A national laboratory of the U.S. Department of Energ y Office of Energy Efficiency & Renewable Energ y 

### 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 Laborator y 1617 Cole Boulevard, Golden, Colorado 80401-3393

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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 br oadening the country’s domestic production of economic, flexible, and secure sources of energy fu els. 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 Ener gy Initiative which outlines significant new investments and policies to change the way we fu el 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” ha s been quantified as $1.07 per gallon ethanol minimum plant gate price.

This process design and tec hnoeconomic evaluation addresses the convers ion 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 wa s 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 et hanol production are very similar to those for methanol production (although the specific process configuration will be di fferent). The general process areas include: feed preparation, gasi fication, 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 be ing developed as part of the OBP research program. Combined, all process, ma rket, and financial targets in th e design represent what must be achieved to obtain the reported $1.01 per gallon, showing that ethanol fr om a thermochemical conversion process has the possibility of being pr oduced 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 manne r. This allows for greater flexibility in converting biomass resources to ma ke stated volume targets by 2030.

iTable 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

iiList 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 Paramete rs, 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 R eaction 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 Inform ation 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 br oadening the country’s domestic production of economic, flexible, and secure sources of energy fu els. 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 EnergyInitiative 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” ha s 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 w ould 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 show n: 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.

10 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 ethanol a

Conceptual process designs and associated design reports have previously been done by NREL for converting cellulosic biomass feedstock to et hanol 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 manageme nt for two main reasons. First of all, they enable comparison of research and development projects . A conceptual process design helps to direct research by establishi ng a benchmark to which other process configurations can be compared. The anticipated results of proposed resear ch can be translated into design changes; the economic impact of these changes can then be dete rmined 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 th at 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 c omparison 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 es timates 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.

2applicable engineering, construction, and opera ting practices for facilities of this type. The complete process (including not only industry-st andard process components but also the newly researched areas) must be designed and their costs determined.

Following the methodology of the biochemical desi gn reports, this process design and technoeconomic evaluation addresses the conversion of biomass to etha nol 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. (S ee 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 conve rsion assessments done by NREL and others. This assessment direct ly builds upon an initial analysis for the TC production of ethanol and ot her 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 feedst ock, switchgrass, and it considers a variety of thermochemical product options, including etha nol, power and Fischer-Tropsch liquids [ 11 ]. Other notable gasification studies have been co mpleted by Larsen at Princeton University, including a study examining the bioproduct potential of Kraft mill black liquor based upon gasification [ 12 ].

Indirect steam gasification wa s 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 spec ific process configura tion will be different). The general process areas include: feed prep aration, gasification, gas cleanup and conditioning, and alcohol synthesis & purification.

Gasification involves the devolatiliz ation 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 char s. 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 larg e plant sizes to improve economics [ 14 ].

3The second general class, steam gasifiers (indirectly-heated gasifi ers), 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 prod uct gas can be combusted with air (external to the gasifier itself) to provide the energy requi red for gasification. Steam gasifiers have the advantage of not requiring oxygen; but since most operate at low pressu re they require product gas compression for downstream purification a nd synthesis unit operations. The erosion of refractory due to circulating hot solids in an i ndirect 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. Th ese 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 th e 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. Th is figure shows that steam gasification capital intensity is comparable or lower than POX gasification. The estimates

indicate that both steam gasification and POX gasification pro cesses should be evaluated, but if the processes need to be evaluated sequentia lly, choosing steam gasification for the first

evaluation is reasonable.

4Capital 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 th e 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 biom ass could be added as fuel direct ly to the heat and power system, however, this would increase the process bey ond 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 lo wers 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 el ectricity sales to the grid nor electricity purchases. The only exception to this would be if other operating sp ecifications 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 cr edit). This resulted in 28% of the unconditioned syngas being diverted to power the process. Mode l calculations show that if none of the syngas was diverted in this manner, a nd 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 to tal production of all alcohols as high as 130.3

5gal/dry ton. However, the minimum ethanol plant gate price increases in this scenario because of the cost of the natural gas required to m eet the energy demands of the process.

2.1. Analysis Approach

The general approach used in the developmen t of the process desig n, process model, and economic analysis is depicted in Figure 3 . The first step was to assemble a general process flow

schematic or more detail ed 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 equipmen t 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 cas h flow rate of return (DCFROR) analysis was used to determine the ethanol minimum plant gate price necessary to meet an n th plant hurdle rate (IRR) of 10%.

Figure 3. Approach to process analysis

6This TC conversion process wa s 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 th e mass and energy balances for the process. The operations were separated in to 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 compon ent properties databa se. The raw biomass feedstock, ash, and char components were mode led 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 r eactor). Because this design pro cesses three different phases of matter (solid phase, gas phase, and liquid pha se), 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 be havior. The non-random two-liquid “NRTL” option with ideal gas properties was used for alcohol separation ca lculations. 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, lo nger start-ups, and other costs associated with first-of-a-kind plan ts 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 th e actual installation of the purchased equipment but also instrumentation and controls, piping, el ectrical systems, buildings, yard improvements,

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

7The purchased component equipmen t 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 expone ntial 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 (typical ly 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 ex changer (appropriate if the log-mean temperature difference is known not to change greatly). Genera lly 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 multip le 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 200 2 (essentially zero infla tion) 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

8Chemical 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 we re determined, we a pplied 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 analys is 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.

92.2. Process Design Overview

Figure 5. Block flow diagram

A simple block flow diagram of th e 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 supp lied by circulating hot synthetic olivine a “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, H 2 , CO 2 , CH 4 , etc.), tars, and a solid “char” that is mainly the fixed carbon residual from the biomass plus carbon (coke) deposited on the sa nd. 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 be d 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 com bustor by the combustion gases and separated from the hot gases using another pair of cyclon es. 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 (MgSiO 3 ), Forsterite (Mg 2 SiO 3 ), and Hematite (Fe 2 O 3 ). 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.

10cyclone. The second cyclone is designed to capture the ash and any sand passing through the first cyclone. The hot sand captured by the fi rst cyclone flows by gravity back into the gasifier to provide the heat for the gasifi cation 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 H 2 ; syngas cooling/quench; and acid gas (CO 2 and H 2 S)

removal with subse quent reduction of H 2 S 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 additiona l 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 CO 2 and H 2 S. The H 2 S is reduced to elemental sulfur and stockpiled for disposal. The CO 2 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 unc onverted 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 recy cled back to the entrance of the alcohol synthesis reactor, this recycle is not done in this process design because CO 2 concentrations in the recycle loop would increase beyond acceptable limits of the catalyst. Added cost would be incurred if this CO 2 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 ro ugh separation is performed in a flash separator; the evolve d syngas is recycled to the Gas Cleanup & Conditioning section, mostly as feed to th e tar reformer. The depressurized alcohol stream is dehydrated using vapor-phase mol ecular sieves. The dehydr ated alcohol stream is introduced to the main alcohol separation column that splits methanol and ethanol from the higher molecular weight alcohols. The ove rheads 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 h eat (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 generato rs, and super-heaters are integrated within the process design to create the steam. The steam will run through turbines to drive compressors, generate elec tricity or be withdraw n at various pressure

11levels 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 de termine 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 T on Vision study addressed short and long term availability issues for biomass feedstocks wit hout 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 Cr ops”) are projections of availability that will depend upon changes to agricultural practices and th e creation of a new energy crop industry. In the target year of 2012 it is most probable th at the amounts labeled “Existing & Unexploited Resources” will be the only ones that can be c ounted 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.

120 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 av ailability of biomass

Past analyses have used hybrid poplar wood chip s 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 basis 29 )

Carbon 50.88 Hydrogen 6.04 Nitrogen 0.17

Sulfur 0.09 Oxygen 41.90

Ash 0.92

Heating value c (Btu/lb): 8,671 HHV d

8,060 LHV e

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 f eedstock 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

13forest 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 stud y is based upon the current operation and R&D performance goals for the catalytic tar destructio n and heteroatom removal work at NREL and alcohol synthesis work at NREL and PNNL. This ta rget 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 pr ocessing in the gasifier. The design is based upon a woody feedstock. It is expected that a fe ed 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 ea ch truck capacity is about 25 tons [ 31 ], this means that if the wood, at a moisture content of 50%, was deliver ed 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 st orage pile. From the storage pile, the wood chips are conveyed (C-102) through a magnetic se parator (S-101) and scre ened (S-102). Particles larger than 2 inches are sent through a ha mmer 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, paralle l feed handling and drying trains. The wet wood

14chips 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 combus tor (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 circ ulating olivine that is heated by the combustion of the char downstream of the gasifier. The steam primarily acts as a flui dizing medium in the gasifier and also participates in certain reacti ons when high gasifier temperatures are reached.

From the feed handling and drying section, the drie d 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 corr elations 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 (MgSiO 3 ), Forsterite (Mg 2 SiO 3 ), and Hematite (Fe 2 O 3 ), used as a heat transfer solid for various applications. A small amount of MgO must be

added to the fresh olivine to avoid the formati on 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, K 2 SiO 4 , with the silica in the system. K 2 SiO 4 has a low melting point (~930°F, 500°C) and its formation will cause the bed media to become sticky, agglomer ate, 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 ga sifier/combustor cyclones is also significantly

reduced. The ash content of the feed is assu med 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 oliv ine 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.

15Both the gasifier and the char combustor temperatures are allowed to “float” and are dictated from the energy balances around the gasifier and co mbustor. 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 equ ilibrium position. For the design case the resulting gasifier temperature is 1,633°F (889°C) a nd 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 bot h) is separated in the primary gasifier cyclone (S-201) and gravity-fed to th e 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% separati on in S-204) followed by an electrostatic precipitator (S-205) which remove s the remaining residual amount of solid particles. The sand and ash mixture from the secondary flue gas cycl one 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. Firs t 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)

H 2 15.0 25.1

CO 2 7.4 12.4

CO 25.1 41.9

H 2 O 40.2 -- CH 4 9.0 15.1

C 2 H 2 0.3 0.4

C 2 H 4 2.5 4.1

C 2 H 6 0.1 0.2

C 6 H 6 0.1 0.1 tar (C 10 H 8 ) 0.1 0.2

NH 3 0.2 0.3

H 2 S 0.04 0.07

H 2 :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 S ection 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.

163.4. 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 ar e dictated by the require ments of the synthesis catalyst:

• The tars in the syngas are reformed to additional CO and H 2

.

• Particulates are removed by quenching.

• Acid gases (CO 2 and H 2 S) 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 H 2 while NH 3 is

converted to N 2 and H 2 . 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 be en achieved at NREL [ 32 ] and the conversions used in the simulati on corresponding to the 2012 research targets.

