A national laboratory of the U.S. Department of Energy

Office of Energy Efficiency & Renewable Energy

National Renewable Energy Laboratory 

Innovation for Our Energy Future 

Thermochemical Ethanol via 

Indirect Gasification and Mixed 

Alcohol Synthesis of 

Lignocellulosic Biomass 

S. Phillips, A. Aden, J. Jechura, and D. Dayton 

National Renewable Energy Laboratory 

T. Eggeman 

Neoterics International, Inc.

Technical Report 

NREL/TP-510-41168 

April 2007 

NREL is operated by Midwest Research Institute ● Battelle Contract No. DE-AC36-99-GO10337 
Thermochemical Ethanol via 

Indirect Gasification and Mixed 

Alcohol Synthesis of 

Lignocellulosic Biomass 

S. Phillips, A. Aden, J. Jechura, and D. Dayton 

National Renewable Energy Laboratory 

T. Eggeman 

Neoterics International, Inc.

Prepared under Task No. BB07.3710 

Technical Report 

NREL/TP-510-41168 

April 2007 

National Renewable Energy Laboratory

1617 Cole Boulevard, Golden, Colorado 80401-3393 

303-275-3000 • www.nrel.gov 

Operated for the U.S. Department of Energy 

Office of Energy Efficiency and Renewable Energy 

by Midwest Research Institute • Battelle 

Contract No. DE-AC36-99-GO10337
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1. Executive Summary 

This work addresses a policy initiative by the Federal Administration to apply United States 

Department of Energy (DOE) research to broadening the country’s domestic production of 

economic, flexible, and secure sources of energy fuels. President Bush stated in his 2006 State of 

the Union Address: “America is addicted to oil.” To reduce the Nation’s future demand for oil, 

the President has proposed the Advanced Energy Initiative which outlines significant new 

investments and policies to change the way we fuel our vehicles and change the way we power 

our homes and businesses. The specific goal for biomass in the Advanced Energy Initiative is to 

foster the breakthrough technologies needed to make cellulosic ethanol cost-competitive with 

corn-based ethanol by 2012. 

In previous biomass conversion design reports by the National Renewable Energy Laboratory 

(NREL), a benchmark for achieving production of ethanol from cellulosic feedstocks that would 

be “cost competitive with corn-ethanol” has been quantified as $1.07 per gallon ethanol 

minimum plant gate price. 

This process design and technoeconomic evaluation addresses the conversion of biomass to 

ethanol via thermochemical pathways that are expected to be demonstrated at the pilot-unit level 

by 2012. This assessment is unique in its attempt to match up: 

• Currently established and published technology. 

• Technology currently under development or shortly to be under development from DOE 

Office of Biomass Program funding. 

• Biomass resource availability in the 2012 time frame consistent with the Billion Ton 

Vision study. 

Indirect steam gasification was chosen as the technology around which this process was 

developed based upon previous technoeconomic studies for the production of methanol and 

hydrogen from biomass. The operations for ethanol production are very similar to those for 

methanol production (although the specific process configuration will be different). The general 

process areas include: feed preparation, gasification, gas cleanup and conditioning, and alcohol 

synthesis & purification. 

The cost of ethanol as determined in this assessment was derived using technology that has been 

developed and demonstrated or is currently being developed as part of the OBP research 

program. Combined, all process, market, and financial targets in the design represent what must 

be achieved to obtain the reported $1.01 per gallon, showing that ethanol from a thermochemical 

conversion process has the possibility of being produced in a manner that is “cost competitive 

with corn-ethanol” by 2012. This analysis has demonstrated that forest resources can be 

converted to ethanol in a cost competitive manner. This allows for greater flexibility in 

converting biomass resources to make stated volume targets by 2030. 

i 
Table of Contents 

1. Executive Summary....................................................................................................... i 

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

2.1. Analysis Approach........................................................................................................ 6 

2.2. Process Design Overview ........................................................................................... 10 

2.3. Feedstock and Plant Size ............................................................................................ 12 

3. Process Design.............................................................................................................14 

3.1. Process Design Basis .................................................................................................. 14 

3.2. Feed Handling and Drying – Area 100 ....................................................................... 14 

3.3. Gasification – Area 200 .............................................................................................. 15 

3.4. Gas Cleanup and Conditioning – Area 300 ................................................................ 17 

3.5. Alcohol Synthesis – Area 400..................................................................................... 20 

3.6. Alcohol Separation – Area 500................................................................................... 25 

3.7. Steam System and Power Generation – Area 600 ...................................................... 26 

3.8. Cooling Water and Other Utilities – Area 700 ........................................................... 28 

3.9. Additional Design Information................................................................................... 29 

3.10. Pinch Analysis ............................................................................................................ 29 

3.11. Energy Balance ........................................................................................................... 30 

3.12. Water Issues................................................................................................................ 34 

4. Process Economics.......................................................................................................35 

4.1. Capital Costs............................................................................................................... 35 

4.2. Operating Costs........................................................................................................... 38 

4.3. Value of Higher Alcohol Co-Products........................................................................ 41 

4.4. Minimum Ethanol Plant Gate Price ............................................................................ 42 

5. Process Economics, Sensitivity Analyses, and Alternate Scenarios............................43 

5.1. Financial Scenarios..................................................................................................... 45 

5.2. Feedstocks................................................................................................................... 46 

5.3. Thermal Conversion.................................................................................................... 50 

5.4. Clean-Up & Conditioning........................................................................................... 50 

5.5. Fuels Synthesis............................................................................................................ 50 

5.6. Markets ....................................................................................................................... 50 

6. Conclusions..................................................................................................................51 

7. Future Work.................................................................................................................51 

8. References....................................................................................................................53 

ii 
List of Figures 

Figure 1. U.S. list prices for ethanol............................................................................................. 2 

Figure 2. Estimated capital intensities for biomass-to-methanol processes................................. 5 

Figure 3. Approach to process analysis........................................................................................ 6 

Figure 4. Chemical Engineering Magazine’s plant cost indices.................................................. 9 

Figure 5. Block flow diagram..................................................................................................... 10 

Figure 6. Expected availability of biomass................................................................................ 13 

Figure 7. Pinch analysis composite curve .................................................................................. 30 

Figure 8. Cost contribution details from each process area ....................................................... 43 

Figure 9. Effect of cost year on MESP....................................................................................... 44 

Figure 10. Results of sensitivity analyses .................................................................................... 45 

Figure 11. Sensitivity analysis of biomass ash content................................................................ 47 

Figure 12. Sensitivity analysis of biomass moisture content ....................................................... 48 

Figure 13. Sensitivity analysis of raw syngas diverted for heat and power due to biomass 

moisture content.......................................................................................................... 49 

List of Tables 

Table 1. Chemical Engineering Magazine’s Plant Cost Indices................................................. 8 

Table 2. Ultimate Analysis of Hybrid Poplar Feed................................................................... 13 

Table 3. Gasifier Operating Parameters, Gas Compositions, and Efficiencies......................... 16 

Table 4. Current and Target Design Performance of Tar Reformer......................................... 17 

Table 5. Target Design Tar Reformer Conditions and Outlet Gas Composition...................... 18 

Table 6. Process Conditions for Mixed Alcohols Synthesis..................................................... 21 

Table 7. System of Reactions for Mixed Alcohol Synthesis .................................................... 23 

Table 8. Mixed Alcohol Reaction Performance Results........................................................... 23 

Table 9. Mixed Alcohol Product Distributions......................................................................... 24 

Table 10. Plant Power Requirements.......................................................................................... 27 

Table 11. Utility and Miscellaneous Design Information........................................................... 29 

Table 12. Overall Energy Analysis (LHV basis)........................................................................ 33 

Table 13. Process Water Demands for Thermochemical Ethanol .............................................. 34 

Table 14. General Cost Factors in Determining Total Installed Equipment Costs..................... 35 

Table 15. Cost Factors for Indirect Costs ................................................................................... 36 

Table 16. Feed Handling & Drying and Gasifier & Gas Clean Up Costs from the Literature 

Scaled to 2,000 tonne/day plant.................................................................................. 36 

Table 17. System Design Information for Gasification References ........................................... 37 

Table 18. Variable Operating Costs............................................................................................ 38 

Table 19. Labor Costs................................................................................................................. 39 

Table 20. Other Fixed Costs ....................................................................................................... 40 

Table 21. Salary Comparison...................................................................................................... 41 

Table 22. Economic Parameters.................................................................................................. 42 

iii 
2. Introduction 

This work addresses a policy initiative by the Federal Administration to apply United States 

Department of Energy (DOE) research to broadening the country’s domestic production of 

economic, flexible, and secure sources of energy fuels. President Bush stated in his 2006 State of 

the Union Address: “America is addicted to oil.” [1] To reduce the Nation’s future demand for 

oil, the President has proposed the Advanced Energy Initiative [2] which outlines significant 

new investments and policies to change the way we fuel our vehicles and change the way we 

power our homes and businesses. The specific goal for biomass in the Advanced Energy 

Initiative is to foster the breakthrough technologies needed to make cellulosic ethanol costcompetitive with corn-based ethanol by 2012. 

In previous biomass conversion design reports by the National Renewable Energy Laboratory 

(NREL), a benchmark for achieving production of ethanol from cellulosic feedstocks that would 

be “cost competitive with corn-ethanol” has been quantified as $1.07 per gallon ethanol 

minimum plant gate price [3] (where none of these values have been adjusted to a common cost 

year). The value can be put in context with the historic ethanol price data as shown in Figure 1

[4]. The $1.07 per gallon value represents the low side of the historical fuel ethanol prices. Given 

this historical price data, it is viewed that cellulosic ethanol would be commercially viable if it 

was able to meet a minimum return on investment selling at this price. 

This is a cost target for this technology; it does not reflect NREL’s assessment of where the 

technology is today. Throughout this report, two types of data will be shown: results which have 

been achieved presently in a laboratory or pilot plant, and results that are being targeted for 

technology improvement several years into the future. Only those targeted for the 2012 

timeframe are included in this economic evaluation. Other economic analyses that attempt to 

reflect the current “state of technology” are not presented here. 

 1 
0

50

100

150

200

250

300

350

1960 1965 1970 1975 1980 1985 1990 1995 2000 2005 2010

¢ per gallon

Fuel Alcohol

Ethyl Alchohol

Specially Denatured Alcohol

$1.07 Reference

Figure 1. U.S. list prices for ethanola

Conceptual process designs and associated design reports have previously been done by NREL 

for converting cellulosic biomass feedstock to ethanol via Biochemical pathways. Two types of 

biomass considered have been yellow poplar [5] and corn stover. [3] These design reports have 

been useful to NREL and DOE program management for two main reasons. First of all, they 

enable comparison of research and development projects. A conceptual process design helps to 

direct research by establishing a benchmark to which other process configurations can be 

compared. The anticipated results of proposed research can be translated into design changes; the 

economic impact of these changes can then be determined and this new design can be compared 

to the benchmark case. Following this procedure for several proposed research projects allows 

DOE to make competitive funding decisions based on which projects have the greatest potential 

to lower the cost of ethanol production. Complete process design and economics are required for 

such comparisons because changes in performance in one research area may have significant 

impacts in other process areas not part of that research program (e.g., impacts in product 

recovery or waste treatment). The impacts on the other areas may have significant and 

unexpected impacts on the overall economics. 

Secondly, they enable comparison of ethanol production to other fuels. A cost of production has 

also been useful to study the potential ethanol market penetration from technologies to convert 

lignocellulosic biomass to ethanol. The cost estimates developed must be consistent with 

 a

 The curve marked “Ethyl Alcohol” is for 190 proof, USP, tax-free, in tanks, delivered to the East Coast. That 

marked “Specially Denatured Alcohol” is for SDA 29, in tanks, delivered to the East Coast, and denatured with 

ethyl acetate. That marked “Fuel Alcohol” is for 200 proof, fob works, bulk, and denatured with gasoline. 

 2 
applicable engineering, construction, and operating practices for facilities of this type. The 

complete process (including not only industry-standard process components but also the newly 

researched areas) must be designed and their costs determined. 

Following the methodology of the biochemical design reports, this process design and technoeconomic evaluation addresses the conversion of biomass to ethanol via thermochemical (TC) 

pathways that are expected to be demonstrated at the pilot-unit level by 2012. This assessment is 

unique in its attempt to match up: 

• Currently established and published technology. 

• Technology currently under development or shortly to be under development from DOE 

Office of Biomass Program (OBP) funding. (See Appendix B for these research targets 

and values.) 

• Biomass resource availability in the 2012 time frame consistent with the Billion Ton 

Vision study [6]. 

This process design and associated report provides a benchmark for the Thermochemical 

Platform just as the Aden et al. report [3] has been used as a benchmark for the Biochemical 

Platform since 2002. It is also complementary to gasification-based conversion assessments done 

by NREL and others. This assessment directly builds upon an initial analysis for the TC 

production of ethanol and other alcohol co-products [7, 8], which, in turn, was based upon a 

detailed design and economic analysis for the production of hydrogen from biomass.[9] This 

design report is also complementary to other studies being funded by the DOE OBP, including 

the RBAEF (Role of Biomass in America’s Energy Future) study [10]. However, the RBAEF 

study differs in many ways from this study. For example, RBAEF is designed for a further time 

horizon than 2012. It is based on a different feedstock, switchgrass, and it considers a variety of 

thermochemical product options, including ethanol, power and Fischer-Tropsch liquids [11]. 

Other notable gasification studies have been completed by Larsen at Princeton University, 

including a study examining the bioproduct potential of Kraft mill black liquor based upon 

gasification [12]. 

Indirect steam gasification was chosen as the technology around which this process was 

developed based upon previous technoeconomic studies for the production of methanol and 

hydrogen from biomass [13]. The sub-process operations for ethanol production are very similar 

to those for methanol production (although the specific process configuration will be different). 

The general process areas include: feed preparation, gasification, gas cleanup and conditioning, 

and alcohol synthesis & purification. 

Gasification involves the devolatilization and conversion of biomass in an atmosphere of steam 

and/or oxygen to produce a medium-calorific value gas. There are two general classes of 

gasifiers. Partial oxidation (POX) gasifiers (directly-heated gasifiers) use the exothermic 

reaction between oxygen and organics to provide the heat necessary to devolatilize biomass and 

to convert residual carbon-rich chars. In POX gasifiers, the heat to drive the process is generated 

internally within the gasifier. A disadvantage of POX gasifiers is that oxygen production is 

expensive and typically requires large plant sizes to improve economics [14]. 

 3 
The second general class, steam gasifiers (indirectly-heated gasifiers), accomplish biomass 

heating and gasification through heat transfer from a hot solid or through a heat transfer surface. 

Either byproduct char and/or a portion of the product gas can be combusted with air (external to 

the gasifier itself) to provide the energy required for gasification. Steam gasifiers have the 

advantage of not requiring oxygen; but since most operate at low pressure they require product 

gas compression for downstream purification and synthesis unit operations. The erosion of 

refractory due to circulating hot solids in an indirect gasifier can also present some potential 

operational difficulties. 

A number of POX and steam gasifiers are under development and have the potential to produce a 

synthesis gas suitable for liquid fuel synthesis. These gasifiers have been operated in the 4 to 350 

ton per day scale. The decision as to which type of gasifier (POX or steam) will be the most 

economic depends upon the entire process, not just the cost for the gasifier itself. One indicator 

for comparing processes is “capital intensity,” the capital cost required on a per unit product 

basis. Figure 2 shows the capital intensity of methanol processes [15, 16, 17, 18, 19, 20] based 

on indirect steam gasification and direct POX gasification. This figure shows that steam 

gasification capital intensity is comparable or lower than POX gasification. The estimates 

indicate that both steam gasification and POX gasification processes should be evaluated, but if 

the processes need to be evaluated sequentially, choosing steam gasification for the first 

evaluation is reasonable. 

 4 
Capital Investment ($/annual gal methanol)

0.00

1.00

2.00

3.00

4.00

5.00

Indirect Steam

 Gasification

Oxygen Gasification

w Catalytic

Reforming

w/o Catalytic

Reforming

Slagging

Dry Ash

Sources:

Wyman, et al., 1993

Williams, et al., 1995

Hamelinck & Faaij, 2001

2,000 tpd biomass nominal size

$2002

Figure 2. Estimated capital intensities for biomass-to-methanol processes 

Another philosophy applied to the process development was the idea to make the process energy 

self-sufficient. It was recognized that the heat and power requirements of the process could not 

be met just with char combustion and would require additional fuel. Several options were 

considered. Additional biomass could be added as fuel directly to the heat and power system, 

however, this would increase the process beyond 2,000 tonne/day. Fossil fuels (coal or natural 

gas) could also be added directly to provide the additional fuel. Alternately syngas could be 

diverted from liquid fuel production to heat and power production. This option makes the design 

more energy self-sufficient, but also lowers the overall process yield of alcohols. 

It was decided that (1) no additional fuel would be used for heat and power and (2) only enough 

syngas would be diverted so that the internal heat and power requirements would be exactly met. 

Thus, there would neither be electricity sales to the grid nor electricity purchases. The only 

exception to this would be if other operating specifications were such that syngas could no 

longer be backed out of the heat and power system but there is still excess electricity (that could 

then be sold to the grid for a co-product credit). This resulted in 28% of the unconditioned 

syngas being diverted to power the process. Model calculations show that if none of the syngas 

was diverted in this manner, and all of it was used for mixed alcohols production, the ethanol and 

higher alcohols yields would increase by 38%. Thus, the baseline ethanol yield of 80.1 gal/dry 

ton could rise as high as 110.9 gal/ton, with total production of all alcohols as high as 130.3 

 5 
gal/dry ton. However, the minimum ethanol plant gate price increases in this scenario because of 

the cost of the natural gas required to meet the energy demands of the process. 

2.1. Analysis Approach 

The general approach used in the development of the process design, process model, and 

economic analysis is depicted in Figure 3. The first step was to assemble a general process flow 

schematic or more detailed process flow diagrams (PFDs). (See Appendix H for the associated 

PFDs for this design). From this, detailed mass and energy balance calculations were performed 

around the process. For this design, Aspen Plus software was used. Data from this model was 

then used to properly size all process equipment and fully develop an estimate of capital and 

operating costs. These costs could have potentially been used in several types of economic 

analysis. For this design however, a discounted cash flow rate of return (DCFROR) analysis was 

used to determine the ethanol minimum plant gate price necessary to meet an nth plant hurdle rate 

(IRR) of 10%. 

Figure 3. Approach to process analysis 

 6 
This TC conversion process was developed based upon NREL experience performing conceptual 

designs for biomass conversion to ethanol via biochemical means [3], biopower applications, and 

biomass gasification for hydrogen production.[9] Specific information for potential subprocesses were obtained as a result of a subcontract with Nexant Inc. [21, 22, 23, 24] 

Aspen Plus version 2004.1 was used to determine the mass and energy balances for the process. 

The operations were separated into seven major HIERARCHY areas: 

• Feed Handling and Drying (Area 100) 

• Gasification (Area 200) 

• Cleanup and Conditioning (Area 300) 

• Alcohol Synthesis (Area 400) 

• Alcohol Separation (Area 500) 

• Steam Cycle (Area 600) 

• Cooling Water (Area 700) 

Overall, the Aspen simulation consists of about 300 operation blocks (such as reactors, flash 

separators, etc.), 780 streams (mass, heat, and work), and 65 control blocks (design specs and 

calculator blocks). Many of the gaseous and liquid components were described as distinct 

molecular species using Aspen’s own component properties database. The raw biomass 

feedstock, ash, and char components were modeled as non-conventional components. There was 

more detail and rigor in some blocks (e.g., distillation columns) than others (e.g., conversion 

extent in the alcohol synthesis reactor). Because this design processes three different phases of 

matter (solid phase, gas phase, and liquid phase), no single thermodynamics package was 

sufficient. Instead, four thermodynamics packages were used within the Aspen simulation to give 

more appropriate behavior. The “RKS-BM” option was used throughout much of the process for 

high temperature, high pressure phase behavior. The non-random two-liquid “NRTL” option 

with ideal gas properties was used for alcohol separation calculations. The 1987 Steam Table 

properties were used for the steam cycle calculations. Finally, the ELECNRTL package was used 

to model the electrolyte species potentially present within the quench water system. 

The process economics are based on the assumption that this is the “nth” plant, meaning that 

several plants using this same technology will have already been built and are operating. This 

means that additional costs for risk financing, longer start-ups, and other costs associated with 

first-of-a-kind plants are not included. 

The capital costs were developed from a variety of sources. For some sub-processes that are well 

known technology and can be purchased as modular packages (i.e. amine treatment, acid gas 

removal), an overall cost for the package unit was used. Many of the common equipment items 

(tanks, pumps, simple heat exchangers) were costed using the Aspen Icarus Questimate costing 

software. Other more specific unit operations (gasifier, molecular sieve, etc) used cost estimates 

from other studies and/or from vendor quotes. As documented in the hydrogen design report [9], 

the installed capital costs were developed using general plant-wide factors. The installation costs 

incorporated cost contributions for not only the actual installation of the purchased equipment 

but also instrumentation and controls, piping, electrical systems, buildings, yard improvements, 

etc. These are also described in more detail in Section 3. 

 7 
The purchased component equipment costs reflect the base case for equipment size and cost 

year. The sizes needed in the process may actually be different than what was specifically 

designed. Instead of re-costing in detail, an exponential scaling expression was used to adjust the 

bare equipment costs: 

( ) New Size New Cost Base Cost

Base Size

⎛ ⎞ = ⎜ ⎟ ⎝ ⎠

n

where is a characteristic scaling exponent (typically in the range of 0.6 to 0.7). The sizing 

parameters are based upon some characteristic of the equipment related to production capacity, 

such as inlet flow or heat duty in a heat exchanger (appropriate if the log-mean temperature 

difference is known not to change greatly). Generally these related characteristics are easier to 

calculate and give nearly the same result as resizing the equipment for each scenario. The scaling 

exponent can be inferred from vendor quotes (if multiple quotes are given for different sizes), 

multiple estimates from Questimate at different sizes, or a standard reference (such as Garrett, 

[

n

n

25] Peters and Timmerhaus, [26] or Perry et al. [27]). 

Since a variety of sources were used, the bare equipment costs were derived based upon different 

cost years. Therefore, all capital costs were adjusted with the Chemical Engineering (CE) 

magazine’s Plant Cost Index [28] to a common basis year of 2005: 

( ) Cost Index in New Year New Cost Base Cost

Cost Index in Base Year

⎛ ⎞ = ⎜ ⎟ ⎝ ⎠. 

The CE indices used in this study are listed in Table 1 and depicted in Figure 4. Notice that the 

indices were very nearly the same for 2000 to 2002 (essentially zero inflation) but take a very 

sharp increase after 2003 (primarily due a run-up in worldwide steel prices). 

Table 1. Chemical Engineering Magazine’s Plant Cost Indices 

Year Index Year Index 

1990 357.6 1998 389.5 

1991 361.3 1999 390.6 

1992 358.2 2000 394.1 

1993 359.2 2001 394.3 

1994 368.1 2002 395.6 

1995 381.1 2003 402.0 

1996 381.7 2004 444.2 

1997 386.5 2005 468.2 

 8 
Chemical Engineering Plant Cost Index

350

375

400

425

450

475

500

525

1985 1990 1995 2000 2005 2010

Figure 4. Chemical Engineering Magazine’s plant cost indices 

Once the scaled, installed equipment costs were determined, we applied overhead and 

contingency factors to determine a total plant investment cost. That cost, along with the plant 

operating expenses (generally developed from the ASPEN model’s mass and energy balance 

results) was used in a discounted cash flow analysis to determine the ethanol plant gate price, 

using a specific discount rate. For the analysis done here, the ethanol minimum plant gate price is 

the primary value used to compare alternate designs. 

 9 
2.2. Process Design Overview 

Figure 5. Block flow diagram 

A simple block flow diagram of the current design is depicted in Figure 5. The detailed process 

flow diagrams (PFDs) are in Appendix H. The process has the following steps: 

• Feed Handling & Preparation. The biomass feedstock is dried from the as-received 

moisture to that required for proper feeding into the gasifier using flue gases from the 

char combustor and tar reformer catalyst regenerator. 

• Gasification. Indirect gasification is considered in this assessment. Heat for the 

endothermic gasification reactions is supplied by circulating hot synthetic olivinea “sand” 

between the gasifier and the char combustor. Conveyors and hoppers are used to feed the 

biomass to the low-pressure indirectly-heated entrained flow gasifier. Steam is injected 

into the gasifier to aid in stabilizing the entrained flow of biomass and sand through the 

gasifier. The biomass chemically converts to a mixture of syngas components (CO, H2, 

CO2, CH4, etc.), tars, and a solid “char” that is mainly the fixed carbon residual from the 

biomass plus carbon (coke) deposited on the sand. Cyclones at the exit of the gasifier 

separate the char and sand from the syngas. These solids flow by gravity from the 

cyclones into the char combustor. Air is introduced to the bottom of the reactor and 

serves as a carrier gas for the fluidized bed plus the oxidant for burning the char and 

coke. The heat of combustion heats the sand to over 1800°F. The hot sand and residual 

ash from the char is carried out of the combustor by the combustion gases and separated 

from the hot gases using another pair of cyclones. The first cyclone is designed to capture 

mostly sand while the smaller ash particles remain entrained in the gas exiting the 

 a

 Calcined magnesium silicate, primarily Enstatite (MgSiO3), Forsterite (Mg2SiO3), and Hematite (Fe2O3). This is 

used as a sand for various applications. A small amount of magnesium oxide (MgO) is added to the fresh olivine to 

prevent the formation of glass-like bed agglomerations that would result from biomass potassium interacting with 

the silicate compounds. 

 10 
cyclone. The second cyclone is designed to capture the ash and any sand passing through 

the first cyclone. The hot sand captured by the first cyclone flows by gravity back into the 

gasifier to provide the heat for the gasification reaction. Ash and sand particles captured 

in the second cyclone are cooled, moistened to minimize dust and sent to a land fill for 

disposal. 

• Gas Cleanup & Conditioning. This consists of multiple operations: reforming of tars and 

other hydrocarbons to CO and H2; syngas cooling/quench; and acid gas (CO2 and H2S) 

removal with subsequent reduction of H2S to sulfur. Tar reforming is envisioned to occur 

in an isothermal fluidized bed reactor; de-activated reforming catalyst is separated from 

the effluent syngas and regenerated on-line. The hot syngas is cooled through heat 

exchange with the steam cycle and additional cooling via water scrubbing. The scrubber 

also removes impurities such as particulates and ammonia along with any residual tars. 

The excess scrubber water is sent off-site to a waste-water treatment facility. The cooled 

syngas enters an amine unit to remove the CO2 and H2S. The H2S is reduced to elemental 

sulfur and stockpiled for disposal. The CO2 is vented to the atmosphere in this design. 

• Alcohol Synthesis. The cleaned and conditioned syngas is converted to alcohols in a fixed 

bed reactor. The mixture of alcohol and unconverted syngas is cooled through heat 

exchange with the steam cycle and other process streams. The liquid alcohols are 

separated by condensing them away from the unconverted syngas. Though the 

unconverted syngas has the potential to be recycled back to the entrance of the alcohol 

synthesis reactor, this recycle is not done in this process design because CO2

concentrations in the recycle loop would increase beyond acceptable limits of the 

catalyst. Added cost would be incurred if this CO2 were separated. Instead the 

unconverted syngas is recycled to the Gas Cleanup & Conditioning section, mostly as 

feed to the tar reformer. 

• Alcohol Separation. The alcohol stream from the Alcohol Synthesis section is 

depressurized in preparation of dehydration and separation. Another rough separation is 

performed in a flash separator; the evolved syngas is recycled to the Gas Cleanup & 

Conditioning section, mostly as feed to the tar reformer. The depressurized alcohol 

stream is dehydrated using vapor-phase molecular sieves. The dehydrated alcohol stream 

is introduced to the main alcohol separation column that splits methanol and ethanol from 

the higher molecular weight alcohols. The overheads are topped in a second column to 

remove the methanol to ASTM sales specifications. The methanol leaving in the 

overheads is used to flush the adsorbed water from the molecular sieves. This 

methanol/water mixture is recycled back to the entrance of the alcohol synthesis reactor 

in order to increase the yield of ethanol and higher alcohols. 

• Heat & Power. A conventional steam cycle produces heat (as steam) for the gasifier and 

reformer operations and electricity for internal power requirements (with the possibility 

of exporting excess electricity as a co-product). The steam cycle is integrated with the 

biomass conversion process. Pre-heaters, steam generators, and super-heaters are 

integrated within the process design to create the steam. The steam will run through 

turbines to drive compressors, generate electricity or be withdrawn at various pressure 

 11 
levels for injection into the process. The condensate will be sent back to the steam cycle, 

de-gassed, and combined with make-up water. 

A cooling water system is also included in the Aspen Plus model to determine the requirements 

of each cooling water heat exchanger within the biomass conversion process as well as the 

requirements of the cooling tower. 

Previous analyses of gasification processes have shown the importance of properly utilizing the 

heat from the high temperature streams. A pinch analysis was performed to analyze the energy 

network of this ethanol production process. The pinch concept offers a systematic approach to 

optimize the energy integration of the process. Details of the pinch analysis will be discussed in 

Section 3.10. 

2.3. Feedstock and Plant Size 

Based upon expected availability per the Billion Ton Vision [6] study, the forest resources were 

chosen for the primary feedstock. The Billion Ton Vision study addressed short and long term 

availability issues for biomass feedstocks without giving specific time frames. The amounts are 

depicted in Figure 6. The upper sets of numbers (labeled “High Yield Growth With Energy 

Crops” and “High Yield Growth Without Energy Crops”) are projections of availability that will 

depend upon changes to agricultural practices and the creation of a new energy crop industry. In 

the target year of 2012 it is most probable that the amounts labeled “Existing & Unexploited 

Resources” will be the only ones that can be counted on to supply a thermochemical processing 

facility. Notice that the expected availability of forest resources is nearly the same as that of 

agricultural resources. Prior studies for biochemical processing have largely focused on using 

agricultural resources. It makes sense to base thermochemical processing on the forest resources. 

TC processing could fill an important need to provide a cost-effective technology to process this 

major portion of the expected biomass feedstock. 

 12 
0 50 100 150 200 250 300 350 400 450 500

Existing & Unexploited

Resources

High Yield Growth Without

Energy Crops

High Yield Growth With

Energy Crops

Million Tons Annually

Forest Resources Total

Grains & Manure Sub-Total

Ag Residues (non Energy Crops)

Perennial (Energy) Crops

Figure 6. Expected availability of biomass 

Past analyses have used hybrid poplar wood chips delivered at 50 wt% moisture to model forest 

resources [9]; the same will be done here. The ultimate analysis for the feed used in this study is 

given in Table 2. Performance and cost effects due to composition and moisture content were 

examined as part of the sensitivity analysis and alternate scenarios. 

