0$0 $497,135$0FT catalysts$0 $7,686,720$0$0 $7,686,720$0$0 $7,686,720$0$0 $7,686,720$0Other Variable Costs$13,403,264 $13,403,264 $13,403,264 $13,403,264 $13,403,264$13,403,264 $13,403,264 $13,403,264 $13,403,264 $13,403,264 $13,403,264 $13,403,264Fixed Operating Costs$14,345,785 $14,345,785 $14,345,785 $14,345,785 $14,345,785$14,345,785 $14,345,785 $14,345,785 $14,345,785 $14,345,785 $14,345,785 $14,345,785Total Product Cost$79,056,643 $87,779,529 $79,056,643 $79,056,643 $87,779,529 $79,056,643 $79,056,643$87,779,529 $79,056,643 $79,056,643 $87,779,529 $79,056,643Annual DepreciationGeneral PlantDDBSLRemaining ValueActualSteam PlantDDB$3,119,057 $2,885,128 $2,668,743 $2,468,587 $2,283,443 $2,112,185 $1,953,771 $1,807,238 $1,671,696 $1,546,318 $1,430,344 $1,323,069SL$3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414Remaining Value$38,468,370 $35,583,242 $32,914,499 $30,445,912 $28,162,468 $26,050,283 $24,096,512 $22,289,273 $20,617,578 $19,071,260 $17,640,915 $16,317,847Actual$3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414Net Revenue$100,868,733 $92,145,848 $100,868,733 $100,868,733 $92,145,848 $100,868,733 $100,868,733 $92,145,848 $100,868,733 $100,868,733 $92,145,848 $100,868,733LossesForward$0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0Taxable Income$100,868,733 $92,145,848 $100,868,733 $100,868,733 $92,145,848 $100,868,733 $100,868,733 $92,145,848 $100,868,733 $100,868,733 $92,145,848 $100,868,733Income Tax$39,338,806 $35,936,881 $39,338,806 $39,338,806 $35,936,881 $39,338,806 $39,338,806 $35,936,881 $39,338,806 $39,338,806 $35,936,881 $39,338,806Annual Cash Income$64,988,341 $59,667,381 $64,988,341 $64,988,341 $59,667,381 $64,988,341 $64,988,341 $59,667,381 $64,988,341 $64,988,341 $59,667,381 $64,988,341Discount Factor0.424097618 0.385543289 0.350493899 0.3186308180.28966438 0.263331254 0.239392049 0.217629136 0.197844669 0.17985879 0.163507991 0.148643628Annual Present Value$27,561,401 $23,004,358 $22,778,017 $20,707,288 $17,283,515 $17,113,461 $15,557,692 $12,985,361 $12,857,597 $11,688,724 $9,756,094 $9,660,103Total Capital Investment + Interest($11,747,144.32)Net Present Worth69Low-Temperature ScenarioTable B-9. Discounted Cash Flow Sheet for Construction Period and Years 1-8 of LT ScenarioDCFROR WorksheetYear -2 -1 0 1 2 3 4 5 6 7 8Fixed Capital Investment $41,888,460 $259,981,649 $138,656,880Working Capital $64,995,412Loan Payment$0 $0 $0 $0 $0 $0 $0 $0Loan Interest Payment $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0Loan Principal $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0GGE (Gallon of Gasoline Equiv.) Sales $117,025,289 $156,033,719 $156,033,719 $156,033,719 $156,033,719 $156,033,719 $156,033,719 $156,033,719Diesel Sales$0 $0 $0 $0 $0 $0 $0 $0By-Product Credit $4,945,498 $6,593,998 $6,593,998 $6,593,998 $6,593,998 $6,593,998 $6,593,998 $6,593,998Plant Performance1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00Total Annual Sales$121,970,788 $162,627,717 $162,627,717 $162,627,717 $162,627,717 $162,627,717 $162,627,717 $162,627,717Annual Manufacturing CostRaw Materials$44,894,145 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594WGS catalysts$104,732 $0 $0 $104,732 $0 $0 $104,732 $0Steam reforming catalysts$103,412 $0 $0 $103,412 $0 $0 $103,412 $0ZnO$424,410 $0 $0 $424,410 $0 $0 $424,410 $0Pressure Swing Adsorption Packing$497,135 $0 $0 $497,135 $0 $0 $497,135 $0FT catalysts$6,127,680 $0 $0 $6,127,680 $0 $0 $6,127,680 $0Other Variable Costs$11,238,097 $12,843,539 $12,843,539 $12,843,539 $12,843,539 $12,843,539 $12,843,539 $12,843,539Fixed Operating Costs$12,404,834 $12,404,834 $12,404,834 $12,404,834 $12,404,834 $12,404,834 $12,404,834 $12,404,834Total Product Cost$75,794,444 $76,555,967 $76,555,967 $83,813,336 $76,555,967 $76,555,967 $83,813,336 $76,555,967Annual DepreciationGeneral PlantDDB$104,833,121 $74,880,801 $53,486,286 $38,204,490 $27,288,922 $19,492,087 $13,922,919SL$52,416,561 $43,680,467 $37,440,400 $33,428,929 $31,837,075 $31,837,075 $31,837,075Remaining Value$262,082,803 $187,202,002 $133,715,716 $95,511,226 $68,222,304 $48,730,217 $34,807,298Actual$104,833,121 $74,880,801 $53,486,286 $38,204,490 $31,837,075 $31,837,075 $31,837,075Steam PlantDDB$4,979,012 $4,605,586 $4,260,167 $3,940,654 $3,645,105 $3,371,722 $3,118,843 $2,884,930SL$3,319,341 $3,231,990 $3,155,679 $3,090,709 $3,037,588 $2,997,087 $2,970,327 $2,958,903Remaining Value$61,407,813 $56,802,227 $52,542,060 $48,601,405 $44,956,300 $41,584,577 $38,465,734 $35,580,804Actual$4,979,012 $4,605,586 $4,260,167 $3,940,654 $3,645,105 $3,371,722 $3,118,843 $2,958,903Net Revenue($63,635,790) $6,585,363 $28,325,297 $36,669,236 $50,589,570 $50,862,953 $43,858,462 $83,112,848Losses Forward($63,635,790) ($57,050,426) ($28,725,129) $0 $0 $0 $0Taxable Income($63,635,790) ($57,050,426) ($28,725,129) $7,944,107 $50,589,570 $50,862,953 $43,858,462 $83,112,848Income Tax$0 $0 $0 $3,098,202 $19,729,932 $19,836,551 $17,104,800 $32,414,011Annual Cash Income$46,176,343 $86,071,750 $86,071,750 $75,716,179 $66,341,818 $66,235,199 $61,709,581 $53,657,740Discount Factor 1.21 1.1 1 0.909090909 0.826446281 0.751314801 0.683013455 0.620921323 0.56447393 0.513158118 0.46650738Annual Present Value $530,655,988 $41,978,494 $71,133,678 $64,666,980 $51,715,169 $41,193,049 $37,388,043 $31,666,772 $25,031,732Total Capital Investment + Interest $50,685,036 $285,979,814 $203,652,292Net Present Worth$070Table B-10. Discounted Cash Flow Sheet for Years 9-20 of LT ScenarioDCFROR WorksheetYear 9 10 11 12 13 14 15 16 17 18 19 20Fixed Capital InvestmentWorking Capital($64,995,412)Loan Payment $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0Loan Interest Payment $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0Loan Principal $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0GGE (Gallon of Gasoline Equiv.) Sales $156,033,719 $156,033,719 $156,033,719 $156,033,719 $156,033,719 $156,033,719 $156,033,719 $156,033,719 $156,033,719 $156,033,719 $156,033,719 $156,033,719Diesel Sales $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0By-Product Credit $6,593,998 $6,593,998 $6,593,998 $6,593,998 $6,593,998 $6,593,998 $6,593,998 $6,593,998 $6,593,998 $6,593,998 $6,593,998 $6,593,998Plant Performance 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00Total Annual Sales $162,627,717 $162,627,717 $162,627,717 $162,627,717 $162,627,717 $162,627,717 $162,627,717 $162,627,717 $162,627,717 $162,627,717 $162,627,717 $162,627,717Annual Manufacturing CostRaw Materials $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594WGS catalysts $0 $104,732 $0 $0 $104,732 $0 $0 $104,732 $0 $0 $104,732 $0Steam reforming catalysts $0 $103,412 $0 $0 $103,412 $0 $0 $103,412 $0 $0 $103,412 $0ZnO $0 $424,410 $0 $0 $424,410 $0 $0 $424,410 $0 $0 $424,410 $0Pressure Swing Adsorption Packing $0 $497,135 $0 $0 $497,135 $0 $0 $497,135 $0 $0 $497,135 $0FT catalysts $0 $6,127,680 $0 $0 $6,127,680 $0 $0 $6,127,680 $0 $0 $6,127,680 $0Other Variable Costs $12,843,539 $12,843,539 $12,843,539 $12,843,539 $12,843,539 $12,843,539 $12,843,539 $12,843,539 $12,843,539 $12,843,539 $12,843,539 $12,843,539Fixed Operating Costs $12,404,834 $12,404,834 $12,404,834 $12,404,834 $12,404,834 $12,404,834 $12,404,834 $12,404,834 $12,404,834 $12,404,834 $12,404,834 $12,404,834Total Product Cost $76,555,967 $83,813,336 $76,555,967 $76,555,967 $83,813,336 $76,555,967 $76,555,967 $83,813,336 $76,555,967 $76,555,967 $83,813,336 $76,555,967Annual DepreciationGeneral PlantDDBSLRemaining ValueActualSteam PlantDDB $2,668,560 $2,468,418 $2,283,287 $2,112,040 $1,953,637 $1,807,115 $1,671,581 $1,546,212 $1,430,246 $1,322,978 $1,223,755 $1,131,973SL $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903Remaining Value $32,912,244 $30,443,825 $28,160,538 $26,048,498 $24,094,861 $22,287,746 $20,616,165 $19,069,953 $17,639,706 $16,316,728 $15,092,974 $13,961,001Actual $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903Net Revenue $83,112,848 $75,855,478 $83,112,848 $83,112,848 $75,855,478 $83,112,848 $83,112,848 $75,855,478 $83,112,848 $83,112,848 $75,855,478 $83,112,848Losses Forward $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0Taxable Income $83,112,848 $75,855,478 $83,112,848 $83,112,848 $75,855,478 $83,112,848 $83,112,848 $75,855,478 $83,112,848 $83,112,848 $75,855,478 $83,112,848Income Tax $32,414,011 $29,583,637 $32,414,011 $32,414,011 $29,583,637 $32,414,011 $32,414,011 $29,583,637 $32,414,011 $32,414,011 $29,583,637 $32,414,011Annual Cash Income $53,657,740 $49,230,744 $53,657,740 $53,657,740 $49,230,744 $53,657,740 $53,657,740 $49,230,744 $53,657,740 $53,657,740 $49,230,744 $53,657,740Discount Factor 0.424097618 0.385543289 0.350493899 0.318630818 0.28966438 0.263331254 0.239392049 0.217629136 0.197844669 0.17985879 0.163507991 0.148643628Annual Present Value $22,756,120 $18,980,583 $18,806,710 $17,097,009 $14,260,393 $14,129,760 $12,845,236 $10,714,044 $10,615,898 $9,650,816 $8,049,620 $7,975,881Total Capital Investment + Interest($9,661,153.89)Net Present Worth

### 71 

# Appendix C. Scenario Modeling Details

Property Method The model setup includes a particle size distribu tion in order to better estimate the solids simulation in the grinding and cyclone operati ons. It operates globally with the Redlich-KwongSoave with Boston-Mathias modification (RKS -BM) property method, which is recommended for medium-temperature refining and gas pr ocessing operations incl uding combustion and gasification. During acid-gas absorption and st ripping, another property method, ELECNRTL, is

used for more accurate simulation. For solids hand ling, such as in the pr etreatment area and cyclones, the SOLIDS property method is used.

Stream/Block Nomenclature

All streams and blocks within the model follow a specific al phanumeric notation with the purpose of clarity and consistenc y across scenarios and across platforms. Each area within the model (e.g., Area 200 gasification) ha s a two letter abbreviation (e .g., gasification is GS). These abbreviations are used for naming streams as we ll as blocks. In addition to serving purposes mentioned above, the notation is descriptive (e.g., th e notation REAC desc ribes a block as a reactor). Another example is SGAS , which describes a stream that contains syngas. Aspen Plus

software limits block and stream names to be eight characters.

Figure C-1 shows the pattern of notation for a syngas stream in the gasification area.

Area Number Description 

# G S 0 1 S G A S

Figure C-1. Stream nomenclature used in model



### Similarly, the notation for the first reactor block in the gasification area is shown in Figure C-2.

Area Description Number 

# G S R E A C 0 1

Figure C-2. Block nomenclature used in model



### Table C-1 contains the abbrevia tions for areas, unit operation bl ock descriptions, and stream descriptions.

### 72 Table C-1. Detailed Description of Stream and Block Nomenclature

Area Description Name Block Name Stream Name

Plant All Areas PL Reactor REAC Biomass BMAS

A100 Pretreatment PR Mixer MIX Steam STM

A100CHOP Chopping CH Heat Mixer QMX Flue Gas FLUE

A100DRY Drying DR Work Mixer WMX Syngas SGAS

A100GRIN Grinding GR Splitter SPL Ash ASH

A200 Gasification GS,SL Separator SEP Carbon Dioxide CO 2 A200CYC Cyclones CY Cyclones CYC Air AIR
A200COMB Combustion CB Flash Drum DRUM Hydrogen HYD

A300 Syngas Cleaning CL Column COL FT Products FT
A300AGR Acid Gas Removal AG Distillation DIST Water WAT
A300SUL Sulfur Recovery SU Grinder GRIN Oxygen OX

A400 Fuel Synthesis FS Dryer DRY Sulfur SUL

A400COND Syngas Conditioning CD Heater HEAT Fuel FUEL
A400MTG Methanol to Gasoline MG Heat Exchanger HX Tar TAR

A500 Hydrocracking HY Tank/Hopper TANK Char CHAR

A500 Fuel Separation SE Pump PMP Acid Gas AG

A600 Power Generation PG Compressor COMP Lean MEA Soln. MEAL

A700 Air Separation Unit Turbine TURB Rich MEA Soln. MEAR Light Gases LGAS Nitrogen NTGN



### A special notation is used for heat and work streams. In the case that the first reactor in the gasification area includes a heat stream leaving the unit, it follows the nomenclature shown in Figure C-3.

Q or W Area Block Description Number 

# Q - G S R E A 1

Figure C-3. Heat and work stream nomenclature used in model



### The Q or W sets the stream apart as a heat or work stream. The block de scription is limited to

three characters and the number is limited to one character.

### 73 Aspen Plus Calculator Block Descriptions

High-Temperature Scenario

AIRCOMB

This block calculates the nitrogen that accompan ies the oxygen in the air inlet for the combustion of unconverted syngas. Molar nitrogen flow (in kmol/h) is calculated as follows:

ܯ

ሶ ேଶ 

### ൌ ൬ 0.79

0.21 ൰ · ܯ ሶ ைଶ 

### (eqn. 20)

where ܯ ሶ ைଶ 

### is molar flow of oxygen in kmol/h.

AMINE

This block calculates the mola r flow of monoethanolamine (MEA ) needed for the required acid gas removal (CO 2 

### and H 2 

### S) arriving from syngas quench a nd FT unconverted syngas recycle stream. The MEA is able to capture 0.35 moles acid gas per mole MEA. Additionally, the MEA is diluted as expl ained in DILUTH2O.

Molar MEA flow (in kmo l/h) is calculated by

ܯ

ሶ ொ஺ 

### ൌ ሺܯ ሶ ஼ைଶ,௦௬௡ 

### ൅ ܯ ሶ ஼ைଶ,௥௘௖ 

### ൅ ܯ ሶ ுଶௌ,௦௬௡ 

### ሻ/0.35 

### (eqn. 21)

where ܯ ሶ ஼ைଶ,௦௬௡ 

### is molar flow of CO 2 

### from the syngas after the syngas quench, ܯ ሶ ஼ைଶ,௥௘௖ 

### is the

molar flow of CO 2 

### from the unconverted syngas recycle after th e FT synthesis, and ܯ ሶ ுଶௌ,௦௬௡ 

### is

the molar flow rate of H 2 

### S from the syngas quench.

Because the MEA solution in the amine absorption un it is to be 20 wt% con centrated with water,

the flow of water must be calculated. Mole flow of water is calculated as

ܯ

ሶ ுଶை 

### ൌ ܯ ሶ ஺ொ 

### ܹܯ כ ஺ொ 

### /0.20

ܹܯ ுଶை 

### (eqn. 22)

BIOELEM

Because the high-temperature gasifier is mode led at equilibrium, the simulation software requires that all components in th e input are located in the conven tional stream. Therefore, this block splits the biomass into the following compounds based on its ultimate analysis: carbon, hydrogen, oxygen, sulfur, nitrogen, and ash. Water in the biomass is not affected because it is already a conventional component. Biomass in the exit stream is set to zero.

### 74 FTDISTR

This calculator block calculates an alpha chain growth parameter using the equation by Song et al. (2004) for cobalt catalyst. Inlet and outlet streams are defined and calculated. FT products include paraffins from C1 through C 20. FT waxes are paraffins at C30.

FT reaction is as follows:

CO ൅ 2.1 כ H ଶ 

### െ െ൐ െሺܥH ଶ 

### ሻെ ൅H ଶ 

### O 

### (eqn. 23)

Section 100 sets the CO conversion.

Section 200 calculates th e reaction extent (in lbmol) ba sed on an alpha value of 0.9.

-------------Section 100

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

Percent conversion of CO is calculated as follo ws, and then the molar amount of converted CO (COCONV) is calculated knowing the molar amount of CO entering (COIN).

PERCEN = 40

CONV=PERCEN/100.0

COCONV=COIN*CONV

------------Section 200--

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

R1, R2, R3, etc. represent the mola r reaction extent (lbmo l/h) that is utilized in the FT reactor for each reaction (i.e., CO + 3*H2  CH4 + H2O, 2*CO + 5*H2  C2H6 + 2*H2O, etc.). The

coefficients of each reaction extent are calcula ted by solving a set of 21 equations shown below

in Table C-2 and as described in the Detail ed Calculations section of this Appendix.

### 75 Table C-2. Reaction Extent Equations for Each Alkane Hydrocarbon

Alkane Component Equation

1C 0.01 כ ܱܸܰܥܱܥ ൌ ܴ1

2C 0.018/2 כ ܱܸܰܥܱܥ ൌ ܴ2

3C 0.0243/3 כ ܱܸܰܥܱܥ ൌ ܴ3

4C 0.02916/4 כ ܱܸܰܥܱܥ ൌ ܴ4

5C 0.0328/5 כ ܱܸܰܥܱܥ ൌ ܴ5

6C 0.03543/6 כ ܱܸܰܥܱܥ ൌ ܴ6

7C 0.0372/7 כ ܱܸܰܥܱܥ ൌ ܴ7

8C 0.03826/8 כ ܱܸܰܥܱܥ ൌ ܴ8

9C 0.03874/9 כ ܱܸܰܥܱܥ ൌ ܴ9

10C 0.03874/10 כ ܱܸܰܥܱܥ ൌ ܴ10

11C 0.03835/11 כ ܱܸܰܥܱܥ ൌ ܴ11

12C 0.03766/12 כ ܱܸܰܥܱܥ ൌ ܴ12

13C 0.03672/13 כ ܱܸܰܥܱܥ ൌ ܴ13

14C 0.03559/14 כ ܱܸܰܥܱܥ ൌ ܴ14

15C 0.03432/15 כ ܱܸܰܥܱܥ ൌ ܴ15

16C 0.03294/16 כ ܱܸܰܥܱܥ ൌ ܴ16

17C 0.0315/17 כ ܱܸܰܥܱܥ ൌ ܴ17

18C 0.03002/18 כ ܱܸܰܥܱܥ ൌ ܴ18

19C 0.02852/19 כ ܱܸܰܥܱܥ ൌ ܴ19

20C 0.02702/20 כ ܱܸܰܥܱܥ ൌ ܴ20

30C 0.36473/30 כ ܱܸܰܥܱܥ ൌ ܴ21



### GRIND

This block calculates the power requirement (i n kW) for grinding the biomass from the chop size

of 15 mm to the final size of 1 mm. This power requirement data is found in Mani et al. and is for 12% exiting moisture. The correlation was ch anged from a polynomial (quadratic) regression, which Mani et al. used, to a power regressi on because the power regression more accurately

matched the data. ܵ ௖௨௧ 

### is the final grind size in the units of millimeters.

ܲ

௚௥௜௡ௗ 

### ൌ ൫28.76 ڄ ܵ ௖௨௧ ି଴.଼ଵ 

### ൯ڄ݉

ሶ ௕௜௢௠௔௦௦ 

### (eqn. 24)

HRSG

This calc block totals th e heat transfer areas of all the heat exchangers in A600 Power Generation for use in Aspen Icarus software costing of a he at recovery steam generator, which is estimated

as a waste heat boiler.