Table 4. Current and Target Design Performance of Tar Reformer

Compound Experimental
Conversion to CO & H 2 Target Conversion to

CO & H 2 Methane (CH 4 ) 20% 80% Ethane (C 2 H 6 ) 90% 99% Ethylene (C 2 H 4 ) 50% 90% Tars (C 10+ ) 95% 99.9% Benzene (C 6 H 6 ) 70% 99% Ammonia (NH 3 ) f 70% 90%

In the Aspen simulation the tar reformer operates isothermally at 1,633oF. An implicit assumption in this mode of opera tion 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, a nd operate the reformer adiabatically with a resulting temperature drop across the bed and a lowe r 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 N 2 and H 2

.

17Table 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)

H 2 37.4 43.0

CO 2 9.9 11.4

CO 37.4 43.0

H 2 O 13.0 --- CH 4 1.2 1.4

C 2 H 2 0.01 0.01

C 2 H 4 0.11 0.13

C 2 H 6 10.8 ppmv 12.4 ppmv

C 6 H 6 2.7 ppmv 3.1 ppmv tar (C 10 H 8 ) 0.5 ppmv 0.6 ppmv

NH 3 0.01 0.01

H 2 S 0.02 0.02

N 2 0.72 0.83

H 2 :CO molar ratio 1.00

Prior to the quench step, the hot syngas is cool ed 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), s hown in PFD-P800-A302. The scr ubber also removes impurities such as particulates, residual a mmonia, and any residual tars. The scrubbing system consists of a venturi scrubber (M-302) and que nch 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 treatme nt 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 temper ature of 140°F (60°C). The syngas is thencompressed 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 outle t 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 to lerated in the syngas. For example, sulfur

concentrations as H 2 S 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 H 2 S in the syngas to maintain catalyst activity. Because the syngas exiting the gasifier contains almost 400 ppmv of H 2 S, 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 CO 2 depends on the specific catalyst

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

of CO 2 while others, such as the sulfided molybdenum catalysts can tolerate relatively high CO 2 levels compared to the sulfur species. CO 2 is a major component of th e gasification product, so significant amounts of CO 2 may 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 CO 2 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 chos en. An amine system capable of selectively removing CO 2 and H 2 S 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 simp lified model of SEP bloc ks and specifying the

amount of CO 2 and H 2 S needing to be removed to meet design specifications of 50 ppmv H 2 S

and 5 mol% CO 2 at the synthesis reactor inlet, including any recycle streams to that unit operation. The amine system heating and cooling du ties were calculated us ing information taken from section 21 of the GPSA Data Handbook [ 34 ]. This method gave a heat duty of 2660 Btu per pound of CO 2 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 valu es 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 CO 2 removed 2,660 Btu/lb

If a highly CO 2 -tolerant alcohol synthesis catalyst is used, it may become possible to use other syngas conditioning processes or methods to selectively remove H 2 S, with less energy and possibly at a significantly lower capital cost.

The acid gases removed in the amine scrubber ar e then stripped to regenerate the sorbent and sent through a sulfur removal operation using a liquid phase oxid ation process as shown in PFDP800-A305. The combined Amine/ LO-CAT process will remove the sulfur and CO 2 to the levels desired for the selected molysulfide catalyst [ 35 ]. Although, there are several liquid-phase oxidation processes for H 2 S removal and conversi on available today, the LO-CAT process wasselected 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 cool er (H-305). Elemental sulfur is produced in

19the 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 H 2 S to a concentration of 10 ppmv in the CO 2 vent effluent from the amine scrubber. The air flow rate for re-oxidizing the LO-C AT solution was included in the simulation and calculated based on the requi rement of 2 moles of O 2 per mole of H 2 S. 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 proce ss is equivalent to 120  ̊ F. This degree of superheating is required for the LO-CAT system. The CO 2 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 ma keup 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 ac id gas concentrations (H 2 S, CO 2 ) reduced. After that, it is further compre ssed 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 s ubsequently cooled, allo wing the alcohols to condense and separate from the unconverted syngas. The liquid alcohols are th en 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 literatu re search was conducted. This search and its findings are described in more de tail 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 stat e 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 technol ogy evaluation, a modified Fischer-Tropsch catalyst was used for this process design, sp ecifically a molybdenum-disulfide-based (MoS 2 ) 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 metha nol catalysts). This particular catalyst uses high surface area MoS 2 promoted with alkali metal salts (e.g. potassium ca rbonate) and cobalt (CoS). These promoters shift th e product slate from hydrocarbons to alcohols, and can either be supported on alumina or activated carbon, or be used unsupported.

20Table 6 lists several process and sy ngas 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

H 2 /CO ratio 1.0 – 1.2 1.0

CO 2 concentration (mol%) 0% - 7% 5.0% Sulfur concentration (ppmv) 50 - 100 50

Though the synthesis reactor is modeled as ope rating 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. Te mperature has a significant impact on the alcohol selectivity and product distribution. High pre ssures are typically required to ensure the production of alcohols. MoS 2 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 alc ohol production. However, compression requirements for achieving these pressu res can be quite substantial. Thus, targeting a catalyst that achieves optimal performance at lower pressures can potentially provide significant cost savings.

The CO 2 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 CO 2 in the synthesis gas retarded the catalyst activity. Further study showed that increasing the CO 2 concentration to 30 vol% decreased the CO conversion but did not significantly alter the alcohol:hydrocarbon ratio of the product. With CO 2 concentrations up to 6.7 vol%, the extent of CO conversion is not affected; however, higher chain alcohol yield relati ve to methanol does tend to decrease. This is why CO 2 concentrations were reduced to 5 mol% in the model using the amine system as part of syngas conditioning. The effect of CO 2 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 H 2 S is maintained. Concentrations above 100 ppmv inhibit the reac tion rate and higher alcohol selectivity.

The overall stoichiometric reaction for al cohol synthesis can be summarized as:

n CO + 2 n H 2 Æ C

n H 2 n +1 OH + ( n -1) H 2 O

Stoichiometry suggests an optimum H 2 :CO ratio of 2.0. However this catalyst maintains significant water-gas shift activity and will generate its own H 2 from CO and H 2 O:

21CO + H 2 O Æ H 2 + CO 2.

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

CO 2 . 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 (assu ming 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), a nd sent to the reactor. The capital cost for the compressor was developed using Questimate.

The mixed alcohol synthesis reactor is a fixe d-bed reactor system that contains the MoS 2 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 es timated 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-bas ed 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 co sts 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 de mand goes up, and quite s ubstantially 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 co sts for these catalysts are diffi cult to predict at the present time since so few providers of mixed alcohol s catalyst currently exist (and will likely be negotiated). Nexant also provide d 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 w ill 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.

22Table 7. System of Reactions for Mixed Alcohol Synthesis

Water-Gas Shift CO + H 2 O ↔ H 2 + CO 2 Methanol CO + 2H 2 Æ CH 3 OH

Methane CH 3 OH + H 2 Æ CH 4 + H 2 O

Ethanol CO + 2H 2 + CH 3 OH Æ C 2 H 5 OH + H 2 O

Ethane C 2 H 5 OH + H 2 Æ C 2 H 6 + H 2 O Propanol CO + 2H 2 + C 2 H 5 OH Æ C 3 H 7 OH + H 2 O Propane C 3 H 7 OH + H 2 Æ C 3 H 8 + H 2 O

n-Butanol CO + 2H 2 + C 3 H 7 OH Æ C 4 H 9 OH + H 2 O

Butane C 4 H 9 OH + H 2 Æ C 4 H 10 + H 2 O

Pentanol+ CO + 2H 2 + C 4 H 9 OH Æ C 5 H 11 OH + H 2 O

Pentane+ C 5 H 11 OH + H 2 Æ C 5 H 12 + H 2 O

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 (CO 2 -free basis) 70% - 80% 90%

Gas Hourly Space Velocity (hr -1 ) 1600 – 12,000 4000Catalyst Alcohol Productivity (g/kgcatalyst/hr ) g 150 – 350 600

The individual target values are less important th an 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 CO 2 , 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 mu ch 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 de scribed 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/ft 3 , 600 g/kg-catalyst/hr = 615 g/L-catalyst/hr.

23Table 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 purifica tion and methanol recycle

The most significant differences between the NR EL 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 me thanol 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 litera ture, 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 decompos ition functionality which

indicates that methanol in the feed will not be de trimental to the reaction. The effects of recycled methanol will be examined experimentally as research progresses.

A kinetic model was used to guide these conversi on 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 predic ted 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 comp rised of unconverted syngas, CO 2 , and methane. Aspen Plus contains other physical propert y packages that model non-ideal liquid systems much better than the Redlich-Kwong-Soave (RKS) e quation 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

24the 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. On e particular variation would be to recycle the unconverted syngas to the throat of the synthesis reactor in stead of to the tar reformer. This would save money on upstream equipment costs because of lo wer process throughput, but would also lower yields because the CO 2 would build up in the recycle loop. The limit to the amount of unconverted syngas that could be recycled to th e reactor is less than 50% because this would

cause the H 2 :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 CO 2, CH 4 , and other compounds. Alternate reactor designs will also be examined. For example, FT technology largely ha s switched to slurry reactors instead of fixedbed reactors because the slurry fluidization achie ves better heat and mass tr ansfer properties that allow, in turn, for higher conversions. Such imp rovements could help to achieve the conversion targets outlined above and reduce th e 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 high er-molecular weight alcohols.

The methanol stream is used to back-flush the mo lecular 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 al cohol 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 st ream, especially at the high system pressure. These gases are removed by depressurizing fro m 970 to 60 psia. Most of the dissolved gasses separate from the alcohols in th e knock-out vessel S-502. This gas stream is made up primarily of carbon dioxide with some small amounts of hydr ocarbons 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 ba sed 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 wa ter (92.5 wt% ethanol). The adsorb ed 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 en tire 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 w ith methanol. This methanol/water mixture is then recycled back to the Alcohol Synthesis section (A400).

25The molecular sieve units requi re a superheated vapor. The li quid mixed alcohol stream is vaporized, superheated, and then fed to one of two parallel adso rption columns. The adsorption column preferentially removes water and a sma ll amount of the alcohols. While one adsorption bed is adsorbing water, the other is regenera ting. The water is desorbed from the bed during regeneration by applying a vacuum and fl ushing with dry methanol from D-505. This methanol/water mixture is recycled back to the Alcohol Synthesis section (A400). This

methanol/water mixture is cool ed 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 si eve 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 sepa rated 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 cool ed and sent to storage. The methanol/ethanol overhead stream from D-504 goes to a second di stillation column, D-505, for further processing.