Table 2. Ultimate Analysis of Hybrid Poplar Feed 

Component (wt%, dry basis29) 

 Carbon 50.88 

 Hydrogen 6.04 

 Nitrogen 0.17 

 Sulfur 0.09 

 Oxygen 41.90 

 Ash 0.92 

Heating valuec (Btu/lb): 8,671 HHVd

8,060 LHVe

The design plant size for this study was chosen to match that of the Aden et al. biochemical 

process [3], 2,000 dry tonne/day (2,205 dry ton/day). With an expected 8,406 operating hours per 

year (96% operating factor) the annual feedstock requirement is 700,000 dry tonne/yr (772,000 

dry ton/yr). As can be seen in Figure 6 this is a small portion of the 140 million dry ton/yr of 

 c

 Calculated using the Aspen Plus Boie correlation. 

d

 Higher Heating Value 

e

 Lower Heating Value 

 13 
forest resources potentially available. Cost effects due to plant size were examined as part of the 

sensitivity analysis. 

The delivered feedstock cost was chosen to match recent analyses done at Idaho National 

Laboratory (INL) [30] to target $35 per dry ton by 2012. Cost effects due to feedstock cost were 

also examined as part of the sensitivity analysis. 

3. Process Design 

3.1. Process Design Basis 

The process design developed for this study is based upon the current operation and R&D 

performance goals for the catalytic tar destruction and heteroatom removal work at NREL and 

alcohol synthesis work at NREL and PNNL. This target design shows the effect of meeting these 

specific research and development (R&D) goals. 

The process broadly consists of the following sections: 

• Feed handling and drying 

• Gasification 

• Gas clean up and conditioning 

• Alcohol synthesis 

• Alcohol separation 

• Integrated steam system and power generation cycle 

• Cooling water and other utilities 

3.2. Feed Handling and Drying – Area 100 

This section of the process accommodates the delivery of biomass feedstock, short term storage 

on-site, and the preparation of the feedstock for processing in the gasifier. The design is based 

upon a woody feedstock. It is expected that a feed handling area for agricultural residues or 

energy crops would be very similar. 

The feed handling and drying section are shown in PFD-P800-A101 and PFD-P800-A102. Wood 

chips are delivered to the plant primarily via trucks. However, it is envisioned that there could be 

some train transport. Assuming that each truck capacity is about 25 tons [31], this means that if 

the wood, at a moisture content of 50%, was delivered to the plant via truck transport only, then 

176 truck deliveries per day would be required. As the trucks enter the plant they are weighed 

(M-101) and the wood chips are dumped into a storage pile. From the storage pile, the wood 

chips are conveyed (C-102) through a magnetic separator (S-101) and screened (S-102). Particles 

larger than 2 inches are sent through a hammer mill (T-102/M-102) for further size reduction. 

Front end loaders transfer the wood chips to the dryer feed bins (T-103). 

Drying is accomplished by direct contact of the biomass feed with hot flue gas. Because of the 

large plant size there are two identical, parallel feed handling and drying trains. The wet wood 

 14 
chips enter each rotary biomass dryer (M-104) through a dryer feed screw conveyor (C-104). 

The wood is dried to a moisture content of 5 wt% with flue gas from the char combustor (R-202) 

and tar reformer’s fuel combustor (R-301). The exhaust gas exiting the dryer is sent through a 

cyclone (S-103) and baghouse filter (S-104) to remove particulates prior to being emitted to the 

atmosphere. The stack temperature of the flue gas is set at 62° above the dew point of the gas, 

235°F (113°C). The stack temperature is controlled by cooling the hot flue gas from the char 

combustor and the tar reformer with two steam boilers (H-286B and H-311B) prior to entering 

the dryer. This generated steam is added to the common steam drum (T-604) (see section on 

Steam System and Power Generation – Area 600). The dried biomass is then conveyed to the 

gasifier train (T-104/C-105). 

Equipment costs were derived from the biochemical design report that utilized poplar as a 

feedstock. [5] 

3.3. Gasification – Area 200 

This section of the process converts a mixture of dry feedstock and steam to syngas and char. 

Heat is provided in an indirect manner — by circulating olivine that is heated by the combustion 

of the char downstream of the gasifier. The steam primarily acts as a fluidizing medium in the 

gasifier and also participates in certain reactions when high gasifier temperatures are reached. 

From the feed handling and drying section, the dried wood enters the gasifier section as shown in 

PFD-P800-A201. Because of the plant size, it is assumed that there are two parallel gasifier 

trains. The gasifier (R-201) used in this analysis is a low-pressure indirectly-heated circulating 

fluidized bed (CFB) gasifier. The gasifier was modeled using correlations based on run data from 

the Battelle Columbus Laboratory (BCL) 9 tonne/day test facility (see Appendix I). 

Heat for the endothermic gasification reactions is supplied by circulating a hot medium between 

the gasifier vessel and the char combustor (R-202). In this case the medium is synthetic olivine, a 

calcined magnesium silicate, primarily Enstatite (MgSiO3), Forsterite (Mg2SiO3), and Hematite 

(Fe2O3), used as a heat transfer solid for various applications. A small amount of MgO must be 

added to the fresh olivine to avoid the formation of glass-like bed agglomerations that would 

result from the biomass potassium interacting with the silicate compounds. The MgO titrates the 

potassium in the feed ash. Without MgO addition, the potassium will form glass, K2SiO4, with 

the silica in the system. K2SiO4 has a low melting point (~930°F, 500°C) and its formation will 

cause the bed media to become sticky, agglomerate, and eventually defluidize. Adding MgO 

makes the potassium form a high melting (~2,370°F, 1,300°C) ternary eutectic with the silica, 

thus sequestering it. Potassium carry-over in the gasifier/combustor cyclones is also significantly 

reduced. The ash content of the feed is assumed to contain 0.2 wt% potassium. The MgO flow 

rate is set at two times the molar flow rate of potassium. 

The gasifier fluidization medium is steam that is supplied from the steam cycle (Steam System 

and Power Generation – Area 600). The steam-to-feed ratio is 0.4 lb of steam/lb of bone dry 

biomass. The gasifier pressure is 23 psia. The olivine circulating flow rate is 27 lb of olivine/lb 

of bone dry wood. Fresh olivine is added at a rate of 0.01% of the circulating rate to account for 

losses. The char combustor is operated with 20% excess air. 

 15 
Both the gasifier and the char combustor temperatures are allowed to “float” and are dictated 

from the energy balances around the gasifier and combustor. In general, the more char created, 

the higher the char combustor temperature; but the higher the char combustor temperature, the 

higher the resulting gasifier temperature, resulting in less char. In this way the gasifier and char 

combustor temperatures tend to find an equilibrium position. For the design case the resulting 

gasifier temperature is 1,633°F (889°C) and the char combustor is 1,823°F (995°C). The 

composition of the outlet gas from the gasifier is shown in Table 3. 

Particulate removal from the raw syngas exiting the gasifier is performed using two-stage 

cyclone separators. Nearly all of the olivine and char (99.9% of both) is separated in the primary 

gasifier cyclone (S-201) and gravity-fed to the char combustor. A secondary cyclone (S-202) 

removes 90% of any residual fines. The char that is formed in the gasifier is burned in the 

combustor to reheat the olivine. The primary combustor cyclone (S-203) separates the olivine 

(99.9%) from the combustion gases and the olivine is gravity-fed back to the gasifier. Ash and 

any sand particles that are carried over in the flue gas exiting the combustor are removed in the 

secondary combustor cyclone (99.9% separation in S-204) followed by an electrostatic 

precipitator (S-205) which removes the remaining residual amount of solid particles. The sand 

and ash mixture from the secondary flue gas cyclone and precipitator are land filled but prior to 

this the solids are cooled and water is added to the sand/ash stream for conditioning to prevent 

the mixture from being too dusty to handle. First the ash and sand mixture is cooled to 300°F 

(149°C) using the water cooled screw conveyor (M-201) then water is added directly to the 

mixture until the mixture water content is 10 wt%. 

Table 3. Gasifier Operating Parameters, Gas Compositions, and Efficiencies 

Gasifier Variable Value 

Temperature 1,633oF (890oC) 

Pressure 23 psia (1.6 bar) 

Gasifier outlet gas composition mol% (wet) mol% (dry) 

H2 15.0 25.1 

CO2 7.4 12.4 

CO 25.1 41.9 

H2O 40.2 -- 

CH4 9.0 15.1 

C2H2 0.3 0.4 

C2H4 2.5 4.1 

C2H6 0.1 0.2 

C6H6 0.1 0.1 

tar (C10H8) 0.1 0.2 

NH3 0.2 0.3 

H2S 0.04 0.07 

H2:CO molar ratio 0.60 

Gasifier Efficiency 76.6% HHV basis 

76.1% LHV basis 

Capital costs for the equipment in this section are described in detail in Section 3 of this report. 

The operating costs for this section are listed in Appendix E and consist of makeup MgO and 

olivine, and sand/ash removal. 

 16 
3.4. Gas Cleanup and Conditioning – Area 300 

This section of the process cleans up and conditions the syngas so that the gas can be synthesized 

into alcohol. The type and the extent of cleanup are dictated by the requirements of the synthesis 

catalyst: 

• The tars in the syngas are reformed to additional CO and H2. 

• Particulates are removed by quenching. 

• Acid gases (CO2 and H2S) are removed. 

• The syngas is compressed. 

The gas from the secondary gasifier cyclone is sent to the catalytic tar reformer (R-303). In this 

bubbling fluidized bed reactor the hydrocarbons are converted to CO and H2 while NH3 is 

converted to N2 and H2. In the Aspen simulation, the conversion of each compound is set to 

match targets that are believed to be attainable through near-term research efforts. Table 4 gives 

the current experimental conversions (for deactivated catalyst) that have been achieved at NREL 

[32] and the conversions used in the simulation corresponding to the 2012 research targets. 

Table 4. Current and Target Design Performance of Tar Reformer 

Compound Experimental 

Conversion to CO & H2

Target Conversion to 

CO & H2

Methane (CH4) 20% 80% 

Ethane (C2H6) 90% 99% 

Ethylene (C2H4) 50% 90% 

Tars (C10+) 95% 99.9% 

Benzene (C6H6) 70% 99% 

Ammonia (NH3)

f 70% 90% 

In the Aspen simulation the tar reformer operates isothermally at 1,633oF. An implicit 

assumption in this mode of operation is that the energy needed for the endothermic reforming 

reactions can be transferred into the catalyst bed. Although conceptual reactor designs are readily 

created for providing the heat of reaction from the fuel combustion area directly into the 

reformer catalyst bed, in practice this may be a difficult and prohibitively expensive design 

option requiring internal heat transfer tubes operating at high temperatures. An alternate 

approach, not used in this study, would be to preheat the process gas upstream of the reformer 

above the current reformer exit temperature, and operate the reformer adiabatically with a 

resulting temperature drop across the bed and a lower exit gas temperature. In this configuration, 

the required inlet and exit gas temperatures would be set by the extent of conversion, the kinetics 

of the reforming reactions, and the amount of catalyst in the reactor. 

The composition of the gas from the tar reformer can be seen in Table 5. 

 f

 Converts to N2 and H2. 

 17 
Table 5. Target Design Tar Reformer Conditions and Outlet Gas Composition 

Tar Reformer Variable Value 

Tar reformer inlet temperature 1,633oF (890oC) 

Tar reformer outlet temperature 1,633oF (890oC) 

Tar reformer outlet gas composition mol% (wet) mol% (dry) 

H2 37.4 43.0 

CO2 9.9 11.4 

CO 37.4 43.0 

H2O 13.0 --- 

CH4 1.2 1.4 

C2H2 0.01 0.01 

C2H4 0.11 0.13 

C2H6 10.8 ppmv 12.4 ppmv 

C6H6 2.7 ppmv 3.1 ppmv 

tar (C10H8) 0.5 ppmv 0.6 ppmv 

NH3 0.01 0.01 

H2S 0.02 0.02 

N2 0.72 0.83 

H2:CO molar ratio 1.00 

Prior to the quench step, the hot syngas is cooled to 300°F (149°C) with heat exchangers (H301A-C) that are integrated in the steam cycle (see section Steam System and Power Generation 

– Area 600). After this direct cooling of the syngas, additional cooling is carried out via water 

scrubbing (M-302 and M-301), shown in PFD-P800-A302. The scrubber also removes impurities 

such as particulates, residual ammonia, and any residual tars. The scrubbing system consists of a 

venturi scrubber (M-302) and quench chamber (M-301). The scrubbing system quench water is a 

closed recirculation loop with heat rejected to the cooling tower and a continuous blow down 

rate of approximately 2.3 gallons per minute (gpm) that is sent to a waste water treatment 

facility. The quench water flow rate is determined by adjusting its circulation rate until its exit 

temperature from the quench water recirculation cooler (H-301) is 110°F (43°C). Any solids that 

settle out in T-301 are sent off-site for treatment as well. For modeling purposes, the water 

content of the sludge stream was set at 50 wt%. 

The quench step cools the syngas to a temperature of 140°F (60°C). The syngas is then 

compressed using a five-stage centrifugal compressor with interstage cooling as shown in PFDP800-A303. The compressor was modeled such that each section has a polytropic efficiency of 

78% and intercooler outlet temperatures of 140°F (60°C). The interstage coolers are forced air 

heat exchangers. 

Depending on the specific catalysts being used downstream of the tar reformer, varying 

concentrations of acid gas compounds can be tolerated in the syngas. For example, sulfur 

concentrations as H2S are required to be below 0.1 ppm for copper based synthesis catalysts. 

This design is based upon sulfided molybdenum catalysts which actually require up to 100 ppm 

of H2S in the syngas to maintain catalyst activity. Because the syngas exiting the gasifier 

contains almost 400 ppmv of H2S, some level of sulfur removal will be required by any of the 

synthesis catalysts currently of interest. 

Carbon dioxide is the other acid gas that needs to be removed in the syngas conditioning process. 

Similar to the sulfur compounds, the acceptable level of CO2 depends on the specific catalyst 

 18 
being used in the synthesis reactor to make alcohols. Some synthesis catalysts require low levels 

of CO2 while others, such as the sulfided molybdenum catalysts can tolerate relatively high CO2

levels compared to the sulfur species. CO2 is a major component of the gasification product, so 

significant amounts of CO2may need to be removed upstream of the synthesis reactor. 

Since the catalyst selected for this study is a sulfided catalyst that is tolerant of sulfur up to 100 

ppmv and CO2 up to 7 mol% (see Appendix J for more detail), a design that can provide for the 

removal of both sulfur and carbon dioxide was chosen. An amine system capable of selectively 

removing CO2 and H2S from the main process syngas stream is used. The amine assumed for this 

study is monoethanol amine (MEA), based on the recommendation by Nexant [33]. 

The acid gas scrubber was simulated using a simplified model of SEP blocks and specifying the 

amount of CO2 and H2S needing to be removed to meet design specifications of 50 ppmv H2S 

and 5 mol% CO2 at the synthesis reactor inlet, including any recycle streams to that unit 

operation. The amine system heating and cooling duties were calculated using information taken 

from section 21 of the GPSA Data Handbook [34]. This method gave a heat duty of 2660 Btu 

per pound of CO2 removed, with a similar magnitude cooling duty provided by forced-air cooling 

fans. Power requirements for pumping and fans were also calculated using GPSA recommended 

values. The acid gas scrubber operating values for the base case are given below. 

Acid Gas Removal Parameter Value 

Amine Used Monoethanol amine (MEA) 

Amine Concentration 35 wt% 

Amine Circ. Rate 1,945 gpm 

Amine Temp. @ Absorber 110°F 

Absorber Pressure 450 psia 

Stripper Condenser Temperature 212°F 

Stripper Reboiler Temperature 237°F 

Stripper Pressure 65 psia 

Stripper Reboiler Duty 140.1 MMBTU/hr 

Stripper Condenser Duty 93.4 MMBTU/hr 

Amine Cooler Duty 46.7 MMBTU/hr 

Heat Duty per Pound CO2 removed 2,660 Btu/lb 

If a highly CO2 -tolerant alcohol synthesis catalyst is used, it may become possible to use other 

syngas conditioning processes or methods to selectively remove H2S, with less energy and 

possibly at a significantly lower capital cost. 

The acid gases removed in the amine scrubber are then stripped to regenerate the sorbent and 

sent through a sulfur removal operation using a liquid phase oxidation process as shown in PFDP800-A305. The combined Amine/ LO-CAT process will remove the sulfur and CO2 to the 

levels desired for the selected molysulfide catalyst [35]. Although, there are several liquid-phase 

oxidation processes for H2S removal and conversion available today, the LO-CAT process was 

selected because of its progress in minimizing catalyst degradation and its environmentallybenign catalyst. LO-CAT is an iron chelate-based process that consists of a venturi precontactor 

(M-303), liquid-filled absorber (M-304), air-blown oxidizer (R-301), air blower (K-302), 

solution circulation pump (P-303) and solution cooler (H-305). Elemental sulfur is produced in 

 19 
the oxidizer and, since there is such a small amount (1.3 ton/day), it is stockpiled either for 

eventual disposal or sold as an unconditioned product. The LO-CAT process was modeled to 

remove the H2S to a concentration of 10 ppmv in the CO2 vent effluent from the amine scrubber. 

The air flow rate for re-oxidizing the LO-CAT solution was included in the simulation and 

calculated based on the requirement of 2 moles of O2 per mole of H2S. Prior to entering the LOCAT system the gas stream is superheated to 10°F (5.6°C ) above its dew point in preheater (H304), which in this process is equivalent to 120 ̊F. This degree of superheating is required for the 

LO-CAT system. The CO2 from the LO-CAT unit is vented to the atmosphere. 

The capital costs for the equipment in this section are described in further detail in the 

Appendices. The operating costs consist of makeup reforming catalyst, LO-CAT and amine 

chemical makeup, as well as reforming catalyst disposal cost and WWT. These are described in 

further detail in Section 3. 

3.5. Alcohol Synthesis – Area 400 

The alcohol synthesis reactor system is the heart of the entire process. Entering this process area, 

the syngas has been reformed, quenched, compressed and treated to have acid gas concentrations 

(H2S, CO2) reduced. After that, it is further compressed and heated to the synthesis reaction 

conditions of 1,000 psia and 570°F (300°C). The syngas is converted to the alcohol mixture 

across a fixed bed catalyst. The product gas is subsequently cooled, allowing the alcohols to 

condense and separate from the unconverted syngas. The liquid alcohols are then sent to alcohol 

separation and purification (Area 500). The residual gas stream is recycled back to the tar 

reformer with a small purge to fuel combustion (5%). 

Research on alcohol synthesis catalysts has waxed and waned over many decades for a variety of 

reasons. In order to review the status of mixed alcohol technology and how it has developed over 

the past 20 years, two activities were initiated. First, a literature search was conducted. This 

search and its findings are described in more detail in Appendix J, along with a discussion on 

specific terminology, such as “yield”, “selectivity”, and “conversion”. These terms will be used 

throughout the remainder of this document. Second, an engineering consulting company 

(Nexant) was hired to document the current state of technology with regards to mixed alcohols 

production and higher alcohol synthesis. Their results are published in an NREL subcontract 

[36] report. 

Based on the results of this background technology evaluation, a modified Fischer-Tropsch 

catalyst was used for this process design, specifically a molybdenum-disulfide-based (MoS2) 

catalyst. The former Dow/UCC catalyst was chosen as the basis because of its relatively high 

ethanol selectivity and because its product slate is a mixture of linear alcohols (as opposed to the 

branched alcohols that result from modified methanol catalysts). This particular catalyst uses 

high surface area MoS2 promoted with alkali metal salts (e.g. potassium carbonate) and cobalt 

(CoS). These promoters shift the product slate from hydrocarbons to alcohols, and can either be 

supported on alumina or activated carbon, or be used unsupported. 

 20 
Table 6 lists several process and syngas conditioning requirements for this synthesis reaction. 

These include both experimentally verified conditions typical of those found in literature, as well 

as targeted conditions from the OBP-funded research plan used in the model. 

Table 6. Process Conditions for Mixed Alcohols Synthesis 

Parameter “State of Technology” 

Conditions [41]

Target Conditions 

Used in Process Design & 

Aspen Model 

Temperature (°C) ~ 300 300 

Pressure (psia) 1500 - 2000 1000 

H2/CO ratio 1.0 – 1.2 1.0 

CO2 concentration (mol%) 0% - 7% 5.0% 

Sulfur concentration (ppmv) 50 - 100 50 

Though the synthesis reactor is modeled as operating isothermally, it is recognized that 

maintaining a constant temperature in a fixed bed reactor system would be difficult, especially 

since these reactions are highly exothermic. Temperature has a significant impact on the alcohol 

selectivity and product distribution. High pressures are typically required to ensure the 

production of alcohols. MoS2 catalysts are efficient Fischer-Tropsch (FT) catalysts at ambient or 

low pressures. However, significantly raising the pressure (in addition to promoting with alkali) 

helps to shift the pathways from hydrocarbon production towards alcohol production. However, 

compression requirements for achieving these pressures can be quite substantial. Thus, targeting 

a catalyst that achieves optimal performance at lower pressures can potentially provide 

significant cost savings. 

The CO2 concentration requirements for the syngas are less well-known. Herman [37] states that 

in the first Dow patent application, the presence of larger amounts of CO2 in the synthesis gas 

retarded the catalyst activity. Further study showed that increasing the CO2 concentration to 30 

vol% decreased the CO conversion but did not significantly alter the alcohol:hydrocarbon ratio 

of the product. With CO2 concentrations up to 6.7 vol%, the extent of CO conversion is not 

affected; however, higher chain alcohol yield relative to methanol does tend to decrease. This is 

why CO2 concentrations were reduced to 5 mol% in the model using the amine system as part of 

syngas conditioning. The effect of CO2 concentration on alcohol production will be studied in 

future laboratory experiments. 

One of the benefits of this catalyst is its sulfur tolerance. It must be continuously sulfided to 

maintain its activity; thus an inlet gas concentration of 50 ppmv H2S is maintained. 

Concentrations above 100 ppmv inhibit the reaction rate and higher alcohol selectivity. 

The overall stoichiometric reaction for alcohol synthesis can be summarized as: 

n CO + 2n H2 Æ CnH2n+1OH + (n-1) H2O 

Stoichiometry suggests an optimum H2:CO ratio of 2.0. However this catalyst maintains 

significant water-gas shift activity and will generate its own H2 from CO and H2O: 

 21 
CO + H2O Æ H2 + CO2.

This shifts the optimal ratio closer to 1.0 and also shifts the primary byproduct from water to 

CO2. Experiments [38] have been typically conducted using ratios in the range of 1.0 to 1.2. 

The compressor (K-410) in this area is a 3-stage steam-driven compressor that takes the syngas 

from 415 psia to 1000 psia, requiring 9,420 HP (assuming a polytropic efficiency of 78%). The 

outlet syngas from the compressor is then mixed with recycled methanol from Alcohol 

Purification (Area 500), heated to 570°F (300°C), and sent to the reactor. The capital cost for the 

compressor was developed using Questimate. 

The mixed alcohol synthesis reactor is a fixed-bed reactor system that contains the MoS2

catalyst. Because this is a net exothermic reaction system, water is cross exchanged with the 

reactor to produce steam for the process while helping to maintain a constant reactor 

temperature. Questimate was used to develop the reactor capital cost. 

The purchase price of the catalyst itself was estimated at $5.25/lb based on conversations NREL 

researchers had with CRITERION [39], a petroleum/hydrocarbon catalyst provider. This 

represents a generalized cost of Molybdenum-based catalyst at around $5/lb being sulfided for an 

additional $0.25/lb. In addition, NREL was able to speak with Dow catalyst experts [40] who 

said that in today’s market the raw material costs for producing such a catalyst system would run 

about $20/lb. Adding more cost for the catalyst preparation would bring that cost between $22-

40/lb. However, these costs could go down as demand goes up, and quite substantially if it gets 

to large enough scale. 

In reality, each company developing a process like this will have their own proprietary catalyst 

and associated formulation. The costs for these catalysts are difficult to predict at the present 

time since so few providers of mixed alcohols catalyst currently exist (and will likely be 

negotiated). Nexant also provided information on general catalyst metals price ranges in their 

report. They reported Molybdenum ranging from $2 – 40/lb. 

The lifetime of the catalyst was assumed to be 5 years. While existing mixed alcohols catalysts 

have not been tested for this long, they have operated for over 8,000 hours (roughly 1 year of 

continuous operating time) with little or no loss in performance. 

The reactor was modeled as a simple conversion-specified reactor using a series of alcohol and 

hydrocarbon production reactions as shown in Table 7. The propane, butane, and pentane+ 

reactions are set to zero because the catalyst will likely not favor these reactions. The specific 

conversions of each of the other reactions were set in order to reach catalyst performance targets, 

see Table 9. Those targets are shown in Table 8 along with values for those parameters typically 

found in literature. 

 22 
Table 7. System of Reactions for Mixed Alcohol Synthesis 

Water-Gas Shift CO + H2O ↔ H2 + CO2

Methanol CO + 2H2 Æ CH3OH 

Methane CH3OH + H2 Æ CH4 + H2O 

Ethanol CO + 2H2 + CH3OH Æ C2H5OH + H2O 

Ethane C2H5OH + H2 Æ C2H6 + H2O 

Propanol CO + 2H2 + C2H5OH Æ C3H7OH + H2O 

Propane C3H7OH + H2 Æ C3H8 + H2O 

n-Butanol CO + 2H2 + C3H7OH Æ C4H9OH + H2O 

Butane C4H9OH + H2 Æ C4H10 + H2O 

Pentanol+ CO + 2H2 + C4H9OH Æ C5H11OH + H2O 

Pentane+ C5H11OH + H2 Æ C5H12 + H2O 

Table 8. Mixed Alcohol Reaction Performance Results 

Result “State of Technology” 

Value Ranges [37, 41]

Target Results 

Used in Process Design & 

Aspen Model 

Total CO Conversion (per-pass) 10% - 40% 60% 

Total Alcohol Selectivity 

(CO2-free basis) 70% - 80% 90% 

Gas Hourly Space Velocity (hr-1) 1600 – 12,000 4000 

Catalyst Alcohol Productivity (g/kgcatalyst/hr )g 150 – 350 600 

The individual target values are less important than the net result of the entire collection. For 

example, a catalyst system can have a high CO conversion well above 40%, but if most of that 

CO is converted to methane or CO2, then the alcohol selectivities would be very low and the 

entire process economics would suffer. Likewise, if the catalyst had a high CO conversion and 

selectivity, but had very low productivity, a much larger reactor would have to be built to 

accommodate the volume of catalyst required. The set of targets shown above are improvements 

over current literature values, but were chosen as targets believed to be achievable through 

catalyst research and development. There is precedent for these results from other catalyst 

systems. For example, FT catalysts are currently capable of CO conversions above 70% [42]. 

Also commercial methanol catalysts have productivities over 1000 g/kg-catalyst/hr [37]. 

The reaction conversions were also set to achieve a certain product distribution of alcohols. The 

mixed alcohol products described in literature are often high in methanol, but contain a wide 

distribution of several different alcohols. The product distributions described by Dow and SRI 

are shown in Table 9 along with the relative product concentrations calculated by the model. 

 g

 Based on assumed catalyst density of 64 lb/ft3, 600 g/kg-catalyst/hr = 615 g/L-catalyst/hr. 

 23 
Table 9. Mixed Alcohol Product Distributions 

Alcohol Dow [43] 

(wt %) 

SRI [44]

(wt%) 

NREL Model 

 (wt%)* 

Methanol 30-70% 30.77% 5.01% 

Ethanol 34.5% 46.12% 70.66% 

Propanol 7.7% 13.3% 10.07% 

Butanol 1.4% 4.14% 1.25% 

Pentanol + 1.5% 2.04% 0.17% 

Acetates (C1 & C2) 2.5% 3.63% 

Others 10.98% 

Water 2.4% 1.86% 

Total 100% 100% 100% 

* Prior to alcohol purification and methanol recycle 

The most significant differences between the NREL model product distribution and those shown 

in literature are with regards to the methanol and ethanol distributions. This is primarily due to 

the almost complete recycle of methanol within this process. In the alcohol purification section 

downstream, virtually all methanol is recovered via distillation and recycled back to mix with the 

compressed syngas. This is done in order to increase the production of ethanol and higher 

alcohols. This concept has been proposed in literature, but data from testing in an integrated 

setting has not been seen. In literature, experiments are often conducted on closed or batch 

systems and do not examine the potential impacts of recycled compounds or other integration 

issues. However, this catalyst is known to have methanol decomposition functionality which 

indicates that methanol in the feed will not be detrimental to the reaction. The effects of recycled 

methanol will be examined experimentally as research progresses. 

A kinetic model was used to guide these conversion assumptions to help predict how the catalyst 

may perform as a result of significant methanol recycle. Very few kinetic models have been 

developed for this catalyst system [45, 46, 47]. Of these, only Gunturu examined the possibility 

of methanol recycle. Therefore NREL reproduced this kinetic model using Polymath software. 

This kinetic model predicted that methanol entering the reactor would largely be converted to 

ethanol and methane. This model also predicts that maintaining high partial pressures of 

methanol in the reactor would further reduce the production of alcohols higher than ethanol. 

More detailed discussion on the kinetic model can be found in Appendix K. 

After the reactor, the effluent is cooled to 110°F (43°C) through a series of heat exchangers 

while maintaining high pressure. First, the reacted syngas is cross exchanged with cooler process 

streams, lowering the temperature to 200°F (93°C). Air-cooled exchangers then bring the 

temperature down to 140°F (60°C). The final 30°F (17°C) drop is provided by cooling water. A 

knock-out drum (S-501) is then used to separate the liquids (primarily alcohols) from the 

remaining gas, which is comprised of unconverted syngas, CO2, and methane. Aspen Plus 

contains other physical property packages that model non-ideal liquid systems much better than 

the Redlich-Kwong-Soave (RKS) equation of state used throughout the model. Therefore, the 

Non-Random Two-Liquid (NRTL) package was used to model the alcohol condensation. 

From here, the liquid crude alcohols are sent to product purification while the residual syngas is 

superheated to 1500°F (816°C) and sent through an expander to generate additional power for 

 24 
the process. The pressure is dropped from 970 to 35 psia prior to being recycled to the tar 

reformer. A 5% purge stream is sent to fuel combustion. 

Alternate configurations will be discussed later in this report as will the economic sensitivity of 

certain synthesis parameters. One particular variation would be to recycle the unconverted 

syngas to the throat of the synthesis reactor instead of to the tar reformer. This would save 

money on upstream equipment costs because of lower process throughput, but would also lower 

yields because the CO2 would build up in the recycle loop. The limit to the amount of 

unconverted syngas that could be recycled to the reactor is less than 50% because this would 

cause the H2:CO ratio to grow well above 1.2. 

Future experiments and analysis will examine the impacts of methanol recycle, and of variations 

in concentration of CO2, CH4, and other compounds. Alternate reactor designs will also be 

examined. For example, FT technology largely has switched to slurry reactors instead of fixedbed reactors because the slurry fluidization achieves better heat and mass transfer properties that 

allow, in turn, for higher conversions. Such improvements could help to achieve the conversion 

targets outlined above and reduce the costs of major equipment items. 

3.6. Alcohol Separation – Area 500 

The mixed alcohol stream from Area 400 is sent to Area 500 where it is de-gassed, dried, and 

separated into three streams: methanol, ethanol, and mixed higher-molecular weight alcohols. 

The methanol stream is used to back-flush the molecular sieve drying column and then recycled, 

along with the water removed during back flushing, to the inlet of the alcohol synthesis reactor in 

Area 400. The ethanol and mixed alcohol streams are cooled and sent to product storage tanks. 