HUMIDITY

This block sets humidity of the ai r entering the air separation unit.

### 76 HV-101, HV-203, HV-445

This block calculates the lower and higher heatin g values of the following streams: biomass, syngas, and FT products.

LOCKHOP

This block calculates the CO 2 

### required for pressurizing the lockhopper. Higman et al. reports 0.09 kg of pressurization gas is required per kg of biomass.

ܵܣܯܱܫܤ כ 0.09 ൌ ܱ2ܥܺ 

### (25 .eqn (

MEATEMP

This block sets the temperature of the in coming monoethanolamine solution entering the

absorber column in the AGR area.

MOISTURE

This block sets the moisture content of the en tering biomass to the preprocessing area and sets the biomass moisture content exiting the biomass dryer. Also, the steam loop flow rate for drying

the biomass is set at 9 times the amount of moisture removed during the drying process.

Moisture content (% wet basis) of entering biomass feed, XMOIS1 ൌ 25 

### . Inlet mass flow of moisture, WATERI 

### , is computed.

ሻܵ1/100ܫܱܯܺ െ 1ሺܵ1/100/ܫܱܯܺ כ ܦܧܧܨ ൌ ܫܴܧܶܣܹ 

### (26 .eqn (

Moisture content (% wet basis) of biomass exit ing the dryer, XMOIS2 ൌ 10 

### . Mass flow of moisture, WATERO 

### , is computed.

WATERO ൌ FEED כ XMOIS2/100/ሺ1 െ XMOIS2/100ሻ 

### (eqn. 27)

Specify steam required to remove moisture, STEAMI



### .

STEAMI ൌ 9 כ ሺWATERI െ WATEROሻ 

### (eqn. 28)

O2COMB

Oxygen is required to combust th e char and syngas that provides the energy necessary for drying the biomass. A system of stoichiometric com bustion reactions is set up to sum all the oxygen required to fully combust the unconverted syngas purge from the FT synthesis outlet. The

reactions are shown in Table C-3.

### 77 Table C-3. Combustion Reactions to Determine Required Oxygen

Component Reaction

CO ܱ2ܥ ՜ ܱ2 · 0.5 ൅ ܱܥ

2H 2ܱܪ ՜ ܱ2 · 0.5 ൅ 2ܪ

4CH ܱ2ܥ ൅ 2ܱܪ2 ՜ 2ܱ2 ൅ 4ܪܥ

6H2C ܱ2ܥ2 ൅ 2ܱܪ3 ՜ ܱ2 · 3.5 ൅ 6ܪ2ܥ

4H2C ܱ2ܥ2 ൅ 2ܱܪ2 ՜ 3ܱ2 ൅ 4ܪ2ܥ

2H2C ܱ2ܥ ൅ 2ܱܪ ՜ 2.5ܱ2 ൅ 2ܪ2ܥ

8H3C ܱ2ܥ3 ൅ 2ܱܪ4 ՜ 5ܱ2 ൅ 8ܪ3ܥ

10H4C ܱ2ܥ4 ൅ 2ܱܪ5 ՜ 6.5ܱ2 ൅ 10ܪ4ܥ

12H5C ܱ2ܥ5 ൅ 2ܱܪ6 ՜ 8ܱ2 ൅ 12ܪ5ܥ

14H6C ܱ2ܥ6 ൅ 2ܱܪ7 ՜ 9.5ܱ2 ൅ 14ܪ6ܥ

16H7C ܱ2ܥ7 ൅ 2ܱܪ8 ՜ 11ܱ2 ൅ 16ܪ7ܥ

18H8C ܱ2ܥ8 ൅ 2ܱܪ9 ՜ 12.5ܱ2 ൅ 18ܪ8ܥ

20H9C ܱ2ܥ9 ൅ 2ܱܪ10 ՜ 14ܱ2 ൅ 20ܪ9ܥ

Tar 10ܪ14ܥ ሺ ݎܽݐ ሻ ܱ2ܥ14 ൅ 2ܱܪ5 ՜ 16.5ܱ2 ൅

H2S ܪ2ܵ ൅ 1.5ܱ2 ՜ ܪ2ܱ ൅ ܱܵ2

3NH ܱܰ2 ൅ 2ܱܪ1.5 ՜ 1.75ܱ2 ൅ 3ܪܰ



### The molar flow rate of oxygen entering the comb ustor is summed and multiplied by a factor of
1.25 in order to combust with 25% exce ss air as shown in equation 29 below.

ܯ

ሶ ைଶ,௜௡ 

### ൌ 1.25 · ሺܯ ሶ ஼ு஺ோ,௜௡ 

### ൅ 0.5ܯ ሶ ஼ை,௜௡ 

### + 0.5ܯ ሶ ுଶ,௜௡ 

### ൅ 2ܯ ሶ ஼ுସ,௜௡ 

### ൅ 3.5ܯ ሶ ஼ଶு଺,௜௡ 

### ൅

ܯ3 ሶ ஼ଶுସ,௜௡ 

### ൅ 2.5ܯ ሶ ஼ଶுଶ,௜௡ 

### ൅ 5ܯ ሶ ஼ଷு଼,௜௡ 

### ൅ 6.5ܯ ሶ ஼ସுଵ଴,௜௡ 

### ൅ 8ܯ ሶ ஼ହுଵଶ,௜௡ 

### ൅

ܯ9.5 ሶ ஼଺ுଵସ,௜௡ 

### ൅ 11ܯ ሶ ஼଻ுଵ଺,௜௡ 

### ൅ 12.5ܯ ሶ ஼଼ுଵ଼,௜௡ 

### ൅ 14ܯ ሶ ஼ଽுଶ଴,௜௡ 

### ൅

ܯ16.5 ሶ ்஺ோ,௜௡ 

### + 1.5ܯ ሶ ுଶௌ,௜௡ 

### ൅ 1.75ܯ ሶ ேுଷ,௜௡ 

### ሻ 

### (eqn. 29)

O2TURB

This block calculates the molar flow rate of air (oxygen and nitrogen ) required to combust syngas obtained from FT synthesis and the fuel gas obtained from Area 500 in the gas turbine of
Area 600. A value of 25% excess air is assumed. The calculations are similar to the methodology

in O2COMB

.

OXYSET

This block sets the entering oxygen at 0.35 lb oxy gen per lb dry biomass into the gasifier.

݉ሶ ைଶ,௚௔௦ 

### ൌ 0.35/100 · ݉ ሶ ௕௜௢௠௔௦௦ 

### (eqn. 30)

SWGSSTM

### 78 This block sets the steam flow into the sour water-g as-shift reactor to be at a ratio of 3:1 water to carbon monoxide. This ratio ensures enough water-gas -shift activity occurs within the reactor.

݉ሶ ௌ்ெ,௔ௗௗ௜௧௜௢௡ 

### ൌ 3.0 · ݉ ሶ ஼ை 

### െ ݉ ሶ ுଶை 

### (eqn. 31)

### 79 Low-Temperature Scenario

AMINE

This block calculates the mola r flow of monoethanolamine (MEA ) needed for the required acid gas removal (CO 2 

### and H 2 

### S) arriving from syngas quench a nd FT unconverted syngas recycle stream. The MEA is able to capture 0.35 moles acid gas per mole MEA. Additionally, the MEA is diluted as explained in DILUTH2O.

Molar MEA flow (in kmo l/h) is calculated by

ܯ

ሶ ொ஺ 

### ൌ ሺܯ ሶ ஼ைଶ,௦௬௡ 

### ൅ ܯ ሶ ஼ைଶ,௥௘௖ 

### ൅ ܯ ሶ ுଶௌ,௦௬௡ 

### ሻ/0.35 (eqn. 32)

where ܯ ሶ ஼ைଶ,௦௬௡ 

### is molar flow of CO 2 

### from the syngas after the syngas quench, ܯ ሶ ஼ைଶ,௥௘௖ 

### is the

molar flow of CO 2 

### from the unconverted syngas recycle after th e FT synthesis, and ܯ ሶ ுଶௌ,௦௬௡ 

### is

the molar flow rate of H 2 

### S from the syngas quench.

Because the MEA solution in the amine absorption un it is to be 20 wt% con centrated with water,

the flow of water must be calculated.

Mole flow of water is calculated as

ܯ

ሶ ுଶை 

### ൌ ܯ ሶ ஺ொ 

### ܹܯ כ ஺ொ 

### /0.20

ܯܹ ுଶை 

### (eqn. 33)

BIOELEM

Same as for the HT scenario.

DILUTH2O

This block sets the MEA solution to be 20% concentrated with water.

FTDISTR

Same as for the high-temperature scenario.

GASYIELD

The following model describes how the fluidized-bed gasifier keep s an elemental mass balance. Experiments performed at Iowa State University provide the initial gasi fier product distribution,

### 80 and the model adjusts the yields of those expe riments in order to balance carbon, hydrogen, sulfur, nitrogen, oxygen, and ash.

The approach taken to balance each element across the gasifier is by “floating” a component of each element. The “floating” component for carb on is the char. All sulfur and nitrogen not found in the char is assumed to form hydrogen sulfide and ammonia, respectively. Therefore, sulfur and nitrogen are balanced. Next, elemental hydr ogen is adjusted in the model by either converting diatomic hydrogen to steam or deco mposing steam to diatomic hydrogen. Oxygen balance is more complex. Because gasification operates at fuel-rich conditions, diatomic oxygen should not be present in the s yngas leaving the gasifier. Theref ore, diatomic oxygen cannot be the “floating” component. Instead, oxygen is ba lanced by adjusting the carbon monoxide or carbon dioxide in the exiting syngas. Because th ere is one oxygen difference between those two components, the oxygen can be adjusted to help close the balance.

Carbon balance follows the flow chart shown in Figure C-4. If there is less gaseous carbon out than total carbon in, then the difference is made up of char carbon, CCARB. Char is assumed to be composed of 68% carbon with the rest as H, O, N, and S. Ash is considered apart from the

char and is considered inert in the model. Because the char is now fixed, the only pathway for sulfur and nitrogen to take is to form hydrogen sulfide and ammonia. Therefore, the sulfur and nitrogen are balanced.

Figure C-4. Decision diagram for carbon balance



### Next, as shown in Figure C-5, hydrogen is balanced. Knowing hydrogen in the char and in gaseous products, the hydrogen required (HREQD) is calculated as the sum of those two components. If the hydrogen required is less than hydrogen available (HAVAIL), made up of hydrogen in steam, biomass moisture, and in th e biomass itself (THYD), then there is enough hydrogen available to balance. To balance hydrogen, the product yield swings toward either steam or diatomic hydrogen.

### 81 Figure C-5. Decision diagram for hydrogen balance



### The only element left to balance is oxygen, which is accomplished by forcing creation of carbon

monoxide or creation of carbon dioxide as show n in Figure C-6. The required oxygen (OREQD), made up of oxygen in char and oxygen in syngas, is checked against th e available oxygen found in the entering oxygen, steam, and biomass. If ther e is more oxygen availabl e than required, then the option is to move the excess oxygen to CO 2 

### by decreasing CO. If ther e is still oxygen present when CO is decreased to zero, then the yields ne ed to be adjusted beca use excess oxygen is still present. If there is an oxy gen deficit (OREQD > OAVAIL), then CO is increased and CO 2 

### is decreased. After that, if there is still an oxygen deficit, then insufficient oxygen is present and yields need to be adjusted. When all these step s are completed and no errors are generated, there is an elemental mass balance across the gasifier.

### 82 Figure C-6. Decision diagram for oxygen balance



### GRIND

This block calculates the power requirement (i n kW) for grinding the biomass from the chop size

of 12 mm to the final size of 6 mm. This power requirement data is found in Mani et al. and is for 12% exiting moisture. The correlation has been changed from a polynomial regression (which Mani et al. used) to a power regression because the power regressi on fit the data better.

ܵ ௖௨௧ 

### is in millimeters.

ܲ ௚௥௜௡ௗ 

### ൌ ൫28.76 ڄ ܵ
௖௨௧ ି଴.଼ଵ 

### ൯ڄ݉

ሶ ௕௜௢௠௔௦௦ 

### (eqn. 34)

HUMIDITY

This block sets humidity of the ai r entering the air separation unit.

HV-101, HV-203, HV-445

This block calculates the lower and higher heatin g values of the following streams: biomass, syngas, and FT products.

MOISTURE

This block is the same as found in the HT scenario.

### 83 O2COMB

This block is the same as found in the HT scenario.

O2TURB

This block is the same as found in the HT scenario.

OXYSET

This block sets the entering oxidizing agents, oxygen and steam, into the gasifier. A linear correlation with temperature, ܶ ௚௔௦ 

### (in Fahrenheit), adapted from Bain for oxygen, is used because as oxygen increases in the gasifier, the temperature increases. Mass flow of oxygen,

݉ሶ ைଶ,௚௔௦ 

### , is in percentage of dry feedstock.

݉

ሶ ைଶ,௚௔௦ 

### ൌ ሺെ11.567 ൅ 0.02375 · ܶ ௚௔௦ 

### ሻ/100 · ݉ ሶ ௕௜௢௠௔௦௦ 

### (eqn. 35)

The steam feed rate is set at 0.66 lb steam per lb oxygen.

݉

ሶ ௦௧௘௔௠,௚௔௦ 

### ൌ 0.66 · ݉ ሶ ைଶ,௚௔௦ 

### (eqn. 36)

Because 95% purity oxygen is produced in the air separation unit, argon mass flow is set at 5% of molar oxygen flow.

݉ ሶ ௔௥௚௢௡ 

### ൌ 0.05 · ൬ ݉ ሶ ைଶ,௚௔௦


### ܹܯ ைଶ 

### ൰ ܹܯ כ ஺௥ 

### (37 .eqn(

### 84 Aspen Plus Model Design Specifications

High-Temperature Scenario

DS-1

The exiting temperature of air in the heat ex changer used to pre-cool the air entering the cryogenic distillation column is varied until a net duty of zero is observed.

FSSPL02

This design specification varies th e fraction of unconverted syngas that is piped to Area 200 for the combustion of syngas. The syngas, in turn, provides the heat require d to dry the biomass.

H2SPLIT

This design spec calculates the required hydrogen that needs to be reserved by the PSA unit for use in Area 500: Hydrocracking. A typical yiel d from hydrocracking is shown in Table C-4. Because the FT products are hydrogen deficient re lative to the final blend, makeup hydrogen is required. The syngas purge amount going to the pre ssure swing adsorption (PSA) unit is varied

so that the calculated deli vered hydrogen matches the required hydrogen to Area 500. Without showing the detailed calculations, the basic st eps are first calculati ng the carbon and hydrogen content in the FT product stream. The carbon mass fl ow is the same as that of the final blend stream flow. Using the blend fractions in Table C-4, the amount of hydrogen is calculated in the final blend and the difference in hydrogen is de termined. The difference is multiplied by 1.1 to obtain the delivered hydrogen mass fl ow rate to hydr ocracking area.

Table C-4. Hydroprocessing Product Blend

Component Mass Fraction Fuel Gas (methane) 0.034 LPG (propane) 0.088 Gasoline (n-octane) 0.261 Diesel (n-hexadecane) 0.617



### O2-101, O2-203, O2-445

These design specifications vary the amount of oxygen inlet to the heating value blocks (HV101, HV-203, HV-445) so as to be stoichiometric in the combustion of the duplicate stream.

O2-SULF

This design specification varies the amount of oxygen into the LO -CAT oxidizer unit to fully

oxidize the H 2 

### S into solid sulfur.

SGSTEMP

### 85 The temperature of operation in the sour water- gas-shift reactor is varied until the exiting equilibrium molar ratio of H 2 

### /CO is just above the optimal FT ratio (2.1). A small amount of hydrogen is captured in the PSA unit bringing that ratio down to th e optimum for FT synthesis.

Low-Temperature Scenario

DS-1

This design specification is the same as for the HT scenario.

H2SPLIT

This design specification is the same as for the HT scenario.

O2-101, O2-203, O2-445

These design specifications are the same as in the HT scenario.

O2-SULF

This design specification is the same as for the HT scenario.

STMRECOV

Heat can be recovered from the combustion of syngas and char. This specification varies the steam flow rate (stream 280) to bring the combustion flue gas (stream 252) down to 200°C via heat exchanging.

WGSTEMP

The temperature of operation in the water-gas-shift reactor is varied until the exiting equilibrium

molar ratio of H 2 

### /CO is just above the optimal FT ra tio (2.1). A small amount of hydrogen is captured in the PSA unit, bringing that ratio down to the optimum for FT synthesis.

### 86 Detailed Calculations

Aspen Plus Model Calculations and Notes

Outline Defining Units

Plant Input

Plant Output

Carbon Efficiency to Fuels

Energy Content

FT Reaction Conversion Solver

Equipment Sizing

Dryer

Lockhoppers

Slag/Char Collection

PSA Unit

Fuel Storage

LT Gasifier Cost

FT Reactor Cost

Acid Gas Removal Area Cost

A500 Hydroprocessing Area Cost

Reactors and Catalysts

Natural Gas Utility Usage MJ 10 6 J := MMcf 10 6 ft 3

:=

kPa 10 3 P a
⋅ := Cp 100poise :=

ρ water 1000 kg

m

3 := MW H2O 18.02 gm

mol :=

kmol 1000mo l := MMBTU 10 6 BT U :=

lbmol kmol

2.2 :=

bbl 42gal :=

ρ gas 737.22 kg

m 3 := 100 kg

m

3 6.243 lb

ft 3 =

ρ diesel 840 kg

m

3 := therm 100000 BTU :=

MMgal 10 6 ga l := dekatherm 10therm :=

kJ 1000 J := P ref 1atm ≡

T ref 298 K ≡ bpsd 42gal

day :=

PJ 10 15 J := GJ 10 9 J := HHV stover 7.588 10 3 × BTU

lb = HHV stover 17.65 MJ

kg :=

### 87 Plant Input

Biomass

Elemental Composition

Carbon

Oxygen

Hydrogen

Sulfur

Nitrogen

Ash



### Elemental Mass Flow

m

dot_biomass 2000 tonne
day := Availability 310 day := Load 7446 hr :=

Frac C_biomass 0.4728 ≡ MW C 12.01 gm

mol :=

Frac O_biomass 0.4063 ≡ MW O 16. gm

mol :=

Frac H_biomass 0.0506 ≡ MW H 1.01 gm

mol :=

Frac S_biomass 0.0022 ≡ MW S 32.07 gm

mol :=

Frac N_biomass 0.008 ≡ MW N 14.01 gm

mol :=

Frac A_biomass 0.0600 ≡

m

dot_C_in m dot_biomass Frac C_biomass ⋅ := m dot_C_in 945.6 tonne
day =

m

dot_O_in m dot_biomass Frac O_biomass ⋅ := m dot_O_in 812.6 tonne
day =

m

dot_H_in m dot_biomass Frac H_biomass ⋅ := m dot_H_in 101.2 tonne
day =

m

dot_S_in m dot_biomass Frac S_biomass ⋅ := m dot_S_in 4.4 tonne
day =

m

dot_N_in m dot_biomass Frac N_biomass ⋅ := m dot_N_in 16 tonne
day =

m

dot_A_in m dot_biomass Frac A_biomass ⋅ := m dot_A_in 120 tonne
day =

### 88 Elemental Mole Flow

Biomass Moisture

Source: Kaliyan and Morey, 2005 for 0.66-0.8 mm sized particles n dot_C_in m dot_C_in

MW C := n dot_C_in 911.278 mol

s =

n dot_O_in m dot_O_in

MW O := n dot_O_in 587.818 mol

s =

n dot_H_in m dot_H_in

MW H := n dot_H_in 1160 mol

s =

n dot_S_in m dot_S_in

MW S := n dot_S_in 1.588 mol

s =

n dot_N_in m dot_N_in

MW N := n dot_N_in 13.218 mol

s =

moist in 0.25 := moist dried 0.10 :=

m

dot_moist_in moist in m dot_biomass ⋅

1 moist in

− := m dot_moist_in 666.667 tonne
day =

m

dot_moist_dried moist dried m dot_biomass ⋅

1 moist dried

− := m dot_moist_dried 222.222 tonne
day =

ρ bulk_stover 100 kg

m

3 :=

### 89 HT Gasifier Steam/Oxygen Addition Source: Probstein and Hicks, 2006

Stoichiometric/thermo-neutral requirement for synthesis gas according to following equation:

1.34C + 0.34 O 2 + H 2 O --> 0.34CO 2 + CO + H 2

Oxygen to Carbon: 0.25
Steam to Carbon: 0.75

Steam addition ratio is then three times that of oxygen minus the moisture in the biomass m

dot_O2_in 0.35 m dot_biomass ⋅ := m dot_O2_in 700 tonne
day =

n dot_O2_in m dot_O2_in

2 MW O

⋅ := n dot_O2_in 253.183 mol

s =

Ratio O2_to_C n dot_O2_in

n dot_C_in := Ratio O2_to_C 0.278 =

n dot_H2O_in 3 Ratio O2_to_C ⋅ n dot_C_in ⋅ m dot_moist_dried

MW H2O − :=

m dot_H2O_in n dot_H2O_in MW H2O ⋅ := n dot_H2O_in 616.817 mol

s =

m

dot_H2O_in 960 tonne
day =

### 90 Plant Output

HT Fuel Production

LT Fuel Production m

dot_gasHT 112.78 tonne
day := m dot_dieselHT 266.5 tonne
day :=

v dot_gasHT m dot_gasHT

ρ gas := v dot_dieselHT m dot_dieselHT

ρ diesel :=

v dot_gasHT 40413 gal

day = v dot_dieselHT 83812 gal

day =

v dot_gasHT 962 bbl

day = v dot_dieselHT 1996 bbl

day =

v dot_gasoline_per_year v dot_gasHT Loa d

⋅ :=

v dot_diesel_per_year v dot_dieselHT Loa d

⋅ :=

m

dot_gasLT 87.12 tonne
day := m dot_dieselLT 205.86 tonne
day :=

v dot_gasLT m dot_gasLT

ρ gas := v dot_dieselLT m dot_dieselLT

ρ diesel :=

v dot_gasLT 31218 gal

day = v dot_dieselLT 64741 gal

day =

v dot_gasLT 743 bbl

day = v dot_dieselLT 1541 bbl

day =

v dot_gasoline_per_yearLT v dot_gasLT Loa d

⋅ := v dot_gasoline_per_yearLT 9.685 MMgal =

v dot_diesel_per_yearLT v dot_dieselLT Loa d

⋅ := v dot_diesel_per_yearLT 20.086 MMgal = v dot_gasoline_per_year 12.538 MMgal =

v dot_diesel_per_year 26.003 MMgal =

### 91 Carbon Efficiency to Fuels

HT Scenario

Gasoline Carbon

Diesel Carbon m

dot_gasHT 112.78 tonne
day =

m

dot_C_gasHT Frac C_gasoline m dot_gasHT ⋅ := m dot_C_gasHT 94.835 tonne
day =

m

dot_dieselHT 266.5 tonne
day =

Frac C_diesel 16 12.01

⋅

16 12.01

⋅ 34 1.01 ⋅+ := Frac C_diesel 0.848 =

m

dot_C_dieselHT Frac C_diesel m dot_dieselHT ⋅ := m dot_C_dieselHT 226.096 tonne
day =

m

dot_C_outHT m dot_C_gasHT m dot_C_dieselHT + :=

m

dot_C_outHT 320.931 tonne
day = C_effHT m dot_C_outHT

m

dot_C_in := Frac C_gasoline 8 12.01

⋅

8 12.01

⋅ 18 1.01 ⋅+ := Frac C_gasoline 0.841 =

C_effHT 0.339 =

### 92 LT Scenario

Gasoline Carbon

Diesel Carbon m

dot_gasLT 87.12 tonne
day =

m

dot_C_gasLT Frac C_gasoline m dot_gasLT ⋅ := m dot_C_gasLT 73.258 tonne
day =

m

dot_dieselLT 205.86 tonne
day =

m

dot_C_dieselLT Frac C_diesel m dot_dieselLT ⋅ := m dot_C_dieselLT 174.649 tonne
day =

m

dot_C_outLT m dot_C_gasLT m dot_C_dieselLT + :=

m

dot_C_outLT 247.908 tonne
day = C_effLT m dot_C_outLT

m

dot_C_in := C_effLT 0.262 =

### 93 E gasifierloss_LT 96.796 MW = E gasifierloss_HT 47.111 MW = 

### Energy Content

This section acquires the energy content (on a LH V basis) from the Aspen Plus data and converts to megawatts for use in developing an energy balance



### Biomass

Fuel

Char/Tar

Raw Syngas

Energy Loss across the Gasifier

Energy lost across the gasifier is calculated as difference in energy in the biomass and energy in the raw syngas and char (only in LT scenario) E biomass 1400313 MJ

hr := E biomass 388.976 MW =

E fuelHT 695598 MJ

hr := E fuelHT 193.222 MW =

E fuelLT 539292 MJ

hr := E fuelLT 149.803 MW =

E char_LT 87792 MJ

hr := E char_LT 24.387 MW =

E tar_LT 16980 MJ

hr := E tar_LT 4.717 MW =

E rawsyngas_HT 1230712 MJ

hr := E rawsyngas_HT 341.864 MW =

E rawsyngas_LT 964054 MJ

hr := E rawsyngas_LT 267.793 MW =

E

gasifierloss_HT E biomass E rawsyngas_HT − :=

E gasifierloss_LT E biomass E rawsyngas_LT − E char_LT − :=

### 94 Unconverted Syngas used in A600 Power Generation

Fuel Gas from A500 used in A600 Power Generation

Fischer-Tropsch Product

Electricity Generated

Net Electricity (exported)

Power Generation Loss

The loss is the difference between electric generation out and the gas energy in

E syngasA600_HT 129332 MJ

hr := E syngasA600_HT 35.926 MW =

E syngasA600_LT 109708 MJ

hr := E syngasA600_LT 30.474 MW =

E fuelgas_HT 104114 MJ

hr := E fuelgas_HT 28.921 MW =

E fuelgas_LT 80718 MJ

hr := E fuelgas_LT 22.422 MW =

E FTliquids_HT 782894 MJ

hr := E FTliquids_HT 217.471 MW =

E FTliquids_LT 606801 MJ

hr := E FTliquids_LT 168.556 MW =

E elecgenOUT_HT 48.55 MW := E elecgenOUT_LT 40.73 MW :=

E elecnet_HT 13.8 MW := E elecnet_LT 16.3 MW :=

E

A600losses_HT E syngasA600_HT E fuelgas_HT + E elecgenOUT_HT − :=

E A600losses_HT 16.296 MW =

E A600losses_LT E syngasA600_LT E fuelgas_LT + E elecgenOUT_LT − :=

E A600losses_LT 12.166 MW =

### 95 Loss across FT Reactor

Unconverted Syngas used for Biomass Drying

Only in HT scenario E FTreactorlosses_HT 226737 MJ

hr := E FTreactorlosses_HT 62.983 MW =

E FTreactorlosses_LT 175128 MJ

hr := E FTreactorlosses_LT 48.647 MW =

E biomass_drying_HT 24663 MJ

hr := E biomass_drying_HT 6.851 MW =

### 96 Fischer-Tropsch Reaction Conversion Solver

This section solves for the reaction fractional conversion for each reaction in the FischerTropsch reactor. A set of equations is developed and solved. The resulting ε values ( ε 1 - ε 30) are used directly in the Aspen Plus conversion re actor block. The reactions in the reactor block
are defined as molar extent.

Depending on the alpha chain growth probab ility, the reactor forms different product composition.

Step 1: choose the expected alpha chain growth value

Step 2: using the α FT chain growth, the mole fraction of each hydrocarbon in the FT product is

calculated.

All hydrocarbons greater than C20 make up the balance and are modeled using C30. α FT 0.9 :=

M1 α FT 1 1

− 1 α FT − 

# ( )

⋅ := M1 0.1 = M11 α FT 11 1

− 1 α FT − 

# ( )

⋅ :=

M2 α FT 2 1

− 1 α FT − 

# ( )

⋅ := M2 0.09 = M12 α FT 12 1

− 1 α FT − 

# ( )

⋅ :=

M3 α FT 3 1

− 1 α FT − 

# ( )

⋅ := M3 0.081 = M13 α FT 13 1

− 1 α FT − 

# ( )

⋅ :=

M4 α FT 4 1

− 1 α FT − 

# ( )

⋅ := M4 0.073 = M14 α FT 14 1

− 1 α FT − 

# ( )

⋅ :=

M5 α FT 5 1

− 1 α FT − 

# ( )

⋅ := M5 0.066 = M15 α FT 15 1

− 1 α FT − 

# ( )

⋅ :=

M6 α FT 6 1

− 1 α FT − 

# ( )

⋅ := M6 0.059 = M16 α FT 16 1

− 1 α FT − 

# ( )

⋅ :=

M7 α FT 7 1

− 1 α FT − 

# ( )

⋅ := M7 0.053 = M17 α FT 17 1

− 1 α FT − 

# ( )

⋅ :=

M8 α FT 8 1

− 1 α FT − 

# ( )

⋅ := M8 0.048 = M18 α FT 18 1

− 1 α FT − 

# ( )

⋅ :=

M9 α FT 9 1

− 1 α FT − 

# ( )

⋅ := M9 0.043 = M19 α FT 19 1

− 1 α FT − 

# ( )

⋅ :=

M10 α FT 10 1

− 1 α FT − 

# ( )

⋅ := M10 0.039 = M20 α FT 20 1

− 1 α FT − 

# ( )

⋅ :=

M30 1 M1 M2 + M3 + M4 + M5 + M6 + M7 + M8 + M9 + M10 +
M11 M12 + M13 + M14 + M15 + M16 + M17 + M18 + M19 + M20 + + ... 

###     − :=

M30 0.122 = M13 0.028 =

M14 0.025 = M11 0.035 =

M12 0.031 =

M15 0.023 =

M16 0.021 =

M17 0.019 =

M18 0.017 =

M19 0.015 =

M20 0.014 =

### 97 Step 3: Set up a series of equations to solve along with guess values (required for Mathcad)

For a nominal 1000 moles of CO input, the expected CO output is 600 moles because 40% is
converted.

<----------------- 40% conversion of CO

<----- This value to be varied until COconv is equal to desired.

A nominal 400 moles of CO are converted in the FT reactor. The sum of the exiting amount of moles in the FT product distribution will not be 400, because moles are not conserved. Mass is conserved, however. Therefore, the variable "D " represents a factor that adjusts all the conversions ( ε 1, ε 2, etc.).

The resulting value of D is 0.1 meaning that 40 moles of FT products exit the reactor. Known CO out 600 := CO in 1000 :=

Guess ε 1 20 := ε 2 20 := ε 3 20 := ε 4 20 := ε 5 20 :=

ε 6 20 := ε 7 20 := ε 8 20 := ε 9 20 := ε 10 20 :=

ε 11 20 := ε 12 20 := ε 13 20 := ε 14 20 := ε 15 20 :=

ε 16 20 := ε 17 20 := ε 18 20 := ε 19 20 := ε 20 20 :=

ε 30 20 := D 0.1 :=

Give n

D ε 1 1

2 ε

2 + 1

3 ε

3 + 1

4 ε

4 + 1

5 ε5 + 1

6 ε 6 + 1

7 ε

7 + 1

8 ε 8 + 1

9 ε

9 + 1

10 ε 10 + 1

11 ε 11 + 1

12 ε 12 +

1

13 ε 13 1

14 ε 14 + 1

15 ε 15 + 1

16 ε 16 + 1

17 ε 17 + 1

18 ε 18 + 1

19 ε 19 + 1

20 ε 20 + 1

30 ε 30 + + ... 

###          

M1 ε 1

D M2 1

2 ε

2

D M3 1

3 ε

3

D M4 1

4 ε

4

D M5 1

5 ε 5

D M6 1

6 ε

6

D

M7 1

7 ε

7

D M8 1

8 ε 8

D M9 1

9 ε

9

D M10 1

10 ε 10

D M11 1

11 ε 11

D M12 1

12 ε 12

D

M13 1

13 ε 13

D M14 1

14 ε 14

D M15 1

15 ε 15

D M16 1

16 ε 16

D M17 1

17 ε 17

D M18 1

18 ε 18

D

M19 1

19 ε 19

D M20 1

20 ε 20

D M30 1

30 ε 30

D

Solve Find ε 1 ε 2 , ε 3 , ε 4 , ε 5 , ε 6 , ε 7 , ε 8 , ε 9 , ε 10 , ε 11 , ε 12 , ε 13 , ε 14 , ε 15 , ε 16 , ε 17 , ε 18 , ε 19 , ε 20 , ε 30

, 

## ( ) :=

Solve 0

0

1 0.01

0.018 =

### 98 Step 4: The guess value of D is varied unt il the sum of all reaction conversions ( ε 1, ε 2, etc.)
sum to 1.0 as seen below. This means that a ll 400 moles of CO are converted as expected.

Step 5: Each value for ε is imported into Aspen Plus e1 Solve 0 := e1 0.01 = e11 Solve 10 := e11 0.03835 = D 0.1 =

e2 Solve 1 := e2 0.018 = e12 Solve 11 := e12 0.03766 =

e3 Solve 2 := e3 0.0243 = e13 Solve 12 := e13 0.03672 =

e4 Solve 3 := e4 0.02916 = e14 Solve 13 := e14 0.03559 =

e5 Solve 4 := e5 0.0328 = e15 Solve 14 := e15 0.03432 =

e6 Solve 5 := e6 0.03543 = e16 Solve 15 := e16 0.03294 =

e7 Solve 6 := e7 0.0372 = e17 Solve 16 := e17 0.0315 =

e8 Solve 7 := e8 0.03826 = e18 Solve 17 := e18 0.03002 =

e9 Solve 8 := e9 0.03874 = e19 Solve 18 := e19 0.02852 =

e10 Solve 9 := e10 0.03874 = e20 Solve 19 := e20 0.02702 =

e30 Solve 20 := e30 0.36473 =

CO conv e1 e2 + e3 + e4 + e5 + e6 + e7 + e8 + e9 + e10 + e11 + e12 + e13 + e14 + e15 + e16 +
e17 e18 + e19 + e20 + e30 + + ... :=

CO conv 1 =

### 99 Equipment Sizing

Rotary Dryer Source: Process Engineering Economics by James Couper, 2003

Typical rpm of rotary dryers

Typical product of equals 15-25. Assume value of 25 for larger end

Typical residence times are 5-90 minutes and hold up of solids is 7%-8%. Assume 5 minutes
and 8%.

Typical exit gas temperature is 10°-20°C above the entering solids.

Feed rate into plant is 2000 ton/day with bul k density of stover equal to 100kg/m^3. Water density is accounted for as well.

Volume of solids in dryer

Volume of solids and steam

Length of theoretical dryer

Surface area of theoretical dryer

Max surface area as reported by Aspen Icarus is 185 m 2 ; therefore approximately 10 dryers are required.

Feed throughput in each dryer (used for Icarus input) rpm dryer 4 :=

rpm diameter feet ( )

⋅

D dryer 25ft

rpm dryer := D dryer 6.25 ft =

t res 5min := holdup 0.08 :=

m

dot_feed 2000 tonne
day := m dot_moist_in 666.667 tonne
day =

ρ bulk_stover 100 kg

m

3 = ρ water 1000 kg

m

3 =

V

solids m dot_feed

ρ bulk_stover m dot_moist_in

ρ water + 

###       t res

⋅ :=

V

dryer_total V solids

holdup :=

length dryer V dryer_total

D dryer 2 π

4 ⋅ :=

A surf_dryer length dryer π⋅ D drye r ⋅ :=

m dot_feed m dot_moist_in +

10 24495.8 lb

hr = V

solids 71.759m 3

=

V

dryer_total 896.991m 3

=

length dryer 314.708 m =

A surf_dryer 1883.4m 2

=

### 100 Lockhopper System

Source: CE IGCC Repowering Project Bins and Lockhoppers, Combustion Eng. 1993

Note: this report's feedstock is coal

Assumptions from report -A receiving bin is situated before the lockhopper with a 40 minute residence time -Design pressure is for 50 psia. -Cycle time for lockhopper system is designed for 10 minutes resulting in approximately 50,000 cycles per year -Storage volume for lockhopper and feed bin is assumed to be 10 minutes -Approximate lockhopper and feed bin vessel thickness is 1.5 inches and design pressure is for 450 psia
-Volume is theoretical + 33%



### Residence Time

biomass receiving bin

biomass lockhopper

biomass feed bin

Density of feed



### HT Scenario Lockhopper System (1 train)

Volume of biomass receiving bin

Volume of biomass lockhopper

Volume of biomass feed bin t res_rbin 40min := ε void 25% :=

t res_lock 10min :=

t res_fbin 10min :=

m dot_feed_lock m dot_feed m dot_moist_dried + :=

m

dot_feed_lock 2222 tonne
day =

ρ stover_10%moist ρ bulk_stover 2000 ⋅ ρ water 222 ⋅+

2222 :=

ρ stover_10%moist 189.919 kg

m

3 =

V

r_bin t res_rbin m dot_feed_lock ⋅

ρ stover_10%moist 1

1 ε void

− ⋅ :=

V

lock t res_lock m dot_feed_lock ⋅

ρ stover_10%moist 1

1 ε void

− ⋅ :=

V

f_bin t res_fbin m dot_feed_lock ⋅

ρ stover_10%moist 1

1 ε void

− ⋅ := V r_bin 433m 3

=

V

lock 108m 3

=

V

f_bin 108m 3

=

### 101 LT Scenario Lockhopper System (7 trains)

Volume of biomass receiving bin

Volume of biomass lockhopper

Volume of biomass feed bin

Source: Combustion Engineering 1993 V

r_binLT V r_bin

7 :=

V

lockLT V lock

7 :=

V

f_binLT V f_bin

7 := V r_binLT 61.909m 3

=

V

lockLT 15.477m 3

=

V

f_binLT 15.477m 3

=

### 102 Lockhopper Power Consumption

Source: Techno-Economic Analysis of Hydrogen Production by Gasification of Biomass by Lau et al. [2002]

Specific power of lockhopper, kW/tonne/day

Biomass inlet to gasifier



### Fly Ash Collection Storage Tank (assume 7 days storage)

(assumed)



### Slag Separation Drum (5 minut e residence time, 20% volume)

Slag Collection Storage Tank (7 days storage)

Note: the resulting volumes are used to assist in costing using Aspen Icarus 

### Char Collection Storage Bin (1-d ay residence time, 80% volume)

assume 20% voidage SP lock 0.082 kW

tonne

day :=

m dot_gasifier m dot_biomass m dot_moist_dried + :=

Power lock SP lock m dot_gasifier ⋅ := Power lock 182.222 kW =

ρ ash 700 kg

m

3 := m dot_ash 5.88 tonne
day :=

V

tank m dot_ash

ρ ash 7

⋅ da y := V tank 58.8m 3

= V tank 2.077 10 3 × ft 3

=

ρ slag 2700 kg

m

3 := m dot_slag 114 tonne
day := ε void_slag 0.8 :=

V

drum m dot_slag

ρ slag 5

⋅ min 1

1 ε void_slag

− ⋅ := V drum 0.733m 3

=

V

chardrum m dot_char

ρ char 1

⋅ day 1

1 ε void_char

− ⋅ := V slag_storage 295.6m 3

=

V

slag_storage m dot_slag

ρ slag 7

⋅ da y :=

ε void_char 0.2 := ρ char 2700 kg

m

3 := m dot_char 214 tonne
day := V chardrum 99.074m 3

=

### 103 Pressure Swing Adsorption Unit Sizing

The adsorption unit is 1/3 molsieve and 2/3 activated carbon

(a) (b) (b)

Molsieve 13X

Determine dry volumetric flow rate of the syngas stream at atmospheric pressure and 25 deg C

Mole fraction of components that are adsorbed

Actual flow rate of components adsorbed

Adsorbent capacity

(d) SCF/lb corrected for P and T to actual

cm 3 /gm; PSA occurs at ambient temperature

Mass of molsieve required References in parentheses are given at the end of this section. Pi 3.1415 := nm 10 9