D-505 separates the methanol fro m the binary methyl/ethyl alcohol mixture. The ethanol recovery in D-505 is 99% of the incoming etha nol and has a maximum me thanol concentration of 0.5 mole percent to meet product specifications for fuel etha nol. 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 a nd 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 pro cess; 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 pr oduced 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 show n 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 tu rbine condensate and make -up water. The total

26condensate stream is heated to the saturation temp erature 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 co llected in the steam drum (T-604). To prevent solids build up, water must be periodically disc harged from the steam drum. The blowdown rate is equal to 2% of water circul ation rate. The saturated steam fro m 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 s lipstream 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 comp ressors in this process (K301, K410) are steamdriven. All other drives for pumps, fans, etc ar e 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 pow er requirement much less than it would be if comp ression 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 plan t 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 th e 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

273.8. 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 exchange rs in the plant. The tower utilizes large fans to force air through circulated wa ter. 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 c ondenses the steam exiting the steam turbine

Make-up water for the cooling tower is supp lied at 14.7 psia and 60°F (16°C). Water losses include evaporation, drift (water entrained in the cooling towe r exhaust air), and tower basin

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

blowdown were calculated based on informa tion 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 poi nt 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.

283.9. Additional Design Information

Table 11 contains some additional information us ed 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

T Dry Bulb : 90°F

T Wet Bulb : 80°F Composition (mol%) : N 2 : 75.7% O 2 : 20.3% Ar: 0.9% CO 2 : 0.03% H 2 O: 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. N itrogen value adjusted slightly to force mole fraction closure using only N 2 , O 2 , Ar, and CO 2 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., thos e 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 th is 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.

290 500 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/hrTemperature, 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 intercha nges or process-steam interchanges. The minimum vertical

distance between the curves is Δ T min , 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 Δ T min 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 practic e 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 networ k 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

30recovered is required. Detailed energy balances around the major process areas were derived using data from the Aspen Plus simulation. Co mparing 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 somewh at different from what

was done for the biochemical etha nol 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 thermo chemical ethanol production process differs significantly in that most of the process str eams 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 te mperature (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 c ooling 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 etha nol process was analyzed, the contributions for the HHVs, the sensible heating effects, and the la tent heat effects were directly computed and

combined. The calculations of th e sensible and latent heat effe cts 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 en ergy potential. The enthalpies calculated and reported by Aspen Plus are actually based upon a h eat of formation for the energy potential of a stream. So, the reported enthalpies already include the sensib le, latent, and non-ideal mixing effects. If certain constants in Aspen’s enthalpy expressions could be m odified 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. Ho wever, since the constants cannot be easily changed, the reported enthalpy values were adjusted instead as part of a spreadsheet calculation.

31The factors used to adjust the reported enthal pies were calculated from the difference between each component’s heat of combustion (LHV) a nd the reported pure component enthalpy at

combustion conditions.

This process for thermochemical conversion of ce llulosic 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-produc t either. The only saleable products are the fuel ethanol and a higher molecular we ight 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 th e primary energy input, however othe r 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 boile r 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.

32Table 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 CO 2 ) 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 importa nt 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 biom ass. Cooling tower eva porative losses, excess CO 2 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 ambien t heating effects and mechanical work (pump, compressor) efficiency losses.

Some of this lost heat could potentially be rec overed 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 c onsideration. Additional h eat integration with process streams could also be examined, however , there comes a point where this becomes too complex and costly for a cost-effec tive 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 recovere d in the two alcohol products. Improvements in these energy efficiencies could potentially result in additional cost savings to the process.

333.12. Water Issues

Water is required as a reactant, a fluidizing agen t, 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 importa nt 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 concerns 51 . For several years, significant areas of water stress have been reported during the gr owing 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, wh ich draws from the local aquifers that are not readily recharged. This is driven by the need fo r high quality water in the boiler system. Studies have shown that water usage by today’s corn et hanol 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 inject ed into the gasifier, but other system losses (blowdown) also exist. The overall water dema nd 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 et hanol 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 e quipment cost) as well as variable and fixed operating costs is developed first. With these cost s, 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 co mplete the discounted cash flow analysis.

4.1. Capital Costs

The following sections discuss the methods and so urces 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 po ssible on the design work done by Spath et al . for the hydrogen design report [ 9 ] and Aden et al. for the bioche mical 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 As pen 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 th e TIC and the total indirect costs).

35Table 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 we re examined for the BCL gasifier. HGCU = hot gas clean up.

36Table 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, comp ressors, blowers and pumps were estimated for a “base” size using Questimate and then scaled us ing material and energy balance results from the Aspen Plus simulation. The reactors were size d based on a gas hourly space velocity (GHSV),

where GHSV is measured at standard te mperature 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 most ly developed based on the required surface area as calculated from the heat transfer equa tion 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 

### T Δ

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.

37For 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 hi gher of the operating pressure plus 25 psi or the opera ting 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 dis tillation 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 account ed for in this analysis. The variables, information about them, and costs associ ated 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: $2.07/100ft 3 [ 70 ]

38Previous biomass gasification st udies 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 re flect 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 ente ring 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

39Since the salaries listed above are not fully lo aded (i.e. do not include benefits), a general overhead factor was used. This also covers genera l 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 sa laries $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 Statis tics data and severa l 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.

40Table 21. Salary Comparison

Job Title in

Biomass Plant Corresponding Job Title in Salary Estimating Tool [ 72 ] Salary Range (17 th to 67 th
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 mark et. 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 an y 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 co sts 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 va lue 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 hea ting value), this translates to $0.85 per gallon.

41For 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 woul d 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 eco nomic 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

1 st

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 1 st
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 previ ous 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, ma rket, and financial targets in th e design represent what must be achieved to obtain the reported $1.01 per gallo n. 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 tw o different co-product credits: alcohols from the Alcohol Synthesis area and el ectricity from the Steam Syst em & 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 significantl y 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 an alysis. 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 tenta tive, 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 $1.20Mininum Ethanol Selling Price, MESP($ 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 Se lling 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 MESP h

The process costs (as indicated by the MESP) are determined from various assumptions on technology (based upon 2012 research targets), market s (such as the value of the higher alcohol co-products), and various financ ial assumptions (such as require d Return on Investment, ROI). When any research target cannot be obtained, or a market or financia l assumption does not hold, then the MESP is affected to varying degrees. In addition, uncertainty about equipment design

and installation and cons truction costs will impact the economic s. The key is to understand the impact of those types of parameters that are likel y to vary, and how they might be controlled to a definable range. Discussed here are pr ocess targets that had been identified a priori as key ones to understand and achieve. (As can be seen fro m 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 demons trate the effect of technology advancement (or lack of) on the economic viability of the process.

h Note that the relative splits between feedstock and conversio n costs have been scaled to attribute some of the costs to the mixed alcohol co-products. So, th e feedstock contribution appears to be di fferent than what is depicted in the

cost contribution chart for the different areas.

44The 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). Th ere 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 furt her 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 demonstr ation projects would, at best , ease the ROI requirements of corporations and/or lending institutions and reduce the required MESP toward the baseline case

45in this report. Also, the baseline of 0% debt fi nancing is not a very re alistic scenario, but does represent the conservative endpoint. Many projects of this nature are often financed by some mixture of debt and equity fina ncing. 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) give s 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 plant i . 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 proces s 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 equipm ent 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 re duced 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.

460 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 $2.50Minimum Ethanol Selling Price($ 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 th e case of processing wet ensiled agricultural residues or energy crops; however, these feedstoc ks 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 hi gher alcohol yields. This is depicted in Figure 12

and Figure 13 . As the moisture content increases, th e 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 yiel ds; this is because flue gas is used for the drying and other operating specifications dictate th e amount of raw syngas diverted for heat and power, not the feedstock drying.

470 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 $1.50Minimum Ethanol Selling Price($ per gal) Total Alcohols

Ethanol

Figure 12. Sensitivity analysis of biomass moisture content

480 20 40 60 80 100 120

0% 10% 20% 30% 40% 50% 60%

Raw Syngas to Fuel & PowerAlcohol Yield (gal/dry ton)

$0.90 $1.00 $1.10 $1.20 $1.30 $1.40 $1.50Minimum Ethanol Selling Price($ 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. Th e 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 elect ricity 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.

495.3. Thermal Conversion

Two gasification scenarios were examined. The first was to explore the impact of increasing the olivine cost which could come about from cataly tic 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 meth ane production in the gasi fier to reflect a case where the gasifier does not operate within the gi ven 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 an d the hydrogen that would have gone to the methane instead went to the char (and was lo st for further processing). This required modifications to the operations to keep the H 2 :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 ef fects due to the material and energy balances. Since the amount of tar is small compared to th e 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 perform ance, the cost effects are majo r. 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-sul fide. 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, tota l CO conversion and alcohol selectivity (CO 2 -free basis) have the largest impact on the overall ME SP. 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 th e lower fuel oil value increases the MESP by about 6%; this is still reasonable to meet the qua litative “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 (a bout 20%). This shows that the first couple thermochemical conversion plants could get a si gnificant economic advantage in their early life by being able to do this.

50

### 6. Conclusions

This analysis shows that biomass-derived etha nol from a thermochemical conversion process has the possibility of being produced in a manner th at is “cost competitive with corn-ethanol” by 2012. This thermochemical conversion process woul d make use of many sub-processes that are currently used commercially (such as acid ga s removal) but also requires the successful demonstration of R&D targets be ing funded by the DOE’s OBP.

This analysis has demonstrated that forest re sources can be converted to ethanol in a cost competitive manner allowing greater flexibility in converting biomass resour ces to achieve stated volume targets by 2030.



### 7. Future Work

Future R&D work to develop and demonstrate refo rming and synthesis catalysts is inherent in this study. There many other areas of demons tration and process deve lopment also required:

• Demonstrate gasifier performance on other feed stocks (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 pa rticular importance for the lignin-rich residues is the impact on proce ss 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 po wer usage) vs. organic Rankine cycle for cooling and power production.

• Better understand the trade offs between oper ating conditions in the alcohol synthesis reactor to operating conditions (pressure, temperature, exte nt of reaction, extent of methanol recycle). A “tuned” kinetics ba sed model would be required for this.

• Explore alternate synthesis reactor config urations (slurry phase vs. fixed bed).

• Understand trade offs between an energy neut ral, alcohol production facility to one that could also supply heat and el ectricity to a co-located bi ochemical conversion facility.

• Further explore the potential benefits of in tegrating 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 Et hanol 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 Design and Economics Utilizing Co-Current D ilute Acid Prehydrolysis and Enzymatic Hydrolysis Current and Futuristic Scenarios , NREL report TP-580-26157, July 1999. http://devafdc.nrel.gov/pdfs/3957.pdf

6 Perlack, R.D.; Wright, L.L.; Turhollow, A.F. ; Graham, R.L.; Stokes, B.J.; Erbach, D.C. Biomass as Feedstock for a Bioenergy and Bioproduc ts Industry: the Technical Feasibility of a Billion-Ton Annual Supply . A joint U.S. Department of En ergy and U.S. Department of Agriculture report. DOE/GO-102995-2135 & ORNL/TM-2005/66. April 2005. http://www1.eere.energy.gov/biomass/ pdfs/final_billionton_vision_report2.pdf

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 Batte lle 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-S cale Gasification-Based Co-Production of Fuels and Electricity from Switchgrass.” Draft manuscript for Biomass & Bioenergy . March 2006.