Carbon dioxide is readily absorbed in alcohol. Although the majority of the non-condensable 

gases leaving the synthesis reactor are removed in the separator vessel, S-501, a significant 

quantity of these gases remains in the alcohol stream, especially at the high system pressure. 

These gases are removed by depressurizing from 970 to 60 psia. Most of the dissolved gasses 

separate from the alcohols in the knock-out vessel S-502. This gas stream is made up primarily 

of carbon dioxide with some small amounts of hydrocarbons and alcohols; it is recycled to the 

Tar Reformer in Area 300. After being vaporized by cross exchanging with steam to a 20°F 

(11°C) superheated temperature, the alcohol stream goes to the molecular sieve dehydrator unit 

operation. 

The molecular sieve dehydrator design was based upon previous biochemical ethanol studies [5, 

3]and assumed to have similar performance with mixed alcohols. In the biochemical ethanol 

cases, the molecular sieve is used to dry ethanol after it is distilled to the azeotropic 

concentration of ethanol and water (92.5 wt% ethanol). The adsorbed water is flushed from the 

molecular sieves with a portion of the dried ethanol and recycled to the rectification column. The 

water ultimately leaves out the bottom of the distillation column. In this thermochemical process, 

however, it was determined that drying the entire mixed alcohol stream before any other 

separation would be preferable. The adsorbed water is desorbed from the molecular sieves with a 

combination of depressurization and flushing with methanol. This methanol/water mixture is 

then recycled back to the Alcohol Synthesis section (A400). 

 25 
The molecular sieve units require a superheated vapor. The liquid mixed alcohol stream is 

vaporized, superheated, and then fed to one of two parallel adsorption columns. The adsorption 

column preferentially removes water and a small amount of the alcohols. While one adsorption 

bed is adsorbing water, the other is regenerating. The water is desorbed from the bed during 

regeneration by applying a vacuum and flushing with dry methanol from D-505. This 

methanol/water mixture is recycled back to the Alcohol Synthesis section (A400). This 

methanol/water mixture is cooled to 140°F (60°C) using a forced air heat exchanger, and 

separated from any uncondensed vapor. The gaseous stream is recycled to the Tar Reformer and 

the condensate is pumped to 1,000 psia in P-514, and mixed with high-pressure syngas from 

compressor K-410 in Area 400 upstream of the synthesis reactor pre-heater. 

The dry mixed alcohol stream leaving the mol sieve dehydrator enters into the first of two 

distillation columns, D-504. D-504 is a typical distillation column using trays, overhead 

condenser, and a reboiler. The methanol and ethanol are separated from the incoming stream 

with 99% of the incoming ethanol being recovered in the overhead stream along with essentially 

all incoming methanol. The D-504 bottom stream consists of 99% of the incoming propanol, 1% 

of the incoming ethanol, and all of the butanol and pentanol. The mixed alcohol bottoms is 

considered a co-product of the plant and is cooled and sent to storage. The methanol/ethanol 

overhead stream from D-504 goes to a second distillation column, D-505, for further processing. 

D-505 separates the methanol from the binary methyl/ethyl alcohol mixture. The ethanol 

recovery in D-505 is 99% of the incoming ethanol and has a maximum methanol concentration 

of 0.5 mole percent to meet product specifications for fuel ethanol. The ethanol, which exits from 

the bottom of D-505 is cooled before being sent to product storage. The methanol and small 

quantity of ethanol exiting the overhead of column D-505 is used to flush the mol sieve column 

during its regeneration step as explained above. Currently, all of the methanol from D-505 is 

recycled through the mol sieve dehydrator and then to the synthesis reactor in Area 400. 

3.7. Steam System and Power Generation – Area 600 

This process design includes a steam cycle that produces steam by recovering heat from the hot 

process streams throughout the plant. Steam demands for the process include the gasifier, amine 

system reboiler, alcohol purification reboilers, and LO-CAT preheater. Of these, only the steam 

to the gasifier is directly injected into the process; the rest of the plant heat demands are provided 

by indirect heat exchange of process streams with the steam and have condensate return loops. 

Power for internal plant loads is produced from the steam cycle using an extraction steam 

turbine/generator (M-602). Power is also produced from the process expander (K-412), which 

takes the unconverted syngas from 965 psia to 35 psia before being recycled to the tar reformer. 

Steam is supplied to the gasifier from the low pressure turbine exhaust stage. The plant energy 

balance is managed to generate only the amount of electricity required by the plant. The steam 

system and power generation area is shown in PFD-P800-A601, -A602, and -A603 in Appendix 

H. 

A condensate collection tank (T-601) gathers condensate from the syngas compressors and from 

the process reboilers along with the steam turbine condensate and make-up water. The total 

 26 
condensate stream is heated to the saturation temperature and sent to the deaerator (T-603) to degas any dissolved gases out of the water. The water from the deaerator is first pumped to a 

pressure of 930 psia and then pre-heated to its saturation (bubble point) temperature using a 

series of exchangers. The saturated steam is collected in the steam drum (T-604). To prevent 

solids build up, water must be periodically discharged from the steam drum. The blowdown rate 

is equal to 2% of water circulation rate. The saturated steam from the steam drum is superheated 

with another series of exchangers. The superheated steam temperature and pressure were set as a 

result of pinch analysis. Superheated steam enters the turbine at 900oF and 850 psia and is 

expanded to a pressure of 175 psia. The remaining steam then enters the low pressure turbine and 

is expanded to a pressure of 65 psia. Here a slipstream of steam is removed and sent to the 

gasifier and other exchangers. Finally, the steam enters a condensing turbine and is expanded to 

a pressure of 1.5 psia. The steam is condensed in the steam turbine condenser (H-601) and the 

condensate re-circulated back to the condensate collection tank. 

The integration of the individual heat exchangers can only be seen in the PFDs included in the 

Appendices. To close the heat balance of the system, the Aspen Plus model increases or 

decreases the water flowrate through the steam cycle until the heat balance of the system is met. 

This process design assumes that the two compressors in this process (K301, K410) are steamdriven. All other drives for pumps, fans, etc are electric motors. Additionally, an allowance of 

0.7 MW of excess power is made to total power requirement to account for miscellaneous usage 

and general electric needs (lights, computers, etc). Table 10 contains the power requirement of 

the plant broken out into the different plant sections. Because syngas compression is steam 

driven, it is not a demand on the power system, which makes the total power requirement much 

less than it would be if compression demands were included. The plant power demands and 

power production were designed specifically to be nearly equal. Therefore, no excess power is 

being sold to or purchased from the grid. This plant was designed to be as energy self-sufficient 

as possible. This was accomplished by burning a portion of the “dirty” unreformed syngas in the 

fuel combustor (Section 300). While this does have a negative impact on the overall alcohol 

yields of the process, it does negate the purchase of natural gas or grid power. 

Table 10. Plant Power Requirements 

Plant Section Power Requirement (kW) 

Feed Handling & Drying 742 

Gasification 3,392 

Tar Reforming, Cleanup, & Conditioning 1,798 

Mixed Alcohol Synthesis 119 

Alcohol Separation and Purification 256 

Steam System & Power Generation 431 required 

7,994 generated 

Cooling Water & Other Utilities 529 

Miscellaneous 727 

Total plant power requirement 7,994 

 27 
3.8. Cooling Water and Other Utilities – Area 700 

The cooling water system is shown on PFD-P800-A701. A mechanical draft cooling tower (M701) provides cooling water to several heat exchangers in the plant. The tower utilizes large fans 

to force air through circulated water. Heat is transferred from the water to the surrounding air by 

the transfer of sensible and latent heat. Cooling water is used in the following pieces of 

equipment: 

• the sand/ash cooler (M-201) which cools the sand/ash mixture from the 

gasifier/combustor 

• the quench water recirculation cooler (M-301) which cools the water used in the syngas 

quench step 

• the water-cooled aftercooler (H-303) which follows the syngas compressor and cools the 

syngas after the last stage of compression 

• the LO-CAT absorbent solution cooler (H-305) which cools the regenerated solution that 

circulates between the oxidizer and absorber vessels 

• the reacted syngas cooler (H-414) which cools the gas in order to condense out the liquid 

alcohols 

• the end product finishing coolers (H-591, H-593) for both the higher alcohols co-product 

and the primary ethanol product 

• the blowdown water-cooled cooler (H-603) which cools the blowdown from the steam 

drum 

• the steam turbine condenser (H-601) which condenses the steam exiting the steam turbine 

Make-up water for the cooling tower is supplied at 14.7 psia and 60°F (16°C). Water losses 

include evaporation, drift (water entrained in the cooling tower exhaust air), and tower basin 

blowdown. Drift losses were estimated to be 0.2% of the water supply. Evaporation losses and 

blowdown were calculated based on information and equations in Perry, et al. [27]. The cooling 

water returns to the process at a supply pressure of 65 psia and temperature is 90°F (32°C). The 

cooling water return temperature is 110°F (43°C). 

An instrument air system is included to provide compressed air for both service and instruments. 

The instrument air system is shown on PFD-P800-A701. The system consists of an air 

compressor (K-701), dryer (S-701) and receiver (T-701). The instrument air is delivered at a 

pressure of 115 psia, a moisture dew point of -40°F (-40°C), and is oil free. 

Other miscellaneous items that are taken into account in the design include: 

• a firewater storage tank (T-702) and pump (P-702) 

• a diesel tank (T-703) and pump (P-703) to fuel the front loaders 

• an olivine truck scale with dump (M-702) and an olivine lock hopper (T-705) as well as 

an MgO lock hopper (T-706) 

• a hydrazine storage tank (T-707) and pump (P-705) for oxygen scavenging in the cooling 

water 

This equipment is shown on PFD-P800-A702. 

 28 
3.9. Additional Design Information 

Table 11 contains some additional information used in the Aspen Plus model and production 

design. 

Table 11. Utility and Miscellaneous Design Information 

Item Design Information 

Ambient air conditions (1,2, and 

3)

Pressure: 14.7 psia 

TDry Bulb: 90°F 

TWet Bulb: 80°F 

Composition (mol%): 

N2: 75.7% O2: 20.3% Ar: 0.9% CO2: 0.03% H2O: 3.1% 

Pressure drop allowance Syngas compressor intercoolers = 2 psi 

Heat exchangers and packed beds = 5 psi 

(1) In the GPSA Engineering Data Book [48], see Table 11.4 for typical design values for dry 

bulb and wet bulb temperature by geography. Selected values would cover summertime 

conditions for most of lower 48 states. 

(2) In Weast [49], see F-172 for composition of dry air. Nitrogen value adjusted slightly to force 

mole fraction closure using only N2, O2, Ar, and CO2 as air components. 

(3) In Perry, et al. [27], see psychrometric chart, Figure 12-2, for moisture content of air. 

3.10. Pinch Analysis 

A pinch analysis was performed to analyze the energy network of the biomass gasification to 

ethanol production process. The pinch technology concept offers a systematic approach to 

optimum energy integration of the process. First temperature and enthalpy data were gathered for 

the “hot” process streams (i.e., those that must be cooled), “cold” process streams (i.e., those that 

must be heated), and utility streams (such as steam, flue gas, and cooling water). The minimum 

approach temperature was set at 42.6°F. A temperature versus enthalpy graph (the “composite 

curve”) was constructed for the hot and cold process streams. These two curves are shifted so 

that they touch at the pinch point. From this shifted graph, a grand composite curve is 

constructed which plots the enthalpy differences between the hot and cold composite curves as a 

function of temperature. The composite curve is shown in Figure 7. From this figure the heat 

exchanger network of the system was determined. 

 29 
0

500

1000

1500

2000

2500

0 100,000,000 200,000,000 300,000,000 400,000,000 500,000,000 600,000,000 700,000,000 800,000,000

Heat Flow, Btu/hr

Temperature, F

Hot Process

Cold Process

Figure 7. Pinch analysis composite curve 

The total heating enthalpy equals the total cooling enthalpy because the Aspen model is designed 

to adjust the water flowrate through the steam cycle until the heat balance in the system is met. 

Because no outside utilities were used in this process, all heating and cooling duties are satisfied 

through process-process interchanges or process-steam interchanges. The minimum vertical 

distance between the curves is ΔTmin, which is theoretically the smallest approach needed in the 

exchange network. For this design, the pinch occurs at ~ H = 280,000,000 BTU/hr, and the upper 

and lower pinch temperatures are 570.0oF and 527.4oF, respectively, giving a ΔTmin of 42.6oF. 

Design of the heat exchange network for the above the pinch and below pinch regions are done 

separately. While pinch theory teaches that multiple solutions are possible, this particular 

solution has the advantage that heat released by the alcohol synthesis reactor is dissipated by 

raising steam. This is a standard design practice for removing heat from methanol synthesis and 

other similar reactors. The left-hand side of the composite curve shows the below pinch curves 

are constrained at the pinch and are also nearly pinched at the very left-hand side in the ~ 100oF 

range. This makes heat exchanger network design below the pinch more difficult. 

3.11. Energy Balance 

Energy integration is extremely important to the overall economics and efficiency of this 

process. Therefore a detailed understanding of how and where the energy is utilized and 

 30 
recovered is required. Detailed energy balances around the major process areas were derived 

using data from the Aspen Plus simulation. Comparing the process energy inputs and outputs 

enables the energy efficiency of the process to be quantified. Also, tracing energy transfer 

between process areas makes it possible to identify areas of potential improvement to the energy 

efficiency. 

 

The philosophy of defining the “energy potential” of a stream is somewhat different from what 

was done for the biochemical ethanol process design report [50]. For that analysis the definition 

of the energy potential was based upon the higher heating values (HHVs) of each component. 

This HHV basis is convenient when a process is primarily made up of aqueous streams in the 

liquid phase. Since liquid water at the standard temperature has a zero HHV, the contribution for 

any liquid water is very small, especially as compared to any other combustible material also 

present in the stream. However, the thermochemical ethanol production process differs 

significantly in that most of the process streams are in the gas phase. To remove the background 

contributions of the water, the energy potential is based instead upon the lower heating values 

(LHVs) of each component. 

The total energy potential for a stream has other contributions beyond that of the heating value. 

Other energy contributions are: 

• Sensible heat effect – the stream is at a temperature (and pressure) different from that of 

the standard conditions at which the heating values are defined. 

• Latent heat effect – one or more components in the stream are in a different phase from 

that at which their heating values are defined. 

• Non-ideal mixing effect – any heating or cooling due to blending dissimilar components 

in a mixture. 

The procedure for actually calculating the energy potential of a stream is also different from 

what was done prior. When the biochemical ethanol process was analyzed, the contributions for 

the HHVs, the sensible heating effects, and the latent heat effects were directly computed and 

combined. The calculations of the sensible and latent heat effects were done in an approximate 

manner. For example, the sensible heat effect was estimated from the heat capacity at the 

stream’s temperature, pressure, and composition; it was assumed that this heat capacity remained 

constant over the temperature range between the stream’s temperature and the standard 

temperature. For the relatively low temperatures of the biochemical ethanol process systems, this 

assumption makes sense. However, for this thermochemical process design, this assumption is 

not accurate because of the much larger differences between the process stream temperatures and 

the standard temperature 

The enthalpy values reported by Aspen Plus can actually be adjusted in a fairly simple manner to 

reflect either an HHV or LHV basis for the energy potential. The enthalpies calculated and 

reported by Aspen Plus are actually based upon a heat of formation for the energy potential of a 

stream. So, the reported enthalpies already include the sensible, latent, and non-ideal mixing 

effects. If certain constants in Aspen’s enthalpy expressions could be modified to be based on 

either the components’ HHVs or LHVs instead of the heats of formation then Aspen Plus would 

report the desired energy potential values. However, since the constants cannot be easily 

changed, the reported enthalpy values were adjusted instead as part of a spreadsheet calculation. 

 31 
The factors used to adjust the reported enthalpies were calculated from the difference between 

each component’s heat of combustion (LHV) and the reported pure component enthalpy at 

combustion conditions. 

This process for thermochemical conversion of cellulosic biomass was designed with the goal of 

being as energy self-sufficient. Natural gas inputs that could be used to fire the char combustor 

and fuel combustor have been eliminated. Instead, a slipstream of “dirty” unreformed syngas is 

used to meet the fuel demand. The downside to this is a decrease in ethanol yield. In addition, the 

process was designed to require no electricity be purchased from the grid. Instead, the integrated 

combined heat and power system supplies all steam and electricity needed by the plant. 

Consequently no electricity is sold as a co-product either. The only saleable products are the fuel 

ethanol and a higher molecular weight mixed alcohol co-product. 

The major process energy inputs and outlets are listed in Table 12, along with their energy 

flowrates. Each input and output is also ratioed to the biomass energy entering the system. The 

biomass is of course the primary energy input, however other energy inputs are required. Air is 

required for both the fuel combustor as well as the char combustor; however it remains a minor 

energy input. Some water is used to wet the ash leaving the gasification system, however, the 

majority of process water is used for boiler feed water makeup and cooling water makeup. A 

large negative energy flow value is associated with this because it enters the process as a liquid. 

 32 
Table 12. Overall Energy Analysis (LHV basis) 

Energy Flow 

(MMBTU/hr, LHV basis) 

Ratio to Feedstock 

Energy Flow 

Energy Inlets 

Wood Chip Feedstock (wet) 1269.7 1.000 

Natural Gas 0.0 0.000 

Air 2.3 0.002 

Olivine 0.0 0.000 

MgO 0.0 0.000 

Water -133.4 -0.105 

Tar Reforming Catalyst 0.0 0.000 

Other 0.0 0.000 

Total 1138.6 0.897 

Energy Outlets 

Ethanol 619.1 0.488 

Higher Alcohols Co-product 122.1 0.096 

Cooling Tower Evaporation 17.0 0.013 

Flue Gas 46.2 0.036 

Sulfur 0.4 0.000 

Compressor Heat 178.3 0.140 

Heat from Air-cooled Exchangers 222.0 0.175 

Vents to Atmosphere 

(including excess CO2) 

0.8 0.001 

Sand and Ash 16.4 0.013 

Catalyst Purge 0.0 0.000 

Wastewater -1.2 -0.001 

Other -82.5 -0.064 

Total 1138.6 0.897 

Besides the saleable alcohol products, other important process energy outlets also exist. There 

are two sources of flue gas: the char combustor and the reformer fuel combustor. Together, they 

total about 4% of the energy in the raw biomass. Cooling tower evaporative losses, excess CO2

vent to the atmosphere, wastewater, and ash streams are also minor process energy outlets. 

However, two of the larger energy outlets come from air-cooled interstage cooling of the 

compressors, and from several other air-cooled heat exchangers. Together, these two loss 

categories represent over 30% of the energy that is not recovered within the process. The “other” 

category consists primarily of other losses from the cooling tower system (drift and blowdown), 

but also accounts for energy losses due to ambient heating effects and mechanical work (pump, 

compressor) efficiency losses. 

Some of this lost heat could potentially be recovered by using cooling water instead of air-cooled 

exchangers. However, this would require additional makeup water, and limiting water usage 

throughout the process was a primary design consideration. Additional heat integration with 

process streams could also be examined, however, there comes a point where this becomes too 

complex and costly for a cost-effective design and practical operation. 

Overall, the TC process is approximately 46% efficient on an LHV basis for moisture-free 

biomass, as shown in the Appendices. Table 12 shows that approximately 58% of the energy in 

the wet raw biomass is recovered in the two alcohol products. Improvements in these energy 

efficiencies could potentially result in additional cost savings to the process. 

 33 
3.12. Water Issues 

Water is required as a reactant, a fluidizing agent, and a cooling medium in this process. As a 

reactant, it participates in reforming and water gas shift reactions. Using the BCL gasifier, it also 

acts as the fluidizing agent in the form of steam. Its cooling uses are outlined in Section 3.8. 

Water usage is becoming an increasingly important aspect of plant design, specifically with 

regards to today’s ethanol plants. Most ethanol plants reside in the Midwest where many places 

are experiencing significant water supply concerns51. For several years, significant areas of 

water stress have been reported during the growing season, while livestock and irrigation 

operations compete for the available resources. 

Today’s dry mill ethanol plants have a high degree of water recycle. In fact many plants use what 

is known as a “zero discharge” design where no process water is discharged to wastewater 

treatment. The use of centrifuges and evaporators enables this recycle of process water. 

Therefore, much of the consumptive water demand of an ethanol plant comes from the 

evaporative losses from the cooling tower and utility systems. Oftentimes well water is used to 

supply the water demands of the ethanol plants, which draws from the local aquifers that are not 

readily recharged. This is driven by the need for high quality water in the boiler system. Studies 

have shown that water usage by today’s corn ethanol plants range from 3-7 gallons per gallon of 

ethanol produced. This means that a 50 MM gal/yr dry mill will use between 150-350 MM 

gallons/yr of water that is essentially a non-renewable resource. This ratio however has 

decreased over time from an average of 5.8 gal/gal in 1998 to 4.2 gal/gal in 2005. 

Therefore, a primary design consideration for this process was the minimization of fresh water 

requirements, which therefore meant minimizing the cooling water demands and recycling 

process water as much as possible. Air-cooling was used in several areas of the process in place 

of cooling water (e.g. distillation condensers, compressor interstage cooling, etc). However there 

are some instances where cooling water is required to reach a sufficiently low temperature that 

air-cooling can not reach. 

Table 13 quantifies the particular water demands of this design. Roughly 71% of the fresh water 

demand is from cooling tower makeup, with most of the remainder needed as makeup boiler feed 

water. Some of this water is directly injected into the gasifier, but other system losses 

(blowdown) also exist. The overall water demand is considerably less than today’s ethanol 

plants. This design requires less than 2 gallons of fresh water for each gallon of ethanol 

produced. It may be worthwhile for the entire ethanol industry to more thoroughly investigate 

efficiency gains that are possible within these utility systems. 

Table 13. Process Water Demands for Thermochemical Ethanol 

Fresh Water Demands lb per hour 

Cooling Tower Makeup 84,672 

Boiler Feed Makeup 34,176 

Sand/ash Wetting 243 

Total 119,091 

Overall Water Demand (gal water / gal 

ethanol) 

1.94 

 34 
4. Process Economics 

The total project investment (based on total equipment cost) as well as variable and fixed 

operating costs is developed first. With these costs, a discounted cash flow analysis was used to 

determine the production cost of ethanol when the net present value of the project is zero. This 

section describes the cost areas and the assumptions made to complete the discounted cash flow 

analysis. 

4.1. Capital Costs 

The following sections discuss the methods and sources for determining the capital cost of each 

piece of equipment within the plant. A summary of the individual equipment costs can be found 

in Appendix D. 

The capital cost estimates are based as much as possible on the design work done by Spath et al . 

for the hydrogen design report [9] and Aden et al. for the biochemical conversion design report 

[3]. The majority of the Spath et al. costs came from literature and Questimate (an equipment 

capital cost estimating software tool by Aspen Tech), not from vendor quotes. For these 

estimated costs, the purchased cost of the equipment was calculated and then cost factors were 

used to determine the installed equipment cost. This method of cost estimation has an expected 

accuracy of roughly +30% to -10%. The factors used in determining the total installed cost (TIC) 

of each piece of equipment are shown in Table 14 [52]. The Aden et al. cost estimates came 

from a variety of sources (including vendor quotes); the installation factors for these estimates 

may be significantly different from what is in Table 14. 

Table 14. General Cost Factors in Determining Total Installed Equipment Costs 

% of TPEC 

Total Purchased Equipment Cost (TPEC) 100 

 Purchased equipment installation 39 

 Instrumentation and controls 26 

 Piping 31 

 Electrical systems 10 

 Buildings (including services) 29 

 Yard improvements 12 

Total Installed Cost (TIC) 247 

The indirect costs (non-manufacturing fixed-capital investment costs) were also estimated as per 

Spath et al. using cost factors. The factors are shown in Table 15 [52] and have been put as 

percentages in terms of total purchased equipment cost, total installed cost (TIC), and total 

project investment (TPI, the sum of the TIC and the total indirect costs). 

 35 
Table 15. Cost Factors for Indirect Costs 

Indirect Costs % of TPEC % of TIC % of TPI 

 Engineering 32 13 9 

 Construction 34 14 10 

 Legal and contractors fees 23 9 7 

 Project contingency 7.4 3 2 

 Total Indirect Costs 96.4 39 28 

The biomass handling and drying costs as well as the gasification and gas clean up costs were 

estimated by Spath et al. using several reports by others that documented detailed design and cost 

estimates. Some of the reports gave costs for individual pieces of equipment while others lumped 

the equipment costs into areas. The costs from the reports were amalgamated into: 

• feedstock handling and drying. 

• gasification and clean up. 

Costs from those reports scaled to a 2,000 bone dry tonne/day plant are given in Table 16. Table 

17 gives the basic dryer and gasifier design basis for the references. Spath et al. used an average 

feed handling and drying cost from all of the literature sources and an average gasifier and gas 

clean up cost for the references using the BCL gasifier. 

Table 16. Feed Handling & Drying and Gasifier & Gas Clean Up Costs from the Literature Scaled 

to 2,000 tonne/day plant 

Reference Scaled Feed 

Handling and 

Drying Cost $K 

(2002) 

BCL - Scaled 

Gasifier and Gas 

Clean Up Cost $K 

(2002) 

Breault and Morgan [53] (a) $15,048 $15,801 

Dravo Engineering Companies [54] (a) $14,848 $15,774 

Weyerhaeuser, et al., [55]

 (a) $21,241 $24,063 

Stone & Webster, et al. [56]

 (a) $25,067 --- 

Wan and Malcolm [57] (a) $18,947 (b)

$14,098 (c)

$11,289 (b)

$11,109 (c)

Weyerhaeuser [58] (a) $13,468 $10,224 

Wright and Feinberg [59] (a) $26,048 – BCL 

design 

$21,942 – GTI design 

$12,318 - quench (d)

$26,562 - HGCU (d)

Craig [60] $13,680 --- 

AVERAGE $18,840 $16,392 

(a) From detailed design and cost estimates 

(b) Estimated from a 200 dry ton/day plant design. 

(c) Estimated from a 1,000 dry ton/day plant design. 

(d) Two separate gas clean up configurations were examined for the BCL gasifier. HGCU = hot 

gas clean up. 

 36 
Table 17. System Design Information for Gasification References 

Reference Feed Handling and 

Drying 

BCL Gasifier and Gas 

Clean Up 

Breault and Morgan [53] Rotary dryer Cyclones, heat exchange & 

scrubber 

Dravo Engineering Companies 

[54] 

Rotary drum dryer Cyclones, heat exchange & 

scrubber 

Weyerhaeuser, et al. [55] Steam dryer Cyclones, heat exchange, 

tar reformer, & scrubber 

Stone & Webster, et al. [56] Flue gas dryer --- 

Wan and Malcolm [57] Flue gas dryer Cyclones, heat exchange & 

scrubber 

Weyerhaeuser [58] Flue gas dryer Cyclones, heat exchange & 

scrubber 

Wright and Feinberg [59] Unclear Quench system – details 

are not clear 

Tar reformer system – 

details are not clear 

Craig [60] Rotary drum dryer --- 

In this report, we have further broken apart the gasification and clean up costs into their 

respective areas. Based upon the Utrecht report [19] these were split 50/50 between the two 

areas. 

The cost of reactors, heat exchangers, compressors, blowers and pumps were estimated for a 

“base” size using Questimate and then scaled using material and energy balance results from the 

Aspen Plus simulation. The reactors were sized based on a gas hourly space velocity (GHSV), 

where GHSV is measured at standard temperature and pressure, 60°F and 1 atm [61], and a 

height to diameter ratio of 2. The GHSV for the mixed alcohol reactor and tar reformer were set 

at 4,000/hr and 2475/hr, respectively. These are in agreement with typical values given by Kohl 

and Nielsen [62]. The heat exchanger costs were mostly developed based on the required surface 

area as calculated from the heat transfer equation appropriate for a 1-1 shell and tube heat 

exchanger: 

( ) ( ) lm

lm

Q Q UA T A U T

= Δ ⇒=

Δ

where Q is the heat duty, U is the heat transfer coefficient, is the exchanger surface area, and 

is the log mean temperature difference. The heat transfer coefficients were estimated 

from literature sources, primarily Perry, et al [

A

( )lm ΔT

27]. However, many of the exchangers used in the 

pinch analysis are subsequently scaled from their calculated duties. At present, these duties will 

not change as the process changes, unless the pinch calculations are specifically updated. This is 

acceptable as long as the total cost of the heat exchange network remains a small fraction of the 

overall minimum ethanol plant gate price, and as long as plant scale does not change 

significantly. 

 37 
For the various pieces of equipment, the design temperature is determined to be the operating 

temperature plus 50°F (28°C) [63]. The design pressure is the higher of the operating pressure 

plus 25 psi or the operating pressure times 1.1 [63]. 

The following costs were estimated based on the Aden, et al. design report: [3] 

• cooling tower. 

• plant and instrument air. 

• steam turbine/generator/condenser package. 

• Deaerator. 

• alcohol separation equipment (e.g., the distillation columns and molecular sieve unit). 

Appendix G contains the design parameters and cost references for the various pieces of 

equipment in the plant. 

4.2. Operating Costs 

There are two kinds of operating costs: variable and fixed costs. The following sections discuss 

the operating costs including the assumptions and values for these costs. 

There are many variable operating costs accounted for in this analysis. The variables, 

information about them, and costs associated with each variable are shown in Table 18. 

Table 18. Variable Operating Costs 

Variable Information and Operating Cost 

Tar reformer catalyst To determine the amount of catalyst inventory, the tar reformer was 

sized for a gas hourly space velocity (GHSV) of 2,476/hr based on 

the operation of the tar reformer at NREL’s TCPDU where GHSV is 

measured at standard temperature and pressure [61]. Initial fill then a 

replacement of 1% per day of the total catalyst volume. 

Price: $4.67/lb [64] 

Alcohol Synthesis 

Catalyst 

Initial fill then replaced every 5 years based on typical catalyst 

lifetime. 

Catalyst inventory based on GHSV of 6,000/hr. 

Price: $5.25/lb [7] 

Gasifier bed material Synthetic olivine and MgO. Delivered to site by truck equipped with 

self-contained pneumatic unloading equipment. Disposal by landfill. 

Olivine price: $172.90/ton [65] 

MgO price: $365/ton [66] 

Solids disposal cost Price: $18/ton [67] 

Diesel fuel Usage: 10 gallon/hr plant wide use 

Price: $1.00/gallon [68] 

Chemicals Boiler chemicals – Price: $2.80/lb [3] 

Cooling tower chemicals – Price: $2.00/lb [3] 

LO-CAT chemicals – Price: $150/tonne of sulfur produced [69] 

Waste Water The waste water is sent off-site for treatment. 

Price: $2.07/100ft3

 [70] 

 38 
Previous biomass gasification studies have not looked at fixed operating costs (i.e. salaries, 

overhead, maintenance, etc) in detail, therefore little data were available. As a result, the fixed 

operating costs for a biochemical ethanol facility given in Aden, et al., 2002 [3] were used as a 

starting point to develop fixed costs for this thermochemical design. 

The fixed operating costs used in this analysis are shown in Table 19 (labor costs) and Table 20

(other fixed costs). They are shown in 2002 U.S. dollars. The following changes in base salaries 

and number of employees were made compared to those used in the ethanol plant design in 

Aden, et al., 2002 [3]. 