− m
⋅ :=

BulkDens 43 lb

ft 3 ⋅ := SA 1320 m 2

gm ⋅ := PoreVol 0.51 cm 3

gm ⋅ :=

VolFlowRate 167 1

− ( ) kmol

hr ⋅ 22.414

⋅ m 3

kmol ⋅ 14.696 psi

⋅ 400 psi

⋅ ⋅ 273.15 25 + ( )K

⋅

273.15 K

⋅ ⋅ :=

VolFlowRate 149.211 m 3

hr =

CO 23 := CO2 1 := CH4 1 :=

FlowRateAds VolFlowRate CO CO2 + CH4 +

100 ⋅ :=

FlowRateAds 37.303 m 3

hr =

AdsCap 0.34 ft 3

lb ⋅ 14.696 psi

⋅ 400 psi

⋅ ⋅ 273.15 25 + ( )K

⋅

273.15 K

⋅ ⋅ :=

AdsCap 0.851 cm 3

gm =

CycleTime 5 mi n
⋅ :=

MolSieveMass FlowRateAds CycleTime

⋅ AdsCap := MolSieveMass 3.652 10 3 × kg =

### 104 Determine volume and length of molsieve bed and activated carbon bed

(assumed)

(Just molsieve bed)

(bed is 1/3 molsieve, 2/3 activated carbon)

(a) http://www.sigmaaldrich.com/Brands/Aldrich/ Tech_Bulletins/AL_143/Molecular_Sieves.html

(b) US Pat 6117810

(d) WO/1998/058726 BULK SEPARATION OF CARBON DIOXIDE FROM METHANE
USING NATURAL CLINOPTILOLITE --ex trapolate to partial pressure of
CO 2 +CH 4 +N 2 +CO=32.6%*400 psi BedVolume MolSieveMass

BulkDens := BedVolume 5.302 m 3

=

Diam 4 f t

⋅ := Diam 1.219 m =

Length BedVolume

Pi Diam 2

⋅ := Length 3.725f t =

RxtrLength 3 Length

⋅ :=

RxtrLength 11.175f t = RxtrLength 3.406 m =

RxtrVolume RxtrLength Diam 2

⋅ 0.25

⋅ π

⋅ := RxtrVolume 3.977m 3

= × 4

### 105 HT Scenario Fuel Storage

Gasoline Storage Tank (30 days storage)

Diesel Storage Tank (30 days storage)

Note: the resulting volumes are used to assist in costing using Aspen Icarus software m dot_gasHT 112.78 tonne
day =

v dot_gasHT 4.041 10 4 × gal

day =

V

gas_tankHT v dot_gasHT 30

⋅ day :=

V

gas_tankHT 4589m 3

=

m

dot_dieselHT 266.5 tonne
day =

v dot_dieselHT 8.381 10 4 × gal

day =

V

diesel_tankHT v dot_dieselHT 30

⋅ day :=

V

diesel_tankHT 9518m 3

=

### 106 LT Scenario Fuel Storage

Gasoline Storage Tank (30 days storage)

Diesel Storage Tank (30 days storage)

Note: the resulting volumes are used to assist in costing using Aspen Icarus m

dot_gasLT 87.12 tonne
day =

v dot_gasLT 3.122 10 4 × gal

day =

V

gas_tankLT v dot_gasLT 30

⋅ day :=

V gas_tankLT 3545 m 3

=

m

dot_dieselLT 205.86 tonne
day =

v dot_dieselLT 6.474 10 4 × gal

day =

V

diesel_tankLT v dot_dieselLT 30

⋅ day :=

V

diesel_tankLT 7352 m 3

=

### 107 LT Gasifier Cost

Source: Larson et al. 2005 in 2003$

$MM

Biomass throughput of 300 tpd

The cost ($MM) of one train at 300 ton per day

$MM

Because 2,205 ton/day we need 7 gasifiers but we can apply the multiple train scaling exponent

$MM C 0_gasifier 6.41 10 6

⋅ := S 0_gasifier 41.7 tonne

hr := S max 120 tonne

hr :=

S gasifierLT 300 ton
day := S gasifierLT 11.34 tonne

hr =

C gasifierLT C 0_gasifier S gasifierLT

tonne

hr 1

S 0_gasifier

tonne

hr ⋅ 

###           f

⋅ :=

C gasifierLT 2.576 10 6 ×=

m train 0.9 :=

C gasifierLTtrain C gasifierLT 7 m trai n

⋅ := C gasifierLTtrain 1.484 10 7 ×= f 0.7 :=

### 108 FT Reactor Costing

Source: Larson et al. 2005 in 2003$

$MM



### HT Scenario

Installed cost $MM (assume 3.6 install factor consistent with Peters et al.)



### LT Scenario

Installed cost $MM (assume 3.6 install factor consistent with Peters et al.) C FT_base 10.5 := f FT2 0.72 := S FT_base 2.52 MMcf

hr :=

M dot_FTHT 13829 kmol

hr := V standard_FTHT M dot_FTHT 22.4

⋅ L

mol :=

V

standard_FTHT 10.939 MMcf

hr =

C

FTHT_reac C FT_base V standard_FTHT

S FT_base 

###       f

FT2

⋅ :=

C

FTHT_reac 30.217 =

M dot_FTLT 11400 kmol

hr := V standard_FTLT M dot_FTLT 22.4

⋅ L

mol :=

V

standard_FTLT 9.018 MMcf

hr =

C

FTLT_reac C FT_base V standard_FTLT

S FT_base 

###       f

FT2

⋅ :=

C

FTLT_reac 26.294 =

### 109 Acid-Gas Removal Area Cost

Source: Phillips et al. 2007 in 2005$

Calculated by adding the input synga s streams to the absorber column



### HT Scenario

LT Scenario S AGR_base 332910 lb

hr := f AGR 0.65 := C AGR_base 5446503 :=

S AGR_HT 2965 2308 + ( ) tonne
day ⋅ := S AGR_HT 484374 lb

hr =

C

AGR_HT C AGR_base S AGR_HT

lb

hr 1

S AGR_base

lb

hr ⋅ 

###           f

AG R

⋅ := C AGR_HT 6949808 =

S AGR_LT 2070 2190 + ( ) tonne
day := S AGR_LT 391321 lb

hr =

C

AGR_LT C AGR_base S AGR_LT

lb

hr 1

S AGR_base

lb

hr ⋅ 

###           f

AG R

⋅ := C AGR_LT 6049946 =

### 110 A500 Hydroprocessing Area Cost

Source: Robinson et al. 2007 in 2007$

Note: "bpsd" is barrels per standard day



### HT Scenario

(from Aspen Plus model)



### Power required for A500

LT Scenario

(from Aspen Plus model)



### Power required for A500 (assumed) AreaCost 0 4000

bpsd := S 0_HY 25000 bpsd :=

C 0_HY AreaCost 0 S 0_HY ⋅ := C 0_HY 100000000 =

m

dot_FTL_HT 428 tonne
day := ρ FTL 750 kg

m

3 :=

v dot_FTL_HT m dot_FTL_HT

ρ FTL := v dot_FTL_HT 3.589 10 3 × bpsd =

C

HY_HT C 0_HY v dot_FTL_HT

bpsd 1

S 0_HY

bpsd ⋅ 

###           f

H Y

⋅ := C HY_HT 2.832 10 7 ×=

Power per_bpsd 15kW hr

⋅ bpsd day

⋅ :=

Power areaHT Power per_bpsd v dot_FTL_HT ⋅ := Power areaHT 2.243 MW =

m

dot_FTL_LT 330.42 tonne
day :=

v dot_FTL_LT m dot_FTL_LT

ρ FTL := v dot_FTL_LT 2.771 10 3 × bpsd =

C

HY_LT C 0_HY v dot_FTL_LT

bpsd 1

S 0_HY

bpsd ⋅ 

###           f

H Y

⋅ := C HY_LT 2.394 10 7 ×=

Power areaLT Power per_bpsd v dot_FTL_LT ⋅ := Power areaLT 1.732MW = f HY 0.65 :=

### 111 Reactors and Catalysts

Fischer-Tropsch Reactor and Cobalt Catalyst

FT reactor volume

Using gas hourly space velocity and actual volume tric flow rate, the volume of the reactor is
determined

(assumed)

(from Aspen Plus model)



### FT catalyst cost

(assumed) (assumed)

Replacement cost of cobalt catalyst GHSV ν 0

V

GHSV FT 100 hr 1

−

⋅ :=

ν rate_actHT 6.298 m 3

s :=

V

FTHT ν rate_actHT

GHSV FT := V FTHT 226.728m 3

= V FTHT 8.007 10 3 × ft 3

=

ν rate_actLT 5.021 m 3

s :=

V

FTLT ν rate_actLT

GHSV FT := V FTLT 180.756m 3

= V FTLT 6.383 10 3 × ft 3

=

Co cost 15

lb := ρ Co 64 lb

ft 3 :=

Co vol_cost Co cost ρ C o ⋅ := Co vol_cost 960 1

ft 3 =

Co total_costHT Co vol_cost V FTH T ⋅ := Co total_costHT 7.687 10 6 ×=

Co total_costLT Co vol_cost V FTL T ⋅ := Co total_costLT 6.128 10 6 ×=

### 112 Water Gas Shift Reactor and Catalyst

Sour WGS reactor volume (HT scenario)

Using gas hourly space velocity and actual volume tric flow rate, the volume of the reactor is
determined

(assumed)

(from Aspen Plus)



### WGS reactor volume (LT scenario)

(from Aspen Plus)



### WGS and SWGS Catalyst Cost

Replacement cost of WGS catalyst GHSV ν 0

V

GHSV WGS 1000 hr 1

−

⋅ :=

ν rate_actSWGS 2.008 m 3

s :=

V

SWGS ν rate_actSWGS

GHSV WGS := V SWGS 7.229m 3

= V SWGS 255.283ft 3

=

ν rate_actWGS 1.834 m 3

s :=

V

WGS ν rate_actWGS

GHSV WGS := V WGS 6.602m 3

= V WGS 233.162ft 3

=

CatCost WGS 8

lb := ρ cat_WGS 56 lb

ft 3 := ρ cat_WGS 897.034 kg

m

3 =

CatCost vol_WGS CatCost WGS ρ cat_WGS ⋅ := CatCost vol_WGS 448 1

ft 3 =

TotalCatCost SWGS CatCost vol_WGS V SWGS ⋅ := TotalCatCost SWGS 114367 =

TotalCatCost WGS CatCost vol_WGS V WGS ⋅ := TotalCatCost WGS 104456 =

### 113 Steam Methane Reformer Reactor and Catalyst (LT Scenario)

(assumed)

(from Aspen Plus model)



### SMR Catalyst Cost

Replacement cost of SMR catalyst GHSV SMR 2600hr 1

− :=

ν rate_actSMR 7.082 m 3

s :=

V

SMR ν rate_actSMR

GHSV SMR := V SMR 9.806m 3

= V SMR 346.29ft 3

=

CatCost SMR 4.67

lb := ρ cat_SMR 64 lb

ft 3 := ρ cat_SMR 1.025 10 3 × kg

m

3 =

CatCost vol_SMR CatCost SMR ρ cat_SMR ⋅ := CatCost vol_SMR 298.88 1

ft 3 =

TotalCatCost SMR CatCost vol_SMR V SM R ⋅ := TotalCatCost SMR 103499 =

### 114 Natural Gas Utility Consumption

Annual natural gas requirement at 5% of yearly operating hours

Natural gas properties



### HT Scenario

(from Aspen Plus model, includes power required for gas turbine air compressor)

(assumed)

Annual natural gas requirement

Average flow rate of natural gas Price (Source: Energy Information Administration)

Cost ng 6.4

1000 ft 3 :=
HHV ng 54 MJ

kg := MW ng 16.04 gm

mol := ρ ng 22.4 L

mol :=

Cost ng ρ ng

MW ng ⋅ 286.335 1

ton =

P required_plantHT 32.813MW :=

Eff ng_to_power 0.35 :=

m

dot_ngHT P required_plantHT

Eff ng_to_power

HHV ng := m dot_ngHT 1.378 10 4 × lb

hr =

M ngHT m dot_ngHT Availability

⋅ 0.05

⋅ := M ngHT 2563 ton =

m

dot_ng_5%HT M ngHT

8760hr := m dot_ng_5%HT 585.14 lb

hr =

### 115 LT Scenario

(from Aspen Plus model, includes power required for gas turbine air compressor)

Annual natural gas requirement

Average flow rate of natural gas P required_plantLT 24.3MW :=

m

dot_ngLT P required_plantLT

Eff ng_to_power

HHV ng := m dot_ngLT 1.02 10 4 × lb

hr =

M ngLT m dot_ngLT Availability

⋅ 0.05

⋅ := M ngLT 1898 ton =

m dot_ng_5%LT M ngLT

8760 hr := m dot_ng_5%LT 433.331 lb

hr =116Appendix D. Process Flow Diagrams117High-Temperature ScenarioFigure D-1. Overall plant area process flow diagram for HT scenario118Figure D-2. Preprocessing area process flow diagram for HT scenario119Figure D-3. Gasification area process flow diagram for HT scenario120Figure D-4. Syngas cleaning area process flow diagram for HT scenario121Figure D-5. Acid gas removal area process flow diagram for HT scenario122Figure D-6. Sulfur recovery area process flow diagram for HT scenario123Figure D-7. Fuel synthesis area process flow diagram for HT scenario124Figure D-8. Hydroprocessing area process flow diagram for HT scenario125Figure D-9. Power generation area process flow diagram for HT scenario126Figure D-10. Air separation unit process flow diagram for HT scenario127Low-Temperature ScenarioFigure D-11. Overall plant area process flow diagram for LT scenario128Figure D-12. Preprocessing area process flow diagram for LT scenario129Figure D-13. Gasification area process flow diagram for LT scenario130Figure D-14. Syngas cleaning area process flow diagram for LT scenario131Figure D-15. Acid gas removal area process flow diagram for LT scenario132Figure D-16. Sulfur recovery process flow diagram for LT scenario133Figure D-17. Fuel synthesis area process flow diagram for LT scenario134Figure D-18. Syngas conditioning area process flow diagram for LT scenario135Figure D-19. Hydroprocessing area process diagram for LT scenario136Figure D-20. Power generation area process flow diagram for LT scenario137Figure D-21. Air separation unit process flow diagram for LT scenario138Appendix E. Stream Data139High-Temperature ScenarioTable E-1. Overall Plant Stream Data for HT ScenarioHT Overall Plant P L 00BMAS P L 08BMAS PL09

C AI R PL17SLAG PL21SGAS PL

3 4SGAS PL4

2 CO2 PL4

3 CO2 PL 4 7SGAS PL 4 8SGAS PL 4 9SGAS PL5

0F

T PL

5 2F

L UE PL

65FGAS PL

71 GASO P L8

1D

I ES P L8 1STM PL8

1W

AT PL

83 SU

L PL8

4S

T M PL

8 8F

L U E P L8 9AI

R PL

90 H YD PL

90OX PL9

0S

T M PL

92NT

G N PL

9 2WAT PL98STM P L 99AI

RTemperature (C)25 90 25 50 1300 62 250 53 45 45 45 35 220 35 36 37 200 40 50 120 273 32 30 149 190 -179 120 200 30Pressure (bar)1.01 1.01 1.01 26.62 26.62 22.75 28.00 3.45 23.58 23.58 23.58 22.20 1.03 22.20 1.03 1.03 1.98 24.82 1.93 1.98 1.00 1.01 1.00 28.00 10.00 1.20 1.98 28.00 1.01Vapor Fraction0.00 0.00 1.00 0.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 0.00 1.00 1.00 0.00 0.00 1.00 0.000.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 0.00 1.00Volume Flow** (m3/sec)0.43 0.14 2.33 0 11.23 3.45 0.07 3.32 0.02 1.640.11 0.01 4.33 0.02 0 0.01 50.53 0.02 0 42.34 45.37 29.07 0.63 0.34 1.30 5.76 4.70 0.02 22.37Mole Flow** (kmol/hr) 1542 513.97 343.36 0 8192 10089 175.05 1541 71.82 5260 355.46 89.49 393.42 74.55 41.08 48.97 9251 3276 13.46 9251 3595 4177 90.27 957.07 1271 3255 1028 2220 3237Mass Flow (tonnes/day) 2667 2222 237.82 114.00 3825 3377 180.00 1585 29.08 2130 143.95 427.14 266.91 52.77 112.62 266.11 4000 1501 7.20 4000 2439 2903 4.37 743.70 549.62 2189 444.44 960.00 2242H2O 666.67 222.22 0 0 988.43 45.53 2.50 21.97 0 0 0 0 27.89 0 0 0 0 1348 4.04 0 233.43 0 0 0 549.62 0 444.44 960.00 0CO0 0 0 0 1457 1818 0.47 4.15 11.24 823.54 55.65 0 0 0 0 0 0 22.10 0 0 0 0 0 0 00000H20 0 0 0 122.88 288.34 0 0 1.83 134.08 9.06 0 0 0 0 0 0 0 0 0 0 0 4.37 0 00000CO20 0 0 0 1184 190.01 175.21 1543 2.40 175.74 11.88 0 40.00 0 0 0 0 128.26 0 0 352.24 0 0 0 00000O20 0 55.97 0 0 0 0 0 0 0 0 0 10.28 0 0 0 0 0 0 0 104.25 672.19 0 700.00 0 0 0 0 527.27N20 0 181.85 0 17.68 0 1.78 15.63 0 0 0 0 181.85 0 0 0 0 0.26 0 0 1715 2194 0 0 0 2189 0 0 1715CH40 0 0 0 0.02 63.41 0.01 0.07 0.87 63.47 4.29 0 0 14.94 0 0 0 0 0 0 0 0 0 0 00000C2H60 0 0 0 0 106.90 0.01 0.10 1.46 107.01 7.23 0 0 0 0 0 0 0 0 0 0 0 0 0 00000C2H400000000000000000000000000000C2H200000000000000000000000000000C30 0 0 0 0 141.45 0.01 0.10 1.93 141.56 9.57 0 0 37.83 0 0 0 0 0 0 0 0 0 0 00000C40 0 0 0 0 167.38 0.02 0.15 2.29 167.55 11.32 0 0 0 0 0 0 0 0 0 0 0 0 0 00000H2S0 0 0 0 4.50 0.03 0 0 0 0 0 0 0 0 0 0 0 1.02 0 0 0 0 0 0 00000NH30 0 0 0 0.11 0 0 0 0 0 0 0 0 0 0 0 0 0.10 0 0 0 0 0 0 00000TAR00000000000000000000000000000SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 3.16 0 0 0 0 0 00000CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 00000CHAR00000000000000000000000000000STEAM0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 4000 0 0 4000 0 0 0 0 00000SO20 0 0 0 0 0 0 0 0 0 0 0 0.02 0 0 0 0 0 0 0 0.08 0 0 0 00000NO200000000000000000000000000000MEA00000000000000000000000000000AIR00000000000000000000000000000WAXES0 0 0 0 0 0 0 0 0 0 0 170.22 0 0 0 0 0 0 0 0 0 0 0 0 00000C50 0 0 0 0 2.29 0 0 0.03 2.29 0.15 15.49 0 0 0 0 0 0 0 0 0 0 0 0 00000C60 0 0 0 0 2.46 0 0 0.03 2.46 0.17 16.65 0 0 0 0 0 0 0 0 0 0 0 0 00000C70 0 0 0 0 1.15 0 0 0.02 1.15 0.08 17.54 0 0 0 0 0 0 0 0 0 0 0 0 00000C80 0 0 0 0 1.18 0 0 0.02 1.18 0.08 17.99 0 0 112.62 0 0 0 0 0 0 0 0 0 00000C90 0 0 0 0 1.11 0 0 0.02 1.19 0.08 18.10 0 0 0 0 0 0 0 0 0 0 0 0 00000C100 0 0 0 0 1.16 0 0 0.02 1.19 0.08 18.12 0 0 0 0 0 0 0 0 0 0 0 0 00000C110 0 0 0 0 0.83 0 0 0.02 1.16 0.08 17.62 0 0 0 0 0 0 0 0 0 0 0 0 00000C120 0 0 0 0 0.13 0 0 0.01 0.54 0.04 17.25 0 0 0 0 0 0 0 0 0 0 0 0 00000C130 0 0 0 0 0.02 0 0 0.01 0.52 0.04 16.70 0 0 0 0 0 0 0 0 0 0 0 0 00000C140 0 0 0 0 0 0 0 0.01 0.50 0.03 16.16 0 0 0 0 0 0 0 0 0 0 0 0 00000C150 0 0 0 0 0 0 0 0.01 0.49 0.03 15.57 0 0 0 0 0 0 0 0 0 0 0 0 00000C160 0 0 0 0 0 0 0 0.01 0.47 0.03 14.93 0 0 0 266.11 0 0 0 0 0 0 0 0 00000C170 0 0 0 0 0 0 0 0 0 0 14.75 0 0 0 0 0 0 0 0 0 0 0 0 00000C180 0 0 0 0 0 0 0 0 0 0 14.05 0 0 0 0 0 0 0 0 0 0 0 0 00000C190 0 0 0 0 0 0 0 0 0 0 13.35 0 0 0 0 0 0 0 0 0 0 0 0 00000C200 0 0 0 0 0 0 0 0 0 0 12.64 0 0 0 0 0 0 0 0 0 0 0 0 00000COS0 0 0 0 0.30 1.20 0 0 0.02 1.12 0.08 0 0 0 0 0 0 0.13 0 0 0 0 0 0 00000AR0 0 0 0 43.70 544.37 0 0 6.87 502.93 33.99 0 6.87 0 0 0 0 1.03 0 0 33.99 37.28 0 43.70 0 0.12000BIOMASS 2000 2000 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 00000ASH00000000000000000000000000000SOOT0 0 0 0 6.00 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 00000SLAG0 0 0 114.00 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 00000**Volumetric and Mole flow values do not include biomass, ash, soot, or slag140Table E-2. Preprocessing Area Stream Data for HT ScenarioHT A100 CH0