5312 E.D. Larson, G.W. McDonald, W. Yang, W.J. Fr ederick, K. Iisa, T.G. Kreutz, E.W. Malcom, and C.A. Brown, “A Cost-Benefit Assessment of Black Liquor Gasifier/Combined Cycle Technology Integrated into a Kraft Pulp Mill,” Tappi Journal, 83(6): 57-, June 2000.

13 Bain, R. L., K. R. Craig and R. P. Overe nd (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 Fr ance, 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 Pa cific 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 Interna tional 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 Prosp ects for Production of Methanoland 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 Re port No. 292, Princeton University, Princeton, NJ, July.

21 Nexant Inc. Equipment Design and Cost Estimation fo r Small Modular Biomass Systems, Synthesis Gas Cleanup, and Oxygen Separation Equi pment, 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 fo r 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

5423 Nexant Inc. Equipment Design and Cost Estimation fo r 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 fo r Small Modular Biomass Systems, Synthesis Gas Cleanup, and Oxygen Separation E quipment, 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 , 5 th

Edition, McGraw-Hill, Inc., New York. 2003.

27 Perry, R.H., Green, D.W., Maloney, J.O. Perry’s Chemical Engineers’ Handbook , 7 th 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 T echnology 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 Labor atory, 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 fo r 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, 12 th ed., Tulsa,

OK. 2004

35 A Claus unit for reducing sulfur was briefly cons idered as an alternative to LO-CAT, however

the size of the stream and H 2 S concentration made this a fairly impractical choice.

5536 Nexant, Inc. Equipment Design and Cost Estimation fo r Small Modular Biomass Systems, Synthesis Gas Cleanup, and Oxygen Separation Equi pment. Task 9: Mixed Alcohols from Syngas – State of Technology. NREL/SR-510-39947. Performed by Nexa nt 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 – Classi cal 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 Fu els from Biomass; Technical Options, Process Analysis, and Optimisation, and Development Potential. Utrecht University, Copernicus Institute. Ma rch 2003. ISBN 90-393-3342-4.

43 Quarderer, G.J. “Mixed Alcohols from Synthesis Gas.” Proceedings from the 78 th Spring National AIChE Meeting. April 6-10, 1986, New Orleans, LA.

44 Nirula, S. ”Dow / Union Carbide Pro cess for Mixed Alcohols from Syngas” SRI

International. March 1986. PEP Re view No 85-1-4. Menlo Park, CA.

45 Park, T.; et al.; “Kinetic Analysis of Mixed Alcohol Synthesis from Syngas over K/MoS 2 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, 12 th ed., Tulsa,

OK. 2004

49 Weast, R.C., ed. CRC Handbook of Chemistry and Physics , 62 nd ed., CRC Press, Boca Raton,

FL. 1981

5650 Aden, et .al. “Lignocellulosic Biomass to Ethanol Process Design and Economics Utilizing Co-Current Dilute Acid Prehydrolysis and En zymatic 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/publicati ons.cfm?accountID=258&refID=89449

52 Peters, M.S.; K.D. Timmerhaus. Plant Design and Economics for Chemical Engineers , 5 th

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 Esti mation. 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 , 5 th Edition. Gulf Publishing Company. 1997

63 Walas, S.M. Chemical Process Equipm ent Selection and Design . Butterworth-Heinemann.

1988.

5764 Leiby, S.M. (1994) Operations for Refinery Hydrogen. SRI Report No. 212. Menlo Park, CA.

65 Jeakel, D. (2004). Price quote fro m 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_per mits_&_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/petr oleum/info_glance/petroleum.html

58

## Appendix A

List of AcronymsASME 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

CH 4 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)

CO 2 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

H

2 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

MoS 2 Molybdenum disulfide

## Appendix B

OBP Thermochemical Platform Research TargetsBackground/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 fo r the production of biofuels. Moving forward, the role of thermochemical c onversion 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 thermoche mical process can more easily convert lowcarbohydrate or “non-fermentabl e” biomass materials such as forest and wood residues to alcohol fuels, which adds technol ogy robustness to efforts to achieve the 30 x 30 goal (Foust, et al., 2006). This Appendix describes the R&D need ed to achieve the market target production price in 2012 for a stand-alone bi omass gasification/mixed alcohol process. Future advanced technology scenarios rely on considerable biof uel yield enhancements achieved by combining

biochemical and thermochemical conversion technologies into an integrated biorefinery that implements mixed alcohol producti on from gasification of lignin-ri ch bioconversion residues to maximize the liquid fuel yield per delivered ton of biomass.

Biomass gasification can convert a heterogeneous supply of bioma ss feedstock into a consistent gaseous intermediate that can then be reliably co nverted 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, H 2 , CO 2 , H 2 O, N 2 , and hydrocarbons. Minor components of the syngas in clude tars, sulfur and nitrogen oxides, alkali metals, and particulates. Thes e minor components of the syng as potentially threaten the successful application of downstr eam syngas conversion steps.

Commercially available and near -commercial syngas conversion processes were evaluated on technological, environmental, a nd economic bases (Spath and Da yton, 2003). This design report provides the basis for identifying promising, co st-effective fuel synthesis technologies that maximize the impact of biomass gasification fo r 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 MoS 2 catalyst, a variant of Fischer-Tropsch synt hesis, was selected as the conversion technology of choice because high yields of ethanol are possible w ith targeted R&D technology advancements. The

MoS 2 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% sele ctivity. 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 goal s 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 tec hnology scenario for doing this. The conceptual process design and ethanol produc tion cost estimate quantify the benefits of meeting the R&Dgoals for tar reforming and improved mixed al cohol 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 bi omass fuel properties that impact the long-term technical and

economic success of a thermoche mical conversion process: mois ture content, fixed carbon and volatiles content, impurity (sulfur, nitrogen, chlorine) concentrations, and ash content. High

moisture and ash contents reduce the usable fr action of delivered biom ass fuels proportionally. Therefore, maximum system effici encies 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

CO 2 removal

H 2 /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

CO 2 removal

H 2 /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 yiel d a raw syngas. The raw gas composition and quality are dependent on a wide range of factors includi ng 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 th e 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 polis hing to remove trace levels of remaining H 2 S and water-gas shift to adjust the final H 2 :CO ratio for optimized fuel synthesis.Comprehensive cleanup and conditioning of the raw biomass gasification product gas yields a “clean” syngas comprised of essentially CO and H 2, in a given ratio that can converted to a mixed alcohol product. Separation of ethanol from this product yi elds 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 tec hnical barrier areas

to meet the established 2012 t echnical target for thermochemic al ethanol production are outlined

in Table 1 . The rigorous engineering analysis of th e 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 integrat ed biomass gasification mixed alcohol synthesis. Process data collected in th e integrated pilot-scale testing will provide the basis for process optimization and cost estimates that will guide deployment of the technology.Table 1. Thermochemical Ethanol (Gasification/Mixed Alcohols) R&D Targets to meet the 2012 Cost-CompetitiveThermochemical Ethanol Cost TargetR&D Area Current 2007 2008 2009 2010 2011 2012FeedstockInterface$30/dry ton wood chips 50%moisture dried to 12% - 2000tpd plant$30/dry ton biorefineryresidues based on $45/dry toncorn stover. 50% moisturedried to 12% - 2000 tpd plantThermochemicalConversion -GasificationWood chips (model) - Indirect(atm) gasification – 78%syngas efficiency:H2/CO = 1.0-1.5CH4≤15vol%Tars≤30 g/Nm3benzene≤1vol%H2S = 50-600 ppmNH3and HCl to bedeterminedBiorefinery residues - Indirect(atm) gasification : cornstover; switchgrass; wheatstraw; lignin - 78% syngasefficiency:H2/CO = 1.0-1.5CH4≤15vol%Tars≤30 g/Nm3; benzene≤1vol%H2S = 50-600 ppmNH3and HCl to bedeterminedDemonstrate biomassgasification for$6.88/MMBtu syngas costbased on 2007Indirect (atm) gasification –78% syngas efficiency:H2/CO = 1.0-1.5CH4≤8vol%Tars≤10 g/Nm3; benzene≤0.1vol%; H2S≤20 ppm;NH3and HCl to bedeterminedDemonstrate biomassgasification for$5.25/MMBtu syngas costIndirect (atm) catalyticgasification – 78% syngasefficiency:H2/CO = 1.0CH4≤5vol%Tars≤1 g/Nm3; benzene≤0.04 vol%; H2S≤20 ppm;NH3and HCl to bedeterminedCleanup andConditioningCyclone particulate removalH2S≥50 ppm (based onfeedstock) with no S removalTar Reformer EfficiencyCH4≥20%Benzene≥70%heavy tars≥95%(79% CH4conversion inseparate SMR)Sorbent injection to maintainH2S levels≤50 ppm forsyngas from biomass toreduce sulfur deactivation oftar reforming catalysts.Tar Reformer EfficiencyCH4≥50%Benzene≥90%heavy tars≥97%(79% CH4conversion inseparate SMR)Improve tar reformingcatalyst performance -Regen/TOS ratio≤600Tar Reformer EfficiencyCH4≥80%Benzene≥99%heavy tars≥99.9%Eliminate SMR; highestactivity re-gained byregenerating deactivatedcatalystImprove tar reformingcatalyst performance -Regen/TOS ratio≤250Integrated operations forsyngas cleanup andconditioning targetcomposition for fuel;synthesis:CH4≤3vol%Benzene≤10 ppmHeavy tars≤0.1 g/Nm3H2S≤1 ppmNH3≤10 ppmHCl≤10 ppbCatalytic FuelsSynthesis(MixedAlcohols)H2/CO = 1.2Pressure≤2000 psiaProductivity = 100-400gMA/kg(cat)/hrEtOH Selectivity≥70% (CO2-free)H2/CO≤1.2Pressure≤2000 psiaProductivity≥150gMA/kg(cat)/hrEtOH Selectivity≥70% (CO2-free)Demonstrate 500 hourscatalyst lifetime at 2007performance with bottledsyngas for mixed alcoholcatalyst cost of≤$0.50/galEtOHH2/CO≤1.0Pressure≤1500 psiaProductivity≥300gMA/kg(cat)/hrEtOH Selectivity≥75%(CO2-free)Demonstrate 500 hourscatalyst lifetime at 2009performance. with biomasssyngas for mixed alcoholcatalyst cost of≤$0.22/galEtOHH2/CO≤1.0Pressure≤1000 psiaProductivity≥600gMA/kg(cat)/hrEtOH Selectivity≥80% (CO2-free)Demonstrate 1000 hourscatalyst lifetime at 2009performance. with biomasssyngasIntegration andModelingResearch state-of-technology- 56 gal/dry ton EtOH$2.02/gal minimum EtOHselling price (higher alcoholssold at 85% of market value)at $2.71/gal installed capitalcosts.Biomass Gasification/MixedAlcohol Design Report-Establishes a cost and qualitybaseline for technologyimprovements for $1.07/galthermochemical ethanol by2012 from indirect biomassgasification through a cleansyngas intermediate.Improved hydrocarbonconversion efficiencyyields- 56 gal/dry ton EtOH$1.73/gal minimum EtOHselling price (higheralcohols priced as gasolineon an energy adjusted basis- $1.15/gal) at $2.69/galinstalled capital costs.Validated $1.73/gal EtOHfor integrated Cleanup &Conditioning + MixedAlcohol synthesisDemonstrate feasibility ofsystem (8000 hr on streamwith≤10% catalyst lossesper year) based onregenerating fluidizable tarreforming catalyst toeliminate SMRValidated $1.35/gal EtOH forintegrated Cleanup &Condtioning + Mixed AlcoholsynthesisDemonstrate mixed alcoholyields of 89 gal/ton (76gal/dry ton EtOH) viaindirect biomassgasification at pilot-scalefor “$1.07” minimumETOH selling price (higheralcohols priced as gasolineon an energy adjusted basis- $1.15/gal). Total installedcapital costs are$2.31/annual gallon ofethanol.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-fermen table” 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 requirement s to yield biomass feedstocks at $35 per ton delivered to the thermochemical process. Additi onal challenges will be associated with feeding the delivered biomass into developing pressuri zed biomass gasification systems. In all cases, demonstrating biomass feed systems beyond the p ilot-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 feedstoc ks 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 tec hnology chosen for the basis of this analysis is the Battelle Columbus Laboratory indirectly heat ed gasifier. Other gasi fier 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, a nd 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 a nd 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 go al case conversions we re selected to yield an economically viable clean syngas that is suitable for use in a catalytic fuel synthesis process without further hydrocarbon conversion steps.Table 2. Tar Reformer Performance – % Conversion