• Plant manager salary raised from $80,000 to $110,000 

• Shift supervisor salary raised from $37,000 to $45,000 

• Lab technician salary raised from $25,000 to $35,000 

• Maintenance technician salary raised from $28,000 to $40,000 

• Shift operators salaries raised from $25,000 to $40,000 

• Yard employees salaries raised from $20,000 to $25,000 and number reduced from 32 to 

12. 

• General manager position eliminated 

• Clerks and secretaries salaries raised from $20,000 to $25,000 and number reduced from 

5 to 3. 

The number of yard employees was changed to reflect a different feedstock and feed handling 

system compared to Aden, et al., 2002 [3]. Handling baled stover requires more hands-on 

processing when compared to a wood chip feedstock. Based on a 4-shift system, 3 yard 

employees were estimated to be needed, mostly to run the front end loaders. The general 

manager position was eliminated because a plant manager would likely be sufficient for this type 

of facility. Biomass gasification plants are more likely to be operated by larger companies 

instead of operating like the dry mill ethanol model of farmer co-ops. Finally, the number of 

clerks and secretaries was reduced from 5 to 3. The estimate of three comes from needing 1 to 

handle the trucks and scales entering and leaving the facility, 1 to handle accounting matters, and 

1 to answer phones, do administrative work, etc. 

Table 19. Labor Costs 

Position Salary Number Total Cost 

Plant manager $110,000 1 $110,000 

Plant engineer $65,000 1 $65,000 

Maintenance supervisor $60,000 1 $60,000 

Lab manager $50,000 1 $50,000 

Shift supervisor $45,000 5 $225,000 

Lab technician $35,000 2 $70,000 

Maintenance technician $40,000 8 $320,000 

Shift operators $40,000 20 $800,000 

Yard employees $25,000 12 $300,000 

Clerks & secretaries $25,000 3 $75,000 

Total salaries (2002 $) $2,080,000 

 (2005 $) $2,270,000 

 39 
Since the salaries listed above are not fully loaded (i.e. do not include benefits), a general 

overhead factor was used. This also covers general plant maintenance, plant security, janitorial 

services, communications, etc. The 2003 PEP yearbook [71] lists the national average loaded 

labor rate at $37.66 per hr. Using the salaries in Table 19 above along with the 60% general 

overhead factor from Aden, et al. [3] gave an average loaded labor rate of $30 per hr. To more 

closely match the PEP yearbook average, the overhead factor was raised to 95%. The resulting 

average loaded labor rate was $36 per hr. 

Table 20. Other Fixed Costs 

Cost Item Factor Cost 

General overhead 95% of total salaries $2,155,000 

Maintenance [52] 2% of total project investment $3,817,000 

Insurance & taxes [52] 2% of total project investment $3,817,000 

The updated salaries in Table 19 above were examined against salaries from a free salary 

estimation tool [72] which uses Bureau of Labor Statistics data and several other sources. 

Because the biomass analysis does not reflect a specific site in the United States, National 

Average Salaries for 2003 were used. With such an extensive listing of job titles in the salary 

estimation tool, a general position such as “clerks and secretaries” could be reflected by multiple 

job titles. In these instances, care was taken to examine several of the possible job titles that were 

applicable. A list of the job positions at the production plant and the corresponding job titles in 

the salary estimation tool [72] is shown in Table 21. Overall, the salaries used in the biomass-tohydrogen production plant design are close to the U.S. national average values given in column 

4. 

 40 
Table 21. Salary Comparison 

Job Title in 

Biomass Plant 

Corresponding Job 

Title in Salary 

Estimating Tool [72] 

Salary Range 

(17th to 67th

percentile) 

Average 

Salary (U.S. 

national 

average) 

Salary used 

in Biomass 

Plant Design 

(see Table 19) 

Plant manager Plant manager 

(experience) 

$81,042-

$220,409 

$106,900 $110,000 

Plant engineer Plant engineer $36,213-$66,542 $58,324 $65,000 

Maintenance crew 

supervisor 

$35,036-$53,099 $45,191 

Supervisor 

maintenance 

$34,701-$56,097 $47,046 

Maintenance 

supervisor 

Supervisor 

maintenance & 

custodians 

$23,087-$45,374 $39,924 

$60,000 

Lab manager Laboratory manager $38,697-$70,985 $51,487 $50,000 

Shift supervisor Supervisor production $32,008-$51,745 $43,395 $45,000 

Lab technician Laboratory technician $25,543-$41,005 $34,644 $35,000 

Maintenance 

technician 

Maintenance worker $27,967-$46,754 $39,595 $40,000 

Shift operators Operator control room $33,983-$61,362 $49,243 $40,000 

Yard employees Operator front end 

loader 

$24,805-$39,368 $31,123 $25,000 

Administrative clerk $19,876-$25,610 $26,157 

Secretary $20,643-$31,454 $26,534 

Clerks & 

secretaries 

Clerk general $15,984-$25,610 $22,768 

$25,000 

Overall, Aden, et al.[3] lists fixed operating costs totaling $7.54MM in $2000. Using the labor 

indices, this equates to $7.85MM in $2002. On the other hand, the mixed alcohols design report 

has fixed operating costs totaling $12.06MM in $2005. 

4.3. Value of Higher Alcohol Co-Products 

The alcohol synthesis process will create higher molecular weight alcohols. How this co-product 

is valued will depend upon its end market. There were two extreme cases envisioned. At the high 

end, these might be sold into the chemical market. This could command a high value for this coproduct, upwards to $3.70 to $4.20 per gallon [7]. However, it is unlikely that the market would 

support more than one or two biomass plants to support these prices. Because of this, the 

biomass process did not include any detailed separation or clean-up of the separate alcohols. It is 

envisioned that if this co-product was sold for this purpose, it would be transferred “over the 

fence” as is and the buyer would take on the costs of separation and clean-up. So, even at the 

high end, the highest value would be some fraction of the chemical market value. 

At the low end, the co-product could command a value for a fuel with minimal ASTM standards 

on its specifications. This would be priced similar to a residual fuel oil. Historically, this is about 

80% of gasoline price [73]. Using the ethanol minimum plant gate price as a scaled reference 

gasoline price (adjusted for ethanol’s lower heating value), this translates to $0.85 per gallon. 

 41 
For the baseline case, a middle ground was chosen. It is anticipated that the higher alcohols 

would make an excellent gasoline additive or gasoline replacement in its own right – engine 

testing and certification would be required. If this is done, then it should command a price 

similar to that of gasoline. Again using the ethanol minimum plant gate price as a scaled 

reference gasoline price and adjusting to n-propanol’s heating value (the major constituent of the 

higher alcohol stream), then its value should be $1.25 per gallon. However, since no special 

efforts were taken in the process design to clean up this stream to meet anticipated specs, its 

value is discounted to $1.15 per gallon. 

4.4. Minimum Ethanol Plant Gate Price 

Once the capital and operating costs were determined, a minimum ethanol selling price (MESP) 

was determined using a discounted cash flow rate of return analysis. The methodology used is 

identical to that used in Aden, et al., (2002) [3]. The MESP is the selling price of ethanol that 

makes the net present value of the process equal to zero with a 10% discounted cash flow rate of 

return over a 20 year plant life. The base case economic parameters used in this analysis are 

given in Table 22. A sensitivity analysis was performed to examine the MESP for different 

financial scenarios. These are discussed in Section 4. 

Table 22. Economic Parameters 

Assumption Value 

Internal rate of return (after-tax) 10% 

Debt/equity 0%/100% 

Plant life 20 years 

General plant depreciation 200% DDB 

General plant recovery period 7 years 

Steam plant depreciation 150% DDB 

Steam plant recovery period 20 years 

Construction period 

1st 6 months expenditures 

Next 12 months expenditures 

Last 12 months expenditures 

2.5 years 

8% 

60% 

32% 

Start-up time 

Revenues 

Variable costs 

Fixed costs 

6 months 

50% 

75% 

100% 

Working capital 5% of Total Capital Investment 

Land 6% of Total Purchased Equipment Cost 

(Cost taken as an expense in the 1st

construction year) 

Note: The depreciation amount was determined using the same method as that documented in 

Aden, et al.[3] using the IRS Modified Accelerated Cost Recovery System (MACRS). 

 42 
5. Process Economics, Sensitivity Analyses, and Alternate Scenarios 

The cost of ethanol as determined in the previous section was derived using technology that has 

been developed and demonstrated or is currently being developed as part of the OBP research 

program. Combined, all process, market, and financial targets in the design represent what must 

be achieved to obtain the reported $1.01 per gallon. A summary of the breakdown of costs are 

depicted in Figure 8 and further tabulated in Appendix F. 

Feedstock

Feed Handling & Drying

 Gasification

 Tar Reforming; Acid 

Gas & Sulfur Removal

 Alcohol Synthesis - 

Compression

 Alcohol Synthesis - 

Other

 Alcohol Separation

 Steam System & 

Power Generation

 Cooling Water & Other 

Utilities

-$0.30 -$0.20 -$0.10 $0.00 $0.10 $0.20 $0.30 $0.40 $0.50

Capital Recovery Charge Catalysts, Raw Materials, & Waste Process Electricity

Electricity Generated Co-Product Credits Fixed Costs

43.7¢

13.2¢

10.4¢

31.3¢

8.8¢

-18.0¢ (Net)

4.1¢

4.5¢ (Net)

3.2¢ $1.01 MESP

Figure 8. Cost contribution details from each process area 

This cost contribution chart appears to show two different co-product credits: alcohols from the 

Alcohol Synthesis area and electricity from the Steam System & Power Generation area. 

However, the process was adjusted so the electricity generated is balanced by the electricity 

required by all other areas, so there is no net credit for electricity generation. 

The cost year chosen for the analysis had a significant effect on the results. As discussed in 

Section 1.1, capital costs increased significantly after 2003 primarily because of the large 

increase in steel costs worldwide. Figure 9 depicts how the MESP for this process would change 

depending on the cost year chosen for the analysis. Notice that between the years 2000 to 2003 

the MESP would be much lower, $0.89 to $0.91 per gallon ethanol, instead of the $1.01 

determined for 2005. The values for 2006 are tentative, since all factors necessary for the MESP 

calculation have not yet been published. 

 43 
$0.00

$0.20

$0.40

$0.60

$0.80

$1.00

$1.20

Mininum Ethanol Selling Price, MESP 

($ per gal ethanol)

Conversion Costs $0.54 $0.54 $0.54 $0.55 $0.61 $0.65 $0.70

Feedstock Contribution $0.35 $0.35 $0.36 $0.36 $0.36 $0.36 $0.37

Minimum Ethanol Selling Price ($/gal) $0.89 $0.89 $0.90 $0.91 $0.97 $1.01 $1.07

2000 2001 2002 2003 2004 2005 2006

Figure 9. Effect of cost year on MESPh

The process costs (as indicated by the MESP) are determined from various assumptions on 

technology (based upon 2012 research targets), markets (such as the value of the higher alcohol 

co-products), and various financial assumptions (such as required Return on Investment, ROI). 

When any research target cannot be obtained, or a market or financial assumption does not hold, 

then the MESP is affected to varying degrees. In addition, uncertainty about equipment design 

and installation and construction costs will impact the economics. The key is to understand the 

impact of those types of parameters that are likely to vary, and how they might be controlled to a 

definable range. Discussed here are process targets that had been identified a priori as key ones 

to understand and achieve. (As can be seen from the sensitivity results, many items examined 

had much less affect on the MESP than had been thought.) In most cases, values used for the 

sensitivities are picked from current experimental data, to demonstrate the effect of technology 

advancement (or lack of) on the economic viability of the process. 

 h

 Note that the relative splits between feedstock and conversion costs have been scaled to attribute some of the costs 

to the mixed alcohol co-products. So, the feedstock contribution appears to be different than what is depicted in the 

cost contribution chart for the different areas. 

 44 
The results for the sensitivity analysis discussed in the following sections are depicted in Figure 

10; those sensitivities directly impacted by research programs are shown first. Nearly all of these 

ranges represent variations of a single variable at a time (e.g., ash content while holding the ratio 

of the non-ash elements constant). There are a couple exceptions to this: 

• The feedstock comparison of corn stover to lignin necessitated varying the ultimate 

elemental analysis, ash content, and moisture content simultaneously. 

• The Combined Tar Reformer Conversions incorporated all of the ranges listed for the 

methane, benzene, and tar simultaneously. 

Note that all items in the chart have values associated with them. If a bar is not readily seen, then 

the MESP effect over the range listed is insignificant. 

($0.50) ($0.25) $0.00 $0.25 $0.50 $0.75 $1.00 $1.25 $1.50

Co-Product Values (69% Chemical Market:baseline:Fuel Oil Value)

Plant Size (10,000:2000:600 dry tonnes/day).

Stream Factor (.98; .96; .90)

Total Project Investment (-10%:baseline:+30%)

Average Installation Factors (-10%:baseline:+30%)

Contingency (0%:3%:15% of TIC)

Return on Investment (0%:10%:30%)

Loan vs. Equity Financing (100% debt @ 7.5%:100% Equity:100%

Feedstock Cost ($10:$35:$53 per dry ton)

Feedstocks (Lignin:wood:Corn Stover)

Feedstock Quality - Ash (1%:1%:12%)

Feedstock Quality - Ash (1%:1%:12%) Feed Cost Adj for ash

Sulfur Content (Baseline to 4X)

Feed Moisture Content (15%:50%:70%)

Reduced CH4 to CO (baseline:baseline:25%)

Olivine cost (1/10:baseline:10X)

Tar Reformer Methane Conversion (95%:80%:50%)

Tar Reformer Benzene Conversion (99.9%:99%:90%)

Tar Reformer Tar Conversion (99.9%:99.9%:95%)

Combined Tar Reformer Conversions

Tar Reformer Equipment Costs (-10%:baseline:+100%)

Acid Gas Removal Equip Costs (-10%:baseline:+100%)

Level of CO2 removal (10:5:0.1 mol%)

Single Pass CO conversion (80%:60%:30%)

CO Selectivity to Alcohols (95%:90%:70%)

Operating Pressure (800:1000:2,000 psia)

Total Alcohol Catalyst Productivity (1,000:600:200 g/kg-cat/hr) 

Catalyst Lifetime (10 yrs:5yrs: 1 yr)

Catalyst Poison Allowability (100:50:10 ppm)

Catalyst cost ($2.50:$5.25:$2,250 per lb)

Recycling Unconverted Syngas to Synthesis Reactor (25%:0%:0%)

Change to MESP ($ per gallon Ethanol)

Research

Financial / Market

Figure 10. Results of sensitivity analyses 

All analyses are discussed further in the following sections. 

5.1. Financial Scenarios 

These parameters have the greatest effect on the MESP but R&D has the smallest direct effect on 

them. In particular, the required ROI for the project could more than double the calculated 

MESP. Successful R&D and demonstration projects would, at best, ease the ROI requirements of 

corporations and/or lending institutions and reduce the required MESP toward the baseline case 

 45 
in this report. Also, the baseline of 0% debt financing is not a very realistic scenario, but does 

represent the conservative endpoint. Many projects of this nature are often financed by some 

mixture of debt and equity financing. However, the magnitude of this parameter’s effect on 

MESP is quite small in comparison to many of the other financial and market parameters. 

A conceptual design like this is normally thought to give accuracy in the capital requirements of 

-10% to +30%. Using this range for the TPI (Total Project Investment) gives an MESP range of -

6% to +20%. 

5.2. Feedstocks 

Because this process has been designed for utilization of forest resources there may be little 

control over the feedstock quality coming to the planti

. The two most important feedstock q

parameters that can most impact the process economics are moisture and ash content. 

uality 

 

The high range of the ash content examined here are more indicative of agricultural residues 

(from fertilizer) or lignin-rich biochemical process residues; forest resources should have ash 

contents near that of this baseline case (about 1%). It was originally thought that the cost effects 

of high ash content could be damped by basing feedstock payments on a dry and ash-free basis, 

not just a dry basis. However, Figure 11 shows that this is not the case. Increased ash in the 

feedstock results in larger ash handling equipment and power requirements, especially in the 

electrostatic precipitator (ESP) used to remove ash fines from the flue gas. These higher power 

requirements are met by diverting more syngas to the fuel system to generate electricity. Keeping 

the feedstock cost constant on a moisture and ash free (“maf”) basis decreases the MESP for 

high-ash feeds by reducing the cost per pound of biomass. However, at a constant mass feedrate 

to the process, there is inherently less carbon available for conversion to alcohols and therefore 

smaller revenues. The reduced revenues together with increased capital and operating costs result 

in an overall increase in MESP despite the lower feedstock cost. 

The operating costs due to ash disposal may be reduced by finding an alternate use for the ash. 

One potential use may be as a soil amendment to replace minerals lost from the soil. The ash 

collected from gasification in this case should be comparable to the minerals removed from the 

soil during the plant growth. More study would be needed to determine the best and most 

economic method for using the ash as a soil amendment. 

i

 At least less so than using agricultural residues or energy crops that can be bred for specific properties in these 

lignocellulosic materials. 

 46 
0

10

20

30

40

50

60

70

80

90

100

0.00 5.00 10.00 15.00 20.00

Biomass Ash Content (wt%)

Alcohol Yield (gal/dry ton)

$1.00

$1.25

$1.50

$1.75

$2.00

$2.25

$2.50

Minimum Ethanol Selling Price

($ per gal)

Ethanol

Total Alcohols

$35.00 per dry ton

$35.32 per maf ton

Figure 11. Sensitivity analysis of biomass ash content 

The biomass feed’s moisture content is a problem if it is higher than the baseline 50%. This is 

not envisioned as being very likely except in the case of processing wet ensiled agricultural 

residues or energy crops; however, these feedstocks are more envisioned to be processed by 

biochemical means, not thermochemical means. Drier feedstocks will have lower MESPs 

because of decreased heat requirements to dry the incoming feedstock directly relate to lower 

raw syngas diversion to heat and power and higher alcohol yields. This is depicted in Figure 12

and Figure 13. As the moisture content increases, the alcohol yield will decrease because more 

raw syngas must be diverted for heat. Note that very low moisture contents do not give 

corresponding increased alcohol yields; this is because flue gas is used for the drying and other 

operating specifications dictate the amount of raw syngas diverted for heat and power, not the 

feedstock drying. 

 47 
0

20

40

60

80

100

120

0% 10% 20% 30% 40% 50% 60% 70% 80%

Moisture Inlet Biomass (wt%)

Alcohol Yield (gal/dry ton)

$0.90

$1.00

$1.10

$1.20

$1.30

$1.40

$1.50

Minimum Ethanol Selling Price

($ per gal)

Total Alcohols

Ethanol

Figure 12. Sensitivity analysis of biomass moisture content 

 48 
0

20

40

60

80

100

120

0% 10% 20% 30% 40% 50% 60%

Raw Syngas to Fuel & Power

Alcohol Yield (gal/dry ton)

$0.90

$1.00

$1.10

$1.20

$1.30

$1.40

$1.50

Minimum Ethanol Selling Price

($ per gal)

Total Alcohols

Ethanol

Figure 13. Sensitivity analysis of raw syngas diverted for heat and power due to biomass moisture 

content 

Two combined scenarios were analyzed for two different kinds of feedstocks: corn stover and 

lignin-rich residues from a biochemical process. The compositions of both are consistent with the 

Aden et al. design report [3]. Corn stover gives rise to a higher MESP even though its elemental 

analysis is very similar to wood and its moisture content is very low. The overwhelming effect is 

due to its higher ash content. Lignin-rich residues have a much lower MESP. Lignin-rich 

residues also have the virtue of making more electricity than the process needs, so it is exported, 

even if the raw syngas to the fuel system is minimized while still achieving all other operating 

specifications. This is a very positive sign that incorporating a thermochemical conversion unit 

with a biochemical conversion unit and make the heat and power for the entire complex will be 

cost effective. The feed handling system may have to be different, however, since lignin tends to 

get very powdery when dried; direct contact with the flue gas for drying would very likely lead 

to high losses of the feedstock. Drying with indirect contact of the heating medium must be 

investigated. 

 49 
5.3. Thermal Conversion 

Two gasification scenarios were examined. The first was to explore the impact of increasing the 

olivine cost which could come about from catalytic modification. Increasing the olivine cost by 

an order of magnitude could increase the MESP by 6%. This is not significant. The second 

scenario examined the effect of reducing methane production in the gasifier to reflect a case 

where the gasifier does not operate within the given correlations. This gave an unexpected result. 

It was expected that the MESP would decrease when it actually increased a nearly insignificant 

amount (0.1%). There are two reasons for this. One is an artifact of the way in which the 

decrease was modeled – more CO was formed and the hydrogen that would have gone to the 

methane instead went to the char (and was lost for further processing). This required 

modifications to the operations to keep the H2:CO ratio to the alcohol synthesis reactor above 1.0 

and, in doing so, increased the MESP. 

5.4. Clean-Up & Conditioning 

These scenarios appeared to have an imperceptible effect on the MESP. However, this is 

misleading. The scenarios show primarily cost effects due to the material and energy balances. 

Since the amount of tar is small compared to the amount of CO and H2, these effects are small. 

In reality, Clean-Up and Conditioning is absolutely required for acceptable performance of gas 

compressors, waste water treatment, and alcohol synthesis catalysts. Excessive tars in the syngas 

would significantly impact compressors and waste water treatment with severe consequences to 

equipment and increased operating costs that are not rigorously modeled here. So, not meeting 

these targets would give poor performance, leading to greater cost effects than reflected by the 

sensitivity analysis for this area. 

5.5. Fuels Synthesis 

These scenarios show the importance of the R&D for the synthesis catalysts. Poor performance 

could increase MESP by 25% or more. Whether this is due to actual non-target catalyst 

formulations or due to poor performance in Clean-Up and Conditioning that leads to poor 

alcohol synthesis catalyst performance, the cost effects are major. The catalyst cost sensitivity 

range was extremely large, from $2.50/lb to over $2,250/lb. This was done to bracket a variety of 

potential catalyst systems, not just cobalt moly-sulfide. Exotic metals such as rhodium (Rh) or 

ruthenium (Ru) can add considerable cost to a catalyst system even at relatively low 

concentrations. At low catalyst costs, total CO conversion and alcohol selectivity (CO2-free 

basis) have the largest impact on the overall MESP. The catalyst productivity (g/kg/hr) did not 

show much impact over the sensitivity range chosen. In reality, all of these catalyst performance 

indicators are tightly linked. It is unlikely that research could change one without affecting the 

others. 

5.6. Markets 

Crediting the co-product higher alcohols with the lower fuel oil value increases the MESP by 

about 6%; this is still reasonable to meet the qualitative “cost competitive” target. Of even more 

significance is that selling these higher alcohols for even 69% of their chemical market value 

will lead to a significant reduction of MESP (about 20%). This shows that the first couple 

thermochemical conversion plants could get a significant economic advantage in their early life 

by being able to do this. 

 50 
6. Conclusions 

This analysis shows that biomass-derived ethanol from a thermochemical conversion process has 

the possibility of being produced in a manner that is “cost competitive with corn-ethanol” by 

2012. This thermochemical conversion process would make use of many sub-processes that are 

currently used commercially (such as acid gas removal) but also requires the successful 

demonstration of R&D targets being funded by the DOE’s OBP. 

This analysis has demonstrated that forest resources can be converted to ethanol in a cost 

competitive manner allowing greater flexibility in converting biomass resources to achieve stated 

volume targets by 2030. 

7. Future Work 

Future R&D work to develop and demonstrate reforming and synthesis catalysts is inherent in 

this study. There many other areas of demonstration and process development also required: 

• Demonstrate gasifier performance on other feedstocks (agricultural residues such as corn 

stover, energy crops such as switchgrass, and lignin-rich residues that would be available 

from a co-located biochemical conversion process). Of particular importance for the 

lignin-rich residues is the impact on process performance of trace amounts of chemicals 

used in the biochemical processing that might negatively impact the thermochemical 

conversion process. 

• Compare the relative merits of direct oxygen blown gasifiers to the indirect steam 

gasifier upon which this study is based. 

• Examine the trade-offs of the greater use of water cooling (greater water losses in the 

cooling tower) vs. air cooling (greater power usage) vs. organic Rankine cycle for 

cooling and power production. 

• Better understand the trade offs between operating conditions in the alcohol synthesis 

reactor to operating conditions (pressure, temperature, extent of reaction, extent of 

methanol recycle). A “tuned” kinetics based model would be required for this. 

• Explore alternate synthesis reactor configurations (slurry phase vs. fixed bed). 

• Understand trade offs between an energy neutral, alcohol production facility to one that 

could also supply heat and electricity to a co-located biochemical conversion facility. 

• Further explore the potential benefits of integrating biochemical and thermochemical 

technologies. 

 51 
• Examine potential for decreased heat integration complexity and increased overall 

energy efficiency. 

• Better understand the kinetics of catalytic tar reforming and deactivation, and the 

necessary regeneration kinetics to achieve a sustainable tar reforming process. 

• Examine the emissions profile from the plant and explore alternate emissions control 

equipment. 

 52 
8. References 

1 2006 State of the Union address, January 31, 2006. 

http://www.whitehouse.gov/stateoftheunion/2006/ 

2 http://www.whitehouse.gov/news/releases/2006/01/20060131-6.html 

3 Aden, A.; Ruth, M.; Ibsen, K.; Jechura, J.; Neeves, K.; Sheehan, J.; Wallace, B.; Montague, L.; 

Slayton, A.; Lukas, J. Lignocellulosic Biomass to Ethanol Process Design and Economics 

Utilizing Co-Current Dilute Acid Prehydrolysis and Enzymatic Hydrolysis for Corn Stover. 

NREL Report No. TP-510-32438. June 2002. http://www.nrel.gov/docs/fy02osti/32438.pdf 

4 CEH Marketing Research Report - Ethyl Alcohol, Chemical Economics Handbook — SRI 

International, pp 50-51. 2002 

5 Wooley, R.; Ruth, M.; Sheehan, J.; Ibsen, K. Lignocellulosic Biomass to Ethanol Process 

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6Perlack, R.D.; Wright, L.L.; Turhollow, A.F.; Graham, R.L.; Stokes, B.J.; Erbach, D.C. 

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7 Aden, A.; Spath, P.; Atherton, B. The Potential of Thermochemical Ethanol Via Mixed 

Alcohols Production. NREL Milestone Report. September 30, 2005. 

http://devafdc.nrel.gov/bcfcdoc/9432.pdf 

8 Bain, R.L. The Role of Thermochemical Processing in an Integrated Sugars/Thermochemical 

Biorefinery: Conceptual Process Technoeconomic Analysis. NREL Technical Memo, February 

9, 2005. http://devafdc.nrel.gov/bcfcdoc/8967.pdf 

9 Spath, P.; Aden, A.; Eggeman, T.; Ringer, M.; Wallace, B.; Jechura, J. Biomass to Hydrogen 

Production Detailed Design and Economics Utilizing the Battelle Columbus Laboratory 

Indirectly-Heated Gasifier. NREL Report No. Number NREL/TP-510-37408. May 30, 2005. 

http://www.nrel.gov/docs/fy05osti/37408.pdf 

10 http://engineering.dartmouth.edu/rbaef/ Final report under development. 

11 Larson, E.D.; Jin, H.; Celika, F.E. “Large-Scale Gasification-Based Co-Production of Fuels 

and Electricity from Switchgrass.” Draft manuscript for Biomass & Bioenergy. March 2006. 

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12 E.D. Larson, G.W. McDonald, W. Yang, W.J. Frederick, K. Iisa, T.G. Kreutz, E.W. Malcom, 

and C.A. Brown, “A Cost-Benefit Assessment of Black Liquor Gasifier/Combined Cycle 

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13 Bain, R. L., K. R. Craig and R. P. Overend (2000). “Gasification for Heat and Power, 

Methanol, and Hydrogen,” Chapter 9.2 of Industrial Uses of Biomass Energy, ed. F. RosilloCalle, et al., Taylor and France, London, UK, ISBN-0-7484-0884-3. 

14 Wyman, C.E., R. L. Bain, N.D. Hinman, and D.J. Stevens, (1993) "Ethanol and Methanol 

from Cellulosic Materials, "Chapter 21 of Renewable Energy: Sources for Fuels and Electricity, 

ed T.B. Johansson, et al, Island Press. 

15 Chem Systems "Assessment of Costs and Benefits of Flexible and Alternative Fuel Use in the 

U.S. Transportation Sector", by Chem Systems, Tarrytown, NY, for the U.S. Department of 

Energy, Washington, D.C., Report No. DOE/PE--0093, 46 pp, November 1989. 

16 Feldmann, H.F., Paisley, M.A., Appelbaum, H.R., and D.R. Taylor, "Conversion of Forest 

Residues to a Methane-Rich Gas in a High- Throughput Gasifier", prepared by Battelle 

Columbus Division, Columbus, Ohio, for the Pacific Northwest Laboratory, Richland, WA, PNL 

Report no. PNL-6570, May 1988. 

17 Wan, E. I. and M. D. Fraser, "Economic Assessment of Advanced Biomass Gasification 

Systems", by Science Applications International Corporation, McLean, VA, IGT Biomass 

Conference, New Orleans, LA, Feb 1989. 

18 Wyman, C.E., R. L. Bain, N.D. Hinman, and D.J. Stevens, (1993) "Ethanol and Methanol 

from Cellulosic Materials, "Chapter 21 of Renewable Energy: Sources for Fuels and Electricity, 

ed T.B. Johansson, et al, Island Press. 

19 Hamelinck, C. N. and A. P. C. Faaij (2001). “Future Prospects for Production of Methanol 

and Hydrogen from Biomass,” Utrecht University, Utrecht, the Netherlands, Report No. NWS-E2001-49, ISBN-90-73958-84-9 

20 Williams, R. H., E. D. Larson, R. E. Katofsky, and J. Chen (1995). “Methanol and Hydrogen 

from Biomass for Transportation, with Comparisons to Methanol and Hydrogen from Natural 

Gas and Coal,” PU/CEES Report No. 292, Princeton University, Princeton, NJ, July. 

21 Nexant Inc. Equipment Design and Cost Estimation for Small Modular Biomass Systems, 

Synthesis Gas Cleanup, and Oxygen Separation Equipment, Task 1: Cost Estimates of Small 

Modular Systems. Subcontract Report NREL/SR-510-39943. May 2006. 

http://www.nrel.gov/docs/fy06osti/39943.pdf 

22 Nexant Inc. Equipment Design and Cost Estimation for Small Modular Biomass Systems, 

Synthesis Gas Cleanup, and Oxygen Separation Equipment ,Task 2: Gas Cleanup Design and 

Cost Estimates – Wood Feedstock. Subcontract Report NREL/SR-510-39945. May 2006. 

http://www.nrel.gov/docs/fy06osti/39945.pdf 

 54 
23 Nexant Inc. Equipment Design and Cost Estimation for Small Modular Biomass Systems, 

Synthesis Gas Cleanup, and Oxygen Separation Equipment, Task 2.3: Sulfur Primer. 