0B

MA

S CH02BMAS CH03

B MA

S CH90BMAS DR05

B MA

S DR81STM DR84

S TM D R9

2WA

T G R0

6B

MA

S GR0

8 BMAS G R9

0B

MA

STemperature (C)25 25 25 25 90 200 120 120 90 90 90Pressure (bar) 1.01 1.01 1.01 1.01 1.01 1.98 1.98 1.98 1.01 1.01 1.01Vapor Fraction0.00 0.00 0.00 0.000.00 1.00 1.00 1.00 0.00 0.00 0.00Volume Flow** (m3/sec)0.43 0.43 0.43 0 0.14 50.53 42.34 4.70 0.14 0.14 0Mole Flow** (kmol/hr) 1542 1542 1542 0 513.97 9251 9251 1028 513.97 513.97 0Mass Flow (tonnes/day) 2667 2709 2667 42.20 2222 4000 4000 444.44 2240 2222 17.60H2O 666.67 666.67 666.67 0 222.22 0 0 444.44 222.22 222.22 0CO 0 0 0 0 0 0 0 0 0 0 0H2 0 0 0 0 0 0 0 0 0 0 0CO2 0 0 0 0 0 0 0 0 0 0 0O2 0 0 0 0 0 0 0 0 0 0 0N2 0 0 0 0 0 0 0 0 0 0 0CH4 0 0 0 0 0 0 0 0 0 0 0C2H6 0 0 0 0 0 0 0 0 0 0 0C2H4 0 0 0 0 0 0 0 0 0 0 0C2H2 0 0 0 0 0 0 0 0 0 0 0C3 0 0 0 0 0 0 0 0 0 0 0C4 0 0 0 0 0 0 0 0 0 0 0H2S 0 0 0 0 0 0 0 0 0 0 0NH3 0 0 0 0 0 0 0 0 0 0 0TAR 0 0 0 0 0 0 0 0 0 0 0SULFUR 0 0 0 0 0 0 0 0 0 0 0CARBON 0 0 0 0 0 0 0 0 0 0 0CHAR 0 0 0 0 0 0 0 0 0 0 0STEAM 0 0 0 0 0 4000 4000 0 0 0 0SO2 0 0 0 0 0 0 0 0 0 0 0NO2 0 0 0 0 0 0 0 0 0 0 0MEA 0 0 0 0 0 0 0 0 0 0 0AIR 0 0 0 0 0 0 0 0 0 0 0WAXES 0 0 0 0 0 0 0 0 0 0 0C5 0 0 0 0 0 0 0 0 0 0 0C6 0 0 0 0 0 0 0 0 0 0 0C7 0 0 0 0 0 0 0 0 0 0 0C8 0 0 0 0 0 0 0 0 0 0 0C9 0 0 0 0 0 0 0 0 0 0 0C10 0 0 0 0 0 0 0 0 0 0 0C11 0 0 0 0 0 0 0 0 0 0 0C12 0 0 0 0 0 0 0 0 0 0 0C13 0 0 0 0 0 0 0 0 0 0 0C14 0 0 0 0 0 0 0 0 0 0 0C15 0 0 0 0 0 0 0 0 0 0 0C16 0 0 0 0 0 0 0 0 0 0 0C17 0 0 0 0 0 0 0 0 0 0 0C18 0 0 0 0 0 0 0 0 0 0 0C19 0 0 0 0 0 0 0 0 0 0 0C20 0 0 0 0 0 0 0 0 0 0 0COS 0 0 0 0 0 0 0 0 0 0 0AR 0 0 0 0 0 0 0 0 0 0 0BIOMASS 2000 2042 2000 42.20 2000 0 0 0 2018 2000 17.60ASH 0 0 0 0 0 0 0 0 0 0 0SOOT 0 0 0 0 0 0 0 0 0 0 0SLAG 0 0 0 0 0 0 0 0 0 0 0**Volumetric and Mole flow values do not include biomass, ash, soot, or slag141Table E-3. Gasification Area Stream Data for HT ScenarioHT A200 G S 02 S L A G GS08BMAS GS09CAIR GS11BMAS G S 4 2CO

2 G S5

2 F L UE G S 54 S GA

S G S 8 1S

T M G S 8 4S

T M G S 9 0 OX G S 9 2CO

2 G S 9 8S

TM S 0 6S

G A S SL

0 1S

L AG S L 02 S G ASTemperature (C)50 90 25 90 250 220 45 200 120 149 92 200 203 1300 1300Pressure (bar)26.62 1.01 1.01 1.01 28.00 1.03 23.58 1.98 1.98 28.00 28.00 28.00 25.93 26.62 26.62Vapor Fraction0.00 0.00 1.00 0.53 1.00 1.00 1.00 1.00 1.00 1.00 1.00 0.00 1.00 0.00 1.00Volume Flow** (m3/sec)0 0.14 2.33 30.59 0.07 4.33 0.02 50.53 42.34 0.34 0.04 0.02 3.50 0 11.23Mole Flow** (kmol/hr) 0 513.97 343.36 6974 175.05 393.42 71.82 9251 9251 957.07 158.25 2220 8308 0 8192Mass Flow (tonnes/day) 114.00 2222 237.82 2222 180.00 266.91 29.08 4000 4000 743.70 166.90 960.00 3869 114.00 3825H2O 0 222.22 0 222.22 2.50 27.89 0 0 0 0 0 960.00 1038 0 988.43CO 0 0 0 0 0.47 0 11.24 0 0 0 0.45 0 1457 0 1457H2 0 0 0 101.20 0 0 1.83 0 0 0 0 0 122.88 0 122.88CO2 0 0 0 0 175.21 40.00 2.40 0 0 0 166.45 0 1184 0 1184O2 0 0 55.97 812.60 0 10.28 0 0 0 700.00 0 0 0 0 0N2 0 0 181.85 16.00 1.78 181.85 0 0 0 0 0 0 17.68 0 17.68CH4 0 0 0 0 0.01 0 0.87 0 0 0 0 0 0.02 0 0.02C2H6 0 0 0 0 0.01 0 1.46 0 0 0 0 0 0 0 0C2H4 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C2H2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C3 0 0 0 0 0.01 0 1.93 0 0 0 0 0 0 0 0C4 0 0 0 0 0.02 0 2.29 0 0 0 0 0 0 0 0H2S 0 0 0 0 0 0 0 0 0 0 0 0 4.50 0 4.50NH3 0 0 0 0 0 0 0 0 0 0 0 0 0.11 0 0.11TAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0SULFUR 0 0 0 4.40 0 0 0 0 0 0 0 0 0 0 0CARBON 0 0 0 945.60 0 0 0 0 0 0 0 0 0 0 0CHAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0STEAM 0 0 0 0 0 0 04000 4000 000000SO2 0 0 0 0 0 0.02 0 0 0 0 0 0 0 0 0NO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0MEA 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0AIR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0WAXES000000000000000C5 0 0 0 0 0 0 0.03 0 0 0 0 0 0 0 0C6 0 0 0 0 0 0 0.03 0 0 0 0 0 0 0 0C7 0 0 0 0 0 0 0.02 0 0 0 0 0 0 0 0C8 0 0 0 0 0 0 0.02 0 0 0 0 0 0 0 0C9 0 0 0 0 0 0 0.02 0 0 0 0 0 0 0 0C10 0 0 0 0 0 0 0.02 0 0 0 0 0 0 0 0C11 0 0 0 0 0 0 0.02 0 0 0 0 0 0 0 0C12 0 0 0 0 0 0 0.01 0 0 0 0 0 0 0 0C13 0 0 0 0 0 0 0.01 0 0 0 0 0 0 0 0C14 0 0 0 0 0 0 0.01 0 0 0 0 0 0 0 0C15 0 0 0 0 0 0 0.01 0 0 0 0 0 0 0 0C16 0 0 0 0 0 0 0.01 0 0 0 0 0 0 0 0C17 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C18 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C19 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C20 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0COS 0 0 0 0 0 0 0.02 0 0 0 0 0 0.30 0 0.30AR 0 0 0 0 0 6.87 6.87 0 0 43.70 0 0 43.70 0 43.70BIOMASS 0 2000 0 0 0 0 0 0 0 0 0 0 0 0 0ASH 0 0 0 120.00 0 0 0 0 0 0 0 0 0 0 0SOOT 0 0 0 0 0 0 0 0 0 0 0 0 0 0 6.00SLAG 114.00 0 0 0 0 0 0 0 0 0 0 0 0 114.00 0**Volumetric and Mole flow values do not include biomass, ash, soot, or slag142Table E-4. Syngas Cleaning Area Stream Data for HT ScenarioHT A300 C L 2 5ACG CL25SGAS CL

2 6S

G A S CL

28 S G A S CL29SGAS CL

3 2SGA

S CL

4 0CO2 CL

4 1CO

2 CL

42

CO

2 CL

4 9SGA

S C L81W

A T C L 8 2W

A T CL8

3 SUL C L 83 W AT CL90STM CL

9 2A

I R CL95A I R G S 84STM SG91

S TM S L0

1 H2O SL02SGAS S L 03

H2

O S L 80

H2OTemperature (C)50 203 291 240 60 40 53 53 250 45 30 40 50 60 190 25 50 62 250 203 1300 203 30Pressure (bar)3.45 25.93 24.82 24.82 24.82 24.82 3.45 3.45 28.00 23.58 24.82 24.82 1.93 24.82 10.00 1.01 1.93 22.75 25.86 26.62 26.62 25.93 26.62Vapor Fraction1.00 1.00 1.00 1.00 0.69 1.00 1.00 1.00 1.00 1.00 0.00 0.00 0.00 0.00 1.00 1.00 0.93 1.00 1.00 0.00 1.00 0.00 0.00Volume Flow** (m3/sec)3.69 1.08 2.01 4.54 2.07 1.86 3.700.38 0.07 1.64 0 0.02 0 0.02 1.30 0.01 0.01 3.45 0.54 0.07 11.23 0 0Mole Flow** (kmol/hr) 1728 2576 3847 9579 9579 6419 1716 175.05 175.05 5260 115.64 3276 13.46 2313 1271 2.07 2.58 10089 1271 9251 8192 0 115.64Mass Flow (tonnes/day) 1773 1199 1749 4418 4418 2967 1765 180.00 180.00 2130 50.00 1501 7.20 1000 549.62 1.59 2.32 3377 549.62 4000 3825 6.00 50.00H2O 27.18 321.91 591.14 1308 1308 9.19 24.46 2.50 2.50 0 50.00 1348 4.04 1000 549.62 0 0.45 45.53 549.62 4000 988.43 0 50.00CO 4.62 451.69 15.74 1021 1021 999.01 4.62 0.47 0.47 823.54 0 22.10 0 0 0 0 0 1818 0 0 1457 0 0H2 0 38.09 69.47 154.26 154.26 154.26 0 0 0 134.080 0 0 0 0 0 0 288.34 0 0 122.88 0 0CO2 1718 367.08 1052 1869 1869 1741 1718 175.21 175.21 175.74 0 128.26 0 0 0 0 0 190.01 0 0 1184 0 0O2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 1.59 0.02 0 0 0 0 0 0N2 17.41 5.48 5.48 17.68 17.68 17.43 17.41 1.78 1.78 0 0 0.26 0 0 0 0 0 0 0 0 17.68 0 0CH4 0.08 0.01 0.01 0.02 0.02 0.02 0.08 0.01 0.01 63.47 0 0 0 0 0 0 0 63.41 0 0 0.02 0 0C2H6 0.11 0 0 0 0 0 0.11 0.01 0.01 107.010 0 0 0 0 0 0 106.90 0 0 0 0 0C2H4 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C2H2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C3 0.11 0 0 0 0 0 0.11 0.01 0.01 141.560 0 0 0 0 0 0 141.45 0 0 0 0 0C4 0.17 0 0 0 0 0 0.17 0.02 0.02 167.550 0 0 0 0 0 0 167.38 0 0 0 0 0H2S 3.39 1.40 1.40 4.50 4.50 3.48 0 0 0 0 0 1.02 0 0 0 0 0.03 0.03 0 0 4.50 0 0NH3 0 0.03 0.03 0.11 0.11 0.01 0 0 0 0 0 0.10 0 0 0 0 0 0 0 0 0.11 0 0TAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 3.16 0 0 0 0 0 0 0 0 0 0CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0CHAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0STEAM 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0SO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0NO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0MEA 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0AIR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0WAXES 0 0 000000000000000000000C5 0 0 0 0 0 0 0 0 0 2.29 0 0 0 0 0 0 0 2.29 0 0 0 0 0C6 0 0 0 0 0 0 0 0 0 2.46 0 0 0 0 0 0 0 2.46 0 0 0 0 0C7 0 0 0 0 0 0 0 0 0 1.15 0 0 0 0 0 0 0 1.15 0 0 0 0 0C8 0 0 0 0 0 0 0 0 0 1.18 0 0 0 0 0 0 0 1.18 0 0 0 0 0C9 0.08 0 0 0 0 0 0 0 0 1.19 0 0 0 0 0 0 0.08 1.11 0 0 0 0 0C10 0.03 0 0 0 0 0 0 0 0 1.19 0 0 0 0 0 0 0.03 1.16 0 0 0 0 0C11 0.23 0 0 0 0 0 0 0 0 1.16 0 0 0 0 0 0 0.23 0.83 0 0 0 0 0C12 0.13 0 0 0 0 0 0 0 0 0.54 0 0 0 0 0 0 0.13 0.13 0 0 0 0 0C13 0.04 0 0 0 0 0 0 0 0 0.52 0 0 0 0 0 0 0.04 0.02 0 0 0 0 0C14 0.01 0 0 0 0 0 0 0 0 0.50 0 0 0 0 0 0 0.01 0 0 0 0 0 0C15 0 0 0 0 0 0 0 0 0 0.49 0 0 0 0 0 0 0 0 0 0 0 0 0C16 0 0 0 0 0 0 0 0 0 0.47 0 0 0 0 0 0 0 0 0 0 0 0 0C17 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C18 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C19 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C20 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0COS 0.08 0.09 0.09 0.30 0.30 0.17 0 0 0 1.12 0 0.13 0 0 0 0 0.08 1.20 0 0 0.30 0 0AR 1.22 13.55 13.55 43.70 43.70 42.66 0 0 0 502.93 0 1.03 0 0 0 0 1.22 544.37 0 0 43.70 0 0BIOMASS 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0ASH 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0SOOT 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 6.00 6.00 0SLAG 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0**Volumetric and Mole flow values do not include biomass, ash, soot, or slag143Table E-5. Acid Gas Removal and Sulfur Recovery Areas Stream Data for HT ScenarioHT A300AGR AG0

2 MEAR AG03MEAR A G 04

ME

A R A G 06

ME

A R AG0

8 MEA

L AG0

9 MEA

L AG10MEAL A G 11 ME

AL AG2

3 ACG A G 2 4 ACG A G 2 5A

CG A G 3 1S

G AS AG32SGAS AG3

3 SGAS A G 4 9 SG

A SA300SUL S U2

5A

CG SU26SUL SU

4 0CO2 S U83

S UL S U9

2A

I R S U 93A

I R SU94AIR SU9

5 AIRTemperature (C)58 62 58 86 50 123 96 96 86 50 50 40 58 62 4550 53 53 50 25 100 50 50Pressure (bar)22.75 22.75 22.75 3.45 3.45 3.45 3.45 20.68 3.45 3.45 3.45 24.82 22.75 22.7523.583.45 3.45 3.45 1.93 1.01 2.07 1.93 1.93Vapor Fraction0.00 0.00 0.00 0.01 0.00 0.00 0.00 0.00 0.98 0.83 1.00 1.00 0.93 1.00 1.001.00 0.49 1.00 0.00 1.00 1.00 1.00 0.93Volume Flow** (m3/sec)0.480 0 0.480 1.590 0 0.480 0.470 0.470 4.820 3.690 3.690 1.860 3.410 3.450 1.6403.690 0.010 3.700 0 0.010 0.010 0.010 0.010Mole Flow** (kmol/hr) 84877 699.000 85576 85576 342.110 83848 83848 83848 2070 2070 1728 6419 10788 10089 52601728 11.910 1716 13.460 2.070 2.070 2.070 2.580Mass Flow (tonnes/day) 42560 308.340 42868 42868 155.170 41096 41096 41096 1928 1928 1773 2967 3685 3377 21301773 7.930 1765 7.200 1.590 1.590 1.5902.320H2O33956 297.460 34254 34254 144.090 34226 34226 34226 171.270 171.270 27.180 9.190 342.990 45.530 027.180 2.720 24.460 4.040 0 0 0 0.450CO0 4.640 4.640 4.640 0.010 0.010 0.010 0.010 4.630 4.630 4.620 999.010 1823 1818 823.5404.620 0 4.62000000H20 0 0 0 0 0 0 0 0 0 0 154.260 288.340288.340 134.0800 0 000000CO21725 1.650 1727 1727 8.690 8.690 8.690 8.690 1727 1727 1718 1741 191.660 190.010 175.7401718 0 171800000O20000000000000000 0 0 0 1.590 1.590 1.590 0.020N217.430 0 17.430 17.430 0.020 0.020 0.020 0.020 17.430 17.430 17.410 17.430 0 0 0 17.410 0 17.410 0 0 0 0 0CH40 0.080 0.080 0.080 0 0 0 0 0.080 0.080 0.080 0.020 63.490 63.410 63.470 0.080 0 0.080 0 0 0 0 0C2H60 0.110 0.110 0.110 0 0 0 0 0.110 0.110 0.110 0 107.010106.900 107.0100.110 0 0.11000000C2H400000000000000000000000C2H200000000000000000000000C30 0.110 0.110 0.110 0 0 0 0 0.110 0.110 0.110 0 141.560141.450 141.5600.110 0 0.11000000C40 0.170 0.170 0.170 0 0 0 0 0.170 0.170 0.170 0 167.550167.380 167.5500.170 0 0.17000000H2S 3.450 0 3.450 3.450 0.060 0.060 0.060 0.060 3.450 3.450 3.390 3.480 0.030 0.030 0 3.390 3.390 0 0 0 0 0 0.030NH3 0.010 0 0.010 0.010 0.010 0.010 0.010 0.010 0.010 0.010 0 0.010 0 0 000000000TAR00000000000000000000000SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 00 0 0 3.160 0 0 0 0CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 000000000CHAR00000000000000000000000STEAM00000000000000000000000SO200000000000000000000000NO200000000000000000000000MEA6858 0 6858 6858 0 6858 6858 6858000000000000000AIR00000000000000000000000WAXES0 0 000000000000000000000C50 0 0 0 0 0 0 0 0 0 0 0 2.290 2.290 2.29000000000C60 0 0 0 0 0 0 0 0 0 0 0 2.460 2.460 2.46000000000C70 0 0 0 0 0 0 0 0 0 0 0 1.150 1.150 1.15000000000C80 0 0 0 0 0 0 0 0 0 0 0 1.180 1.180 1.18000000000C90 0.080 0.080 0.080 0.010 0.010 0.010 0.010 0.080 0.080 0.080 0 1.190 1.110 1.190 0.080 0.080 0 0 0 0 0 0.080C100 0.030 0.030 0.030 0 0 0 0 0.030 0.030 0.030 0 1.190 1.160 1.190 0.030 0.030 0 0 0 0 0 0.030C110 0.330 0.330 0.330 0.100 0.100 0.100 0.100 0.320 0.320 0.230 0 1.160 0.830 1.160 0.230 0.230 0 0 0 0 0 0.230C120 0.410 0.410 0.410 0.280 0.280 0.280 0.280 0.410 0.410 0.130 0 0.540 0.130 0.540 0.130 0.130 0 0 0 0 0 0.130C130 0.500 0.500 0.500 0.460 0.460 0.460 0.460 0.500 0.500 0.040 0 0.520 0.020 0.520 0.040 0.040 0 0 0 0 0 0.040C140 0.500 0.500 0.500 0.490 0.490 0.490 0.490 0.500 0.500 0.010 0 0.500 0 0.500 0.010 0.010 0 0 0 0 0 0.010C150 0.490 0.490 0.490 0.490 0.490 0.490 0.490 0.490 0.490 0 0 0.490 0 0.49000000000C160 0.470 0.470 0.470 0.470 0.470 0.470 0.470 0.470 0.470 0 0 0.470 0 0.47000000000C1700000000000000000000000C1800000000000000000000000C1900000000000000000000000C2000000000000000000000000COS0 0.080 0.080 0.080 0 0 0 0 0.080 0.080 0.080 0.170 1.290 1.200 1.120 0.080 0.080 0 0 0 0 0 0.080AR0 1.220 1.220 1.220 0 0 0 0 1.220 1.220 1.220 42.660 545.590544.370 502.9301.220 1.220 00 0 0 0 1.220BIOMASS 0 0 0 0 0 0 0 0 0 0 0 0 0 0 000000000ASH00000000000000000000000SOOT00000000000000000000000SLAG00000000000000000000000**Volumetric and Mole flow values do not include biomass, ash, soot, or slag144Table E-6. Fuel Synthesis Area Stream Data for HT ScenarioHT A400 FS