Compound Current (2005) Goal (2012) Methane (CH 4 ) 20% 80% Ethane (C 2 H 6 ) 90% 99% Ethylene (C 2 H 4 ) 50% 90% Tars (C 10 +) 95% 99.9% Benzene (C 6 H 6 ) 70% 99% Ammonia (NH 3 ) 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 sepa rately 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 reform ing 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 deactiv ation 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 hydrocar bon 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). Fo r mixed alcohol synthesis to become aneconomical 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 catalyst s could potentially increase alcohol yields and selectivity of ethanol production from clean sy ngas and improve the overall economics of the process through better heat inte gration and control and fewer sy ngas recycling loops. Specific research targets to achieve the co st-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 hi gher 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 sy ngas pressure for alcohol synthesis and less unconverted syngas to recompress and recycl e 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 th e mixed alcohol synthesis reaction to improve the yields and the economics of the process.

Integration/Demonstration

As is the case for any sophisti cated conversion process, combin ing the individual unit operations into a complete, integrated systematic process is a significant challenge. Individual pilot-scale operations to demonstrate the re quired performance of the unit ope rations as well as complete integrated pilot development runs will be required to demonstrate the cost-competitive technology. A specific challenge will be to conti nue to demonstrate pro cess 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 deploy ment 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 effec tive R&D to achieve the 2012 technical target for thermochemical ethanol production.References

Aden, A., Spath,. P.L. (2005) Milestone Comple tion Report “The Potentia l 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.; Me lendez, M.; Hess, J.R.; Kenney, K.; Wang, M.; Snyder, S.; Werpy, T. A National Laboratory Market and Technol ogy Assessment of the 30X30 Scenario

. Draft NREL Report. December 2006.

Fierro, J. L. G. (1993). "Catalysis in C1 chemistr y: 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 Scr eening -- 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

## Appendix C

NREL Biorefinery D esign Database Description and SummaryN REL’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 informatio n 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

N umber Required: B

N umber 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

N otes:

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.EQUIPMENT_NUEQUIPMENT_NAME EQUIPMENT_CATEGOEQUIPMENT_TYPE EQUIPMENT_DESCRIPTION MATERIAL_CONSCOST_BASISC-101 Hopper Feeder CONVEYORVIBRATINGFEEDER Included in overall cost for feed handling & drying taken from several literature sources CS LITERATUREC-102 Screener Feeder Conveyor CONVEYOR BELT Included in overall cost for feed handling & drying taken from several literature sources CS LITERATUREC-103 Radial Stacker Conveyor CONVEYOR BELT Included in overall cost for feed handling & drying taken from several literature sources CS LITERATUREC-104 Dryer Feed Screw Conveyor CONVEYOR SCREW Included in overall cost for feed handling & drying taken from several literature sources CS LITERATUREC-105 Gasifier Feed Screw Conveyor CONVEYOR SCREW Included in overall cost for feed handling & drying taken from several literaturesources 316SS LITERATUREH-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 = 1,410ft^2; fixed tube sheet CS/INCL QUESTIMATEH-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; fixed TS CS/A214 QUESTIMATEH-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; fixed TS CS/316S ICARUSK-101 Flue Gas Blower FAN CENTRIFUGAL Included in overall cost for feed handling & drying taken from several literature sources SS304 LITERATUREM-101 Hydraulic Truck Dump with Scale SCALE TRUCK-SCALE Included in overall cost for feed handling & drying taken from several literature sources LITERATUREM-102 Hammermill SIZE-REDUCTION Included in overall cost for feed handling & drying taken from several literature sources CS LITERATUREM-103 Front End Loaders VEHICLE LOADER Included in overall cost for feed handling & drying taken from several literature sources CS LITERATUREM-104 Rotary Biomass Dryer DRYER ROTARY-DRUM Included in overall cost for feed handling & drying taken from several literature sources CS LITERATURES-101 Magnetic Head Pulley SEPARATOR MAGNET Included in overall cost for feed handling & drying taken from several literature sources CS LITERATURES-102 Screener SEPARATOR SCREEN Included in overall cost for feed handling & drying taken from several literature sources CS LITERATURES-103 Dryer Air Cyclone SEPARATOR GAS CYCLONE Included in overall cost for feed handling & drying taken from several literature sources CS LITERATURES-104 Dryer Air Baghouse Filter SEPARATOR FABRIC-FILTER Included in overall cost for feed handling & drying taken from several literature sources LITERATURET-101 Dump Hopper TANK LIVE-BTM-BIN Included in overall cost for feed handling & drying taken from several literature sources CS LITERATURET-102 Hammermill Surge Bin TANK LIVE-BTM-BIN Included in overall cost for feed handling & drying taken from several literature sources CS LITERATURET-103 Dryer Feed Bin TANK LIVE-BTM-BIN Included in overall cost for feed handling & drying taken from several literature sources CS LITERATURET-104 Dried Biomass Hopper TANK VERTICAL-VESSEL Included in overall cost for feed handling & drying taken from several literature sources CS LITERATUREC-201 Sand/ash Conditioner/Conveyor CONVEYOR SCREW Included in overall cost for gasification & gas clean up taken from several literature sources CS LITERATUREK-202 Combustion Air Blower FAN CENTRIFUGAL Included in overall cost for gasification & gas clean up taken from several literaturesources CS LITERATUREM-201 Sand/ash Cooler MISCELLANEOUS MISCELLANEOUS Included in overall cost for gasification & gas clean up taken from several literature sources LITERATURER-201 Indirectly-heated Biomass Gasifier REACTOR VERTICAL-VESSEL Included in overall cost for gasification & gas clean up taken from several literature sources CS w/refractory LITERATURER-202 Char Combustor REACTOR VERTICAL-VESSEL Included in overall cost for gasification & gas clean up taken from several literaturesources CS w/refractory LITERATURES-201 Primary Gasifier Cyclone SEPARATOR GAS CYCLONE Included in overall cost for gasification & gas clean up taken from several literature sources CS w/refractory LITERATURES-202 Secondary Gasifier Cyclone SEPARATOR GAS CYCLONE Included in overall cost for gasification & gas clean up taken from severalliterature sources CS w/refractory LITERATURES-203 Primary Combustor Cyclone SEPARATOR GAS CYCLONE Included in overall cost for gasification & gas clean up taken from several literature sources CS w/refractory LITERATURES-204 Secondary Combustor Cyclone SEPARATOR GAS CYCLONE Included in overall cost for gasification & gas clean up taken from severalliterature sources CS w/refractory LITERATURES-205 Electrostatic Precipitator SEPARATOR MISCELLANEOUS Included in overall cost for gasification & gas clean up taken from several literature sources CS LITERATURET-201 Sand/ash Bin TANKFLAT-BTMSTORAGE Included in overall cost for gasification & gas clean up taken from several literature sources CS LITERATUREPFD-P800-A101-2PFD-P800-A201EQUIPMENT_NUEQUIPMENT_NAME EQUIPMENT_CATEGOEQUIPMENT_TYPE EQUIPMENT_DESCRIPTION MATERIAL_CONSCOST_BASISH-301 Quench Water Recirculation Cooler HEATX SHELL-TUBE Included in overall cost for gasification & gas clean up taken from several literature sources CS LITERATUREH-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; fixed TS CS/316S ICARUSH-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; fixed TS CS/INCL QUESTIMATEH-301CReformedSyngascooler/SynthesisReactorPreheat#1 HEATX SHELL-TUBE duty =40.0 MMBtu/hr; LMTD = ?? F; U = 90 Btu/hr-ft^2-F; surface area = 1,552 ft^2 A214 QUESTIMATEH-302 Syngas Compressor Intercoolers HEATXAIR-COOLEDEXCHANGER Cost of intercoolers included in cost for syngas compressor, K-301 CS ICARUSH-303 Water-cooled Aftercooler HEATX SHELL-TUBE duty = 2.9 MMBtu/hr; LMTD = 25F; U = 150 Btu/hr-ft^2-F; surface area = 794 ft^2; fixed TS SS304CS/A214 QUESTIMATEH-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^2; fixed TS A285C/CA443 QUESTIMATEH-305 LO-CAT Absorbent Solution Cooler HEATX SHELL-TUBE Included in LO-CAT system cost 304SS VENDORH-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; fixed TS CS/INCL QUESTIMATEH-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^2; fixed TS A285C/CA443 QUESTIMATEK-301 Syngas Compressor COMPRESSOR CENTRIFUGALgas flow rate = 70,000 CFM; 6 impellers; design outlet pressure = 465 psi; 30,000 HP; intercoolers,aftercooler, & K.O.s included A285C QUESTIMATEK-302 LO-CAT Feed Air Blower FAN CENTRIFUGAL Included in LO-CAT system cost CS VENDORK-305 Regenerator Combustion Air Blower FAN CENTRIFUGAL gas flow rate (actual) = 70133 CFM; SS304 QUESTIMATEM-301 Syngas Quench Chamber MISCELLANEOUS Included in overall cost for gasification & gas clean up taken from several literature sources CS LITERATUREM-302 Syngas Venturi Scrubber MISCELLANEOUS Included in overall cost for gasification & gas clean up taken from several literaturesources CS LITERATUREM-303 LO-CAT Venturi Precontactor MISCELLANEOUS Included in LO-CAT system cost 304SS VENDORM-304 LO-CAT Liquid-filled Absorber COLUMN ABSORBER Included in LO-CAT system cost 304SS VENDORP-301 Sludge Pump PUMP CENTRIFUGAL 1.4 GPM; 0.053 brake HP; design pressure = 60 psia CS QUESTIMATEP-302 Quench Water Recirculation Pump PUMP CENTRIFUGAL Included in the cost of the gasification & gas clean up system CS LITERATUREP-303 LO-CAT Absorbent Solution Circulating Pump PUMP CENTRIFUGAL Included in LO-CAT system cost 304SS VENDORR-301A Tar Reformer Catalyst Regenerator REACTOR VERTICAL-VESSEL Taken from literature source CS w/refractory LITERATURER-303 Tar Reformer REACTOR VERTICAL-VESSEL Included in overall cost for gasification & gas clean up taken from several literature sources CS w/refractory LITERATURER-304 LO-CAT Oxidizer Vessel REACTOR VERTICAL-VESSEL Included in LO-CAT system cost 304SS VENDORS-301 Pre-compressor Knock-out SEPARATOR KNOCK-OUT DRUM 18 ft diameter; 36 ft height; design pres = 40 psia; design temp = 197 F CS QUESTIMATES-302 Syngas Compressor Interstage Knock-outs SEPARATOR KNOCK-OUT DRUM Cost of intercoolers K.O.s included in cost for syngas compressor, K-301 CS ICARUSS-303 Post-compressor Knock-out SEPARATOR KNOCK-OUT DRUM 7 ft. diameter; 14 ft height; design pres = 506 psia; design temp = 160 F CS QUESTIMATES-306 Tar Reformer Cyclone SEPARATOR GAS CYCLONE Included in the cost of the tar reformer catalyst renegerator, R-204 CS LITERATURES-307 Catalyst Regenerator Cyclone SEPARATOR GAS CYCLONE Included in the cost of the tar reformer catalyst renegerator, R-204 CS LITERATURES-310 L.P. Amine System COLUMN ABSORBEROTHERT-301 Sludge Settling Tank SEPARATOR CLARIFIER 3 ft diameter; 7 ft height; 431 gal volume; SS304 QUESTIMATET-302 Quench Water Recirculation Tank TANKHORIZONTALVESSEL Included in overall cost for gasification & gas clean up taken from several literature sources CS LITERATUREPFD-P800-A301-5EQUIPMENT_NUEQUIPMENT_NAME EQUIPMENT_CATEGOEQUIPMENT_TYPE EQUIPMENT_DESCRIPTION MATERIAL_CONSCOST_BASISH-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 = 98ft^2; fixed TS A285C/CA443 QUESTIMATEH-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^2; fixed TS A285C/CA443 QUESTIMATEH-411BPostsynthesiscooler#2/DeaeratorWaterPreheater#1 HEATX SHELL-TUBE duty = 20 MMBtu/hr; LMTD = 244 F; U = 100 Btu/hr-ft^2-F; area = 823 ft^2; fixed TS CS/A214 QUESTIMATEH-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; fixed TS CS/A214 QUESTIMATEH-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^2; fixed TS A285C/CA443 QUESTIMATEH-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; fixed TS CS/A214 QUESTIMATEH-412 Post Mixed Alcohol Cooler HEATX SHELL-TUBE duty = 76.5 MMBtu/hr; LMTD = ?? F; U = 100 Btu/hr-ft^2-F; surface area = 4,763 ft^2 A214 QUESTIMATEH-413 Mixed Alcohol first Condenser (air cooled) HEATXAIR-COOLEDEXCHANGERH-414 Mixed Alcohol Condenser HEATX SHELL-TUBE duty = 78 MMBtu/hr; LMTD = 41 F; U = 150 Btu/hr-ft^2-F; surface area = 12,462 ft^2 A214 QUESTIMATEH-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= 1,552 ft^2 A214 QUESTIMATEK-410 Mixed Alcohol Gas Compressor COMPRESSOR CENTRIFUGALgas flow rate = 2,481 CFM; 4 impellers; design outlet pressure = 700 psi; 10,617 HP; intercoolers,aftercooler, & K.O.s included A285C QUESTIMATEK-412 Purge Gas Expander COMPRESSOR CENTRIFUGAL gas flow rate = 144 CFM; design outlet pressure =25 psi; 2740 HP A285C QUESTIMATER-410 Mixed Alcohol Reactor REACTOR VERTICAL-VESSEL Fixed Bed Synthesis Reactor with MoS2-based catalyst. Sized from hourly spacevelocity of 3000 CS w/refractory QUESTIMATES-501 Mixed Alcohols Condensation Knock-out SEPARATOR KNOCK-OUT DRUM H/D = 2; 5 ft diam; 9 ft height; operating pressure = 1993 psia; operating temperature = 110 F A-515 QUESTIMATED-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; fixed TS CS/A214 QUESTIMATEH-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^2; fixed TS A285C/CA443 QUESTIMATEH-504C D-504 condenser (air cooled) HEATXAIR-COOLEDEXCHANGERH-504R Ethanol/Propanol Splitter Reboiler HEATX SHELL-TUBEH-505C D-505 condenser (air cooled) HEATXAIR-COOLEDEXCHANGERH-505R Methanol/Ethanol Splitter Reboiler HEATX SHELL-TUBEH-513 Mol Sieve Flush Condenser (air cooled) HEATXAIR-COOLEDEXCHANGERH-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 = 20ft^2; pre-engineered U-tube A285C/CA443 QUESTIMATEH-591 Higher Alcohol Product Finishing cooler HEATX SHELL-TUBE Included in overall cost for gasification & gas clean up taken fromseveral literature sources CS LITERATUREH-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-engineered U-tube A285C/CA443 QUESTIMATEH-593 ETHANOL Product Finishing cooler HEATX SHELL-TUBE Included in overall cost for gasification & gas clean up taken from severalliterature sources CS LITERATUREP-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 SS304 ICARUSS-503 Molecular Sieve (9 pieces) MISCELLANEOUS PACKAGE Superheater, twin mole sieve columns, product cooler, condenser, pumps, vacuum source. SS VENDORT-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 TANKFLAT-BTMSTORAGET-592 Ethanol Product Storage Tank TANKFLAT-BTMSTORAGEPFD-P800-A501-2PFD-P800-A401-2EQUIPMENT_NUEQUIPMENT_NAME EQUIPMENT_CATEGOEQUIPMENT_TYPE EQUIPMENT_DESCRIPTION MATERIAL_CONSCOST_BASISH-601 Steam Turbine Condenser HEATX SHELL-TUBEIncluded in the cost of the steam trubine/generator (M-602); condenser steam flow rate = 342,283lb/hrADEN, ET. AL.2002H-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-engineered U-tube A285C/CA443 QUESTIMATEH-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; fixedTS A214 QUESTIMATEM-601 Hot Process Water Softener System MISCELLANEOUS PACKAGE scaled cost to 700 gpm flow, 24" dia softener. Includes filters, chemical feeders, piping, valves RICHARDSONM-602 Extraction Steam Turbine/Generator GENERATOR STEAM-TURBINE 25.6 MW generated; 34,308 HPVENDORM-603 Startup Boiler MISCELLANEOUS PACKAGEAssume need steam requirement equal to 1/2 of steam requirement for gasifier at full rate steamrate = 36,560 lb/hrCS QUESTIMATEP-601 Collection Pump PUMP CENTRIFUGAL 513 GPM; 4 brake HP; outlet pressure = 25 psia CS QUESTIMATEP-602 Condensate Pump PUMP CENTRIFUGAL 190 GPM; 4 brake HP; outlet pressure = 25 psia SS304 QUESTIMATEP-603 Deaerator Feed Pump PUMP CENTRIFUGAL 702 GPM; 14 brake HP; outlet pressure = 40 psia CS QUESTIMATEP-604 Boiler Feed Water Pump PUMP CENTRIFUGAL 730 GPM; 759 brake HP; outlet pressure = 1,345 psia CS QUESTIMATES-601 Blowdown Flash Drum TANKHORIZONTALVESSEL H/D = 2; residence time = 5 min; 2 ft diameter; 4 ft height; op press = 1,280 psi; op temp = 575 F CS QUESTIMATET-601 Condensate Collection Tank TANKHORIZONTALVESSEL residence time = 10 minutes; H/D = 2; 8 ft diameter; 17 ft height CS QUESTIMATET-602 Condensate Surge Drum TANKHORIZONTALVESSEL residence time = 10 minutes; H/D = 2; 9 ft diameter; 17 ft height CS QUESTIMATET-603 Deaerator TANKHORIZONTALVESSEL liquid flow rate = 348,266 lb/hr; 150 psig design pressure; 10 min residence time CS;SS316 VENDORT-604 Steam Drum TANKHORIZONTALVESSEL 424 gal, 4.5' x 4'dia, 15 psig CS ICARUSK-701 Plant Air Compressor COMPRESSOR RECIPROCATING 450 cfm, 125 psig outletCS ICARUSM-701 Cooling Tower System COOLING-TOWER INDUCED-DRAFT approx 16,500 gpm, 140 MMBtu/hr FIBERGLASS DELTA-T98M-702 Hydraulic Truck Dump with Scale SCALE TRUCK-SCALE Hydraulic Truck Dumper with Scale CS VENDORM-703 Flue Gas Stack MISCELLANEOUS MISCELLANEOUS 42 inch diameter; 250 deg F A515 QUESTIMATEP-701 Cooling Water Pump PUMP CENTRIFUGAL 16,188 GPM; 659 brake HP; outlet pressure 75 psi CS QUESTIMATEP-702 Firewater Pump PUMP CENTRIFUGAL 2,500 gpm, 50 ft headCS ICARUSP-703 Diesel Pump PUMP CENTRIFUGAL 30 gpm, 150 ft headCS ICARUSP-704 Ammonia Pump PUMP CENTRIFUGAL 8.5 gpm, 22 ft headCS ICARUSP-705 Hydrazine Pump PUMP CENTRIFUGAL 5 gpm, 75 ft headCS DELTA-T98S-701 Instrument Air Dryer DRYER PACKAGE 400 SCFM Air Dryer, -40 F Dewpoint CS RICHARDSONT-701 Plant Air Receiver TANKHORIZONTALVESSEL 900 gal., 200 psigCS ICARUST-702 Firewater Storage Tank TANKFLAT-BTMSTORAGE 600,000 gal, 4 hr res time, 51' dia x 40' high, atmospheric A285C ICARUST-703 Diesel Storage Tank TANKFLAT-BTMSTORAGE 10,667 gal, 120 hr res time, 90% wv, 10' dia x 18.2' high, atmospheric A285C ICARUST-704 Ammonia Storage Tank TANKHORIZONTALSTORAGE Included in the cost of the feed handling step. A515 ICARUST-705 Olivine Lock Hopper TANK VERTICAL-VESSEL Included in the cost of the feed handling step. CS DELTA-T98T-706 MgO Lock Hopper TANK VERTICAL-VESSEL 20' x 20' Bin, Tapering to 3' x 3' at Bottom. Capacity 6,345 cf, two truck loads. CS DELTA-T98T-707 Hydrazine Storage Tank TANK VERTICAL-VESSEL 260 gal, 4.9' x 3'dia., 10psig SS316 ICARUSPFD-P800-A601-3PFD-P800-A701-2