Subcontract Report NREL/SR-510-39946 May 2006. 

http://www.nrel.gov/docs/fy06osti/39946.pdf 

24 Nexant Inc. Equipment Design and Cost Estimation for Small Modular Biomass Systems, 

Synthesis Gas Cleanup, and Oxygen Separation Equipment, Task 9: Mixed Alcohols From 

Syngas — State of Technology. Subcontract Report NREL/SR-510-39947 May 2006. 

http://www.nrel.gov/docs/fy06osti/39947.pdf 

25 Garrett, D.E., Chemical Engineering Economics, Van Nostrand Reinhold, New York, 1989. 

26 Peters, M.S.; K.D. Timmerhaus. Plant Design and Economics for Chemical Engineers, 5th

Edition, McGraw-Hill, Inc., New York. 2003. 

27 Perry, R.H., Green, D.W., Maloney, J.O. Perry’s Chemical Engineers’ Handbook, 7th ed., 

McGraw-Hill. 1997 

28 October 2006 values used from http://www.che.com/inc/CEeconomicIndicators.php. 

29 Craig, K.R.; Mann, M.K. Cost and Performance Analysis of Biomass-Based Integrated 

Gasification Combined-Cycle (BIGCC) Power Systems, NREL Report No. NREL/TP-430-

21657. October 1996. http://www.nrel.gov/docs/legosti/fy97/21657.pdf 

30 Foust, T.; et al., A National Laboratory Market and Technology Assessment of the 30x30 

Scenario. Draft in progress, 2007, DOE Report.

31 Mann, M.K.; Spath, P.L. Life Cycle Assessment of a Biomass Gasification Combined-Cycle 

Power System. National Renewable Energy Laboratory, Golden, CO, TP-430-23076. 1997. 

http://www.nrel.gov/docs/legosti/fy98/23076.pdf 

32 Phillips, S.; Carpenter, D.; Dayton, D.; Feik, C.; French, R.; Ratcliff, M.; Hansen, R.; Deutch, 

S.; and Michener, B. (2004). Preliminary Report on the Performance of Full Stream Tar 

Reformer. Internal NREL Milestone report. 

33 Nexant Inc. Equipment Design and Cost Estimation for Small Modular Biomass Systems, 

Synthesis Gas Cleanup, and Oxygen Separation Equipment ,Task 2: Gas Cleanup Design and 

Cost Estimates – Wood Feedstock. Subcontract Report NREL/SR-510-39945. May 2006. 

http://www.nrel.gov/docs/fy06osti/39945.pdf

34 Gas Processors Suppliers Association. Engineering Data Book, FPS Version, 12th ed., Tulsa, 

OK. 2004 

35 A Claus unit for reducing sulfur was briefly considered as an alternative to LO-CAT, however 

the size of the stream and H2S concentration made this a fairly impractical choice. 

 55 
36 Nexant, Inc. Equipment Design and Cost Estimation for Small Modular Biomass Systems, 

Synthesis Gas Cleanup, and Oxygen Separation Equipment. Task 9: Mixed Alcohols from Syngas 

– State of Technology. NREL/SR-510-39947. Performed by Nexant Inc., San Fransisco, CA. 

Golden, CO: National Renewable Energy Laboratory, May 2006. 

http://www.nrel.gov/docs/fy06osti/39947.pdf

37 Herman, R.G. (1991). “Chapter 7 – Classical and Non-classical Routes for Alcohol 

Synthesis.” New Trends in CO Activation, L. Guczi, ed., Elsevier, New York, pp 265-349. 

38 Consult Appendix J for specific literature experiment. 

39 Personal communication between Yves Parent at NREL and CRITERION (Houston, TX) on 

Friday, April 13, 2006. 

40 Email correspondence between David Dayton, NREL, and Mark Jones, Dow Chemical. 

January 19, 2007. 

41 Forzatti, P., Tronconi, E., and Pasquon, I. (1991). “Higher Alcohol Synthesis.” Catalysis 

Reviews – Science and Engineering, 33(1-2), pp109-168. 

42 Hamelinck, C.; et al. Production of FT Transportation Fuels from Biomass; Technical 

Options, Process Analysis, and Optimisation, and Development Potential. Utrecht University, 

Copernicus Institute. March 2003. ISBN 90-393-3342-4. 

43 Quarderer, G.J. “Mixed Alcohols from Synthesis Gas.” Proceedings from the 78th Spring 

National AIChE Meeting. April 6-10, 1986, New Orleans, LA. 

44 Nirula, S. ”Dow / Union Carbide Process for Mixed Alcohols from Syngas” SRI 

International. March 1986. PEP Review No 85-1-4. Menlo Park, CA. 

45 Park, T.; et al.; “Kinetic Analysis of Mixed Alcohol Synthesis from Syngas over K/MoS2

Catalyst.” Ind. Eng. Chem. Res. Vol 36, 1997. pp 5246 – 5257. 

46 Gunturu, A.; et al. ”A Kinetic Model for the Synthesis of High-Molecular-Weight Alcohols 

over a Sulfided Co-K-Mo/C Catalyst.” Ind. Eng. Chem. Res. Vol. 37, 1998. pp. 2107 – 2115. 

47 Smith, K.; et al. ”Kinetic Modeling of Higher Alcohol Synthesis over Alkali-Promoted 

Cu/ZnO and MoS2 Catalysts.” Chemical Engineering Science, Vol. 45. 1990. pp. 2639 – 2646. 

48 Gas Processors Suppliers Association. Engineering Data Book, FPS Version, 12th ed., Tulsa, 

OK. 2004 

49 Weast, R.C., ed. CRC Handbook of Chemistry and Physics, 62nd ed., CRC Press, Boca Raton, 

FL. 1981 

 56 
50 Aden, et .al. “Lignocellulosic Biomass to Ethanol Process Design and Economics Utilizing 

Co-Current Dilute Acid Prehydrolysis and Enzymatic Hydrolysis for Corn Stover”, NREL/TP510-32438, June 2002. 

51 Keeney, D.; et al. “Water Use by Ethanol Plants Potential Challenges”. Institute for 

Agriculture and Trade Policy, Minneapolis, MN. October 2006. 

http://www.iatp.org/iatp/publications.cfm?accountID=258&refID=89449

52 Peters, M.S.; K.D. Timmerhaus. Plant Design and Economics for Chemical Engineers, 5th

Edition, McGraw-Hill, Inc., New York. 2003. 

53 Breault, R.; Morgan, D. Design and Economics of Electricity Production Form An Indirectlyheated Biomass Gasifier. Report TR4533-049-92. Columbus, Ohio: Battelle Columbus 

Laboratory. 1992. 

54 Dravo Engineering Companies. Gasification Capital Cost Estimation. Obtained from Mark 

Paisley, August, 1994. Battelle Columbus Laboratory. 1987. 

55 Weyerhaeuser, Nexant, and Stone & Webster. Biomass Gasification Combined Cycle. 

Weyerhaeuser Company, Tacoma, WA . DOE DE-FC36-96GO10173. 2000. 

56 Stone & Webster; Weyerhaeuser; Amoco; and Carolina Power & Light. New Bern Biomass to 

Energy Project Phase 1 Feasibility Study. Response to NREL Contract No. LOI No. RCA-3-

13326. NREL Report No. TP-421-7942. June 1995. 

57 Wan, E. I. and Malcolm D. F. "Economic Assessment of Advanced Biomass Gasification 

Systems," in Energy from Biomass and Wastes XIII, Donald L. Klass, ed. Chicago: Institute of 

Gas Technology, pp.791-827. 1990. 

58 Weyerhaeuser. Gasification Capital Cost Estimation. Obtained from Mark Paisley, August, 

1994. Battelle Columbus Laboratory. 1992. 

59 Wright, J.; Feinberg, D. A Comparison of the Production of Methanol and Ethanol from 

Biomass. For the International Energy Agency. Contract no. 23218-1-9201/01-SQ. 1993. 

60 Craig, K. Electric Power Generation Cost - Version 1.11 spreadsheet from Kevin Craig. July 

6, 1994. 

61 Fogler, H.S. Elements of Chemical Reaction Engineering. Second Edition. Prentice Hall. 

Englewood Cliffs, New Jersey. 1992. 

62 Kohl, A.L.; Nielsen, R.B. Gas Purification, 5th Edition. Gulf Publishing Company. 1997

63 Walas, S.M. Chemical Process Equipment Selection and Design. Butterworth-Heinemann. 

1988. 

 57 
64 Leiby, S.M. (1994) Operations for Refinery Hydrogen. SRI Report No. 212. Menlo Park, CA. 

65 Jeakel, D. (2004). Price quote from AGSCO for super sacks or bulk. 

66 Chemical Marketing Reporter. (2004). August 23-30 issues. 

67 Chem Systems. (1994). Biomass to Ethanol Process Evaluation. Prepared for NREL. 

Tarrytown, New York. 

68 Energy Information Agency (EIA). 2003. 

http://www.eia.doe.gov/pub/oil_gas/petroleum/data_publications/petroleum_marketing_monthly

/current/pdf/pmmtab16.pdf. 

69 Graubard, D. Personal correspondence. Gas Technology Products LLC. Schaumburg, IL. 

2004 

70 East Bay municipal utility district. 2004.

http://www.ebmud.com/wastewater/industrial_&_commercial_permits_&_fees/wastewater_rates

/default.htm#non-residential%20rates

71 SRI International. (2003). PEP Yearbook International. Volume 1E. United States. Menlo 

Park, CA. 

72 Baker, Thomsen Associates Insurance Services Inc. (BTA), Salary Expert ePro©, 

www.salaryexpert.com, 2004.

73 Based on historical wholesale prices found on the Energy Information Agency web site, 

http://www.eia.doe.gov/oil_gas/petroleum/info_glance/petroleum.html 

 58 
 

Appendix A 

List of Acronyms 

59
 

ASME American Society of Mechanical 

Engineers 

 MTBE Methyl-tertiary-butyl Ether 

BCL Battelle Columbus Laboratory MW Megawatts 

BFW Boiler Feed Water NREL National Renewable Energy Laboratory 

bpd Barrels per Day NRTL Non-Random Two Liquid activity 

coefficient method 

BTU British Thermal Unit OBP Office of the Biomass Program 

CFM Cubic Feet per Minute PFD Process flow diagram 

CH4 Methane PEFI Power Energy Fuels, Inc. 

CIP Clean-in-place PNNL Pacific Northwest National Laboratory 

CO Carbon Monoxide PPMV Parts per million by volume 

Co Cobalt psia Pounds per square inch (absolute) 

CO2 Carbon Dioxide RKS-BM Redlich-Kwong-Soave equation of state 

with Boston-Mathius modifications 

DCFROR Discounted Cash Flow Rate of 

Return 

 SEHT Snamprogetti, Enichem and Haldor 

Topsoe 

DOE US. Department of Energy SMR Steam Methane Reformer 

EIA Energy Information 

Administration 

 TC EtOH Thermochemical Ethanol 

EtOH Ethanol tpd Short Tons per Day 

FT Fischer-Tropsch TPI Total Project Investment 

FY Fiscal Year UCC Union Carbide Corp. 

GHSV Gas Hourly Space Velocity WGS Water Gas Shift 

GJ GigaJoule WRI Western Research Institute 

gpm Gallons per minute WWT Wastewater Treatment 

H2 Hydrogen 

HAS Higher Alcohol Synthesis 

HHV Higher Heating Value 

IFP Institut Francais du Petrole 

IRR Internal Rate of Return 

kWh Kilowatt-hour 

LHV Lower Heating Value 

MA Mixed Alcohols 

MASP Minimum alcohols selling price 

MeOH Methanol 

MESP Minimum ethanol selling price 

MoS2 Molybdenum disulfide 

60
 

Appendix B 

OBP Thermochemical Platform Research Targets 

61
 

Background/Introduction 

Thermochemical conversion technology options include both gasification and pyrolysis. 

Thermochemical conversion is envisioned to be important in enabling lignocellulosic 

biorefineries and to maximize biomass resource utilization for the production of biofuels. 

Moving forward, the role of thermochemical conversion is to provide a technology option for 

improving the economic viability of the developing bioenergy industry by converting the fraction 

of the biomass resources that are not amenable to biochemical conversion technologies into 

liquid transportation fuels. The thermochemical route to ethanol is synergistic with the 

biochemical conversion route. A thermochemical process can more easily convert lowcarbohydrate or “non-fermentable” biomass materials such as forest and wood residues to 

alcohol fuels, which adds technology robustness to efforts to achieve the 30 x 30 goal (Foust, et 

al., 2006). This Appendix describes the R&D needed to achieve the market target production 

price in 2012 for a stand-alone biomass gasification/mixed alcohol process. Future advanced 

technology scenarios rely on considerable biofuel yield enhancements achieved by combining 

biochemical and thermochemical conversion technologies into an integrated biorefinery that 

implements mixed alcohol production from gasification of lignin-rich bioconversion residues to 

maximize the liquid fuel yield per delivered ton of biomass. 

Biomass gasification can convert a heterogeneous supply of biomass feedstock into a consistent 

gaseous intermediate that can then be reliably converted to liquid fuels. The biomass gasification 

product gas (“synthesis gas” or simply “syngas”) has a low to medium energy content 

(depending on the gasifying agent) and consists mainly of CO, H2, CO2, H2O, N2, and 

hydrocarbons. Minor components of the syngas include tars, sulfur and nitrogen oxides, alkali 

metals, and particulates. These minor components of the syngas potentially threaten the 

successful application of downstream syngas conversion steps. 

Commercially available and near-commercial syngas conversion processes were evaluated on 

technological, environmental, and economic bases (Spath and Dayton, 2003). This design report 

provides the basis for identifying promising, cost-effective fuel synthesis technologies that 

maximize the impact of biomass gasification for transforming biomass resources into clean, 

affordable, and domestically produced biofuels. For the purpose of this report the precommercial mixed alcohols synthesis process implementing an alkali promoted MoS2 catalyst, a 

variant of Fischer-Tropsch synthesis, was selected as the conversion technology of choice 

because high yields of ethanol are possible with targeted R&D technology advancements. The 

MoS2 catalyst is also tolerant of low levels of sulfur gases that are common catalyst poisons. The 

proposed mixed alcohol process does not produce ethanol with 100% selectivity. Production of 

higher normal alcohols (e.g., n-propanol, n-butanol, and n-pentanol) is unavoidable. Fortunately, 

these by-product higher alcohols have value as commodity chemicals, fuel additives, or 

potentially fuels in their own right. 

The schedule for meeting specific research goals for improved tar reforming and mixed alcohol 

synthesis catalyst performance was accelerated by the President’s Advanced Energy Initiative to 

achieve cost-competitive cellulosic ethanol by 2012. This design report provides a rigorous 

engineering analysis to provide a baseline technology scenario for doing this. The conceptual 

process design and ethanol production cost estimate quantify the benefits of meeting the R&D 

62
 

goals for tar reforming and improved mixed alcohol catalyst performance helps establish 

technical R&D targets that need to be overcome by a concerted and directed core research effort. 

Process Description 

Figure 1 shows a block process flow diagram of the cost-competitive target process and the 

major technical barriers that need to be addressed to accomplish this target case. The feedstock 

interface addresses the main biomass fuel properties that impact the long-term technical and 

economic success of a thermochemical conversion process: moisture content, fixed carbon and 

volatiles content, impurity (sulfur, nitrogen, chlorine) concentrations, and ash content. High 

moisture and ash contents reduce the usable fraction of delivered biomass fuels proportionally. 

Therefore, maximum system efficiencies should be possible with dry, low ash biomass fuels. 

Feedstock

Interface

Gasification

Gas Cleanup

&

Conditioning

Fuel Synthesis

Heat

&

Power

Products

By-products

Size Reduction

Storage & Handling

De-watering

Drying

Partial Oxidation

Pressurized Oxygen

Indirect/Steam

Technical Feasibility of

Syngas Quality

Particulate removal

Catalytic Reforming

Tars

Benzene

Light Hydrocarbons

Methane

S, N, Cl mitigation

CO2 removal

H2/CO adjustment

Separations

Recycle

Selectivity

Methanol

n-Propanol

n-Butanol

n-Pentanol

Ethanol

Feedstock

Interface

Gasification

Gas Cleanup

&

Conditioning

Fuel Synthesis

Heat

&

Power

Products

By-products

Size Reduction

Storage & Handling

De-watering

Drying

Partial Oxidation

Pressurized Oxygen

Indirect/Steam

Technical Feasibility of

Syngas Quality

Particulate removal

Catalytic Reforming

Tars

Benzene

Light Hydrocarbons

Methane

S, N, Cl mitigation

CO2 removal

H2/CO adjustment

Separations

Recycle

Selectivity

Methanol

n-Propanol

n-Butanol

n-Pentanol

Ethanol

Figure 1. Process flow diagram with research barriers for cost-competitive 

thermochemical ethanol production 

Biomass gasification is a complex thermochemical process that begins with the thermal 

decomposition of a lignocellulosic fuel followed by partial oxidation of the fuel with a gasifying 

agent, usually air, oxygen, or steam to yield a raw syngas. The raw gas composition and quality 

are dependent on a wide range of factors including feedstock composition, type of gasification 

reactor, gasification agents, stoichiometry, temperature, pressure, and the presence or lack of 

catalysts. 

Gas cleanup is a general term for removing the unwanted impurities from biomass gasification 

product gas and generally involves an integrated, multi-step approach that depends on the end 

use of the product gas. This entails removing or eliminating tars, acid gas removal, ammonia 

scrubbing, alkali metal capture, and particulate removal. Gas conditioning refers to final 

modifications to the gas composition that makes it suitable for use in a fuel synthesis process. 

Typical gas conditioning steps include sulfur polishing to remove trace levels of remaining H2S 

and water-gas shift to adjust the final H2:CO ratio for optimized fuel synthesis. 

63
 

Comprehensive cleanup and conditioning of the raw biomass gasification product gas yields a 

“clean” syngas comprised of essentially CO and H2, in a given ratio that can converted to a 

mixed alcohol product. Separation of ethanol from this product yields a methanol-rich stream 

that can be recycled with unconverted syngas to improve process yield. The higher alcohol-rich 

stream yields by-product chemical alcohols. The fuel synthesis step is exothermic so heat 

recovery is essential to maximize process efficiency. 

R&D Needs To Achieve the 2012 Technical Target for Thermochemical Ethanol 

Essential R&D activities from 2007 through 2011 to overcome identified technical barrier areas 

to meet the established 2012 technical target for thermochemical ethanol production are outlined 

in Table 1. The rigorous engineering analysis of the thermochemical ethanol process conducted 

in this study will help to validate the feasibility of these technical targets and provide focus for 

the technical barriers that provide the largest economic benefit. These R&D activities include 

fundamental kinetic measurements, micro-activity catalyst testing, bench-scale thermochemical 

conversion studies, pilot-scale validation of tar reforming catalyst performance, mixed alcohol 

catalyst development, and pilot-scale demonstration of integrated biomass gasification mixed 

alcohol synthesis. Process data collected in the integrated pilot-scale testing will provide the 

basis for process optimization and cost estimates that will guide deployment of the technology. 

64
Table 1. Thermochemical Ethanol (Gasification/Mixed Alcohols) R&D Targets to 

Thermochemical Ethanol Cost Target 

R&D Area Current 2007 2008 2009 2010 

Feedstock Interface 

$30/dry ton wood chips 50% moisture dried to 12% - 2000 tpd plant 

 

Thermochemical Conversion - Gasification 

Wood chips (model) - Indirect (atm) gasification – 78% syngas efficiency: H2/CO = 1.0-1.5 

CH4≤15vol% 

Tars ≤30 g/Nm3

benzene ≤ 1vol% 

H2S = 50-600 ppm 

NH3 and HCl to be 

determined 

Biorefinery residues - Indirect (atm) gasification : corn stover; switchgrass; wheat straw; lignin - 78% syngas efficiency: H2/CO = 1.0-1.5 

CH4≤15vol% 

Tars ≤30 g/Nm3; benzene ≤

1vol% H2S = 50-600 ppm 

NH3 and HCl to be 

determined 

Demonstrate biomass gasification for $6.88/MMBtu syngas cost based on 2007 

Indirect (atm) gasification –78% syngas efficiency: H2/CO = 1.0-1.5 

CH4≤8vol% 

Tars ≤10 g/Nm3; benzene ≤

0.1vol%; H2S ≤ 20 ppm; NH3 and HCl to be 

determined 

Demonstrate biomass gasification for $5.25/MMBtu syngas cost 

Cleanup and Conditioning 

Cyclone particulate removal H2S ≥ 50 ppm (based on 

feedstock) with no S removal Tar Reformer Efficiency CH4 ≥20% 

Benzene ≥70% 

heavy tars ≥95% 

(79% CH4 conversion in 

separate SMR) 

Sorbent injection to maintain H2S levels ≤ 50 ppm for 

syngas from biomass to reduce sulfur deactivation of tar reforming catalysts. 

Tar Reformer Efficiency CH4 ≥50% 

Benzene ≥90% heavy tars ≥97% 

(79% CH4 conversion in 

separate SMR) 

Improve tar reforming catalyst performance - Regen/TOS ratio ≤ 600 

Tar Reformer Efficiency CH4 ≥80% 

Benzene ≥99% heavy tars ≥99.9% 

Eliminate SMR; highest activity re-gained by regenerating deactivated catalyst 

Catalytic Fuels Synthesis (Mixed Alcohols) 

H2/CO = 1.2 

Pressure ≤ 2000 psia Productivity = 100-400 gMA/kg(cat)/hr EtOH Selectivity ≥70% (CO2-

free) 

H2/CO ≤ 1.2 

Pressure ≤ 2000 psia Productivity ≥ 150 gMA/kg(cat)/hr EtOH Selectivity ≥70% (CO2-

free) 

Demonstrate 500 hours catalyst lifetime at 2007 performance with bottled syngas for mixed alcohol catalyst cost of ≤ $0.50/gal EtOH 

H2/CO ≤ 1.0 

Pressure ≤ 1500 psia Productivity ≥ 300 gMA/kg(cat)/hr EtOH Selectivity ≥75% (CO2-free) 

Demonstrate 500 hours catalyst lifetime at 2009 performance. with biomass syngas for mixed alcohol catalyst cost of ≤ $0.22/gal EtOH 

Integration and Modeling 

Research state-of-technology - 56 gal/dry ton EtOH $2.02/gal minimum EtOH selling price (higher alcohols sold at 85% of market value) at $2.71/gal installed capital costs. 

Biomass Gasification/Mixed 

Alcohol Design Report - 

Establishes a cost and quality baseline for technology improvements for $1.07/gal thermochemical ethanol by 2012 from indirect biomass gasification through a clean syngas intermediate. 

Improved hydrocarbon conversion efficiency yields- 56 gal/dry ton EtOH $1.73/gal minimum EtOH selling price (higher alcohols priced as gasoline on an energy adjusted basis - $1.15/gal) at $2.69/gal installed capital costs. 

Validated $1.73/gal EtOH for integrated Cleanup & Conditioning + Mixed Alcohol synthesis 

Demonstrate feasibility of system (8000 hr on stream with ≤10% catalyst losses 

per year) based on regenerating fluidizable tar reforming catalyst to eliminate SMR 

65
 

Feedstock Interface 

Feedstock handling, processing, and feeding specifically related to the thermochemical 

conversion process will need to be addressed. Because the 30 x 30 scenario envisions mixed 

alcohol conversion for low-grade or “non-fermentable” feedstocks, refinements in dry biomass 

feeder systems for use with gasification will be required to meet cost targets. These refinements 

should reduce upfront feed processing requirements to yield biomass feedstocks at $35 per ton 

delivered to the thermochemical process. Additional challenges will be associated with feeding 

the delivered biomass into developing pressurized biomass gasification systems. In all cases, 

demonstrating biomass feed systems beyond the pilot-scale will be necessary but this is not a 

significant component of the proposed research portfolio. 

Fundamental Gasification Studies R&D Needs 

The thermochemical mixed alcohol synthesis conversion route is envisioned initially for forest 

thinnings and other predominately woody feedstocks and residues. Hence, gasification studies 

will need to be performed to determine how feedstock composition affects syngas composition 

and quality and syngas efficiency. The gasifier technology chosen for the basis of this analysis is 

the Battelle Columbus Laboratory indirectly heated gasifier. Other gasifier technologies are 

under development that could prove more promising. These technologies will need to be tracked 

to ascertain their applicability to the mixed alcohol synthesis process. 

Tar Cleanup and Conditioning R&D Needs 

Previous techno-economic analyses (Aden and Spath, 2005) have shown that achieving the 

research goals for cleanup and conditioning of biomass-derived syngas to remove chemical 

contaminants such as tar, ammonia, chlorine, sulfur, alkali metals, and particulates has the 

greatest impact on reducing the cost of mixed alcohol synthesis. To date, gas cleanup and 

conditioning technologies and systems are unproven in integrated biorefinery applications. The 

goal of this research is to eliminate the tar removal and disposal via water quench, which is 

problematic both from efficiency and waste disposal perspectives, and develop a consolidated tar 

and light hydrocarbon reforming case. 

The current lab-scale demonstration results and target conversions for various impurities 

measured in biomass-derived syngas are listed in Table 2 for the year 2005 “current” state of 

technology case and the year 2012 “goal” case. The goal case conversions were selected to yield 

an economically viable clean syngas that is suitable for use in a catalytic fuel synthesis process 

without further hydrocarbon conversion steps. 

66
 

Table 2. Tar Reformer Performance – % Conversion 

Compound Current

(2005)

Goal 

(2012)

Methane (CH4) 20% 80% 

Ethane (C2H6) 90% 99% 

Ethylene (C2H4) 50% 90% 

Tars (C10+) 95% 99.9% 

Benzene (C6H6) 70% 99% 

Ammonia (NH3) 70% 90% 

The research target will be met when tar and light hydrocarbons are sufficiently converted to 

additional syngas, technically validating the elimination of a downstream steam methane 

reforming unit operation to separately reform methane from the other light hydrocarbons. 

Specific research to generate the required chemical and engineering data to design and 

successfully demonstrate a regenerating tar reforming reactor for long-term, reliable gas cleanup 

and conditioning includes: 

• Performing tar deactivation/regeneration cycle tests to determine activity profiles to 

maintain the required long-term tar reforming catalyst activity 

• Performing fundamental catalyst studies to determine deactivation kinetics and 

mechanisms by probing catalyst surfaces to uncover molecular-level details 

• Determining optimized catalyst formulations and materials at the pilot scale to 

demonstrate catalyst performance and lifetimes as a function of process conditions and 

feedstock 

Although consolidated tar and light hydrocarbon reforming tests performed with Ni-based 

catalysts have demonstrated the technical feasibility of this gas cleanup and conditioning 

strategy, alternative catalyst formulations can be developed to optimize reforming catalyst 

activity and lifetime in addition to expanded functionality. Specific further improvements that 

could be realized in catalyst functionality are: 

• Further process intensification is possible by designing catalysts with higher tolerances 

for sulfur and chlorine poisons. 

• Further reductions in gas cleanup costs could be realized by lowering or eliminating the 

sulfur and chlorine removal cost prior to reforming. 

• Optimizing the water gas shift activity of reforming catalysts could reduce or eliminate 

the need for an additional downstream shift reactor. 

Mixed Alcohol Synthesis R&D Needs 

The ability to produce mixed alcohols from syngas has been known since the beginning of the 

last century; however, the commercial success of mixed alcohol synthesis has been limited by 

poor selectivity and low product yields. Single-pass yields are on the order of 10% syngas 

conversion (38.5% CO conversion) to alcohols, with methanol typically being the most abundant 

alcohol produced (Wender 1996; Herman 2000). For mixed alcohol synthesis to become an 

67
 

economical commercial process, there is a need for improved catalysts that increase the 

productivity and selectivity to higher alcohols (Fierro 1993). 

Improvements in mixed alcohol synthesis catalysts could potentially increase alcohol yields and 

selectivity of ethanol production from clean syngas and improve the overall economics of the 

process through better heat integration and control and fewer syngas recycling loops. Specific 

research targets to achieve the cost-competitive 2012 target case are: 

• Develop improved mixed alcohol catalysts that will increase the single-pass CO 

conversion from 38.5% to 50% and potentially higher and improve the CO selectivity to 

alcohols from 80% to 90%. 

• Develop improved mixed alcohol catalysts with higher activity that will require a lower 

operating pressure (1,000 psia compared with 2,000 psia) to significantly lower process 

operating costs. This combination of lower syngas pressure for alcohol synthesis and less 

unconverted syngas to recompress and recycle has the added benefit of lowering the 

energy requirement for the improved synthesis loop. 

• Alternative mixed alcohol synthesis reactors and catalysts should be explored. Greatly 

improved temperature control of the exothermic synthesis reaction has been demonstrated 

to significantly improve yields and product selectivity. Precise temperature control 

reactor designs need to be developed for the mixed alcohol synthesis reaction to improve 

the yields and the economics of the process. 

Integration/Demonstration 

As is the case for any sophisticated conversion process, combining the individual unit operations 

into a complete, integrated systematic process is a significant challenge. Individual pilot-scale 

operations to demonstrate the required performance of the unit operations as well as complete 

integrated pilot development runs will be required to demonstrate the cost-competitive 

technology. A specific challenge will be to continue to demonstrate process intensification and 

higher yields at pilot scale to reduce capital costs. 

Achieving the technical target for the accelerated path to thermochemical ethanol requires 

meeting the specific research targets as outlined above. Missing or delaying any of these targets 

forfeits the 2012 target and jeopardizes the deployment of technologies in time to meet the 30x30 

goal. The cost implications of missing, hitting or exceeding a target or set of targets are easily 

determined with process uncertainty analysis that will be performed and detailed in the 

upcoming Mixed Alcohol Design Report due in January 2007. Combinations of sensitivity 

analysis can provide several ways to achieve the same cost-competitive target, which reduces the 

overall risk of the process. Quantifying the relative cost savings for process improvements allows 

work to be directed to the most cost effective R&D to achieve the 2012 technical target for 

thermochemical ethanol production. 

68
 

References 

Aden, A., Spath,. P.L. (2005) Milestone Completion Report “The Potential of Thermochemical 

Ethanol Via Mixed Alcohols Production.” September 2005. 

http://devafdc.nrel.gov/bcfcdoc/9432.doc

Bizzari, S.N.; Gubler, R.; Kishi, A. (2002). “Oxo Chemicals.” Chemical Economics Handbook, 

SRI International, Menlo Park, CA. Report number 682.7000 

Foust, T.D.; Wooley, R.; Sheehan, J.; Wallace, R.; Ibsen, K.; Dayton, D.; Himmel, M.; 

Ashworth, J.; McCormick, R.; Melendez, M.; Hess, J.R.; Kenney, K.; Wang, M.; Snyder, S.; 

Werpy, T. A National Laboratory Market and Technology Assessment of the 30X30 Scenario. 

Draft NREL Report. December 2006. 

Fierro, J. L. G. (1993). "Catalysis in C1 chemistry: future and prospect." Catalysis Letters 22(1-

2): 67-91. 

Herman, R. G. (2000). "Advances in catalytic synthesis and utilization of higher alcohols." 

Catalysis Today 55(3): 233-245. 

Wender, I. (1996). "Reactions of synthesis gas." Fuel Processing Technology 48(3): 189-297. 