34 S G A S FS

3 5S

G A S FS

3 6SGA

S FS37SGAS FS40SGAS F S42SGAS FS

43

S G A S FS4

4 FT FS

4 4S

G A S FS45FT FS45SGAS

FS

4 6FT F S46SGAS FS

47

S G A S FS

48

S G A S FS

4 9S

G A S FS50FT FS50SGAS FS51SGAS F S52SGAS F S 6 0W

A T F S 8 9HY

D FS9

0 HYD F S99CONTTemperature (C)62 76 200 200 30 30 417 200 202 35 202 43 0 45 45 45 35 45 50 200 35 30 30 200Pressure (bar)22.75 26.00 26.00 24.97 1.00 1.00 24.97 23.58 24.97 23.58 24.96 23.58 0.00 23.58 23.58 23.58 22.20 23.58 24.96 24.96 22.20 1.00 1.00 24.97Vapor Fraction1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 0.86 1.00 0.00 1.00 1.00 1.00 1.00 0.00 1.00 1.00 1.00 0.00 1.00 1.00 1.00Volume Flow** (m3/sec)3.45 3.16 4.29 4.47 1.26 0.63 0.06 4.84 4.45 2.71 5.86 0.02 2.74 0.02 1.64 0.110.01 1.00 0.96 1.42 0.01 0.63 0.63 0Mole Flow** (kmol/hr) 10089 10089 10089 10089 179.78 89.51 89.51 10438 9998 10438 13198 1575 8887 71.82 5260 355.46 89.49 3199 3199 3199 1483 90.27 90.27 0.04Mass Flow (tonnes/day) 3377 3377 3377 3377 60.18 55.81 55.81 4668 3372 4668 4668 1069 3599 29.08 2130 143.95 427.14 1296 1296 1296 641.34 4.37 4.37 0.03H2O 45.53 45.53 45.53 45.53 0.810.81 0.81 642.07 45.53 642.07 45.53 642.07 0 0 0 0 0 0 0 0 641.34 0 0 0CO 1818 1818 1818 1818 32.40 32.40 32.40 1391 1818 1391 2319 0 1391 11.24 823.54 55.65 0 500.87 500.87 500.87 0 0 0 0H2 288.34 288.34 288.34 288.34 5.14 0.77 0.77 225.34 283.97 225.34 365.52 0 226.52 1.83 134.08 9.06 0 81.55 81.55 81.55 0 4.37 4.37 0CO2 190.01 190.01 190.01 190.01 3.39 3.39 3.39 296.89 190.01 296.89 296.89 0 296.90 2.40 175.74 11.88 0 106.88 106.88 106.88 0 0 0 0O2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0N2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0CH4 63.41 63.41 63.41 63.41 1.13 1.13 1.13 107.33 63.41 107.33 102.02 0 107.23 0.87 63.47 4.29 0 38.60 38.60 38.60 0 0 0 0C2H6 106.90 106.90 106.90 106.90 1.91 1.91 1.91 180.95 106.90 180.95 171.99 0 180.79 1.46 107.01 7.23 0 65.08 65.08 65.08 0 0 0 0C2H4 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C2H2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C3 141.45 141.45 141.45 141.45 2.52 2.52 2.52 239.37 141.45 239.37 227.54 0 239.16 1.93 141.56 9.57 0 86.10 86.10 86.10 0 0 0 0C4 167.38 167.38 167.38 167.38 2.98 2.98 2.98 283.31 167.38 283.31 269.28 0 283.06 2.29 167.55 11.32 0 101.90 101.90 101.90 0 0 0 0H2S 0.03 0.03 0.03 0 0 0 0 0 0 0 0 0 0.00 0 0 0 0 0 0 0 0 0 0 0.03NH3 0 0 0 0 0 0 0 0 0 0 0 0 0.00 0 0 0 0 0 0 0 0 0 0 0TAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0CHAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0STEAM 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0SO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0NO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0MEA 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0AIR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0WAXES 00 0 0 0 0 0 170.22 0 170.22 0170.22 0 0 0 0 170.22 0 0 0 0 0 0 0C5 2.29 2.29 2.29 2.29 0.04 0.04 0.04 19.36 2.2919.36 3.69 15.49 3.87 0.03 2.290.15 15.49 1.39 1.39 1.39 0 0 0 0C6 2.46 2.46 2.46 2.46 0.04 0.04 0.04 20.81 2.4620.81 3.96 16.65 4.16 0.03 2.460.17 16.65 1.50 1.50 1.50 0 0 0 0C7 1.15 1.15 1.15 1.15 0.02 0.02 0.02 19.49 1.1519.49 1.85 17.54 1.95 0.02 1.150.08 17.54 0.70 0.70 0.70 0 0 0 0C8 1.18 1.18 1.18 1.18 0.02 0.02 0.02 19.99 1.1819.99 1.90 17.99 2.00 0.02 1.180.08 17.99 0.72 0.72 0.72 0 0 0 0C9 1.11 1.11 1.11 1.11 0.02 0.02 0.02 20.11 1.1120.11 1.83 18.10 2.01 0.02 1.190.08 18.10 0.72 0.72 0.72 0 0 0 0C10 1.16 1.16 1.16 1.16 0.02 0.02 0.02 20.13 1.16 20.13 1.88 18.12 2.01 0.02 1.19 0.08 18.12 0.72 0.72 0.72 0 0 0 0C11 0.83 0.83 0.83 0.83 0.01 0.01 0.01 19.58 0.83 19.58 1.54 17.62 1.96 0.02 1.16 0.08 17.62 0.71 0.71 0.71 0 0 0 0C12 0.13 0.13 0.13 0.13 0 0 0 18.16 0.13 18.16 0.45 17.25 0.91 0.01 0.54 0.04 17.25 0.33 0.33 0.33 0 0 0 0C13 0.02 0.02 0.02 0.02 0 0 0 17.58 0.02 17.58 0.33 16.70 0.88 0.01 0.52 0.04 16.70 0.32 0.32 0.32 0 0 0 0C14 0 0 0 0 0 0 0 17.01 0 17.01 0.31 16.16 0.85 0.01 0.50 0.03 16.16 0.31 0.31 0.31 0 0 0 0C15 0 0 0 0 0 0 0 16.39 0 16.39 0.30 15.57 0.82 0.01 0.49 0.03 15.57 0.29 0.29 0.29 0 0 0 0C16 0 0 0 0 0 0 0 15.72 0 15.72 0.28 14.93 0.79 0.01 0.47 0.03 14.93 0.28 0.28 0.28 0 0 0 0C17 0 0 0 0 0 0 0 14.75 0 14.75 0 14.75 0 0 0 0 14.75 0 0 0 0 0 0 0C18 0 0 0 0 0 0 0 14.05 0 14.05 0 14.05 0 0 0 0 14.05 0 0 0 0 0 0 0C19 0 0 0 0 0 0 0 13.35 0 13.35 0 13.35 0 0 0 0 13.35 0 0 0 0 0 0 0C20 0 0 0 0 0 0 0 12.64 0 12.64 0 12.64 0 0 0 0 12.64 0 0 0 0 0 0 0COS 1.20 1.20 1.20 1.20 0.02 0.02 0.02 1.88 1.20 1.88 1.88 0 1.88 0.02 1.12 0.08 0 0.68 0.68 0.68 0 0 0 0AR 544.37 544.37 544.37 544.37 9.70 9.70 9.70 850.25 544.37 850.25 850.25 0 849.66 6.87 502.93 33.99 0 305.88 305.88 305.88 0 0 0 0BIOMASS 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0ASH 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0SOOT 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0SLAG 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0**Volumetric and Mole flow values do not include biomass, ash, soot, or slag145Table E-7. Hydroprocessing, Power Generation, and Air Separation Areas Stream Data for HT ScenarioHT A500 HY50FT HY65FGAS HY70GASO HY71GASO H Y 8 0 D I E S H Y 8 1 D IES HY

9 0HY

DHT A600 4 21 602 615 680 CW

2 CWREC HP

S TM1 HP

S TM2HT A700 A I R-1A AIR-A A I R- L N2-OUT O 2 LP

C O2-OUT O2

-O

UT2Temperature (C)35 35 35 36 35 37 3045 1144 273 30 85 73 170 565-170 32 20 16 -177 16 68Pressure (bar)22.20 22.20 22.20 1.03 22.20 1.03 1.0023.58 1.001.00 1.01 7.91 0.30 7.91 173.386.20 1.01 6.30 1.10 1.88 1.10 29.97Vapor Fraction0.00 1.00 0.00 0.00 0.00 0.00 1.001.00 1.00 1.00 1.00 0.00 0.97 0.00 0.001.001.00 1.00 1.00 1.00 0.98 1.00Volume Flow** (m3/sec)0.01 0.02 0 0 0.01 0.01 0.63 0.11 117.68 45.37 22.37 0.01 49.13 0.01 0.01 1.33 29.07 4.49 19.73 1.04 5.58 0.24Mole Flow** (kmol/hr) 89.49 74.55 41.08 41.08 48.97 48.97 90.27355.46 3595 3595 3237 1939 1939 1939 19393968 4177 4177 3255 921.35 921.35 921.35Mass Flow (tonnes/day) 427.14 52.77 112.62 112.62 266.11 266.11 4.37143.95 2439 2439 2242 838.24 838.23 838.23 838.232758 2903 2903 2189 714.26 714.26 714.26H2O0 0 0 0 0 0 0 0 233.43 233.43 0838.24 838.23 838.23 838.230 0 0 0 0 0 0CO0 0 0 0 0 0 0 55.65 0 0 0 0 0 0 0 0 0 0 0 0 0 0H20 0 0 0 0 0 4.37 9.06 0 0 0 0 0 0 0 0 0 0 0 0 0 0CO20 0 0 0 0 0 0 11.88 352.24 352.24 0 0 0 0 0 0 0 0 0 0 0 0O20 0 0 0 0 0 0 0 104.25 104.25 527.27 0 0 0 0638.58 672.19 672.19 0 672.19 672.19 672.19N20 0 0 0 0 0 0 0 1715 1715 1715 0 0 0 0 2084 21942194 2189 4.91 4.91 4.91CH40 14.94 0 0 0 0 0 4.29 0 0 0 0 0 0 0 0 0 0 0 0 0 0C2H60 0 0 0 0 0 0 7.23 0 0 0 0 0 0 0 0 0 0 0 0 0 0C2H40000000 00000000 0000000C2H20000000 00000000 0000000C30 37.83 0 0 0 0 0 9.57 0 0 0 0 0 0 0 0 0 0 0 0 0 0C40 0 0 0 0 0 0 11.32 0 0 0 0 0 0 0 0 0 0 0 0 0 0H2S0000000 00000000 0000000NH30000000 00000000 0000000TAR0000000 00000000 0000000SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0CHAR0000000 00000000 0000000STEAM0000000 00000000 0000000SO20 0 0 0 0 0 0 0 0.08 0.08 0 0 0 0 0 0 0 0 0 0 0 0NO20000000 00000000 0000000MEA0000000 00000000 0000000AIR0000000 00000000 0000000WAXES170.22000000 00000000 0000000C515.49 0 0 0 0 0 0 0.15 0 0 0 0 0 0 0 0 0 0 0 0 0 0C616.65 0 0 0 0 0 0 0.17 0 0 0 0 0 0 0 0 0 0 0 0 0 0C717.54 0 0 0 0 0 0 0.08 0 0 0 0 0 0 0 0 0 0 0 0 0 0C817.99 0 112.62 112.62 0 0 0 0.08 0 0 0 0 0 0 0 0 0 0 0 0 0 0C918.10 0 0 0 0 0 0 0.08 0 0 0 0 0 0 0 0 0 0 0 0 0 0C1018.12 0 0 0 0 0 0 0.08 0 0 0 0 0 0 0 0 0 0 0 0 0 0C1117.62 0 0 0 0 0 0 0.08 0 0 0 0 0 0 0 0 0 0 0 0 0 0C1217.25 0 0 0 0 0 0 0.04 0 0 0 0 0 0 0 0 0 0 0 0 0 0C1316.70 0 0 0 0 0 0 0.04 0 0 0 0 0 0 0 0 0 0 0 0 0 0C1416.16 0 0 0 0 0 0 0.03 0 0 0 0 0 0 0 0 0 0 0 0 0 0C1515.57 0 0 0 0 0 0 0.03 0 0 0 0 0 0 0 0 0 0 0 0 0 0C1614.93 0 0 0 266.11 266.11 0 0.03 0 0 0 0 0 0 0 0 0 0 0 0 0 0C1714.75 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C1814.05 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C1913.35 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C2012.64 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0COS0 0 0 0 0 0 0 0.08 0 0 0 0 0 0 0 0 0 0 0 0 0 0AR0 0 0 0 0 0 0 33.99 33.99 33.99 0 0 0 0 0 35.42 37.28 37.28 0.12 37.16 37.16 37.16BIOMASS 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0ASH0000000 00000000 0000000SOOT0000000 00000000 0000000SLAG0000000 00000000 0000000**Volumetric and Mole flow values do not include biomass, ash, soot, or slag146Low-Temperature ScenarioTable E-8. Overall Plant Stream Data for LT ScenarioHT Overall Plant P L 0 0 BM

A S PL0

8 BMA

S P L 0 9 C A I R PL21SGA S P L 3 4 SGA

S PL 4 1 C O2 PL43CO2 P L 52F

LU

E P L 5 5 SGA

S P L 5 6 F T P L 5 6 SGA

S PL60ASH PL 6 5 FGA

S PL71G A S O P L 8 1 D IES P L 8 1 STM P L81WA T PL83SUL PL84STM P L 88F

LUE P L90HYD PL9

0 O X PL90STM PL 9 2 WA

T PL98STM P L8 9 A I R P L92NTG N P L9 9 A I RTemperature (C)25 90 100 870 62 50 243 200 32 35 32 0 50 50 50 200 40 50 120 344 60 149 300 120 204 32 16 30Pressure (bar)1.01 1.01 1.01 27.55 22.89 3.45 28.00 1.00 22.89 22.89 22.89 1.00 22.89 22.89 22.89 1.98 26.86 2.07 1.98 1.00 1.01 22.00 25.00 1.98 22.001.01 1.10 1.01Vapor Fraction0.00 0.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 0.00 1.00 0.00 1.00 0.00 0.00 1.00 0.00 0.00 1.00 1.00 1.00 1.00 1.00 1.00 0.00 1.00 1.00 1.00Volume Flow** (m3/sec)0.43 0.14 16.89 4.76 2.40 2.84 0.07 22.48 1.50 0.01 0.12 0 0.02 0 0 50.53 0.02 0 42.34 40.84 0.53 0.32 1.15 4.70 0.01 23.17 15.73 17.42Mole Flow** (kmol/hr) 1542 513.97 1985 4939 7066 1334 170.42 2053 4869 69.40 394.75 0 57.67 31.78 37.88 9251 2804 5.56 9251 2863 69.96 722.81 2313 1028814.33 3328 2594 2522Mass Flow (tonnes/day) 2667 2222 1375 2930 2706 1380 180.00 1471 2071 330.42 167.90 118.88 40.83 87.12 205.86 4000 1388 3.10 4000 1955 3.38 561.66 1000 444.44 352.09 2313 1744 1746H2O 666.67 222.22 0 413.42 30.18 20.50 0 27.16 0 0.09 0 0 0 0 0 0 1060 2.40 0 189.23 0 0 1000 444.44 352.09 0 0 0CO 0 0 0 797.86 1575 0 0 0 795.10 0 64.47 0 0 0 0 0 15.27 0 0 0 0 0 0 0 0 0 0 0H2 0 0 0 47.75 168.38 0 0 0 120.71 0 9.79 0 0 0 0 0 0.07 0 0 0 3.38 0 0 0 0 0 0 0CO2 0 0 0 1427 167.80 1359 180.00 239.52 559.11 0 45.33 0 0 0 0 0 263.09 0 0 318.35 0 0 0 0 0 0 0 0O2 0 0 320.19 0 0 0 0 148.99 0 0 0 0 0 0 0 0 0 0 0 80.99 0 528.66 0 0 0 535.65 0 410.78N2 0 0 1055 0 0 0 0 1055 0 0 0 0 0 0 0 0 0 0 0 1336 0 0 0 0 0 1748 1744 1336AR 0 0 0 33.00 411.50 0 0 0 380.61 0 30.86 0 0 0 0 0 0.69 0 0 30.86 0 33.00 0 0 0 29.71 0.10 0CH4 0 0 0 103.81 149.06 0 0 0 52.01 0 4.22 0 11.56 0 0 0 4.58 0 0 0 0 0 0 0 0 0 0 0C2H6 0 0 0 21.82 28.49 0 0 0 13.55 0 1.10 0 0 0 0 0 5.75 0 0 0 0 0 0 0 0 0 0 0C2H4 0 0 0 44.96 36.05 0 0 0 0.36 0 0.03 0 0 0 0 0 8.11 0 0 0 0 0 0 0 0 0 0 0C6H6 0 0 0 4.10 39.50 0 0 0 36.54 0 2.96 0 0 0 0 0 0.98 0 0 0 0 0 0 0 0 0 0 0C3 0 0 0 0 54.18 0 0 0 58.62 0 4.75 0 29.27 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C4 0 0 0 0 34.31 0 0 0 41.73 0 3.38 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0H2S 0 0 0 4.51 0.13 0 0 0 0 0 0 0 0 0 0 0 1.87 0 0 0 0 0 0 0 0 0 0 0NH3 0 0 0 19.09 1.52 0 0 0 0.28 0 0.02 0 0 0 0 0 17.36 0 0 0 0 0 0 0 0 0 0 0TAR 0 0 0 10.57 0 0 0 0 0 0 0 0 0 0 0 0 10.57 0 0 0 0 0 0 0 0 0 0 0SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0STEAM 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 4000 0 0 4000 0 0 0 0 0 0 0 0 0SO2 0 0 0 0 0 0 0 0.30 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0NO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0.06 0 0 0 0 0 0 0 0MEA 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0AIR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0WAXES 0 0 0 0 0 0 0 0 0 131.46 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C5 0 0 0 0 2.21 0 0 0 2.21 11.92 0.18 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C6 0 0 0 0 2.37 0 0 0 2.37 12.82 0.19 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C7 0 0 0 0 1.11 0 0 0 1.11 13.53 0.09 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C8 0 0 0 0 1.14 0 0 0 1.14 13.88 0.09 0 0 87.12 0 0 0 0 0 0 0 0 0 0 0 0 0 0C9 0 0 0 0 1.08 0 0 0 1.15 13.95 0.09 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C10 0 0 0 0 1.13 0 0 0 1.15 13.98 0.09 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C11 0 0 0 0 0.50 0 0 0 0.54 13.86 0.04 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C12 0 0 0 0 0.44 0 0 0 0.53 13.54 0.04 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C13 0 0 0 0 0.34 0 0 0 0.51 13.10 0.04 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C14 0 0 0 0 0.20 0 0 0 0.49 12.57 0.04 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C15 0 0 0 0 0 0 0 0 0.46 11.93 0.04 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C16 0 0 0 0 0.04 0 0 0 0.45 11.48 0.04 0 0 0 205.86 0 0 0 0 0 0 0 0 0 0 0 0 0C17 0 0 0 0 0 0 0 0 0 11.39 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C18 0 0 0 0 0 0 0 0 0 10.85 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C19 0 0 0 0 0 0 0 0 0 10.31 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C20 0 0 0 0 0 0 0 0 0 9.76 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0BIOMASS 2000 2000 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0ASH 0 0 0 1.01 0 0 0 0.01 0 0 0 118.88 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0CHAR 0 0 0 0.81 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0**Volumetric and Mole flow values do not include biomass, ash, or char147Table E-9. Preprocessing Area Stream Data for LT ScenarioHT A100 CH00BM A S CH02BM A S CH03BM A S CH90BM A S DR05