## Appendix D

Individual Equipment Cost SummaryEquipmentNumberNumberRequiredNumberSpares Equipment Name Size RatioOriginal Equip Cost (perunit) Base YearTotal Original Equip Cost(Req'd & Spare) in BaseYearScalingExponentScaled Cost in BaseYear Installation FactorInstalled Cost in BaseYear Installed Cost in 2005$C-101 4 Hopper Feeder1.00$0 2002 $0 0.75 $0 2.47$0 $0C-102 2 Screener Feeder Conveyor1.00$0 2002 $0 0.75 $0 2.47$0 $0C-103 2 Radial Stacker Conveyor1.00$0 2002 $0 0.75 $0 2.47$0 $0C-104 2 Dryer Feed Screw Conveyor1.00$0 2002 $0 0.75 $0 2.47$0 $0C-105 2 Gasifier Feed Screw Conveyor0.93$0 2002 $0 0.75 $0 2.47$0 $0H-286B 1 Flue Gas Cooler/Steam Generator #10.60$347,989 2002 $347,989 0.6 $256,890 2.47$634,519 $750,965H-286C 1 Flue Gas Cooler /Boiler Water Preheater #10.03$20,989 2002 $20,989 0.6 $2,637 2.47$6,512 $7,708H-311B 1 Flue Gas Cooler / Steam Generator #30.69$69,089 2002 $69,089 0.65 $54,035 2.47$133,465 $157,959K-101 2 Flue Gas Blower1.56$0 2002 $0 0.75 $0 2.47$0 $0M-101 4 Hydraulic Truck Dump with Scale1.00$0 2002 $0 0.75 $0 2.47$0 $0M-102 2 Hammermill1.00$0 2002 $0 0.75 $0 2.47$0 $0M-103 3 Front End Loaders1.00$0 2002 $0 0.75 $0 2.47$0 $0M-104 2 Rotary Biomass Dryer1.00$3,813,728 2002 $7,627,455 0.75 $7,627,450 2.47$18,839,801 $22,297,257S-101 2 Magnetic Head Pulley1.00$0 2002 $0 0.75 $0 2.47$0 $0S-102 2 Screener1.00$0 2002 $0 0.75 $0 2.47$0 $0S-103 2 Dryer Air Cyclone1.56$0 2002 $0 0.75 $0 2.47$0 $0S-104 2 Dryer Air Baghouse Filter0.93$0 2002 $0 0.75 $0 2.47$0 $0T-101 4 Dump Hopper1.00$0 2002 $0 0.75 $0 2.47$0 $0T-102 1 Hammermill Surge Bin1.00$0 2002 $0 0.75 $0 2.47$0 $0T-103 2 Dryer Feed Bin1.00$0 2002 $0 0.75 $0 2.47$0 $0T-104 2 Dried Biomass Hopper0.93$0 2002 $0 0.75 $0 2.47$0 $0A100Subtotal$8,065,522 $7,941,0112.47$19,614,298 $23,213,888C-201 1 Sand/ash Conditioner/Conveyor0.33$0 2002 $0 0.65 $0 2.47$0 $0K-202 2 Combustion Air Blower0.97$0 2002 $0 0.65 $0 2.47$0 $0M-201 2 Sand/ash Cooler0.33$0 2002 $0 0.65 $0 2.47$0 $0R-201 2 Indirectly-heated Biomass Gasifier1.00$2,212,201 2002 $4,424,402 0.65 $4,418,389 2.47$10,913,421 $12,916,238R-202 2 Char Combustor1.00$0 2002 $0 0.65 $0 2.47$0 $0S-201 2 Primary Gasifier Cyclone1.00$0 2002 $0 0.65 $0 2.47$0 $0S-202 2 Secondary Gasifier Cyclone0.97$0 2002 $0 0.65 $0 2.47$0 $0S-203 2 Primary Combustor Cyclone1.00$0 2002 $0 0.65 $0 2.47$0 $0S-204 2 Secondary Combustor Cyclone0.95$0 2002 $0 0.65 $0 2.47$0 $0S-205 2 Electrostatic Precipitator0.96$0 2002 $0 0.65 $0 2.47$0 $0T-201 1 Sand/ash Bin0.33$0 2002 $0 0.65 $0 2.47$0 $0A200Subtotal$4,424,402 $4,418,3892.47$10,913,421 $12,916,238H-301 1 Quench Water Recirculation Cooler1.57$0 2002 $0 0.44 $0 2.47$0 $0H-301A 1 Post-tar Reformer Cooler / Steam Generator #21.66$69,089 2002 $69,089 0.65 $96,054 2.47$237,253 $280,793H-301B 1 Reformer Flue Gas Cooler/Steam superheater1.00$196,589 2002 $196,589 0.6 $196,056 2.47$484,259 $573,129H-301C 1 Reformed Syngas cooler / Synthesis Reactor Preheat #10.32$144,006 2002 $144,006 0.44 $87,219 2.47$215,431 $254,966H-302 5 Syngas Compressor Intercoolers1.57$0 2002 $0 0.65 $0 2.47$0 $0H-303 1 Water-cooled Aftercooler1.81$20,889 2002 $20,889 0.44 $27,111 2.47$66,965 $79,254H-304 1 LO-CAT Preheater0.15$4,743 2002 $4,743 0.6 $1,539 2.47$3,800 $4,498H-305 1 LO-CAT Absorbent Solution Cooler0.29$0 2002 $0 0.44 $0 2.47$0 $0H-315D1 1 Recycle Syngas Cooler / Steam Generator #44.20$26,143 2002 $26,143 0.6 $61,841 2.47$152,746 $180,778H-315D2 1 Recycle Syngas cooler #2 / Air preheat #14.08$4,743 2002 $4,743 0.6 $11,020 2.47$27,219 $32,214K-301 1 Syngas Compressor1.73$3,896,834 2002 $3,896,834 0.8 $6,036,915 2.47$14,911,181 $17,647,662K-302 1 LO-CAT Feed Air Blower0.73$0 2002 $0 0.65 $0 2.47$0 $0K-305 1 Regenerator Combustion Air Blower0.94$35,020 2002 $35,020 0.59 $33,806 2.47$83,500 $98,824M-301 1 Syngas Quench Chamber1.57$0 2002 $0 0.65 $0 2.47$0 $0M-302 1 Syngas Venturi Scrubber1.57$0 2002 $0 0.65 $0 2.47$0 $0M-303 1 LO-CAT Venturi Precontactor0.73$0 2002 $0 0.65 $0 2.47$0 $0M-304 1 LO-CAT Liquid-filled Absorber0.29$0 2002 $0 0.65 $0 2.47$0 $0P-301 1 1 Sludge Pump0.08$3,911 2002 $7,822 0.33 $3,351 2.47$8,277 $9,795P-302 1 1 Quench Water Recirculation Pump0.99$0 2002 $0 0.65 $0 2.47$0 $0P-303 1 1 LO-CAT Absorbent Solution Circulating Pump1.57$0 2002 $0 0.65 $0 2.47$0 $0R-301A 1 Tar Reformer Catalyst Regenerator1.62$2,429,379 2002 $2,429,379 0.65 $3,324,994 2.47$8,212,736 $9,719,926R-303 1 Tar Reformer1.57$2,212,201 2002 $2,212,201 0.65 $2,965,912 2.47$7,325,802 $8,670,224R-304 1 LO-CAT Oxidizer Vessel0.73$1,000,000 2002 $1,000,000 0.65 $813,486 2.47$2,009,311 $2,378,057S-301 1 Pre-compressor Knock-out1.73$157,277 2002 $157,277 0.6 $218,395 2.47$539,436 $638,432S-302 4 Syngas Compressor Interstage Knock-outs1.73$0 2002 $0 0.6 $0 2.47$0 $0S-303 1 Post-compressor Knock-out1.86$40,244 2002 $40,244 0.6 $58,421 2.47$144,300 $170,781S-306 1 Tar Reformer Cyclone1.57$0 2002 $0 0.65 $0 2.47$0 $0S-307 1 Catalyst Regenerator Cyclone1.62$0 2002 $0 0.65 $0 2.47$0 $0S-310 1 L.P. Amine System1.26$3,485,685 2002 $3,485,685 0.75 $4,155,524 2.47$10,264,143 $12,147,805T-301 1 Sludge Settling Tank0.00$11,677 2002 $11,677 0.6 $260 2.47$641 $759T-302 1 Quench Water Recirculation Tank1.57$0 2002 $0 0.65 $0 2.47$0 $0A300Subtotal$13,742,341 $18,091,9032.47$44,687,000 $52,887,900H-410B 1 Flue Gas Cool / syngas rxn preheat6.78$4,743 2002 $4,743 0.6 $14,951 2.47$36,928 $43,705H-411A 1 Air preheat #3 / post Reactor Syngas cooling #12.65$4,743 2002 $4,743 0.6 $8,505 2.47$21,007 $24,862H-411B 1 Post synthesis cooler #2/Deaerator Water Preheater #10.77$20,989 2002 $20,989 0.6 $17,979 2.47$44,409 $52,558H-411C 1 Post Synthesis cooler #3/Makeup Water heater0.33$20,989 2002 $20,989 0.6 $10,757 2.47$26,569 $31,445H-411D 1 Post Synthesis cooler #4 / syngas recycle heat #15.20$4,743 2002 $4,743 0.6 $12,758 2.47$31,512 $37,295H-411E 1 Post Synthesis Cooler #5/Mol Sieve preheater #20.04$20,989 2002 $20,989 0.6 $2,991 2.47$7,389 $8,745H-412 1 Post Mixed Alcohol Cooler0.61$90,820 2002 $90,820 0.44 $73,284 2.47$181,011 $214,230H-413 1 Mixed Alcohol first Condenser (air cooled)0.76$42,255 1990 $42,255 1 $42,255 2.47$104,369 $136,649H-414 1 Mixed Alcohol Condenser0.06$338,016 2002 $338,016 0.44 $96,403 2.47$238,116 $281,814H-416B 1 Recycle Syngas Heat #2 / Flue gas Cool0.63$144,006 2002 $144,006 0.44 $117,700 2.47$290,718 $344,070K-410 1 Mixed Alcohol Gas Compressor0.89$851,523 2002 $851,523 0.8 $773,871 2.47$1,911,462 $2,262,251K-412 1 Purge Gas Expander12.01$642,014 2002 $642,014 0.8 $4,689,955 2.47$11,584,188 $13,710,103R-410 1 Mixed Alcohol Reactor0.34$2,026,515 2002 $2,026,515 0.56 $1,101,031 2.47$2,719,545 $3,218,633S-501 1 Mixed Alcohols Condensation Knock-out2.01$55,447 2002 $55,447 0.6 $84,229 2.47$208,045 $246,225A400Subtotal$4,267,792 $7,046,6672.47$17,405,267 $20,612,585D-504 1 Ethanol/Propanol Splitter0.50$478,100 1998 $478,100 1.32 $189,541 2.1$398,035 $478,460D-505 1 Methanol/Ethanol Splitter0.54$478,100 1998 $478,100 1.32 $210,444 2.1$441,933 $531,228H-503A 1 Syngas Cooler #4 / Mol Sieve preheater #11.08$20,989 2002 $20,989 0.6 $21,965 2.47$54,255 $64,211H-503B 1 Mol Sieve Superheater / reformed syngas cool #50.66$4,743 2002 $4,743 0.6 $3,709 2.47$9,161 $10,842H-504C 1 D-504 condenser (air cooled)0.68$39,408 1990 $39,408 1 $39,408 2.47$97,338 $127,442H-504R 1 Ethanol/Propanol Splitter Reboiler0.18$158,374 1996 $158,374 0.68 $49,237 2.1$103,398 $126,830H-505C 1 D-505 condenser (air cooled)0.64$62,938 1990 $62,938 1 $62,938 2.47$155,458 $203,538H-505R 1 Methanol/Ethanol Splitter Reboiler0.26$158,374 1996 $158,374 0.68 $63,019 2.1$132,340 $162,330H-513 1 Mol Sieve Flush Condenser (air cooled)0.19$21,181 1990 $21,181 1 $21,181 2.47$52,318 $68,499H-590 1 MA Product Cooler / Mol Sieve preheater #30.34$3,043 2002 $3,043 0.6 $1,590 2.47$3,928 $4,649H-591 2 Ethanol Product Pump0.21$7,501 1998 $15,002 1.79 $957 3.47$3,321 $3,992H-592 1 ETHANOL Product Cooler / Mol Sieve preheater #41.51$3,043 2002 $3,043 0.6 $3,903 2.47$9,639 $11,408H-593 3 Ethanol Product Pump1.19$7,502 1999 $22,506 2.79 $36,692 4.47$164,012 $196,596P-504B 1 1 Ethanol/Propanol Splitter Bottoms Pump0.04$42,300 1997 $84,600 0.79 $7,120 2.8$19,937 $24,151P-504R 1 1 Ethanol/Propanol Splitter Reflux Pump0.37$1,357 1998 $2,714 0.79 $1,240 2.8$3,471 $4,172P-505B 1 1 Methanol/Ethanol Splitter Bottoms Pump0.08$42,300 1997 $84,600 0.79 $11,343 2.8$31,761 $38,475P-505R 1 1 Methanol/Ethanol Splitter Reflux Pump0.37$1,357 1998 $2,714 0.79 $1,240 2.8$3,471 $4,172P-590 1 Mixed Alcohol Product Pump0.21$7,500 1997 $7,500 0.79 $2,226 2.47$5,499 $6,661P-592 1 Ethanol Product Pump1.19$7,500 1997 $7,500 0.79 $8,614 2.47$21,276 $25,773S-502 1 LP Syngas Separator0.49$55,447 2002 $55,447 0.6 $36,305 2.47$89,673 $106,130S-503 1 Molecular Sieve (9 pieces)1.43$904,695 1998 $904,695 0.7 $1,160,495 2.47$2,866,422 $3,445,594T-504 1 Ethanol/Propanol Splitter Reflux Drum0.37$11,900 1997 $11,900 0.93 $4,731 2.1$9,934 $12,034T-505 1 Methanol/Ethanol Splitter Reflux Drum0.37$11,900 1997 $11,900 0.93 $4,731 2.1$9,934 $12,034T-590 2 Mixed Alcohol Product Storage Tank0.21$165,800 1997 $331,600 0.51 $151,390 2.47$373,933 $452,976T-592 2 Ethanol Product Storage Tank1.19$165,800 1997 $331,600 0.51 $362,599 2.47$895,620 $1,084,940A500Subtotal$3,302,572 $2,456,6172.424499069$5,956,067 $7,207,140H-601 1 Steam Turbine Condenser0.87$0 2002 $0 0.71 $0 2.47$0 $0H-602 1 Blowdown Cooler / Deaerator Water Preheater0.00$3,043 2002 $3,043 0.6 $0 2.47$0 $0H-603 1 Blowdown Water-cooled Cooler4.90$16,143 2002 $16,143 0.44 $32,485 2.47$80,237 $94,962M-601 1 Hot Process Water Softener System1.00$1,031,023 1999 $1,031,023 0.82 $1,028,430 2.47$2,540,222 $3,044,885M-602 1 Extraction Steam Turbine/Generator1.00$4,045,870 2002 $4,045,870 0.71 $4,037,059 2.47$9,971,535 $11,801,498M-603 1 Startup Boiler1.00$198,351 2002 $198,351 0.6 $198,351 2.47$489,927 $579,837P-601 1 1 Collection Pump0.13$7,015 2002 $14,030 0.33 $7,226 2.47$17,847 $21,122P-602 1 1 Condensate Pump0.87$5,437 2002 $10,874 0.33 $10,366 2.47$25,605 $30,304P-603 1 1 Deaerator Feed Pump1.00$8,679 2002 $17,358 0.33 $17,340 2.47$42,831 $50,691P-604 1 1 Boiler Feed Water Pump1.00$95,660 2002 $191,320 0.33 $191,126 2.47$472,082 $558,718S-601 1 Blowdown Flash Drum1.00$14,977 2002 $14,977 0.6 $14,950 2.47$36,926 $43,703T-601 1 Condensate Collection Tank1.00$24,493 2002 $24,493 0.6 $24,448 2.47$60,386 $71,468T-602 1 Condensate Surge Drum1.00$28,572 2002 $28,572 0.6 $28,519 2.47$70,443 $83,371T-603 1 Deaerator1.00$130,721 2002 $130,721 0.72 $130,432 2.47$322,168 $381,292T-604 1 Steam Drum1.00$9,200 1997 $9,200 0.72 $9,180 2.47$22,674 $27,467A600Subtotal$5,735,975 $5,729,9122.47$14,152,883 $16,789,318K-701 2 1 Plant Air Compressor1.00$32,376 2002 $97,129 0.34 $97,129 2.47$239,908 $283,936M-701 1 Cooling Tower System0.66$267,316 2002 $267,316 0.78 $193,008 2.47$476,730 $564,218M-702 1 Hydraulic Truck Dump with Scale1.00$80,000 1998 $80,000 0.6 $80,000 2.47$197,600 $237,526M-703 1 Flue Gas Stack0.31$51,581 2002 $51,581 1 $15,917 2.47$39,315 $46,530P-701 1 1 Cooling Water Pump0.66$158,540 2002 $317,080 0.33 $276,264 2.47$682,373 $807,601P-702 1 1 Firewater Pump1.00$18,400 1997 $36,800 0.79 $36,800 2.47$90,896 $110,110P-703 1 1 Diesel Pump1.00$6,100 1997 $12,200 0.79 $12,200 2.47$30,134 $36,504P-704 1 1 Ammonia Pump1.00$5,000 1997 $10,000 0.79 $10,000 2.47$24,700 $29,921P-705 1 Hydrazine Pump1.00$5,500 1997 $5,500 0.79 $5,500 2.47$13,585 $16,457S-701 1 1 Instrument Air Dryer1.00$8,349 2002 $16,698 0.6 $16,698 2.47$41,244 $48,813T-701 1 Plant Air Receiver1.00$7,003 2002 $7,003 0.72 $7,003 2.47$17,297 $20,472T-702 1 Firewater Storage Tank1.00$166,100 1997 $166,100 0.51 $166,100 2.47$410,267 $496,991T-703 1 Diesel Storage Tank1.00$14,400 1997 $14,400 0.51 $14,400 2.47$35,568 $43,086T-704 1 Ammonia Storage Tank1.00$287,300 1997 $287,300 0.72 $287,300 2.47$709,631 $859,635T-705 1 Olivine Lock Hopper1.00$0 1998 $0 0.71 $0 2.47$0 $0T-706 1 MgO Lock Hopper1.00$0 1998 $0 0.71 $0 2.47$0 $0T-707 1 Hydrazine Storage Tank1.00$12,400 1997 $12,400 0.93 $12,400 2.47$30,628 $37,102A700Subtotal$1,381,507 $1,230,7192.47$3,039,875 $3,601,800Equipment Cost $40,920,110 $46,915,2182.467617439$115,768,810 $137,228,869