Spath, P.L. and Dayton, D. C. “Preliminary Screening -- Technical and Economic Assessment of 

Synthesis Gas to Fuels and Chemicals with Emphasis on the Potential for Biomass-Derived 

Syngas.” 160 pp.; NREL Report No. TP-510-34929. 2003 

69
 

Appendix C 

NREL Biorefinery Design Database Description and Summary 

70
 

NREL’s Process Engineering Team has developed a database of primary information on all of the 

equipment in the benchmark model. This database contains information about the cost, reference year, 

scaling factor, scaling characteristic, design information and back-up cost referencing. The information is 

stored in a secure database and can be directly linked to the economic portion of the model. In addition to 

having all of the cost information used by the model, it has the ability to store documents pertaining to the 

piece of equipment. These include sizing and costing calculations and vendor information when available. 

The following summarizes the important fields of information contained in the database. A partial listing 

of the information is attached for each piece of equipment. Additional information from the database is 

contained in the equipment cost listing in Appendix D. 

Equipment Number:AB 

Equipment Name:AB 

Associated PFD: 

Equipment Category:A 

Equipment Type:A

Equipment Description:A

Number Required:B

Number Spares:B

Scaling Stream:B

Base Cost:B

Cost Basis:A

Cost Year:B

Base for Scaling:B 

Base Type: 

Base Units: 

Installation Factor:B 

Installation Factor Basis: 

Scale Factor Exponent:B

Scale Factor Basis: 

Material of Construction:A

Notes: 

Document: 

Design Date: 

Modified Date: 

Unique identifier, the first letter indicates the equipment type and the first 

number represents the process area, e.g., P-301 is a pump in Area 300 

Descriptive name of the piece of equipment 

PFD number on which the piece of equipment appears, e.g., PFD-P800-A101 

Code indicating the general type of equipment, e.g., PUMP 

Code indicating the specific type of equipment, e.g., CENTRIFUGAL for a 

pump 

Short description of the size or characteristics of the piece of equipment, e.g., 

20 gpm, 82 ft head for a pump 

Number of duplicate pieces of equipment needed 

Number of on-line spares 

Stream number or other characteristic variable from the ASPEN model by 

which the equipment cost will be scaled 

Equipment cost 

Source of the equipment cost, e.g., ICARUS or VENDOR 

Year for which the cost estimate is based 

Value of the scaling stream or variable used to obtain the base cost of the 

equipment 

Type of variable used for scaling, e.g., FLOW, DUTY, etc. 

Units of the scaling stream or variable, e.g., KG/HR, CAL/S 

Value of the installation factor. Installed Cost = Base Cost x Installation 

Factor 

Source of the installation factor value, e.g., ICARUS, VENDOR 

Value of the exponential scaling equation 

Source of the scaling exponent value, e.g., GARRETT, VENDOR 

Material of Construction 

Any other important information about the design or cost 

Complete, multi-page document containing design calculations, vendor 

literature and quotations and any other important information. This is stored 

as an electronic document and can be pages from a spreadsheet other 

electronic sources or scanned information from vendors. 

Original date for the design of this piece of equipment 

The system automatically marks the date in this field whenever any field is 

changed 

A These fields are listed for all pieces of equipment in this Appendix. 

B These fields are part of the equipment cost listing in Appendix D. 

71
EQUIPMENT_NUEQUIPMENT_NAME EQUIPMENT_CATEGOEQUIPMENT_TYPE EQUIPMENT_DESCRIPTION 

C-101 Hopper Feeder CONVEYOR

VIBRATINGFEEDER Included in overall cost for feed handling & drying taken from several literatur

C-102 Screener Feeder Conveyor CONVEYOR BELT Included in overall cost for feed handling & drying taken from several literaturC-103 Radial Stacker Conveyor CONVEYOR BELT Included in overall cost for feed handling & drying taken from several literaturC-104 Dryer Feed Screw Conveyor CONVEYOR SCREW Included in overall cost for feed handling & drying taken from several literaturC-105 Gasifier Feed Screw Conveyor CONVEYOR SCREW Included in overall cost for feed handling & drying taken from several literaturH-286B Flue Gas Cooler/Steam Generator #1 HEATX SHELL-TUBE duty = 155 MMBtu/hr; LMTD = 733 F; U = 150 Btu/hr-ft^2-F; area = 1,410 ft^2H-286C Flue Gas Cooler /Boiler Water Preheater #1 HEATX SHELL-TUBE duty = 20 MMBtu/hr; LMTD = 244 F; U = 100 Btu/hr-ft^2-F; area = 823 ft^2; fiH-311B Flue Gas Cooler / Steam Generator #3 HEATX SHELL-TUBE duty = 47.9 MMBtu/hr; LMTD = 457; area = 698 sq ft; U = 150 Btu/hr-ft^2-F; fK-101 Flue Gas Blower FAN CENTRIFUGAL Included in overall cost for feed handling & drying taken from several literaturM-101 Hydraulic Truck Dump with Scale SCALE TRUCK-SCALE Included in overall cost for feed handling & drying taken from several literaturM-102 Hammermill SIZE-REDUCTION Included in overall cost for feed handling & drying taken from several literaturM-103 Front End Loaders VEHICLE LOADER Included in overall cost for feed handling & drying taken from several literaturM-104 Rotary Biomass Dryer DRYER ROTARY-DRUM Included in overall cost for feed handling & drying taken from several literaturS-101 Magnetic Head Pulley SEPARATOR MAGNET Included in overall cost for feed handling & drying taken from several literaturS-102 Screener SEPARATOR SCREEN Included in overall cost for feed handling & drying taken from several literaturS-103 Dryer Air Cyclone SEPARATOR GAS CYCLONE Included in overall cost for feed handling & drying taken from several literaturS-104 Dryer Air Baghouse Filter SEPARATOR FABRIC-FILTER Included in overall cost for feed handling & drying taken from several literaturT-101 Dump Hopper TANK LIVE-BTM-BIN Included in overall cost for feed handling & drying taken from several literaturT-102 Hammermill Surge Bin TANK LIVE-BTM-BIN Included in overall cost for feed handling & drying taken from several literaturT-103 Dryer Feed Bin TANK LIVE-BTM-BIN Included in overall cost for feed handling & drying taken from several literaturT-104 Dried Biomass Hopper TANK VERTICAL-VESSEL Included in overall cost for feed handling & drying taken from several literaturC-201 Sand/ash Conditioner/Conveyor CONVEYOR SCREW Included in overall cost for gasification & gas clean up taken from several liteK-202 Combustion Air Blower FAN CENTRIFUGAL Included in overall cost for gasification & gas clean up taken from several liteM-201 Sand/ash Cooler MISCELLANEOUS MISCELLANEOUS Included in overall cost for gasification & gas clean up taken from several liteR-201 Indirectly-heated Biomass Gasifier REACTOR VERTICAL-VESSEL Included in overall cost for gasification & gas clean up taken from several liteR-202 Char Combustor REACTOR VERTICAL-VESSEL Included in overall cost for gasification & gas clean up taken from several liteS-201 Primary Gasifier Cyclone SEPARATOR GAS CYCLONE Included in overall cost for gasification & gas clean up taken from several liteS-202 Secondary Gasifier Cyclone SEPARATOR GAS CYCLONE Included in overall cost for gasification & gas clean up taken from several liteS-203 Primary Combustor Cyclone SEPARATOR GAS CYCLONE Included in overall cost for gasification & gas clean up taken from several liteS-204 Secondary Combustor Cyclone SEPARATOR GAS CYCLONE Included in overall cost for gasification & gas clean up taken from several liteS-205 Electrostatic Precipitator SEPARATOR MISCELLANEOUS Included in overall cost for gasification & gas clean up taken from several liteT-201 Sand/ash Bin TANK

FLAT-BTMSTORAGE Included in overall cost for gasification & gas clean up taken from several lite

PFD-P800-A101-2PFD-P800-A201

72
EQUIPMENT_NUEQUIPMENT_NAME EQUIPMENT_CATEGOEQUIPMENT_TYPE EQUIPMENT_DESCRIPTION 

H-301 Quench Water Recirculation Cooler HEATX SHELL-TUBE Included in overall cost for gasification & gas clean up taken from several liteH-301A Post-tar Reformer Cooler / Steam Generator #2 HEATX SHELL-TUBE duty = 47.9 MMBtu/hr; LMTD = 457; area = 698 sq ft; U = 150 Btu/hr-ft^2-F; fH-301B Reformer Flue Gas Cooler/Steam superheater HEATX SHELL-TUBE duty = 94 MMBtu/hr; LMTD = 217 F; U = 150 Btu/hr-ft^2-F; area = 2,900 ft^2; H-301C

Reformed Syngas cooler / Synthesis Reactor Preheat

#1 HEATX SHELL-TUBE duty =40.0 MMBtu/hr; LMTD = ?? F; U = 90 Btu/hr-ft^2-F; surface area = 1,55

H-302 Syngas Compressor Intercoolers HEATX

AIR-COOLED 

EXCHANGER Cost of intercoolers included in cost for syngas compressor, K-301 

H-303 Water-cooled Aftercooler HEATX SHELL-TUBE duty = 2.9 MMBtu/hr; LMTD = 25F; U = 150 Btu/hr-ft^2-F; surface area = 794 H-304 LO-CAT Preheater HEATX SHELL-TUBE duty = 0.8 MMBtu/hr;LMTD = 87 F; U = 90 Btu/hr-ft^2-F; surface area = 98 ft^H-305 LO-CAT Absorbent Solution Cooler HEATX SHELL-TUBE Included in LO-CAT system cost H-315D1 Recycle Syngas Cooler / Steam Generator #4 HEATX SHELL-TUBE duty = 1.37 MMBtu/hr; LMTD = 1,220 F; U = 150 Btu/hr-ft^2-F; area = 7 ft^2; H-315D2 Recycle Syngas cooler #2 / Air preheat #1 HEATX SHELL-TUBE duty = 0.8 MMBtu/hr;LMTD = 87 F; U = 90 Btu/hr-ft^2-F; surface area = 98 ft^K-301 Syngas Compressor COMPRESSOR CENTRIFUGAL

gas flow rate = 70,000 CFM; 6 impellers; design outlet pressure = 465 psi; 30aftercooler, & K.O.s included 

K-302 LO-CAT Feed Air Blower FAN CENTRIFUGAL Included in LO-CAT system cost K-305 Regenerator Combustion Air Blower FAN CENTRIFUGAL gas flow rate (actual) = 70133 CFM; M-301 Syngas Quench Chamber MISCELLANEOUS Included in overall cost for gasification & gas clean up taken from several liteM-302 Syngas Venturi Scrubber MISCELLANEOUS Included in overall cost for gasification & gas clean up taken from several liteM-303 LO-CAT Venturi Precontactor MISCELLANEOUS Included in LO-CAT system cost M-304 LO-CAT Liquid-filled Absorber COLUMN ABSORBER Included in LO-CAT system cost P-301 Sludge Pump PUMP CENTRIFUGAL 1.4 GPM; 0.053 brake HP; design pressure = 60 psia P-302 Quench Water Recirculation Pump PUMP CENTRIFUGAL Included in the cost of the gasification & gas clean up system P-303 LO-CAT Absorbent Solution Circulating Pump PUMP CENTRIFUGAL Included in LO-CAT system cost R-301A Tar Reformer Catalyst Regenerator REACTOR VERTICAL-VESSEL Taken from literature source R-303 Tar Reformer REACTOR VERTICAL-VESSEL Included in overall cost for gasification & gas clean up taken from several liteR-304 LO-CAT Oxidizer Vessel REACTOR VERTICAL-VESSEL Included in LO-CAT system cost S-301 Pre-compressor Knock-out SEPARATOR KNOCK-OUT DRUM 18 ft diameter; 36 ft height; design pres = 40 psia; design temp = 197 F S-302 Syngas Compressor Interstage Knock-outs SEPARATOR KNOCK-OUT DRUM Cost of intercoolers K.O.s included in cost for syngas compressor, K-301 S-303 Post-compressor Knock-out SEPARATOR KNOCK-OUT DRUM 7 ft. diameter; 14 ft height; design pres = 506 psia; design temp = 160 F S-306 Tar Reformer Cyclone SEPARATOR GAS CYCLONE Included in the cost of the tar reformer catalyst renegerator, R-204 S-307 Catalyst Regenerator Cyclone SEPARATOR GAS CYCLONE Included in the cost of the tar reformer catalyst renegerator, R-204 S-310 L.P. Amine System COLUMN ABSORBER T-301 Sludge Settling Tank SEPARATOR CLARIFIER 3 ft diameter; 7 ft height; 431 gal volume; T-302 Quench Water Recirculation Tank TANK

HORIZONTALVESSEL Included in overall cost for gasification & gas clean up taken from several lite

PFD-P800-A301-5

73
EQUIPMENT_NUEQUIPMENT_NAME EQUIPMENT_CATEGOEQUIPMENT_TYPE EQUIPMENT_DESCRIPTION 

H-410B Flue Gas Cool / syngas rxn preheat HEATX SHELL-TUBE duty = 0.8 MMBtu/hr;LMTD = 87 F; U = 90 Btu/hr-ft^2-F; surface area = 98 ft^H-411A Air preheat #3 / post Reactor Syngas cooling #1 HEATX SHELL-TUBE duty = 0.8 MMBtu/hr;LMTD = 87 F; U = 90 Btu/hr-ft^2-F; surface area = 98 ft^H-411B

Post synthesis cooler #2/Deaerator Water Preheater

#1 HEATX SHELL-TUBE duty = 20 MMBtu/hr; LMTD = 244 F; U = 100 Btu/hr-ft^2-F; area = 823 ft^2; fi

H-411C Post Synthesis cooler #3/Makeup Water heater HEATX SHELL-TUBE duty = 20 MMBtu/hr; LMTD = 244 F; U = 100 Btu/hr-ft^2-F; area = 823 ft^2; fiH-411D Post Synthesis cooler #4 / syngas recycle heat #1 HEATX SHELL-TUBE duty = 0.8 MMBtu/hr;LMTD = 87 F; U = 90 Btu/hr-ft^2-F; surface area = 98 ft^H-411E Post Synthesis Cooler #5/Mol Sieve preheater #2 HEATX SHELL-TUBE duty = 20 MMBtu/hr; LMTD = 244 F; U = 100 Btu/hr-ft^2-F; area = 823 ft^2; fiH-412 Post Mixed Alcohol Cooler HEATX SHELL-TUBE duty = 76.5 MMBtu/hr; LMTD = ?? F; U = 100 Btu/hr-ft^2-F; surface area = 4,H-413 Mixed Alcohol first Condenser (air cooled) HEATX

AIR-COOLED EXCHANGER

H-414 Mixed Alcohol Condenser HEATX SHELL-TUBE duty = 78 MMBtu/hr; LMTD = 41 F; U = 150 Btu/hr-ft^2-F; surface area = 12,4

H-416B Recycle Syngas Heat #2 / Flue gas Cool HEATX SHELL-TUBE duty =40.0 MMBtu/hr; LMTD = ?? F; U = 90 Btu/hr-ft^2-F; surface area = 1,55K-410 Mixed Alcohol Gas Compressor COMPRESSOR CENTRIFUGAL

gas flow rate = 2,481 CFM; 4 impellers; design outlet pressure = 700 psi; 10,aftercooler, & K.O.s included 

K-412 Purge Gas Expander COMPRESSOR CENTRIFUGAL gas flow rate = 144 CFM; design outlet pressure =25 psi; 2740 HP R-410 Mixed Alcohol Reactor REACTOR VERTICAL-VESSEL Fixed Bed Synthesis Reactor with MoS2-based catalyst. Sized from hourly sS-501 Mixed Alcohols Condensation Knock-out SEPARATOR KNOCK-OUT DRUM H/D = 2; 5 ft diam; 9 ft height; operating pressure = 1993 psia; operating temD-504 Ethanol/Propanol Splitter COLUMN DISTILLATIOND-505 Methanol/Ethanol Splitter COLUMN DISTILLATIONH-503A Syngas Cooler #4 / Mol Sieve preheater #1 HEATX SHELL-TUBE duty = 20 MMBtu/hr; LMTD = 244 F; U = 100 Btu/hr-ft^2-F; area = 823 ft^2; fiH-503B Mol Sieve Superheater / reformed syngas cool #5 HEATX SHELL-TUBE duty = 0.8 MMBtu/hr;LMTD = 87 F; U = 90 Btu/hr-ft^2-F; surface area = 98 ft^H-504C D-504 condenser (air cooled) HEATX

AIR-COOLED

EXCHANGER

H-504R Ethanol/Propanol Splitter Reboiler HEATX SHELL-TUBEH-505C D-505 condenser (air cooled) HEATX

AIR-COOLED

EXCHANGER

H-505R Methanol/Ethanol Splitter Reboiler HEATX SHELL-TUBEH-513 Mol Sieve Flush Condenser (air cooled) HEATX

AIR-COOLED EXCHANGER

H-590 MA Product Cooler / Mol Sieve preheater #3 HEATX SHELL-TUBE duty = 3 MMBtu/hr; LMTD = 236 F; U = 600 Btu/hr-ft^2-F; area = 20 ft^2; pre-H-591 Higher Alcohol Product Finishing cooler HEATX SHELL-TUBE Included in overall cost for gasification & gas clean up taken from several liteH-592 ETHANOL Product Cooler / Mol Sieve preheater #4 HEATX SHELL-TUBE duty = 3 MMBtu/hr; LMTD = 236 F; U = 600 Btu/hr-ft^2-F; area = 20 ft^2; pre-H-593 ETHANOL Product Finishing cooler HEATX SHELL-TUBE Included in overall cost for gasification & gas clean up taken from several liteP-504B Ethanol/Propanol Splitter Bottoms Pump PUMP CENTRIFUGALP-504R Ethanol/Propanol Splitter Reflux Pump PUMP CENTRIFUGALP-505B Methanol/Ethanol Splitter Bottoms Pump PUMP CENTRIFUGALP-505R Methanol/Ethanol Splitter Reflux Pump PUMP CENTRIFUGALS-502 Methanol Separation Column COLUMN DISTILLATION 13.5' dia, 32 Actual Trays, Nutter V-Grid Trays S-503 Molecular Sieve (9 pieces) MISCELLANEOUS PACKAGE Superheater, twin mole sieve columns, product cooler, condenser, pumps, vaT-504 Ethanol/Propanol Splitter Reflux Drum TANK KNOCK-OUT DRUMT-505 Methanol/Ethanol Splitter Reflux Drum TANK KNOCK-OUT DRUMT-590 Mixed Alcohol Product Storage Tank TANK

FLAT-BTMSTORAGE

T-592 Ethanol Product Storage Tank TANK

FLAT-BTMSTORAGE

PFD-P800-A501-2PFD-P800-A401-2

74
EQUIPMENT_NUEQUIPMENT_NAME EQUIPMENT_CATEGOEQUIPMENT_TYPE EQUIPMENT_DESCRIPTION 

H-601 Steam Turbine Condenser HEATX SHELL-TUBE

Included in the cost of the steam trubine/generator (M-602); condenser steam lb/hr

H-602 Blowdown Cooler / Deaerator Water Preheater HEATX SHELL-TUBE duty = 3 MMBtu/hr; LMTD = 236 F; U = 600 Btu/hr-ft^2-F; area = 20 ft^2; pre-H-603 Blowdown Water-cooled Cooler HEATX SHELL-TUBE duty = 0.6 MMBtu/hr; LMTD = 47 F; U = 225 Btu/hr-ft^2-F; area = 60 ft^2; fixeM-601 Hot Process Water Softener System MISCELLANEOUS PACKAGE scaled cost to 700 gpm flow, 24" dia softener. Includes filters, chemical feedM-602 Extraction Steam Turbine/Generator GENERATOR STEAM-TURBINE 25.6 MW generated; 34,308 HP M-603 Startup Boiler MISCELLANEOUS PACKAGE

Assume need steam requirement equal to 1/2 of steam requirement for gasifrate = 36,560 lb/hr 

P-601 Collection Pump PUMP CENTRIFUGAL 513 GPM; 4 brake HP; outlet pressure = 25 psia P-602 Condensate Pump PUMP CENTRIFUGAL 190 GPM; 4 brake HP; outlet pressure = 25 psia P-603 Deaerator Feed Pump PUMP CENTRIFUGAL 702 GPM; 14 brake HP; outlet pressure = 40 psia P-604 Boiler Feed Water Pump PUMP CENTRIFUGAL 730 GPM; 759 brake HP; outlet pressure = 1,345 psia S-601 Blowdown Flash Drum TANK

HORIZONTALVESSEL H/D = 2; residence time = 5 min; 2 ft diameter; 4 ft height; op press = 1,280 p

T-601 Condensate Collection Tank TANK

HORIZONTALVESSEL residence time = 10 minutes; H/D = 2; 8 ft diameter; 17 ft height 

T-602 Condensate Surge Drum TANK

HORIZONTALVESSEL residence time = 10 minutes; H/D = 2; 9 ft diameter; 17 ft height 

T-603 Deaerator TANK

HORIZONTALVESSEL liquid flow rate = 348,266 lb/hr; 150 psig design pressure; 10 min residence t

T-604 Steam Drum TANK

HORIZONTALVESSEL 424 gal, 4.5' x 4'dia, 15 psig 

K-701 Plant Air Compressor COMPRESSOR RECIPROCATING 450 cfm, 125 psig outlet M-701 Cooling Tower System COOLING-TOWER INDUCED-DRAFT approx 16,500 gpm, 140 MMBtu/hr M-702 Hydraulic Truck Dump with Scale SCALE TRUCK-SCALE Hydraulic Truck Dumper with Scale M-703 Flue Gas Stack MISCELLANEOUS MISCELLANEOUS 42 inch diameter; 250 deg F P-701 Cooling Water Pump PUMP CENTRIFUGAL 16,188 GPM; 659 brake HP; outlet pressure 75 psi P-702 Firewater Pump PUMP CENTRIFUGAL 2,500 gpm, 50 ft head P-703 Diesel Pump PUMP CENTRIFUGAL 30 gpm, 150 ft head P-704 Ammonia Pump PUMP CENTRIFUGAL 8.5 gpm, 22 ft head P-705 Hydrazine Pump PUMP CENTRIFUGAL 5 gpm, 75 ft head S-701 Instrument Air Dryer DRYER PACKAGE 400 SCFM Air Dryer, -40 F Dewpoint T-701 Plant Air Receiver TANK

HORIZONTALVESSEL 900 gal., 200 psig 

T-702 Firewater Storage Tank TANK

FLAT-BTMSTORAGE 600,000 gal, 4 hr res time, 51' dia x 40' high, atmospheric 

T-703 Diesel Storage Tank TANK

FLAT-BTMSTORAGE 10,667 gal, 120 hr res time, 90% wv, 10' dia x 18.2' high, atmospheric 

T-704 Ammonia Storage Tank TANK

HORIZONTALSTORAGE Included in the cost of the feed handling step. 

T-705 Olivine Lock Hopper TANK VERTICAL-VESSEL Included in the cost of the feed handling step. T-706 MgO Lock Hopper TANK VERTICAL-VESSEL 20' x 20' Bin, Tapering to 3' x 3' at Bottom. Capacity 6,345 cf, two truck loadsT-707 Hydrazine Storage Tank TANK VERTICAL-VESSEL 260 gal, 4.9' x 3'dia., 10psig PFD-P800-A601-3PFD-P800-A701-2

75
 

Appendix D 

Individual Equipment Cost Summary 

76
Equipment 

Number

Number 

Required

Number 

Spares Equipment Name Size Ratio

Original Equip Cost (per 

unit) Base Year

Total Original Equip Cost 

(Req'd & Spare) in Base 

Year

Scaling 

Exponent

Scaled Cos

C-101 4 Hopper Feeder 1.00 $0 2002 $0 0.75 C-102 2 Screener Feeder Conveyor 1.00 $0 2002 $0 0.75 C-103 2 Radial Stacker Conveyor 1.00 $0 2002 $0 0.75 C-104 2 Dryer Feed Screw Conveyor 1.00 $0 2002 $0 0.75 C-105 2 Gasifier Feed Screw Conveyor 0.93 $0 2002 $0 0.75 H-286B 1 Flue Gas Cooler/Steam Generator #1 0.60 $347,989 2002 $347,989 0.6 

H-286C 1 Flue Gas Cooler /Boiler Water Preheater #1 0.03 $20,989 2002 $20,989 0.6 H-311B 1 Flue Gas Cooler / Steam Generator #3 0.69 $69,089 2002 $69,089 0.65 

K-101 2 Flue Gas Blower 1.56 $0 2002 $0 0.75 M-101 4 Hydraulic Truck Dump with Scale 1.00 $0 2002 $0 0.75 M-102 2 Hammermill 1.00 $0 2002 $0 0.75 M-103 3 Front End Loaders 1.00 $0 2002 $0 0.75 M-104 2 Rotary Biomass Dryer 1.00 $3,813,728 2002 $7,627,455 0.75 $

S-101 2 Magnetic Head Pulley 1.00 $0 2002 $0 0.75 S-102 2 Screener 1.00 $0 2002 $0 0.75 S-103 2 Dryer Air Cyclone 1.56 $0 2002 $0 0.75 S-104 2 Dryer Air Baghouse Filter 0.93 $0 2002 $0 0.75 T-101 4 Dump Hopper 1.00 $0 2002 $0 0.75 T-102 1 Hammermill Surge Bin 1.00 $0 2002 $0 0.75 T-103 2 Dryer Feed Bin 1.00 $0 2002 $0 0.75 T-104 2 Dried Biomass Hopper 0.93 $0 2002 $0 0.75 

A100 Subtotal $8,065,522 $7

C-201 1 Sand/ash Conditioner/Conveyor 0.33 $0 2002 $0 0.65 K-202 2 Combustion Air Blower 0.97 $0 2002 $0 0.65 M-201 2 Sand/ash Cooler 0.33 $0 2002 $0 0.65 R-201 2 Indirectly-heated Biomass Gasifier 1.00 $2,212,201 2002 $4,424,402 0.65 $4

R-202 2 Char Combustor 1.00 $0 2002 $0 0.65 S-201 2 Primary Gasifier Cyclone 1.00 $0 2002 $0 0.65 S-202 2 Secondary Gasifier Cyclone 0.97 $0 2002 $0 0.65 S-203 2 Primary Combustor Cyclone 1.00 $0 2002 $0 0.65 S-204 2 Secondary Combustor Cyclone 0.95 $0 2002 $0 0.65 S-205 2 Electrostatic Precipitator 0.96 $0 2002 $0 0.65 T-201 1 Sand/ash Bin 0.33 $0 2002 $0 0.65 A200 Subtotal $4,424,402 $4

77
H-301 1 Quench Water Recirculation Cooler 1.57 $0 2002 $0 0.44 

H-301A 1 Post-tar Reformer Cooler / Steam Generator #2 1.66 $69,089 2002 $69,089 0.65 

H-301B 1 Reformer Flue Gas Cooler/Steam superheater 1.00 $196,589 2002 $196,589 0.6 

H-301C 1 Reformed Syngas cooler / Synthesis Reactor Preheat #1 0.32 $144,006 2002 $144,006 0.44 

H-302 5 Syngas Compressor Intercoolers 1.57 $0 2002 $0 0.65 H-303 1 Water-cooled Aftercooler 1.81 $20,889 2002 $20,889 0.44 

H-304 1 LO-CAT Preheater 0.15 $4,743 2002 $4,743 0.6 

H-305 1 LO-CAT Absorbent Solution Cooler 0.29 $0 2002 $0 0.44 H-315D1 1 Recycle Syngas Cooler / Steam Generator #4 4.20 $26,143 2002 $26,143 0.6 

H-315D2 1 Recycle Syngas cooler #2 / Air preheat #1 4.08 $4,743 2002 $4,743 0.6 

K-301 1 Syngas Compressor 1.73 $3,896,834 2002 $3,896,834 0.8 $6

K-302 1 LO-CAT Feed Air Blower 0.73 $0 2002 $0 0.65 K-305 1 Regenerator Combustion Air Blower 0.94 $35,020 2002 $35,020 0.59 

M-301 1 Syngas Quench Chamber 1.57 $0 2002 $0 0.65 M-302 1 Syngas Venturi Scrubber 1.57 $0 2002 $0 0.65 M-303 1 LO-CAT Venturi Precontactor 0.73 $0 2002 $0 0.65 M-304 1 LO-CAT Liquid-filled Absorber 0.29 $0 2002 $0 0.65 P-301 1 1 Sludge Pump 0.08 $3,911 2002 $7,822 0.33 

P-302 1 1 Quench Water Recirculation Pump 0.99 $0 2002 $0 0.65 P-303 1 1 LO-CAT Absorbent Solution Circulating Pump 1.57 $0 2002 $0 0.65 R-301A 1 Tar Reformer Catalyst Regenerator 1.62 $2,429,379 2002 $2,429,379 0.65 $

R-303 1 Tar Reformer 1.57 $2,212,201 2002 $2,212,201 0.65 $2R-304 1 LO-CAT Oxidizer Vessel 0.73 $1,000,000 2002 $1,000,000 0.65 S-301 1 Pre-compressor Knock-out 1.73 $157,277 2002 $157,277 0.6 

S-302 4 Syngas Compressor Interstage Knock-outs 1.73 $0 2002 $0 0.6 S-303 1 Post-compressor Knock-out 1.86 $40,244 2002 $40,244 0.6 

S-306 1 Tar Reformer Cyclone 1.57 $0 2002 $0 0.65 

S-307 1 Catalyst Regenerator Cyclone 1.62 $0 2002 $0 0.65 S-310 1 L.P. Amine System 1.26 $3,485,685 2002 $3,485,685 0.75 $4

T-301 1 Sludge Settling Tank 0.00 $11,677 2002 $11,677 0.6 

T-302 1 Quench Water Recirculation Tank 1.57 $0 2002 $0 0.65 

A300 Subtotal $13,742,341 $18

78
H-410B 1 Flue Gas Cool / syngas rxn preheat 6.78 $4,743 2002 $4,743 0.6 

H-411A 1 Air preheat #3 / post Reactor Syngas cooling #1 2.65 $4,743 2002 $4,743 0.6 

H-411B 1 Post synthesis cooler #2/Deaerator Water Preheater #1 0.77 $20,989 2002 $20,989 0.6 

H-411C 1 Post Synthesis cooler #3/Makeup Water heater 0.33 $20,989 2002 $20,989 0.6 

H-411D 1 Post Synthesis cooler #4 / syngas recycle heat #1 5.20 $4,743 2002 $4,743 0.6 

H-411E 1 Post Synthesis Cooler #5/Mol Sieve preheater #2 0.04 $20,989 2002 $20,989 0.6 H-412 1 Post Mixed Alcohol Cooler 0.61 $90,820 2002 $90,820 0.44 

H-413 1 Mixed Alcohol first Condenser (air cooled) 0.76 $42,255 1990 $42,255 1 

H-414 1 Mixed Alcohol Condenser 0.06 $338,016 2002 $338,016 0.44 

H-416B 1 Recycle Syngas Heat #2 / Flue gas Cool 0.63 $144,006 2002 $144,006 0.44 

K-410 1 Mixed Alcohol Gas Compressor 0.89 $851,523 2002 $851,523 0.8 

K-412 1 Purge Gas Expander 12.01 $642,014 2002 $642,014 0.8 $4

R-410 1 Mixed Alcohol Reactor 0.34 $2,026,515 2002 $2,026,515 0.56 $S-501 1 Mixed Alcohols Condensation Knock-out 2.01 $55,447 2002 $55,447 0.6 