B M AS . DR81

S T M .DR84 S T M .DR92WAT GR06BMAS GR08BM A S GR90BM A STemperature (C)25 25 25 0 90 200 120 120 90 90 0Pressure (bar)1.01 1.01 1.01 0.00 1.01 1.98 1.98 1.98 1.01 1.01 0.00Vapor Fraction0.00 0.00 0.00 0.00 0.00 1.00 1.00 1.00 0.00 0.00 0.00Volume Flow** (m3/sec)0.43 0.43 0.43 0 0.14 50.53 42.34 4.70 0.14 0.14 0Mole Flow** (kmol/hr) 1542 1542 1542 0 513.97 9251 9251 1028 513.97 513.97 0Mass Flow (tonnes/day) 2667 3181 2667 514.02 2222 4000 4000 444.44 3019 2222 796.81H2O 666.67 666.67 666.67 0 222.22 0 0 444.44 222.22 222.22 0CO 0 0 0 0 0 0 0 0 0 0 0H2 0 0 0 0 0 0 0 0 0 0 0CO2 0 0 0 0 0 0 0 0 0 0 0O2 0 0 0 0 0 0 0 0 0 0 0N2 0 0 0 0 0 0 0 0 0 0 0AR 0 0 0 0 0 0 0 0 0 0 0CH4 0 0 0 0 0 0 0 0 0 0 0C2H6 0 0 0 0 0 0 0 0 0 0 0C2H4 0 0 0 0 0 0 0 0 0 0 0C6H6 0 0 0 0 0 0 0 0 0 0 0C3 0 0 0 0 0 0 0 0 0 0 0C4 0 0 0 0 0 0 0 0 0 0 0H2S 0 0 0 0 0 0 0 0 0 0 0NH3 0 0 0 0 0 0 0 0 0 0 0TAR 0 0 0 0 0 0 0 0 0 0 0SULFUR 0 0 0 0 0 0 0 0 0 0 0CARBON 0 0 0 0 0 0 0 0 0 0 0STEAM 0 0 0 0 0 4000 4000 0 0 0 0SO2 0 0 0 0 0 0 0 0 0 0 0NO2 0 0 0 0 0 0 0 0 0 0 0MEA 0 0 0 0 0 0 0 0 0 0 0AIR 0 0 0 0 0 0 0 0 0 0 0WAXES 0 0 0 0 0 0 0 0 0 0 0C5 0 0 0 0 0 0 0 0 0 0 0C6 0 0 0 0 0 0 0 0 0 0 0C7 0 0 0 0 0 0 0 0 0 0 0C8 0 0 0 0 0 0 0 0 0 0 0C9 0 0 0 0 0 0 0 0 0 0 0C10 0 0 0 0 0 0 0 0 0 0 0C11 0 0 0 0 0 0 0 0 0 0 0C12 0 0 0 0 0 0 0 0 0 0 0C13 0 0 0 0 0 0 0 0 0 0 0C14 0 0 0 0 0 0 0 0 0 0 0C15 0 0 0 0 0 0 0 0 0 0 0C16 0 0 0 0 0 0 0 0 0 0 0C17 0 0 0 0 0 0 0 0 0 0 0C18 0 0 0 0 0 0 0 0 0 0 0C19 0 0 0 0 0 0 0 0 0 0 0C20 0 0 0 0 0 0 0 0 0 0 0BIOMASS 2000 2514 2000 514.02 2000 0 0 0 2797 2000 796.81ASH 0 0 0 0 0 0 0 0 0 0 0CHAR 0 0 0 0 0 0 0 0 0 0 0**Volumetric and Mole flow values do not include biomass, ash, or char148Table E-10. Gasification Area Stream Data for LT ScenarioHT A200 C B52FLU E CB61A SH CB62ASH C Y1

4 SGAS CY15CHAR C Y1

6 CH

AR G S0

5 S GAS GS07S G AS G S09C

AI R G S11B

M AS GS19CHAR G S2

0 S GAS GS

4 3 C O2 G S 4 6 CO

2 GS52FLUE GS81STM G S8

4 STM G S90OX G S91STMTemperature (C) 1200 0 0 870 0 0 871 870 100 96 0 870 243 100 200 200 120 149 204Pressure (bar) 1.00 1.00 1.00 27.57 27.57 27.55 28.00 27.58 1.01 1.01 27.55 27.55 28.00 27.58 1.00 1.98 1.98 22.00 22.00Vapor Fraction 1.00 0.00 0.00 1.00 0.00 0.00 1.00 1.00 1.00 0.00 0.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 0.00Volume Flow** (m3/sec)69.99 0 0 2.38 0 0 4.69 4.76 16.89 0 0 4.76 0.07 0.05 22.48 50.53 42.34 0.32 0.01Mole Flow** (kmol/hr) 2053 0 0 2470 0 0 4931 4939 1985 513.97 0 4939 170.42 161.90 2053 9251 9251 722.81 814.33Mass Flow (tonnes/day) 1471 17.82 118.88 1477 95.38 12.09 3136 3145 1375 2222 214.95 2930 180.00 171.00 1471 4000 4000 561.66 352.09H2O 27.16 0 0 206.71 0 0 413.42 413.42 0 222.22 0 413.42 0 0 27.16 0 0 0 352.09CO 0 0 0 398.93 0 0 797.86 797.86 0 0 0 797.86 0 0 0 0 0 0 0H2 0 0 0 23.87 0 0 47.75 47.75 0 0 0 47.75 0 0 0 0 0 0 0CO2 239.52 0 0 713.66 0 0 1418 1427 0 0 0 1427 180.00 171.00 239.52 0 0 0 0O2 148.99 0 0 0 0 0 0 0 320.19 0 0 0 0 0 148.99 0 0 528.66 0N2 1055 0 0 0 0 0 0 0 1055 0 0 0 0 0 1055 0 0 0 0AR 0 0 0 16.50 0 0 33.00 33.00 0 0 0 33.00 0 0 0 0 0 33.00 0CH4 0 0 0 51.90 0 0 103.81 103.81 0 0 0 103.81 0 0 0 0 0 0 0C2H6 0 0 0 10.91 0 0 21.82 21.82 0 0 0 21.82 0 0 0 0 0 0 0C2H4 0 0 0 22.48 0 0 44.96 44.96 0 0 0 44.96 0 0 0 0 0 0 0C6H6 0 0 0 2.05 0 0 4.10 4.10 0 0 0 4.10 0 0 0 0 0 0 0C3 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C4 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0H2S 0 0 0 2.25 0 0 4.51 4.51 0 0 0 4.51 0 0 0 0 0 0 0NH3 0 0 0 9.54 0 0 19.09 19.09 0 0 0 19.09 0 0 0 0 0 0 0TAR 0 0 0 5.29 0 0 10.57 10.57 0 0 0 10.57 0 0 0 0 0 0 0SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0STEAM 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 4000 4000 0 0SO2 0.30 0 0 0 0 0 0 0 0 0 0 0 0 0 0.30 0 0 0 0NO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0MEA 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0AIR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0WAXES 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C5 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C6 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C7 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C8 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C9 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C10 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C11 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C12 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C13 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C14 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C15 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C16 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C17 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C18 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C19 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C20 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0BIOMASS 0 0 0 0 0 0 0 0 0 2000 0 0 0 0 0 0 0 0 0ASH 0.01 17.82 118.88 7.19 52.75 6.69 119.90 119.90 0 0 118.89 1.01 0 0 0.01 0 0 0 0CHAR 0 0 0 5.81 42.63 5.40 96.87 96.87 0 0 96.06 0.81 0 0 0 0 0 0 0**Volumetric and Mole flow values do not include biomass, ash, or char149Table E-11. Syngas Cleaning AreaStream Data for LT ScenarioHT A300 CL01H2O C L 0 2 H 2 O CL

21

SG

AS CL25ACG CL2

5 SG

A S CL25WAT C L 29SGAS C L 40C

O 2 CL4

0 SG

AS CL41CO2 C L 43C

O 2 CL4

9 SG

AS CL80H2O C L 8 0 W AT CL8

1 WAT C L83

SU

L C L 8 3 W AT CL8

4 WAT C L92

AI R CL95AI RTemperature (C)194 150 870 50 194 0 40 50 62 50 243 32 30 30 40 50 98 40 50 50Pressure (bar)27.55 27.55 27.55 3.45 27.55 27.55 26.86 3.45 22.89 3.45 28.00 22.89 27.57 27.57 26.86 2.07 27.55 27.55 2.07 2.07Vapor Fraction0.00 0.00 1.00 1.00 1.00 0.00 1.00 1.00 1.00 1.00 1.00 1.00 0.00 0.00 0.00 0.00 0.00 0.00 1.00 1.00Volume Flow** (m3/sec)0.08 0.08 4.76 3.28 2.03 0 0.98 3.21 2.40 2.84 0.07 1.50 0 0.01 0.02 0 0.07 0.07 0 0.10Mole Flow** (kmol/hr) 13877 13877 4939 1538 5286 0 3638 1505 7066 1334 170.42 4869 346.93 1156 2804 5.56 13877 13877 0.45 27.74Mass Flow (tonnes/day) 6000 6000 2930 1586 3078 1.82 2190 1560 2706 1380 180.00 2071 150.00 500.00 1388 3.10 6000 6000 0.35 22.89H2O 6000 6000 413.42 22.78 563.42 0 3.78 20.50 30.18 20.50 0 0 150.00 500.00 1060 2.40 6000 6000 0 0.27CO 0 0 797.86 3.07 797.86 0 782.59 0 1575 0 0 795.10 0 0 15.27 0 0 0 0 3.07H2 0 0 47.75 0 47.75 0 47.68 0 168.38 0 0 120.71 0 0 0.07 0 0 0 0 0CO2 0 0 1427 1539 1427 0 1164 1539 167.80 1359 180.00 559.11 0 0 263.09 0 0 0 0 0O2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0.35 0N2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0AR 0 0 33.00 1.42 33.00 0 32.31 0 411.50 0 0 380.61 0 0 0.69 0 0 0 0 1.42CH4 0 0 103.81 2.18 103.81 0 99.23 0 149.06 0 0 52.01 0 0 4.58 0 0 0 0 2.18C2H6 0 0 21.82 1.11 21.82 0 16.07 0 28.49 0 0 13.55 0 0 5.75 0 0 0 0 1.11C2H4 0 0 44.96 1.14 44.96 0 36.85 0 36.05 0 0 0.36 0 0 8.11 0 0 0 0 1.14C6H6 0 0 4.10 0.15 4.10 0 3.11 0 39.50 0 0 36.54 0 0 0.98 0 0 0 0 0.15C3 0 0 0 4.29 0 0 0 0 54.18 0 0 58.62 0 0 0 0 0 0 0 4.29C4 0 0 0 6.73 0 0 0 0 34.31 0 0 41.73 0 0 0 0 0 0 0 6.73H2S 0 0 4.51 2.48 4.51 0 2.64 0 0.13 0 0 0 0 0 1.87 0 0 0 0 1.73NH3 0 0 19.09 0.42 19.09 0 1.73 0 1.52 0 0 0.28 0 0 17.36 0 0 0 0 0.42TAR 0 0 10.57 0 10.57 0 0 0 0 0 0 0 0 0 10.57 0 0 0 0 0SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0STEAM 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0SO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0NO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0MEA 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0AIR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0WAXES 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C5 0 0 0 0 0 0 0 0 2.21 0 0 2.21 0 0 0 0 0 0 0 0C6 0 0 0 0 0 0 0 0 2.37 0 0 2.37 0 0 0 0 0 0 0 0C7 0 0 0 0 0 0 0 0 1.11 0 0 1.11 0 0 0 0 0 0 0 0C8 0 0 0 0 0 0 0 0 1.14 0 0 1.14 0 0 0 0 0 0 0 0C9 0 0 0 0.07 0 0 0 0 1.08 0 0 1.15 0 0 0 0 0 0 0 0.07C10 0 0 0 0.02 0 0 0 0 1.13 0 0 1.15 0 0 0 0 0 0 0 0.02C11 0 0 0 0.03 0 0 0 0 0.50 0 0 0.54 0 0 0 0 0 0 0 0.03C12 0 0 0 0.07 0 0 0 0 0.44 0 0 0.53 0 0 0 0 0 0 0 0.07C13 0 0 0 0.10 0 0 0 0 0.34 0 0 0.51 0 0 0 0 0 0 0 0.10C14 0 0 0 0.09 0 0 0 0 0.20 0 0 0.49 0 0 0 0 0 0 0 0.09C15 0 0 0 0 0 0 0 0 0 0 0 0.46 0 0 0 0 0 0 0 0C16 0 0 0 0.02 0 0 0 0 0.04 0 0 0.45 0 0 0 0 0 0 0 0.02C17 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C18 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C19 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C20 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0BIOMASS 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0ASH 0 0 1.01 0 0 1.01 0 0 0 0 0 0 0 0 0 0 0 0 0 0CHAR 0 0 0.81 0 0 0.81 0 0 0 0 0 0 0 0 0 0 0 0 0 0**Volumetric and Mole flow values do not include biomass, ash, or char150Table E-12. Acid Gas Removal and Sulfur Recovery Areas Stream Data for LT ScenarioHT A300AGR AG02MEAR AG

0 3 MEAR AG04MEAR AG

0 6 MEAR AG08MEAL AG

0 9 MEAL AG10MEAL AG

1 1 MEAL AG2

3 AC

G AG

24ACG AG2

5 AC

G AG

2 6 S GAS AG

32S

G AS AG

3 3 S GAS AG49S G ASA300SUL SU25ACG SU26S UL SU

4 0 C O2 SU83S U L SU

9 2 AI R SU95AIRTemperature (C)50 62 50 93 50 123 80 80 93 50 50 40 50 62 3250 50 50 50 50 50Pressure (bar)22.89 22.89 22.89 3.45 3.45 3.45 3.45 26.00 3.45 3.45 3.45 26.86 22.89 22.89 22.893.45 3.45 3.45 2.07 2.07 2.07Vapor Fraction0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 1.00 0.88 1.00 1.00 0.93 1.00 1.001.00 0.82 1.00 0.00 1.00 1.00Volume Flow** (m3/sec)0.31 0 0.31 0.33 0 0.31 0.30 0.30 4.24 3.28 3.28 0.98 2.30 2.40 1.50 3.28 0.06 3.21 0 0 0.10Mole Flow** (kmol/hr) 54067 449.42 54516 54516 207.83 52978 52978 52978 1746 1746 1538 3638 7515 7066 4869 1538 33.30 1505 5.56 0.45 27.74Mass Flow (tonnes/day) 28640 208.71 28848 28848 101.05 27263 27263 27263 1687 1687 1586 2190 2915 2706 2071 1586 25.65 1560 3.10 0.35 22.89H2O 20923 181.16 21104 21104 82.74 21081 21081 21081 105.52 105.52 22.78 3.78 211.34 30.18 0 22.78 2.28 20.50 2.40 0 0.27CO 0 3.07 3.07 3.07 0 0 0 0 3.07 3.07 3.07 782.59 1578 1575 795.10 3.07 3.07 0 0 0 3.07H2 0 0 0 0 0 0 0 0 0 0 0 47.68 168.38 168.38 120.71 0 0 0 0 0 0CO2 1551 4.53 1556 1556 16.15 16.15 16.15 16.15 1556 1556 1539 1164 172.33 167.80 559.11 1539 0 1539 0 0 0O2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0.35 0N2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0AR 0 1.42 1.42 1.42 0 0 0 0 1.42 1.42 1.42 32.31 412.92 411.50 380.61 1.42 1.42 0 0 0 1.42CH4 0 2.19 2.19 2.19 0.01 0.01 0.01 0.01 2.19 2.19 2.18 99.23 151.24 149.06 52.01 2.18 2.18 0 0 0 2.18C2H6 0 1.13 1.13 1.13 0.02 0.02 0.02 0.02 1.13 1.13 1.11 16.07 29.62 28.49 13.55 1.11 1.11 0 0 0 1.11C2H4 0 1.16 1.16 1.16 0.01 0.01 0.01 0.01 1.16 1.16 1.14 36.85 37.21 36.05 0.36 1.14 1.14 0 0 0 1.14C6H6 0 0.15 0.15 0.15 0 0 0 0 0.15 0.15 0.15 3.11 39.65 39.50 36.54 0.15 0.15 0 0 0 0.15C3 0 4.44 4.44 4.44 0.15 0.15 0.15 0.15 4.44 4.44 4.29 0 58.62 54.18 58.62 4.29 4.29 0 0 0 4.29C4 0 7.42 7.42 7.42 0.69 0.69 0.69 0.69 7.41 7.41 6.73 0 41.73 34.31 41.73 6.73 6.73 0 0 0 6.73H2S 2.51 0.01 2.52 2.52 0.04 0.04 0.04 0.04 2.52 2.52 2.48 2.64 0.13 0.13 0 2.48 2.48 0 0 0 1.73NH3 0 0.49 0.49 0.49 0.07 0.07 0.07 0.07 0.49 0.49 0.42 1.73 2.01 1.52 0.28 0.42 0.42 0 0 0 0.42TAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0STEAM 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0SO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0NO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0MEA 6163 0 6163 6163 0 6163 6163 6163 0 0 0 0 0 0 0 0 0 0 0 0 0AIR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0WAXES 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C5 0 0 0 0 0 0 0 0 0 0 0 0 2.21 2.21 2.21 0 0 0 0 0 0C6 0 0 0 0 0 0 0 0 0 0 0 0 2.37 2.37 2.37 0 0 0 0 0 0C7 0 0 0 0 0 0 0 0 0 0 0 0 1.11 1.11 1.11 0 0 0 0 0 0C8 0 0 0 0 0 0 0 0 0 0 0 0 1.14 1.14 1.14 0 0 0 0 0 0C9 0 0.07 0.07 0.07 0 0 0 0 0.07 0.07 0.07 0 1.15 1.08 1.15 0.07 0.07 0 0 0 0.07C10 0 0.02 0.02 0.02 0 0 0 0 0.02 0.02 0.02 0 1.15 1.13 1.15 0.02 0.02 0 0 0 0.02C11 0 0.04 0.04 0.04 0 0 0 0 0.04 0.04 0.03 0 0.54 0.50 0.54 0.03 0.03 0 0 0 0.03C12 0 0.08 0.08 0.08 0.02 0.02 0.02 0.02 0.08 0.08 0.07 0 0.53 0.44 0.53 0.07 0.07 0 0 0 0.07C13 0 0.17 0.17 0.17 0.08 0.08 0.08 0.08 0.17 0.17 0.10 0 0.51 0.34 0.51 0.10 0.10 0 0 0 0.10C14 0 0.29 0.29 0.29 0.20 0.20 0.20 0.20 0.29 0.29 0.09 0 0.49 0.20 0.49 0.09 0.09 0 0 0 0.09C15 0 0.46 0.46 0.46 0.46 0.46 0.46 0.46 0.46 0.46 0 0 0.46 0 0.46 0 0 0 0 0 0C16 0 0.41 0.41 0.41 0.38 0.38 0.38 0.38 0.41 0.41 0.02 0 0.45 0.04 0.45 0.02 0.02 0 0 0 0.02C17 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C18 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C19 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C20 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0BIOMASS 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0ASH 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0CHAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0**Volumetric and Mole flow values do not include biomass, ash, or char151Table E-13. Fuel Synthesis Area Stream Data for LT ScenarioHT A400 FS3

4 SG

AS FS35SGAS FS3

6 SGAS FS37SGAS FS4

0 SGAS FS4

1 G A S FS5

0 SGAS FS51F T FS5

2 F T F S53

SG

AS FS5

4 F T FS5

5 SG

AS FS5

6 F T FS5

6 SG

AS FS57SGAS FS5

8 SG

AS FS59SGAS FS6

0 WAT FS9

0 HYDTemperature (C)62 78 200 200 60 235 200 200 35 32 35 32 35 32 32 41 200 35 60Pressure (bar)22.89 26.50 25.12 25.12 1.01 25.00 24.96 22.89 22.89 22.89 22.89 22.89 22.89 22.89 22.89 25.30 24.96 22.89 1.01Vapor Fraction1.00 1.00 1.00 1.00 0.88 1.00 1.00 1.00 0.72 1.00 0.00 1.00 0.00 1.00 1.00 1.00 1.00 0.00 1.00Volume Flow** (m3/sec)2.40 2.18 4.44 0.07 0.65 0.05 5.02 4.39 2.08 2.02 0.01 1.50 0.01 0.12 0.41 0.38 0.58 0.01 0.53Mole Flow** (kmol/hr) 7066 7066 10153 166.97 97.01 97.01 11399 9268 9268 6579 70.21 4869 69.40 394.75 1316 1316 1316 2619 69.96Mass Flow (tonnes/day) 2706 2706 3705 60.93 57.55 57.55 4261 4261 4261 2798 330.77 2071 330.42 167.90 559.66 559.66 559.66 1132 3.38H2O 30.18 30.18 671.91 11.05 11.05 11.05 671.91 1133 1133 0 0.43 0 0.09 0 0 0 0 1132 0CO 1575 1575 1576 25.92 25.92 25.92 1791 1074 1074 1074 0 795.10 0 64.47 214.89 214.89 214.89 0 0H2 168.38 168.38 242.13 3.98 0.60 0.60 271.37 163.12 163.12 163.12 0 120.71 0 9.79 32.62 32.62 32.62 0 3.38CO2 167.80 167.80 604.44 9.94 9.94 9.94 755.55 755.55 755.55 755.55 0 559.11 0 45.33 151.11 151.11 151.11 0 0O2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0N2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0AR 411.50 411.50 411.50 6.77 6.77 6.77 514.36 514.36 514.36 514.34 0 380.61 0 30.86 102.87 102.87 102.87 0 0CH4 149.06 149.06 52.13 0.86 0.86 0.86 66.19 70.29 70.29 70.29 0 52.01 0 4.22 14.06 14.06 14.06 0 0C2H6 28.49 28.49 7.72 0.13 0.13 0.13 11.38 18.30 18.30 18.30 0 13.55 0 1.10 3.66 3.66 3.66 0 0C2H4 36.05 36.05 0.39 0.01 0.01 0.01 0.49 0.49 0.49 0.49 0 0.36 0 0.03 0.10 0.10 0.10 0 0C6H6 39.50 39.50 39.50 0.65 0.65 0.65 49.38 49.38 49.38 49.38 0 36.54 0 2.96 9.88 9.88 9.88 0 0C3 54.18 54.18 54.18 0.89 0.89 0.89 70.02 79.16 79.16 79.22 0 58.62 0 4.75 15.84 15.84 15.84 0 0C4 34.31 34.31 34.31 0.56 0.56 0.56 45.59 56.43 56.43 56.40 0 41.73 0 3.38 11.28 11.28 11.28 0 0H2S 0.13 0.13 0.01 0 0 0 0.01 0.01 0.01 0.01 0 0 0 0 0 0 0 0 0NH3 1.52 1.52 0.30 0 0 0 0.38 0.38 0.38 0.38 0 0.28 0 0.02 0.08 0.08 0.08 0 0TAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0STEAM 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0SO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0NO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0MEA 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0AIR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0WAXES 0 0 0 0 0 0 0 131.46 131.46 0 131.46 0 131.46 0 0 0 0 0 0C5 2.21 2.21 2.21 0.04 0.04 0.04 2.80 14.91 14.91 2.98 11.92 2.21 11.92 0.18 0.60 0.60 0.60 0 0C6 2.37 2.37 2.37 0.04 0.04 0.04 3.01 16.02 16.02 3.20 12.82 2.37 12.82 0.19 0.64 0.64 0.64 0 0C7 1.11 1.11 1.11 0.02 0.02 0.02 1.41 15.03 15.03 1.50 13.53 1.11 13.53 0.09 0.30 0.30 0.30 0 0C8 1.14 1.14 1.14 0.02 0.02 0.02 1.45 15.42 15.42 1.54 13.88 1.14 13.88 0.09 0.31 0.31 0.31 0 0C9 1.08 1.08 1.08 0.02 0.02 0.02 1.39 15.50 15.50 1.55 13.95 1.15 13.95 0.09 0.31 0.31 0.31 0 0C10 1.13 1.13 1.13 0.02 0.02 0.02 1.44 15.54 15.54 1.55 13.98 1.15 13.98 0.09 0.31 0.31 0.31 0 0C11 0.50 0.50 0.50 0.01 0.01 0.01 0.65 14.59 14.59 0.73 13.86 0.54 13.86 0.04 0.15 0.15 0.15 0 0C12 0.44 0.44 0.44 0.01 0.01 0.01 0.59 14.26 14.26 0.71 13.54 0.53 13.54 0.04 0.14 0.14 0.14 0 0C13 0.34 0.34 0.34 0.01 0.01 0.01 0.47 13.79 13.79 0.69 13.10 0.51 13.10 0.04 0.14 0.14 0.14 0 0C14 0.20 0.20 0.20 0 0 0 0.34 13.23 13.23 0.66 12.57 0.49 12.57 0.04 0.13 0.13 0.13 0 0C15 0 0 0 0 0 0 0.13 12.55 12.55 0.63 11.93 0.46 11.93 0.04 0.13 0.13 0.13 0 0C16 0.04 0.04 0.04 0 0 0 0.16 12.08 12.08 0.60 11.48 0.45 11.48 0.04 0.12 0.12 0.12 0 0C17 0 0 0 0 0 0 0 11.39 11.39 0 11.39 0 11.39 0 0 0 0 0 0C18 0 0 0 0 0 0 0 10.85 10.85 0 10.85 0 10.85 0 0 0 0 0 0C19 0 0 0 0 0 0 0 10.31 10.31 0 10.31 0 10.31 0 0 0 0 0 0C20 0 0 0 0 0 0 0 9.76 9.76 0 9.76 0 9.76 0 0 0 0 0 0BIOMASS 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0ASH 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0CHAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0**Volumetric and Mole flow values do not include biomass, ash, or char152Table E-14. Syngas Conditioning Area Stream Data for LT ScenarioHT A400COND C D35SGAS C D36SGAS C D37SGAS CD38SGAS CD39SGAS CD40SGAS CD41SGAS CD42SGAS CD44SGAS CD45SGAS CD80

ST

M CD81

ST

M CD99CON TTemperature (C)78 150 150 870 870 300 300 474 363 200 300 870 150Pressure (bar)26.50 26.50 26.50 26.50 25.81 25.81 25.81 25.12 25.12 25.12 25.00 27.00 26.50Vapor Fraction1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00Volume Flow** (m3/sec)2.18 2.64 2.64 7.08 10.44 5.24 1.83 2.46 5.98 4.44 1.15 2.26 0Mole Flow** (kmol/hr) 7066 7066 7063 7063 10153 10153 3554 3554 10153 10153 2313 2313 3.12Mass Flow (tonnes/day) 2706 2706 2705 2705 3705 3705 1297 1297 3705 3705 1000 1000 1.34H2O 30.18 30.18 30.18 30.18 850.65 850.65 297.73 118.99 671.91 671.91 1000 1000 0CO 1575 1575 1575 1575 1854 1854 648.81 370.91 1576 1576000H2 168.38 168.38 168.38 168.38 222.13 222.13 77.75 97.75 242.13 242.13000CO2 167.80 167.80 167.80 167.80 167.80 167.80 58.73 495.37 604.44 604.44000O2 0 0 0 0 0 0 0 0 0 0 0 0 0N2 0 0 0 0 0 0 0 0 0 0 0 0 0AR 411.50 411.50 411.50 411.50 411.50 411.50 144.02 144.02 411.50 411.50000CH4 149.06 149.06 149.06 149.06 52.13 52.13 18.25 18.25 52.13 52.13000C2H6 28.49 28.49 28.49 28.49 7.72 7.72 2.70 2.70 7.72 7.72 0 0 0C2H4 36.05 36.05 36.05 36.05 0.39 0.39 0.14 0.14 0.39 0.39 0 0 0C6H6 39.50 39.50 39.50 39.50 39.50 39.50 13.83 13.83 39.50 39.50 0 0 0C3 54.18 54.18 54.18 54.18 54.18 54.18 18.96 18.96 54.18 54.18 0 0 0C4 34.31 34.31 34.31 34.31 34.31 34.31 12.01 12.01 34.31 34.31 0 0 0H2S 0.13 0.13 0.01 0.01 0.01 0.01 0 0 0.01 0.01 0 0 0.12NH3 1.52 1.52 0.30 0.30 0.30 0.30 0.11 0.11 0.30 0.30 0 0 1.21TAR 0 0 0 0 0 0 0 0 0 0 0 0 0SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0STEAM 0 0 0 0 0 0 0 0 0 0 0 0 0SO2 0 0 0 0 0 0 0 0 0 0 0 0 0NO2 0 0 0 0 0 0 0 0 0 0 0 0 0MEA 0 0 0 0 0 0 0 0 0 0 0 0 0AIR 0 0 0 0 0 0 0 0 0 0 0 0 0WAXES 0 0 0 0 0 0 0 0 0 0 0 0 0C5 2.21 2.21 2.21 2.21 2.21 2.21 0.77 0.77 2.21 2.21 0 0 0C6 2.37 2.37 2.37 2.37 2.37 2.37 0.83 0.83 2.37 2.37 0 0 0C7 1.11 1.11 1.11 1.11 1.11 1.11 0.39 0.39 1.11 1.11 0 0 0C8 1.14 1.14 1.14 1.14 1.14 1.14 0.40 0.40 1.14 1.14 0 0 0C9 1.08 1.08 1.08 1.08 1.08 1.08 0.38 0.38 1.08 1.08 0 0 0C10 1.13 1.13 1.13 1.13 1.13 1.13 0.40 0.40 1.13 1.13 0 0 0C11 0.50 0.50 0.50 0.50 0.50 0.50 0.18 0.18 0.50 0.50 0 0 0C12 0.44 0.44 0.44 0.44 0.44 0.44 0.16 0.16 0.44 0.44 0 0 0C13 0.34 0.34 0.34 0.34 0.34 0.34 0.12 0.12 0.34 0.34 0 0 0C14 0.20 0.20 0.20 0.20 0.20 0.20 0.07 0.07 0.20 0.20 0 0 0C15 0 0 0 0 0 0 0 0 0 0 0 0 0C16 0.04 0.04 0.04 0.04 0.04 0.04 0.01 0.01 0.04 0.04 0 0 0C17 0 0 0 0 0 0 0 0 0 0 0 0 0C18 0 0 0 0 0 0 0 0 0 0 0 0 0C19 0 0 0 0 0 0 0 0 0 0 0 0 0C20 0 0 0 0 0 0 0 0 0 0 0 0 0BIOMASS 0 0 0 0 0 0 0 0 0 0 0 0 0ASH 0 0 0 0 0 0 0 0 0 0 0 0 0CHAR 0 0 0 0 0 0 0 0 0 0 0 0 0**Volumetric and Mole flow values do not include biomass, ash, or char153Table E-15. Hydroprocessing, Power Generation, and Air Separation Areas Stream Data for LT ScenarioHT A500 H Y 5 6 FT HY65FGAS H Y 7 0 GA

SO HY7

1 GA

S O H Y 80D

I ES HY8

1 DI E S H Y90H

Y DA600 4 2 1 6 02 615 6 8 0 CW2 CW

R E C H P W1 HPSTMA700 AIR

- 1 A A IR -A AI R -L N 2 - OU

T O2

L P C O2-OU T O 2 -OUT

2Temperature (C)35 50 50 50 50 50 6032 1161 344 30 85 23 170 393-170 32 20 16 35 16 55Pressure (bar)22.89 22.89 22.89 22.89 22.89 22.89 1.0122.89 1.00 1.00 1.01 7.91 0.02 7.91 173.386.20 1.01 6.30 1.10 23.08 1.10 27.00Vapor Fraction0.00 1.00 0.00 0.00 0.00 0.00 1.001.00 1.00 1.00 1.00 0.00 0.89 0.00 1.001.00 1.00 1.00 1.00 1.00 1.00 1.00Volume Flow** (m3/sec)0.01 0.02 0 0 0 0 0.53 0.12 94.88 40.84 17.42 0.01 442.16 0 0.03 1.06 23.17 3.58 15.73 0.22 4.45 0.21Mole Flow** (kmol/hr) 69.40 57.67 31.78 31.78 37.88 37.88 69.96 394.75 2863 2863 2522 1709 1709 1709 1709 3162 3328 3328 2594 734.20 734.20 734.20Mass Flow (tonnes/day) 330.42 40.83 87.12 87.12 205.86 205.86 3.38 167.90 1955 1955 1746 740.10 740.10 740.10 740.10 2198 2313 2313 1744 569.18 569.18 569.18H2O 0.09 0 0 0 0 0 0 0 189.23 189.23 0 740.10 740.10 740.10 740.10 0 0 0 0 0 0 0CO 0 0 0 0 0 0 0 64.47 0 0 0 0 0 0 0 0 0 0 0 0 0 0H2 0 0 0 0 0 0 3.38 9.79 0 0 0 0 0 0 0 0 0 0 0 0 0 0CO2 0 0 0 0 0 0 0 45.33 318.35 318.35 0 0 0 0 0 0 0 0 0 0 0 0O2 0 0 0 0 0 0 0 0 80.99 80.99 410.78 0 0 0 0 508.87 535.65 535.65 0 535.65 535.65 535.65N2 0 0 0 0 0 0 0 0 1336 1336 1336 0 0 0 0 1661 1748 1748 1744 3.91 3.91 3.91AR 0 0 0 0 0 0 0 30.86 30.86 30.86 0 0 0 0 0 28.22 29.71 29.71 0.10 29.61 29.61 29.61CH4 0 11.56 0 0 0 0 0 4.22 0 0 0 0 0 0 0 0 0 0 0 0 0 0C2H6 0 0 0 0 0 0 0 1.10 0 0 0 0 0 0 0 0 0 0 0 0 0 0C2H4 0 0 0 0 0 0 0 0.03 0 0 0 0 0 0 0 0 0 0 0 0 0 0C6H6 0 0 0 0 0 0 0 2.96 0 0 0 0 0 0 0 0 0 0 0 0 0 0C3 0 29.27 0 0 0 0 0 4.75 0 0 0 0 0 0 0 0 0 0 0 0 0 0C4 0 0 0 0 0 0 0 3.38 0 0 0 0 0 0 0 0 0 0 0 0 0 0H2S 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0NH3 0 0 0 0 0 0 0 0.02 0 0 0 0 0 0 0 0 0 0 0 0 0 0TAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0STEAM 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0SO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0NO2 0 0 0 0 0 0 0 0 0.06 0.06 0 0 0 0 0 0 0 0 0 0 0 0MEA 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0AIR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0WAXES 131.46 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C5 11.92 0 0 0 0 0 0 0.18 0 0 0 0 0 0 0 0 0 0 0 0 0 0C6 12.82 0 0 0 0 0 0 0.19 0 0 0 0 0 0 0 0 0 0 0 0 0 0C7 13.53 0 0 0 0 0 0 0.09 0 0 0 0 0 0 0 0 0 0 0 0 0 0C8 13.88 0 87.12 87.12 0 0 0 0.09 0 0 0 0 0 0 0 0 0 0 0 0 0 0C9 13.95 0 0 0 0 0 0 0.09 0 0 0 0 0 0 0 0 0 0 0 0 0 0C10 13.98 0 0 0 0 0 0 0.09 0 0 0 0 0 0 0 0 0 0 0 0 0 0C11 13.86 0 0 0 0 0 0 0.04 0 0 0 0 0 0 0 0 0 0 0 0 0 0C12 13.54 0 0 0 0 0 0 0.04 0 0 0 0 0 0 0 0 0 0 0 0 0 0C13 13.10 0 0 0 0 0 0 0.04 0 0 0 0 0 0 0 0 0 0 0 0 0 0C14 12.57 0 0 0 0 0 0 0.04 0 0 0 0 0 0 0 0 0 0 0 0 0 0C15 11.93 0 0 0 0 0 0 0.04 0 0 0 0 0 0 0 0 0 0 0 0 0 0C16 11.48 0 0 0 205.86 205.86 0 0.04 0 0 0 0 0 0 0 0 0 0 0 0 0 0C17 11.39 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C18 10.85 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C19 10.31 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0C20 9.76 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0BIOMASS 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0ASH 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0CHAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0**Volumetric and Mole flow values do not include biomass, ash, or charF1147

- E( 10 /200 8 ) 

### REPORT DOCUMENTATION PAGE Form Approved
OMB No. 0704

- 0188

The public reporting burden for this collection of information is estimated to average 1 hour per response, including the tim e for reviewing instructions, searching existing data sources, gathering and maintaining the data needed, and completing and reviewing the collection of information. Send comments regarding this bur den estimate or any other aspect of this collection of information, including suggestions for reducing the burden, to Department o f Defense, Executive Services and Communications Directorate (0704

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1. REPORT DATE (DD

- MM

- YYYY) November 2010 2. REPORT TYPE Technical Report 3. DATES COVERED (From

- To)

4. TITLE AND SUBTITLE
Techno

- Economic Analysis of Biofuels Production Based on
Gasification 5a. CONTRACT NUMBER DE

- AC36

- 08

- GO 28308

5b. GRANT NUMBER

5c. PROGRAM ELEMENT NUMBER

6. AUTHOR(S)
R

. M. Swanson, J

. A. Satrio, R

. C. Brown, A

. Platon, and D

. D. Hsu 5d. PROJECT NUMBER NREL/ TP

- 6A2 0

- 4658 7

5e. TASK NUMBER BB07.7510

5f. WORK UNIT NUMBER

7. PERFORMING ORGANIZATION NAME(S) AND ADDRESS(ES)
National Renewable Ener gy Laboratory
1617 Cole Blvd. Golden, CO 80401

- 3393 8. PERFORMING ORGANIZATION
REPORT NUMBER NREL/ TP

- 6A2 0

- 4658 7

9. SPONSORING/MONITORING AGENCY NAME(S) AND ADDRESS(ES) 10. SPONSOR/MONITOR'S ACRONYM(S)
NREL

11. SPONSORING/MONITORING
AGENCY REPORT NUMBER

12. DISTRIBUTION AVAILABILITY STATEMENT National Technical Information Service U.S. Department of Commerce 5285 Port Royal Road Springfield, VA 22161

13. SUPPLEMENTARY NOTES

14. ABSTRACT (Maximum 200 Words) This study compar es capital and production costs of two biomass

- to

- liquid production plants based on gasification. The first biorefinery scenario is an oxygen

- fed, low

- temperature (870°C), non

- slagging, fluidized bed gasifier. The second scenario is an oxygen

- fed, high

- tem perature (1,300°C), slagging, entrained flow gasifier. Both are followed by catalytic Fischer

- Tropsch synthesis and hydroprocessing to naphtha

- range (gasoline blend stock) and distillate

- range (diesel blend stock) liquid fractions. Process modeling softwar e (Aspen Plus) is utilized to organize the mass and energy streams and cost estimation software is used to generate equipment costs. Economic analysis is performed to estimate the capital investment and operating costs. Results show that the total capital investment required for nth plant scenarios is $610 million and $500 million for high

- temperature and low

- temperature scenarios, respectively. Product value (PV) for the high

- temperature and low

- temperature scenarios is estimated to be $4.30 and $4.80 per gallon of gasoline equivalent (GGE), respectively, based on a feedstock cost of $75 per dry short ton. Sensitivity analysis is also performed on process and economic parameters. This analysis shows that total capital investment and feedstock cost are among the most influential parameters affecting the PV.

15. SUBJECT TERMS biomass; biofuels; gasification; Fischer

- Tropsch; techno

- economic analysis; biomass

- to

- liquid; thermochemical

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