A400 Subtotal $4,267,792 $7

D-504 1 Ethanol/Propanol Splitter 0.50 $478,100 1998 $478,100 1.32 

D-505 1 Methanol/Ethanol Splitter 0.54 $478,100 1998 $478,100 1.32 

H-503A 1 Syngas Cooler #4 / Mol Sieve preheater #1 1.08 $20,989 2002 $20,989 0.6 

H-503B 1 Mol Sieve Superheater / reformed syngas cool #5 0.66 $4,743 2002 $4,743 0.6 

H-504C 1 D-504 condenser (air cooled) 0.68 $39,408 1990 $39,408 1 

H-504R 1 Ethanol/Propanol Splitter Reboiler 0.18 $158,374 1996 $158,374 0.68 

H-505C 1 D-505 condenser (air cooled) 0.64 $62,938 1990 $62,938 1 

H-505R 1 Methanol/Ethanol Splitter Reboiler 0.26 $158,374 1996 $158,374 0.68 

H-513 1 Mol Sieve Flush Condenser (air cooled) 0.19 $21,181 1990 $21,181 1 

H-590 1 MA Product Cooler / Mol Sieve preheater #3 0.34 $3,043 2002 $3,043 0.6 

H-591 2 Ethanol Product Pump 0.21 $7,501 1998 $15,002 1.79 

H-592 1 ETHANOL Product Cooler / Mol Sieve preheater #4 1.51 $3,043 2002 $3,043 0.6 H-593 3 Ethanol Product Pump 1.19 $7,502 1999 $22,506 2.79 

P-504B 1 1 Ethanol/Propanol Splitter Bottoms Pump 0.04 $42,300 1997 $84,600 0.79 

P-504R 1 1 Ethanol/Propanol Splitter Reflux Pump 0.37 $1,357 1998 $2,714 0.79 

P-505B 1 1 Methanol/Ethanol Splitter Bottoms Pump 0.08 $42,300 1997 $84,600 0.79 

P-505R 1 1 Methanol/Ethanol Splitter Reflux Pump 0.37 $1,357 1998 $2,714 0.79 

P-590 1 Mixed Alcohol Product Pump 0.21 $7,500 1997 $7,500 0.79 P-592 1 Ethanol Product Pump 1.19 $7,500 1997 $7,500 0.79 

S-502 1 LP Syngas Separator 0.49 $55,447 2002 $55,447 0.6 

S-503 1 Molecular Sieve (9 pieces) 1.43 $904,695 1998 $904,695 0.7 $

T-504 1 Ethanol/Propanol Splitter Reflux Drum 0.37 $11,900 1997 $11,900 0.93 T-505 1 Methanol/Ethanol Splitter Reflux Drum 0.37 $11,900 1997 $11,900 0.93 T-590 2 Mixed Alcohol Product Storage Tank 0.21 $165,800 1997 $331,600 0.51 

T-592 2 Ethanol Product Storage Tank 1.19 $165,800 1997 $331,600 0.51 

A500 Subtotal $3,302,572 $2

79
H-601 1 Steam Turbine Condenser 0.87 $0 2002 $0 0.71 

H-602 1 Blowdown Cooler / Deaerator Water Preheater 0.00 $3,043 2002 $3,043 0.6 H-603 1 Blowdown Water-cooled Cooler 4.90 $16,143 2002 $16,143 0.44 

M-601 1 Hot Process Water Softener System 1.00 $1,031,023 1999 $1,031,023 0.82 $

M-602 1 Extraction Steam Turbine/Generator 1.00 $4,045,870 2002 $4,045,870 0.71 $4M-603 1 Startup Boiler 1.00 $198,351 2002 $198,351 0.6 

P-601 1 1 Collection Pump 0.13 $7,015 2002 $14,030 0.33 

P-602 1 1 Condensate Pump 0.87 $5,437 2002 $10,874 0.33 

P-603 1 1 Deaerator Feed Pump 1.00 $8,679 2002 $17,358 0.33 

P-604 1 1 Boiler Feed Water Pump 1.00 $95,660 2002 $191,320 0.33 

S-601 1 Blowdown Flash Drum 1.00 $14,977 2002 $14,977 0.6 

T-601 1 Condensate Collection Tank 1.00 $24,493 2002 $24,493 0.6 

T-602 1 Condensate Surge Drum 1.00 $28,572 2002 $28,572 0.6 

T-603 1 Deaerator 1.00 $130,721 2002 $130,721 0.72 

T-604 1 Steam Drum 1.00 $9,200 1997 $9,200 0.72 

A600 Subtotal $5,735,975 $5

K-701 2 1 Plant Air Compressor 1.00 $32,376 2002 $97,129 0.34 

M-701 1 Cooling Tower System 0.66 $267,316 2002 $267,316 0.78 

M-702 1 Hydraulic Truck Dump with Scale 1.00 $80,000 1998 $80,000 0.6 

M-703 1 Flue Gas Stack 0.31 $51,581 2002 $51,581 1 

P-701 1 1 Cooling Water Pump 0.66 $158,540 2002 $317,080 0.33 

P-702 1 1 Firewater Pump 1.00 $18,400 1997 $36,800 0.79 

P-703 1 1 Diesel Pump 1.00 $6,100 1997 $12,200 0.79 

P-704 1 1 Ammonia Pump 1.00 $5,000 1997 $10,000 0.79 

P-705 1 Hydrazine Pump 1.00 $5,500 1997 $5,500 0.79 

S-701 1 1 Instrument Air Dryer 1.00 $8,349 2002 $16,698 0.6 

T-701 1 Plant Air Receiver 1.00 $7,003 2002 $7,003 0.72 

T-702 1 Firewater Storage Tank 1.00 $166,100 1997 $166,100 0.51 

T-703 1 Diesel Storage Tank 1.00 $14,400 1997 $14,400 0.51 

T-704 1 Ammonia Storage Tank 1.00 $287,300 1997 $287,300 0.72 

T-705 1 Olivine Lock Hopper 1.00 $0 1998 $0 0.71 T-706 1 MgO Lock Hopper 1.00 $0 1998 $0 0.71 T-707 1 Hydrazine Storage Tank 1.00 $12,400 1997 $12,400 0.93 

A700 Subtotal $1,381,507 $1

Equipment Cost $40,920,110 $46

80
 

Appendix F 

Discounted Cash Flow Rate of Return Summary 

81
DCFROR Worksheet

Year -2 -1 0 1 2 3 4 5 Fixed Capital Investment $18,603,984 $114,503,967 $61,068,782Working Capital $9,541,997Loan Payment $0 $0 $0 $0 $0  Loan Interest Payment $0 $0 $0 $0 $0 $0 $0 $0  Loan Principal $0 $0 $0 $0 $0 $0 $0 $0  Ethanol Sales $46,938,097 $62,584,129 $62,584,129 $62,584,129 $62,584,129  By-Product Credit $9,633,280 $12,844,373 $12,844,373 $12,844,373 $12,844,373 Total Annual Sales $56,571,377 $75,428,503 $75,428,503 $75,428,503 $75,428,503 Annual Manufacturing Cost Raw Materials $23,647,650 $27,025,885 $27,025,885 $27,025,885 $27,025,885  Tar reforming catalysts $808,613 Steam reforming catalysts $0 

 ZnO $0 

 Mixed Alcohol catalysts $542,966 $0 $0 $0 $0  Baghouse Bags $415,430  Other Variable Costs $1,531,320 $1,732,166 $1,732,166 $1,732,166 $1,732,166  Fixed Operating Costs $12,059,682 $12,059,682 $12,059,682 $12,059,682 $12,059,682 Total Product Cost $39,005,661 $40,817,733 $40,817,733 $40,817,733 $40,817,733 Annual Depreciation General Plant DDB $38,064,676 $27,189,054 $19,420,753 $13,871,967 $9,908,548 

 SL $19,032,338 $15,860,282 $13,594,527 $12,137,971 $11,559,972  Remaining Value $95,161,690 $67,972,636 $48,551,883 $34,679,916 $24,771,369  Actual $38,064,676 $27,189,054 $19,420,753 $13,871,967 $11,559,972  Steam Plant DDB $4,321,018 $3,996,942 $3,697,171 $3,419,883 $3,163,392 

 SL $2,880,679 $2,804,872 $2,738,645 $2,682,262 $2,636,160  Remaining Value $53,292,560 $49,295,618 $45,598,446 $42,178,563 $39,015,171 

 Actual $4,321,018 $3,996,942 $3,697,171 $3,419,883 $3,163,392 Net Revenue ($24,819,978) $3,424,773 $11,492,845 $17,318,919 $19,887,405 Losses Forward ($24,819,978) ($21,395,205) ($9,902,360) $0 Taxable Income ($24,819,978) ($21,395,205) ($9,902,360) $7,416,559 $19,887,405 Income Tax $0 $0 $0 $2,892,458 $7,756,088 Annual Cash Income $17,565,716 $34,610,769 $34,610,769 $31,718,311 $26,854,681 Discount Factor 1.21 1.1 1 0.909090909 0.826446281 0.751314801 0.683013455 0.620921323 Annual Present Value $217,657,607 $15,968,833 $28,603,942 $26,003,583 $21,664,033 $16,674,644 Total Capital Investment + Interest $22,510,821.23 $125,954,363.41 $70,610,779.49Net Present Worth $0

82
DCFROR Worksheet

YearFixed Capital InvestmentWorking CapitalLoan Payment Loan Interest Payment Loan Principal Ethanol Sales By-Product CreditTotal Annual SalesAnnual Manufacturing Cost Raw Materials Tar reforming catalysts Steam reforming catalysts ZnO Mixed Alcohol catalysts Baghouse Bags Other Variable Costs Fixed Operating CostsTotal Product CostAnnual Depreciation General Plant DDB SL Remaining Value Actual Steam Plant DDB SL Remaining Value ActualNet RevenueLosses ForwardTaxable IncomeIncome TaxAnnual Cash IncomeDiscount FactorAnnual Present ValueTotal Capital Investment + InterestNet Present Worth

10 11 12 13 14 15 16 17 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 

$62,584,129 $62,584,129 $62,584,129 $62,584,129 $62,584,129 $62,584,129 $62,584,129 $62,584,129 $$12,844,373 $12,844,373 $12,844,373 $12,844,373 $12,844,373 $12,844,373 $12,844,373 $12,844,373 $$75,428,503 $75,428,503 $75,428,503 $75,428,503 $75,428,503 $75,428,503 $75,428,503 $75,428,503 $$27,025,885 $27,025,885 $27,025,885 $27,025,885 $27,025,885 $27,025,885 $27,025,885 $27,025,885 $

$0 $0

$0 $0

$0 $542,966 $0 $0 $0 $0 $542,966 $0 

$415,430 $415,430

$1,732,166 $1,732,166 $1,732,166 $1,732,166 $1,732,166 $1,732,166 $1,732,166 $1,732,166 

$12,059,682 $12,059,682 $12,059,682 $12,059,682 $12,059,682 $12,059,682 $12,059,682 $12,059,682 $$40,817,733 $41,776,129 $40,817,733 $40,817,733 $40,817,733 $40,817,733 $41,776,129 $40,817,733 $

$2,142,208 $1,981,543 $1,832,927 $1,695,457 $1,568,298 $1,450,676 $1,341,875 $1,241,234 $2,567,873 $2,567,873 $2,567,873 $2,567,873 $2,567,873 $2,567,873 $2,567,873 $2,567,873 

$26,420,570 $24,439,027 $22,606,100 $20,910,642 $19,342,344 $17,891,668 $16,549,793 $15,308,559 $

$2,567,873 $2,567,873 $2,567,873 $2,567,873 $2,567,873 $2,567,873 $2,567,873 $2,567,873 

$32,042,896 $31,084,500 $32,042,896 $32,042,896 $32,042,896 $32,042,896 $31,084,500 $32,042,896 $

$0 $0 $0 $0 $0 $0 $0 $0 

$32,042,896 $31,084,500 $32,042,896 $32,042,896 $32,042,896 $32,042,896 $31,084,500 $32,042,896 $$12,496,729 $12,122,955 $12,496,729 $12,496,729 $12,496,729 $12,496,729 $12,122,955 $12,496,729 $$22,114,040 $21,529,419 $22,114,040 $22,114,040 $22,114,040 $22,114,040 $21,529,419 $22,114,040 $0.385543289 0.350493899 0.318630818 0.28966438 0.263331254 0.239392049 0.217629136 0.197844669 

$8,525,920 $7,545,930 $7,046,215 $6,405,650 $5,823,318 $5,293,925 $4,685,429 $4,375,145 

83
 

Minimum Ethanol Selling Price ($/gal) $1.01

EtOH Production at Operating Capacity (MM Gal / year) 61.8

EtOH Product Yield (gal / Dry US Ton Feedstock) 80.1

Mixed Alcohols Production at Operating Capacity (MM Gal / year) 72.6

Mixed Alcohols Product Yield (gal / Dry US Ton Feedstock) 94.1

Delivered Feedstock Cost $/Dry US Ton $35

Internal Rate of Return (After-Tax) 10%

Equity Percent of Total Investment 100%

Capital Costs Operating Costs (cents/gal product)

 Feed Handling & Drying $23,200,000 Feedstock 43.7

 Gasification $12,900,000 Natural Gas 0.0

 Tar Reforming & Quench $38,400,000 Catalysts 0.3

 Acid Gas & Sulfur Removal $14,500,000 Olivine 0.7

 Alcohol Synthesis - Compression $16,000,000 Other Raw Materials 1.6

 Alcohol Synthesis - Other $4,600,000 Waste Disposal 0.4

 Alcohol Separation $7,200,000 Electricity 0.0

 Steam System & Power Generation $16,800,000 Fixed Costs 19.5

 Cooling Water & Other Utilities $3,600,000 Co-product credits -20.7

Total Installed Equipment Cost $137,200,000 Capital Depreciation 15.4

Average Income Tax 11.8

Indirect Costs 53,600,000 Average Return on Investment 28.5

 (% of TPI) 28.1%

 Project Contingency 4,100,000 Operating Costs ($/yr)

Feedstock $27,000,000

Total Project Investment (TPI) $190,800,000 Natural Gas $0

Catalysts $200,000

Installed Equipment Cost per Annual Gallon $2.22 Olivine $400,000

Total Project Investment per Annual Gallon $3.09 Other Raw Matl. Costs $300,000

Waste Disposal $300,000

Loan Rate N/A Electricity $0

Term (years) N/A Fixed Costs $12,100,000

Capital Charge Factor 0.180 Co-product credits @ $1.15 per gal -$12,800,000

Capital Depreciation $9,500,000

Maximum Yields based on carbon content Average Income Tax $7,300,000

 Theoretical Ethanol Production (MM gal/yr) 158.9 Average Return on Investment $17,600,000

 Theoretical Ethanol Yield (gal/dry ton) 205.8

Current Ethanol Yield (Actual/Theoretical) 39% Total Plant Electricity Usage (KW) 7,994

 Electricity Produced Onsite (KW) 7,998

Gasifier Efficiency - HHV % 76.6 Electricity Purchased from Grid (KW) 0

Gasifier Efficiency - LHV % 76.1 Electricity Sold to Grid (KW) 4

Overall Plant Efficiency - HHV % 47.4

Overall Plant Efficiency - LHV % 45.8 Steam Plant + Turboexpander Power Generated (hp) 66,451

 Used for Main Compressors (hp) 55,168

Plant Hours per year 8406 Used for Electricity Generation (hp) 11,283

% 96.0%

Plant Electricity Use (KWh/gal product) 1.5

Gasification & Reforming Steam Use (lb/gal) 9.9

Ethanol from Mixed Alcohols Production Process Engineering Analysis

2,000 Dry Metric Tonnes Biomass per Day

All Values in 2005$

BCL Gasifier, Tar Reformer, Sulfur Removal, MoS2 Catalyst, Fuel Purification, Steam-Power Cycle

2012 Market Target Case

84
 

Appendix G 

Process Parameters & Operation Summary 

85
Energy Efficiencies Tar Reformer Alcohol Synthesis Alco

Gasifier Efficiency - HHV % 76.6 Inlet Molar Flow (MMscf/hr) 6.70 Syngas from Conditioning 279,888 RelaGasifier Efficiency - LHV % 76.1 Space Velocity (hr-1) 2,476 Recycled from initial flash tank 215 Me

Overall Plant Efficiency - HHV % 47.4 Reactor Volume (ft3) 2,705 Recycled from MolSieve Flush 5,026 EtOverall Plant Efficiency - LHV % 45.8 Inlet: Total 285,128 Pr

 Temperature (°F) 1,086 Bu

Dryer Carbon as CO (mol%) 49.0% PeInlet: Carbon as tar (ppmv) 5,758 Conditioned Syngas H2:CO Ratio 1.00 Temperature (°F) 60.0 H2:CO Ratio (mole) 0.87 Recycled Gas H2:CO Ratio 1.06 Flas Moisture Content (wt%) 50.0% TeOutlet: Reformer Conversions: At Reactor Inlet Pr Temperature (°F) 219.7 CO2 --> CO 32.5% Temperature (°F) 570 Moisture Content (wt%) 5.0% Methane --> CO 80.0% Pressure (psia) 991 RelaInlet Flue Gas (°F) 1,206 Ethane --> CO 99.0% H2:CO Molar Ratio 1.00 MeOutlet Flue Gas (°F) 235.9 Ethylene --> CO 90.0% CO2 (mol %) 5.0% Et Dew Point Flue Gas (°F) 173.9 Benzene --> CO 99.0% Methane (mol%) 1.5% Pr Difference 62.0 Tar --> CO 99.9% H2O (wt%) 0.64% Bu

 Ammonia --> CO 90.0% Pe

Gasifier Inlet Molar Flow (MMscf/hr) 6.5Temperature (°F) 1,633 Outlet: Space Velocity (hr-1) 4,000 Vap

Pressure (psia) 23.0 Temperature (°F) 1,633 Reactor Volume (ft3) 1,616 MeH2:CO Molar Ratio After Gasifier 0.60 Carbon as CO (mol%) 75.7% EtMethane (vol%) 9.0% Carbon as tar (ppmv) 43 CO Conversion - Overall 59.5% PrBenzene (vol%) 0.07% H2:CO Ratio (mole) 1.00 CO Conversion - Singlepass 59.4% Bu

Tar (wt%) 0.91% Methane (vol%) 1.2% Conversion To: PeTar (g/Nm3) 9.5 Benzene (ppmv) 2.7 CO2 21.9%Char (wt%) 12.7% Tars (ppmv) 3 Methane 3.4% CleaH2S (ppm) 413 Tars (g/Nm3) 0.01 Ethane 0.3%Residual Heat (MBtu/hr) 0 H2S (ppm) 205 Methanol 0.3% Res

 NH3 (ppm) 80 Ethanol 28.2% Re

Char Combustor Propanol 4.6% To Temperature (°F) 1,823 Quench Butanol 0.6% T

Pressure (psia) 21.4 Benzene (ppmv) 3.1 Pentanol + 0.1% TRatio Actual:Minimum air for combustion 1.20 Tars (ppmv) 4 Total 59.4%Residual Heat (MBtu/hr) 0.0 Tars (g/Nm3) 0.01 Ove

H2S (ppm) 235 Selectivity (CO2 Free) ga

Syngas Usage NH3 (ppm) 83 Alcohols 90.1% gaTo Reformer (lb/hr) 168,120 Hydrocarbons 9.9%To Char Combustor (lb/hr) 0 Acid Gas RemovalTo Fuel System (lb/hr) 64,703 Inlet: At Reactor Outlet

 CO2 (mol/hr) 2,041 Temperature (°F) 570

Fuel System CO2 (mol%) 11.4% Pressure (psia) 986Additional fuel (lb/hr) 0 H2S (mol/hr) 4 CO2 (mol%) 22.7%Raw Syngas (lb/hr) 64,703 H2S (ppmv) 235 Methane (mol%) 4.6%Unconverted Syngas (lb/hr) 10,739 Outlet: H2O (wt%) 0.47%

 CO2 (mol/hr) 846

Into Reformer (°F) 3,680 CO2 (mol%) 5.1% Total Alcohol Productivity (kg/kg/hr) 0.602Out of Reformer (°F) 1,780 Fraction CO2 removed 58.6% Total Ethanol Productivity (kg/kg/hr) 0.488

 H2S (mol/hr) 1 H2S (ppmv) 49 Fraction H2S removed 99.6%

86
 

Appendix H 

Process Flow Diagrams (PFDs) 

87
88
89
90
91
92
93
94
95
96
97
98
99
100
101
102
103
104
 

Appendix I 

Syngas and Char Correlations 

105
 

The gasifier was modeled using correlations based on data from the Battelle Columbus 

Laboratory (BCL) 9 tonne/day test facility. The data and original correlations for the gasifier can 

be found in Bain (1992). The experimental runs were performed for several different wood types 

including Red Oak, Birch, Maple, and Pine chips, sawdust, and other hard and soft wood chips. 

The original pilot plant data for these runs can be found in Feldmann, et al, (1988). The 

temperature range for the data is 1,280 to 1,857oF and the pressure range is 2.4 to 14.4 psig; the 

majority of the data are in the range of 1,500 to 1,672oF. 

The BCL test facility’s gas production data was correlated to gasifier temperature with a 

quadratic function in the form: 

2 X a bT cT =+ +

where the temperature, T, in units of °F. The coefficients a, b, and c, as well as the units for the 

correlated variable are shown in Table 3. Even thought there is a correlation for the char 

formation, it is not used; instead the amount and elemental analysis for the char is determined by 

mass differences between the produced syngas and the converted biomass. 

Table 3. Gasifier Correlation 

Variable a b c Units 

Dry Syngas 28.993 -0.043325 0.000020966 scf gas/lb maf wooda

CO 133.46 -0.1029 0.000028792 mol% dry gas 

CO2 -9.5251 0.037889 -0.000014927 mol% dry gas 

CH4 -13.82 0.044179 -0.000016167 mol% dry gas 

C2H4 -38.258 0.058435 -0.000019868 mol% dry gas 

C2H6 11.114 -0.011667 0.000003064 mol% dry gas 

H2 17.996 -0.026448 0.00001893 mol% dry gas 

C2H2 -4.3114 0.0054499 -0.000001561 mol% dry gas 

Tar 0.045494 -0.000019759 lb/lb dry wood 

The following general procedure is used for the gasifier production: 

• A gasifier temperature T is assumed. 

• The mass and molar amounts of carbon, hydrogen, oxygen, sulfur, nitrogen, and ash (as a 

pseudo-element) are determined from the biomass’s ultimate analysis. 

• The amount of syngas and its composition is determined from the gasifier correlations. 

• The amount of carbon in the syngas and tar is determined. Residual carbon is parsed in 

the char. 

• The amount of oxygen in the syngas is determined. A minimum amount of oxygen is 

required to be parsed to the char (4% of biomass oxygen). If there is a deficit of oxygen, 

then the associated water is decomposed to make sure that this amount of oxygen is 

parsed to the char; if there is excess oxygen, then it is parsed to the char without 

decomposing hydrogen. 

• A set amount of sulfur is parsed to the char (8.3%). All remaining sulfur is set as H2S in 

the syngas. 

 

a

 Scf = standard cubic feet. The standard conditons are 1 atm pressure and 60°F temperature. 

106
 

• A set amount of nitrogen is parsed to the char (6.6%). All remaining nitrogen is set as 

NH3 in the syngas. 

• The amount of hydrogen in the syngas (including tar, H2S, NH3, and decomposed water) 

is determined. All remaining hydrogen is parsed to the char. 

• All ash is parsed to the char. 

• The heat of formation of the char is estimated from the resulting ultimate analysis from 

this elemental material balance. 

• The gasifier temperature is adjusted so the there is no net heat for an adiabatic reaction. 

The syngas amount and composition will be dependent upon the biomass composition and the 

gasifier temperature. As an example, the resulting syngas composition for the woody biomass 

used in this design report can be seen in Figure 2. Note from this figure that the amount of char 

decreases with increasing temperature and that the water does not start to decompose until high 

temperatures (here at 1650°F and higher). 

0%

10%

20%

30%

40%

50%

60%

70%

80%

90%

100%

1500 1550 1600 1650 1700 1750 1800 1850 1900 1950

Temperature (°F)

Gasifier Effluent (wt%)

Char

Benzene

Tar (w/o Bnzn)

NH3

H2S

C2H6

C2H4

C2H2

CH4

CO2

CO

H2

Gasifier Temperature (1,633 °F) H2O

Figure 2. Syngas Composition for Woody Biomass Used in Design Report 

References 

Bain, R. (January 14, 1992). Material and Energy Balances for Methanol from Biomass Using 

Biomass Gasifiers. 

Feldmann, H. F.; Paisley, M. A.; Applebaus, H.R.; Taylor, D.R. (July 1988). Conversion of 

Forest Residues to A Medium-Rich Gas in a High-Throughput Gasifier. 

107
 

Appendix J 

Alcohol Synthesis Catalyst References 

108
 

A literature search was conducted to review existing mixed alcohol technology and how it has 

developed over the past 20 years. This appendix provides a list of literature reviewed, including 

authors and journal names, as well as brief descriptions of the information contained within each 

source. The literature sources examined include a mixture of experimental results and state of 

technology summarizations. Together, over 40 different sources were used. Some 

generalizations garnered from this search are summarized below. 

The term “mixed alcohols” refers to a mixture of C1 – C8 alcohols, with preference towards the 

higher alcohols (C2-C6). Mixed alcohols catalysts are typically categorized into several groups 

based on their composition and/or derivation. Common to all of these catalysts is the addition of 

alkali metals which shift the product slate towards alcohol production. Spath, et al., categorized 

the catalysts into five groupings based on the work of Herman (Herman, 1991): 

• Modified high pressure methanol catalysts. 

• Modified low pressure methanol catalysts. 

• Modified Fischer-Tropsch (FT) catalysts. 

• Alkali-doped sulfides (modified methanation). 

• Other, which includes alternate catalysts, such as Rhodium based catalysts, which are not 

specifically used for mixed alcohols but have been developed for more selective alcohols 

synthesis. 

Others (Smith, et al., 1992) (Forzatti, et al., 1991) group them simply into three categories: 

• Modified methanol catalysts. 

• Modified Fischer-Tropsch catalysts. 

• Others. 

This helps to eliminate confusion that can arise when, for example, molybdenum-based sulfide 

catalysts are promoted with cobalt or other similar FT elements, thus representing both alkalidoped sulfides and modified-FT groupings. 

Since the 1920s scientists have known how to produce mixtures of methanol and other alcohols 

by reacting syngas over certain catalysts. They observed that when methanol catalysts (Zinc or 

Copper based) were promoted with alkali (Na, K, etc.), and certain reaction conditions were met 

(temperature, pressure) a mixture of methanol and higher alcohols resulted. At the same time, 

Fischer and Tropsch observed that hydrocarbon synthesis catalysts produced linear alcohols as 

byproducts. From this they were able to develop the “Synthol” process for producing higher 

alcohols. Some development continued, but it wasn’t until the 1970s oil embargo that significant 

interest re-appeared, and researchers renewed efforts to produce higher alcohols for liquid fuels 

applications. As petroleum prices dropped research declined until the mid-to-late 1980s when 

interest was driven by environmental aspects, specifically oxygenated fuel and octane 

enhancement. 

In 1990, the Clean Air Act mandated the seasonal use of oxygenated compounds in gasoline in 

specific regions of the U.S. Soon after, methyl-tertiary-butyl ether (MTBE) became the 

oxygenate of choice because refiners could cost-effectively produce it using existing products. 

Since then only a few researchers have been active in the field of higher alcohol synthesis. Some 

research in the 90s focused on mixed alcohols as a product of coal gasification. Other work 

continued in Europe, especially by Snamprogetti. Within the past 5 years, however, a desire to 

109
 

find alternatives for petroleum based fuels, and the increasing popularity of ethanol fuels, has 

again brought this research area to life. 

 

Three particular terms require definition before the discussion proceeds. Please refer back to 

these as needed. Within literature, data can often be confusing if these terms aren’t well 

understood or defined. 

Productivity – the amount of product generated, either all alcohols or a specific alcohol, 

per unit time for a certain weight of catalyst loaded. Units are typically g-product / kg 

catalyst / hr. Typical productivities for mixed alcohol production can range from below 

150 to near 400 g/kg/hr. In comparison, methanol synthesis productivities can often be 

over 1000 g/kg/hr, or 1 kg/kg/hr. Less commonly, productivity refers to concentrations 

of liquid product per unit of time, g/L/hr. This is often referred to as “Yield”. 

Conversion – usually the amount of carbon monoxide (CO, molar basis) converted to all 

products. Typically found by: (COinitial – COfinal) / COinitial 

Sometimes researchers present conversions exclusive of CO2 produced but not often. 

Typical single-pass conversions for mixed alcohols range from 10% - 40%. 

Selectivity – selectivities are typically presented on a %-molar basis. Selectivity refers to 

the fraction of CO converted to a specific product. For example, if a reaction achieves a 

20% conversion of CO, and 75% of that CO (or 15% of the total CO) is converted to 

alcohols, then the total alcohols selectivity is said to be 75%. Selectivities are sometimes 

shown on a CO2-exclusive basis. 

Yield – see Productivity. 

The overall stoichiometric reaction for higher alcohol synthesis (HAS) can be summarized as: 

nn n CO 2 H C H OH 1 H O + → +− 2 21 2 n n+ ( )

The value of “n” typically ranges from 1 to 8. The stoichiometry suggests an optimum H2/CO 

ratio of 2, however many of these catalysts also display significant water-gas shift activity. This 

shifts the optimal ratio closer to 1.0 and also shifts the primary byproduct from water to carbon 

dioxide (CO2). The overall reaction is exothermic; therefore, maintaining constant reaction 

temperature is an important design consideration. The reactions become more exothermic for 

greater values of “n”. Secondary reactions and other side products will depend on which catalyst 

system is used. Different kinetic pathways exist for each catalyst system. 

Catalysts 

Modified Methanol Catalysts 

The term “modified” methanol catalyst refers to the addition of an alkali promoter and other 

active elements to a methanol catalyst to shift the product slate from methanol to higher 

110
 

branched primary alcohols. High temperature methanol catalysts typically contain Zinc (Zn) 

Chromium oxides (or manganese chromium oxides), while lower temperature methanol catalysts 

use Copper (Cu) as the active component. The reaction yields primary branched alcohols, 

among which 2-methyl-1-propanol (isobutanol) is a main (and thermodynamically favored) 

component. Aldehydes, esters, ketones, and ethers are also formed, along with large amounts of 

CO2. 

Typical high pressure and low pressure reaction conditions (as provided by Nexant) are listed in 

Table 4. In general there is a trade-off between maximizing CO conversion and maximizing the 

higher alcohol selectivity and yield. 

Table 4. “Typical” Modified Methanol Catalyst Conditions 

H2/CO 

ratio 

Temperature

(°F) 

Pressure 

(psia) 

CO 

conversion 

(per pass) 

Total 

Alcohol 

Yield 

(g/kg/hr) 

C2+OH 

Selectivity 

High Pressure 1 572 to 800 1,810 to 3,625 5 to 20% 203 

Low Pressure* 1.0 to 1.2 482 to 752 725 to 1,450 20 to 60% 41.9 wt% 

 * Lurgi: Octamix 

Snamprogetti (also referred to as SEHT – Snamprogetti, Enichem and Haldor Topsoe) and Lurgi 

were two of the leading technology developers of modified methanol catalysis in the 1980s and 

1990s. SEHT had a MAS (Metanolo piu Alcoli Superiori - methanol plus higher alcohols) 

process and Lurgi developed what they called OCTAMIX, each developing pilot scale plants and 

data. The latest information available to NREL shows each process technology is no longer 

available. One technology developer still involved in this area is the Standard Alcohol Company 

of America. They have a bench-scale process to produce a mixed alcohols product known as 

EnviroleneTM. Envirolene is composed of methanol through octanol, with approximately 50% of 

the product as ethanol. The process uses a modified high pressure methanol catalyst, and the 

company is currently seeking funding for a pilot plant. 

The proposed kinetic pathway for modified methanol catalysts to branched alcohols is through a 

base-catalyzed aldol condensation reaction. Carbon chain growth schemes have been developed 

that describe the product distribution relatively accurately. This is shown in Figure 3. 

111
 

Figure 3. Reaction Network of Smith, et.al., 1991 for the Methanol – Higher 

Alcohol Synthesis over Cu/ZnO-based 

Modified Fischer-Tropsch Catalysts 

Modified Fischer-Tropsch catalysts, on the other hand, are FT catalysts that are alkali-promoted. 

The two most common FT active elements are Iron (Fe) and Cobalt (Co), but Nickel (Ni) is 

considered to have FT activity also. The addition of the alkali promoter helps to shift the product 

slate from hydrocarbons to linear alcohols, although hydrocarbons remain a significant 

byproduct. Typical reaction conditions are 220 – 350°C (430 – 660°F) and 5-20 MPa (725 – 

2,900 psia). One commonly-researched catalyst system is a MoS2-based system that is alkali 

and/or Cobalt-promoted. This has the tendency to increase ethanol and other higher alcohols 

selectivity. CO2 is still a substantial byproduct due to water-gas shift (WGS) activity of the 

catalysts. Other potential byproducts include aldehydes, esters, carboxylic acids, and ketones. 

The primary technology developers for these catalysts were Dow/Union Carbide (UCC) and 

Institut Francais du Petrole (IFP). Dow and UCC jointly developed a sulfided mixed alcohol 

catalyst based on molybdenum (MoS2). Sulfided catalysts have the advantage of being sulfurtolerant (up to 100 ppm) which has the potential to reduce upstream cleanup costs. IFP, in 

112
 

conjunction with Idemitsu Kosan (Japan), developed a process based on Cu-Co and Cu-Ni 

catalyst systems. Dow built a 2-ton-per-day (TPD) demonstration plant in 1990 and IFP built a 

20 barrel-per-day (BPD) pilot plant in Japan. The latest information available to NREL indicates 

that Dow is no longer pursuing the commercial development of their mixed alcohol process and 

IFP has not continued their work since building the pilot plant, and they have no commercial 

interest in pursuing a mixed alcohols process. 

Technology developers that remain active in this area are Power Energy Fuels Inc. (PEFI), 

Western Research Institute (WRI), and Pearson Technologies. PEFI continues to develop the 

EcaleneTM technology and process, which is a modification of Dow’s Sygmal process using 

polysulfite catalyst. According to Nexant, progress has not moved beyond the bench scale and a 

planned 500 gallon/day pilot plant is no longer being pursued. However, 2-3 other pilot plants 

are under funding consideration using various biomass resources. WRI had worked with PEFI in 

the past, however currently they are not, but are conducting their own bench-scale experiments, 

particularly reactor and catalyst testing. Pearson Technologies has developed a 30-ton-per-day 

biomass gasification and alcohols conversion facility in Aberdeen, MS. A project is under 

development by the Worldwide Energy Group and the State of Hawaii to demonstrate 

gasification of sugarcane bagasse and production of ethanol using the Pearson technology on the 

island of Kauai. Sasol (South Africa) is a world leader in FT fuels and chemicals production as 

well as technology development. They currently produce a mixture of alcohols within their 

overall process. However, these are not used for fuels. According to Sasol’s website, 

oxygenates in the aqueous stream from their Sasol Advanced Synthol (SAS) process are 

separated and purified to produce alcohols, acetic acid, and ketones. 

Due to the severe process conditions of higher alcohol synthesis, it would be expected that 

catalyst life would not be significantly longb

 (Nexant, 2005). As a benchmark, it could be 

helpful to recognize that the catalyst life for the typical Fe-Co Fischer-Tropsch catalyst can be 

longer than 5 years. Information on catalyst life is not abundant because most research into HAS 

is bench scale and not commercial. The majority of information on catalyst life comes from the 

earlier commercialization attempts. Dow and UCC, for example, found their catalyst operated 

for over 8 months continuously with little to no process performance degradation. 

The proposed kinetics for modified FT catalysts follow different pathways than for modified 

methanol kinetics. Linear alcohols are formed from a classic CO insertion route for chain 

growth (C-C bond formation) with termination to alcohols and hydrocarbons. This is shown in 

Figure 4. More complex kinetic models have reaction networks that account for the simultaneous 

formation of alcohols, hydrocarbons, and esters. 

 

b

 Meaning, 1 year or less. 

113
 

Figure 4. Kinetic Analysis of HAS with Fischer-Tropsch Mechanism 

Other Catalysts 

Some research has been conducted on alternative reaction systems for mixed alcohols synthesis. 

This includes more exotic catalytic elements (Ruthenium (Ru), Rhodium (Rh), Palladium (Pd)) 

as well as synthesis under supercritical conditions. Rh-based catalysts have been primarily 

developed for selective ethanol synthesis or other oxygenates. One downfall for these catalysts 

is their low catalytic activity which results in the need for high catalyst loadings and more drastic 

reaction conditions. Coupling this with their high cost and limited availability creates limited 

commercialization potential of these processes. For all Group VIII metal catalysts, CO 

conversion to hydrocarbons will be a significant side reaction. It has been observed that the 

selectivity to oxygenates of Rh-based catalysts is highly dependent on the support, promoter, and 

metal precursor used. 

Basis for Catalyst Selection 

Because the focus of this report is thermochemical production of ethanol, a moly-sulfide-based 

(MoS2) system promoted with cobalt and alkali metal salts was chosen as the catalyst system 

because of its ability to produce linear alcohols (as opposed to branched) and its potential for 

higher ethanol selectivities. This is a form of original Dow/UCC technology. 

114
Doc # Title Date Author(s) Company/University Affiliation Journal / Conference 

9542

Synthesis of Higher Alcohols from Syngas over Cu-Co2O3/ZnO, Al2O3 

Catalyst 2005

Mahdavi, V.; M.H. Peyrovi, et.al.

Shahid Beheshti University, Tehran Iran; National Iranian Oil Com pany, Petroleum Research Institute Applied Catalysis, Elsevier

Fafrcog

Mixed Alcohols from Syngas: State of Technology Final Report 2005 Nexant Nexant Subcontract to NREL/DOE

Saa

Gridley Ethanol Dem onstration Project Utilizing Biomass Gasification Technology: Pilot Plant Gasifier and Syngas Conversion Testing 2005 TSS Consultants City of Gridley, CA

NREL Subcontract Report SR510-37581

PGthrem ais re

N/A Production of Mixed Alcohol Fuels 2004 Lucero, Andrew J. W estern Research Institute DOE workshop on BLG

UmcocotaS2

11631

Solvent Effects on Higher Alcohols Synthesis Under Supercritical Conditions: a Therm odynam ic Consideration 2004 Qin, Zhangfeng; et.al.

Institute of Coal Chemistry, Chinese Academy of Sciences

Fuel Processing Technology, Elsevier

Skwsoeefr

9543

Higher-Alcohols Biorefinery: Improvement of Catalyst for Ethanol Conversion 2004 Olson, Edwin S.; et.al. EERC, North Dakota Applied Biochem & Biotech

Rwyi(G

Preliminary Screening - Technical and Econom ic Assessm ent of Synthesis Gas to Fuels and Chemicals with Emphasis on the Potential for Biomassderived Syngas 2003 Spath, P.L.; Dayton, D.C. National Renewable energy Laboratory (NREL) Report TP-510-34929

ocoinm oeg

11635

Palladium-based Catalysts for the Synthesis of Alcohols 2003

M. Josefina Perez-Zurita; Cifarelli, M.; et.al. Universidad Central de Venezuela

Journal of Molecular Catalysis, Elsevier

Cmse

9527

Synthesis of Higher Alcohols in a Slurry Reactor with CS-promoted Zinc Chromite Catalyst in Decahydronaphthalene 2003 Xiaolei Sun, G.W . Roberts North Carolina State Applied Catalysis, Elsevier

Csto

9535

Fuel from the Synthesis Gas - the Role of Process Engineering 2003 Stelm achowski, Marek, et.al. Technical University of Lodz, Poland Applied Energy, Elsevier

Ms

N/A

Alcohol Synthesis over Pre-Reduced Activated Carbon-supported Molybdenum-Based Catalysts 2003

Dadyburjor, D.B., Li, Xianguo; et.al.

W est Virginia University; Ocean University of China Molecules

fooraH

N/A

Co-Production of Fuel Alcohols & Electricity via Refinery Coke Gasification 2003

Ravikumar, Ravi; and Shepard, Paul Fluor Gasification Technologies Conf

UEaaa

115
9519

The Influence of Clay on K2C O 3 /CoMoS2 Catalyst in the Production of 

Higher Alcohol Fuel 2002

Iranm ahboob, Jam shid; H ill, D onald Uconn, M ississippi State

Fuel Processing Technology, Elsevier

9536

Alcohol Synthesis from Syngas over K2C O 3 /CoS/MoS2 on Activated Carbon 2002

Iranm ahboob, Jam shid; H ill, D onald Uconn, M ississippi State Catalysis Letters

9528

Advances in Catalytic Synthesis and Utilization of H igher Alcohols 2000 Herm an, R.G. Lehigh Univ. Catalysis Today, ElsevierThe Econom ical Production of Alcohol Fuels from Coal-derived Synthesis Gas: Final Report, O ctober 1991 - M arch 1998 1999 W est Virginia University

Subm itted to National Energy Technology Lab (NETL), Contract #DE-AC22-91PC91034

25

9545

Synthesis of Higher Alcohols - Enhancem ent by the Addition of M ethanol or Ethanol to the Syngas 1999 Lachowska, M. Polish Academ y of Sciences React.Kinet.Catal.Lett Elsevier

N/A

Form ation of Ethanol and Higher Alcohols from Syngas 1999

Iranm ahboob, Jam shid; D onald Hill M ississippi State University

M ississippi State University, Thesis Report

9515

Selective Synthesis of M ixed Alcohols Catalyzed by Dissolved Base-activated Highly Dispersed Slurried Iron 1999 M ahajan, D. et.al. Brookhaven National Laboratory Fuel (Elsevier)

9538

Synthesis of Short Chain Alcohols over a Cs-prom oted Cu/ZnO/Cr2O 3 catalyst 1998

M ajocchi, L.; Forzatti, Pio; et.al.

Politecnico di M ilano; Snam progetti SpA Applied Catalysis, Elsevier

9511

Higher Alcohols from Synthesis Gas Using Carbon-Supported Doped M olybdenum -Based Catalysts 1998

Xianguo Li; Dadyburjor, Dady B. et.al. W est Virginia University Ind. Eng. Chem . Res. (ACS)

9537

A Kinetic M odel for the Synthesis of 

High-M olecular-W eight Alcohols over a Sulfided Co-K-M o/C Catalyst 1998

G unturu, A.K.; Dadyburjor, D.B., et.al W est Virginia University Ind. Eng. Chem . Res. (ACS)

9529

Developm ent of a Process for Higher Alcohol Production via Synthesis Gas 1998 Beretta, A.; Tronconi, E. et.al. Politecnico di M ilano; Snam progetti SpA Ind. Eng. Chem . Res. (ACS)

9530

High Tem perature Calcined KMoO 3 /gam m a-Al2O 3 Catalysts for M ixed 

Alcohols Synthesis from Syngas: Effects 

of M o Loadings 1998

Bian, G uo-zhu, et.al. (U niv. of S&T, China)

University of Sci & Tech, China; Univ of Tokyo Applied Catalysis, Elsevier

9514

Kinetic Analysis of M ixed Alcohol Synthesis from Syngas over K/M oS2 

Catalyst 1997 Park, T.Y.; Nam , In-sik, et.al. 

Pohang Univ. of Sci. & Tech (Postech), Korea Ind. Eng. Chem Res. (ACS)

9518

Synthesis of M ixed Alcohols from CO 2 

contained syngas on supported M oS2 catalysts 1997 G ang, Lu, et.al. East China University of Sci & Tech Applied Catalysis, Elsevier

116
9524

Development of a Mechanistic Kinetic Model of the Higher Alcohol Synthesis over a Cs-doped Zn/Cr/O Catalyst - 1) Model Derivation and Data Fitting 1996 Beretta, A.; Forzatti, Pio, et.al. Politecnico di Milano; Snamprogetti SpA Ind. Eng. Chem. Res. (ACS)

9525

A Kinetic Model for the Methanol-Higher Alcohol Synthesis from CO/CO2/H2 over 

Cu/ZnO-based Catalysts Including Simultaneous Formation of Methyl Esters and Hydrocarbons 1994

Breman, B.B.; Beenackers, A.C.C.M.; et.al.

University of Groningen, The Netherlands

Chemical Engineering Science, Pergamon

9526

Thermodynamics of Higher Alcohol Synthesis 1993

Mawson, S.; Roberts, G.W .; et.al. North Carolina State Energy & Fuels

11636

SiO2-supported Biometallic Ruthenium Cluster-derived Catalysts for Selectivity and Activity Control in CO Hydrogenation Toward C1-C5 Alcohols 1992 Xiao, F.; Ichikawa, M.; et.al. Hokkaido University

Journal of Molecular Catalysis, Elsevier 

9517

The Thermodynamics of Higher Alcohol Synthesis 1992 Roberts, G.W ., et.al. North Carolina State

Symposium on NG upgrading II, Division of Petroleum Chem, Inc., ACS

9523

An Overview of the Higher Alcohol Synthesis 1992 Smith, K.; Klier, K. Univ. British Columbia; Lehigh

Symposium on NG upgrading II, Division of Petroleum Chem, inc., ACS

9521

Non-Stoichiometric Spinel-type Catalysts for the Synthesis of Methanol and Higher Alcohols 1992 Trifiro, F., et.al.

Dipartim ento di Chimica Industriale e dei materiali, Bologna, Italy

Symposium on Octane and Cetane Enhancement Processes for Reducedemissions Motor Fuels, Division of Petroleum Chemistry, ACS

11632

The Conversion of Synthesis Gas to Higher Oxygenated Fuel on Rh-based Catalysts: Effect of Chemical Additives 1991

Balakos, M.W ., Chaung, S. S. C, University of Akron

Fuel Science and Technology International

9522 Higher Alcohol Synthesis 1991 Forzatti, Pio, et.al. 

Politecnico di Milano; Snamprogetti SpA Catal. Rev. - Sci. Eng

Production of Oxygenates from Synthesis Gas: A Technology Review for US Dept of Energy for Design of Generic Coal conversion Facilities 1991

Pittsburgh Energy Technology Center M.W . Kellogg Company Contract #DE-AC22-91PC89854

Classical and Non-classical Routes for Alcohol Synthesis 1991 Herman, R.G. Lehigh University

Book Chapter, New Trends in CO Activation, L. Guczi, Ed. Elsevier

9539

Catalytic Conversion of Synthesis Gas 

to Methanol and Other Oxygenated Products 1991 Stiles, Alvin, et.al. University of Delaware Ind. Eng. Chem. Res.; ACS

9531

Kinetic Modelling of Higher Alcohol Synthesis over Alkali-promoted Cu/ZnO and MoS2 Catalysts 1990 Smith, K. Herman, R.; Klier, K. Lehigh University Chemical Engineering Science

117
9541

Production of Mixed Alcohols from Synthesis Gas 1990 Uchiyama, S., et.al. Idemitsu Kosan Co. Sekiyu Gakkaishi

9568

Production of Methanol-higher Alcohol Mixtures from Natural Gas via Syngas Chemistry 1990

Coutry, Chaumette, Raimbault, Travers French Petrol Institute (IFP)

European Applied ReConf. On Natural Gas

Production of Methanol-higher Alcohol Mixtures from Natural Gas via Syngas Chemistry 1990 Production of Methanol-higher Alcohol Mixtures from Natural Gas via Syngas Chemistry 1990 Production of Methanol-higher Alcohol Mixtures from Natural Gas via Syngas Chemistry 1990 Production of Methanol-higher Alcohol Mixtures from Natural Gas via Syngas Chemistry 1990

9532

Synthesis of Alcohols from Carbon Oxides and Hydrogen. 4. - Lumped Kinetics for the Higher Alcohol Synthesis over a Zn-Cr-K Oxide Catalyst 1987 Tronconi, E.; Forzatti, P; et.al.

Politecnico di Milano; Snamprogetti SpA Ind. Eng. Chem. Res

9512 Mixed Alcohols from Synthesis Gas 1986 Quarderer, George Dow Chemical AIChE Spring Nationa

SRI PEP #85-1-4 Mixed Alcohols from Syngas 1986 Nirula, Satish SRI PEP Review #85-1-4

118
 

Appendix K 

Alcohol Synthesis Kinetics 

119
 

A more rigorous method for modeling the production of alcohols or other products via synthesis 

is to use kinetics-based reactions to represent the multiple reactions that occur. Used with the 

appropriate reactor type, a good kinetics-based model can more accurately predict the effect of 

varying inlet conditions to the synthesis reactor, especially when unconverted syngas or 

unwanted co-products (i.e., methanol, hydrocarbons) are recycled to the reactor inlet. 

Two major barriers to using kinetics-based models are 1) the need for high-quality experimental 

data to determine the many parameters required for accurate predictions – assuming that a 

suitable model can be found for the system of equations – and 2) even with data suitable for 

estimating model parameters, the resulting model is only truly representative of the catalyst used 

in the experiments over the range of operating conditions explored. Using the model outside the 

range of operating conditions introduces increased uncertainty the more removed the estimate is 

from the experimental conditions used to develop the model. Catalyst performance is very 

sensitive to many factors that can arise during their production. Two seemingly identical 

catalysts based on chemical formulation can vary greatly in performance because of preparation 

techniques and catalyst support characteristics (surface area, crystal structure). 

Heterogeneous catalysts such as those used in most synthesis work today, are very sensitive to a 

number of factors including supports, amounts and types of added promoters or inhibitors that 

help tune the catalyst to promote desired reactions while inhibiting undesirable reactions. An 

“optimum” catalyst is usually a compromise of competing and interacting promoters, inhibitors, 

and supports to give a catalyst with the “best” cumulative properties. When precious metals are 

used, economics also become a major factor in determining the catalyst formulation. 

Very little published data exist for the MoS2-based catalysts that are suitable for developing a 

reasonable model of the alcohol synthesis together with the very important competing reactions 

that reduce the desired product yields. Gunturu (Gunturu, et al., 1998; Gunturu, 1997; Gunturu, 

et al., 1999) published relevant data from thesis work done at the University of West Virginia. 

The data were used to develop a system of Langmuir-Hinshelwood type reactions to describe a 

five-reaction lumped kinetic scheme. Recently, researchers in Italy, using the same data, 

generated a model using the same reaction scheme but with slightly different equations and 

parameters. The procedure for estimating the parameters is not discussed here, but can be found 

in detail in Larson, et al. (Larson, et al., 2006) and are presented here in Table 1. 

120
 

Table 1. Kinetic parameter estimates from Larson, et al. Blank cells do not have 

values for the corresponding rate equation parameter and rate equation. 

Parameter Methanol Ethanol Propanol Hydrocarbons 

Am, Ae, Ap, Ah 14.6233 3.0518 0.2148 9.3856 

Em, Ee, Ep, Eh 143.472 24.986 89.3328 95.416 

nm, ne, np, nh 3 1 1 1 

K1 7.6393E-9 

K2 0.6785 

K3 0.9987 

Ke 0.7367 

Kp 0.6086 

Kh 1.2472 

Kz 0.8359 

The rate equations were created using the typical Langmuir-Hinshelwood kinetic approach found 

in most catalyst kinetics textbooks. Equation 1 is the gross rate of methanol production. It is the 

only rate equation in the set that has a reverse reaction component and therefore an equilibrium 

value. Chemical equilibrium is reached when the partial pressure of methanol is sufficiently 

high to cause the reverse reaction rate, the dissociation of methanol to CO and H2, plus the 

consumption rate of methanol to other products, to equal the forward reaction rate of CO and H2

to make methanol. Given sufficient time in an active catalyst bed, the methanol concentration 

will reach an equilibrium state. The equilibrium value will include the effects of methanol being 

consumed to make ethanol or methane as described by equations 4 and 6, respectively. 

2 3

, 2, 3 ,

3

2 3

123

, 2, 3 ,

2

( / )(1/ 1/ ) 1

1

CO H CH OH m cp

m

CO cp H cp eq CH OH cp

CH OH

CO H CH OH

CO cp H cp CH OH cp

ER T T

gross

nm

pp p A e

p p Kp

r

ppp KKK

ppp

− − ⎛ ⎞ ⎡ ⎤⎡ ⎤ ⎡ ⎤ ⎜ −⎟ ⎢ ⎥⎢ ⎥ ⎢ ⎥ ⎜ ⎟ ⎣ ⎦⎣ ⎦ ⎣ ⎦ ⎝ ⎠ = ⎛ ⎞ ⎡⎤⎡⎤⎡ ⎤ ⎜ ⎟ ++ + ⎢⎥⎢⎥⎢ ⎥ ⎝ ⎠ ⎣⎦⎣⎦⎣ ⎦

 (1) 

where 

( )

3 ,

, 2,

2

CH OH cp

cp

CO cp H cp

p K

p p = (2) 

and 

a

eq

z cp

K K

KK =

 (3) 

121
 

2 5

3

3 ,

3

3 ,

( / )(1/ 1/ )

1

C H OH

CH OH

CH OH cp gross

CH OH

CH OH cp

e cp

e

e

e

ER T T

n

p A e

p

r

p K

p

− − ⎡ ⎤ ⎢ ⎥ ⎣ ⎦ = ⎛ ⎞ ⎡ ⎤ ⎜ ⎟ + ⎢ ⎥ ⎝ ⎠ ⎣ ⎦

 (4) 

2 5

25 ,

3 7

2 5

25 ,

( / )(1/ 1/ )

1

C H OH

C H OH cp

C H OH

C H OH

C H OH cp

p cp

p

gross

p

p

ER T T

n

p A e

p

r

p K

p

− − ⎡ ⎤ ⎢ ⎥ ⎣ ⎦ = ⎛ ⎞ ⎡ ⎤ ⎜ ⎟ + ⎢ ⎥ ⎝ ⎠ ⎣ ⎦ (5) 

3

3 ,

3

3 ,

( / )(1/ 1/ )

1

CH OH

CH OH cp

HC

CH OH

CH OH cp

h cp

h

gross

h

h

ER T T

n

p A e

p

r

p K

p

− − ⎡ ⎤ ⎢ ⎥ ⎣ ⎦ = ⎛ ⎞ ⎡ ⎤ ⎜ ⎟ + ⎢ ⎥ ⎝ ⎠ ⎣ ⎦ (6) 

The other rate equations describe presumed irreversible reactions of methanol to ethanol, ethanol 

to propanol, and methanol to methane. Ethanol can react further with H2 and CO to make 

propanol. Implicit in this set of equations is the simplifying assumption that any products 

generated other than methanol, ethanol or propanol produce methane even though in higher 

molecular weight hydrocarbons are observed experimentally. Small amounts of butanol and 

pentanol, both expected and experimentally observed, have been ignored in the analysis since 

CO conversions were low resulting in minimal production of increasingly higher alcohols. 

Using the rate equations above, plus a rate equation for the water-gas-shift reaction , a system of 

differential equations and initial conditions can be easily written for a plug flow reactor, as in 

this report, or for a Continuous Stirred Tank Reactor (CSTR). The solution of the system of 

ordinary differential equations was programmed into the commercially available equation solver, 

PolyMath, to evaluate the kinetics over a range of conditions. 

Examples of the results given by the model are shown in Figure 1 and Figure 2. In the first case 

the methanol rate reaches a constant value quickly meaning it is consumed at the same rate it is 

produced. In the second case methanol was “added” to the reactor inlet to make 1 mol% 

concentration. CO and H2 inlet flow was reduced equally to keep molar inlet flow the constant. 

The methanol has a negative rate of production initially and then reaches a constant flow rate at a 

value slightly below 2000 lb/hr. The ethanol flow rate is significantly higher than the case with 

no methanol added to the feed. 

122
 

Mass Flows vs CO Conversion

Temp = 570K; Pressure = 68 atm

FH2 = 25000 Kmol/hr; FCO = 25000 Kmol/hr; Fmeoh = 0 Kmol/hr

0

2000

4000

6000

8000

10000

12000

14000

16000

18000

0 2 4 6 8 10 12

CO Conversion

Flow rates, lb/hr

WP WE WH WM

Figure 1. Mass Flow Rates of the Synthesis Reactions vs. CO Conversion in a 

Isothermal, Plug-Flow Reactor at a WHSV of 1000 L-kgcat-1-hr-1

Mass Flows vs CO Conversion

Temp = 570K; Pressure = 68 atm

FH2 = 24750 Kmol/hr; FCO = 24750 Kmol/hr; Fmeoh = 500 Kmol/hr

0

5000

10000

15000

20000

25000

30000

0 2 4 6 8 10 12 14 16

CO Conversion

Flow rates, lb/hr

WP WE WH WM

Figure 2. Mass Flow Rates of the Synthesis Reactions vs. CO Conversion in a 

Isothermal, Plug-Flow Reactor at a WHSV of 1000 L-kgcat-1-hr-1

123
 

For the catalyst used to develop this model, the CO conversion is relatively low at 570K and 68 

atm pressure. With no methanol recycle to the reactor, the CO conversion is 10.7% with an 

ethanol yield of 17000 lb/hr. Adding methanol to the reactor inlet increases the CO conversion 

to 13.5% with an ethanol yield of 25,900 lb/hr. Adding methanol to the synthesis reactor appears 

to improve the catalyst performance for making ethanol. However, there is insufficient methanol 

exiting the reactor to meet the inlet demands. At this temperature and pressure, a lower inlet 

methanol concentration is needed to be sustainable in a process without modifying the process to 

produce more methanol. 

Values of CO conversion and product rates at the reactor outlet for a plug flow reactor at 68, 

with and without methanol added to the inlet, are given in Error! Reference source not found.

for the kinetic model. The first column is for no methanol added to the inlet stream. The second 

column has 1 mol% methanol in the inlet. The GHSV is approximately 1000 L-kgcat-1-hr-1 for 

all cases shown, which is significantly less than the experimental condition at which the kinetic 

parameters were estimated for the kinetic models. 

Table 2. Results From the Kinetic Model at 68 atm Pressure 

 MeOH @ inlet = 0 Kmol/hr MeOH @ inlet = 500 Kmol/hr 

570K XCO = 5.67 

MeOH = 2023 

CH4 = 3237 

EtOH = 9861 

PrOH = 3851 

XCO = 6.38 

MeOH = 1974 

CH4 = 5770 

EtOH = 19600 

PrOH = 6543 

610K XCO = 27.08 

MeOH = 4282 

CH4 = 26400 

EtOH = 25600 

PrOH = 21400 

XCO = 25.94 

MeOH = 4197 

CH4 = 28800 

EtOH = 28300 

PrOH = 23800 

At 570K, adding methanol to the inlet gas causes the rates of ethanol and propanol production to 

increase significantly with little change in the CO conversion. The methanol equilibrium rate is 

approximately 2000 Kmol/hr at this temperature. Increasing the temperature to 610K greatly 

increases the conversion of CO and of the product yields. However, the methane production has 

increased relative to higher alcohols. The effect of adding methanol at 610K is less pronounced 

than at 570K. The CO conversion decreases slightly and the alcohol yields are only slightly 

increased. The methanol equilibrium rate has increased at 610K to approximately 4200 Kmol/hr. 

In the “methanol added” cases shown, there is insufficient methanol exiting the reactor to supply 

the inlet after separating the methanol from the other alcohols. A sustainable level would be 

reached using a lower methanol concentration at the inlet. Higher conversion can also be 

achieved by decreasing the GHSV although this will require a larger reactor and more catalyst. 

The optimum combination involves economics to determine when it is infeasible to increase the 

capital and catalyst costs to increase production. 

The effect of adding methanol on the ethanol yield suggests that other configurations may be 

more effectively used to maximize ethanol production. Since the methanol production at the 

124
 

reactor exit is limited by chemical equilibrium, the amount of methanol that can be recycled to 

the reactor is also limited. Instead of recycling only methanol produced from a mixed alcohol 

catalyst, it may be possible to economically split the main syngas flow with one portion going to 

a methanol synthesis reactor to make methanol and the remaining fraction going to the mixed 

alcohol reactor where the methanol is added to the syngas before entering the reactor. It could 

also be possible to do the reaction in series with the total syngas going through a methanol 

synthesis catalyst first followed the mixed alcohol catalyst. The methanol catalyst section would 

need to be sized to give only partial conversion to methanol, contrary to the way these reactors 

are typically operated when methanol is the desired end product. These alternate designs would 

require other factors to be considered, such as lower sulfur and CO2 concentrations for methanol 

synthesis relative to the mixed alcohol catalyst used. 

As more data become available, the kinetic model can be updated to include different catalysts 

and expanded operation ranges, especially in regards to the amount of methanol and CO2 that can 

be fed to the reactor without adversely affecting product selectivity and CO conversion. 

References 

Gunturu, A.; et.al. ”A Kinetic Model for the Synthesis of High-Molecular-Weight Alcohols over 

a Sulfided Co-K-Mo/C Catalyst.” Ind. Eng. Chem. Res. Vol. 37, 1998. pp. 2107 – 2115. 

Gunturu, A.K., Kugler, E.L., Cropley, J.B., Dadyburjor, D.B., “A Kinetic Model for the 

Synthesis of High-Molecular-Weight Alcohols over a Sulfided Co-K-Mo/C Catalyst”, Chapter 5 

in “The Economical Production of Alcohol Fuels from Coal-Derived Synthesis Gas”, Final 

Report, Contract # DE-AC22-91PC91034-25, West Virginia University Research Corporation on 

behalf of West Virginia University, March 1999. 

Gunturu, A.K., “Higher-Alcohol Synthesis from Carbon Monoxide and Hydrogen: Kinetic 

Studies Over MoS2-Based Catalysts”, M.S. Thesis in Chemical Engineering, West Virginia 

University, Morgantown, West Virginia, 1997. 

Larson, E.D., Consonni, S., Katofsky, R.E., Iisa, K. Frederick, J., “A Cost-Benefit Assessment of 

Gasification-Based Biorefining in the Kraft Pulp and Paper Industry, Volume 2, Detailed 

Biorefinery Design and Performance Simulation”, Final Report under contract DE-FC26-

04NT42260 with the U.S. Department of Energy and with cost-sharing by the American Forest 

and Paper Association, December 2006. 

125
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