## Appendix F

Discounted Cash Flow Rate of Return SummaryDCFROR WorksheetYear-2 -1 0 1 2 3 4 5 6 7 8 9Fixed Capital Investment $18,603,984 $114,503,967 $61,068,782Working Capital$9,541,997Loan Payment$0 $0 $0 $0 $0 $0 $0 $0 $0Loan Interest Payment $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0Loan Principal $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0Ethanol Sales $46,938,097 $62,584,129 $62,584,129 $62,584,129 $62,584,129 $62,584,129 $62,584,129 $62,584,129 $62,584,129By-Product Credit $9,633,280 $12,844,373 $12,844,373 $12,844,373 $12,844,373 $12,844,373 $12,844,373 $12,844,373 $12,844,373Total Annual Sales$56,571,377 $75,428,503 $75,428,503 $75,428,503 $75,428,503 $75,428,503 $75,428,503 $75,428,503 $75,428,503Annual Manufacturing CostRaw Materials$23,647,650 $27,025,885 $27,025,885 $27,025,885 $27,025,885 $27,025,885 $27,025,885 $27,025,885 $27,025,885Tar reforming catalysts$808,613Steam reforming catalysts$0$0ZnO$0$0Mixed Alcohol catalysts$542,966 $0 $0 $0 $0 $542,966 $0 $0 $0Baghouse Bags$415,430$415,430Other Variable Costs$1,531,320 $1,732,166 $1,732,166 $1,732,166 $1,732,166 $1,732,166 $1,732,166 $1,732,166 $1,732,166Fixed Operating Costs$12,059,682 $12,059,682 $12,059,682 $12,059,682 $12,059,682 $12,059,682 $12,059,682 $12,059,682 $12,059,682Total Product Cost$39,005,661 $40,817,733 $40,817,733 $40,817,733 $40,817,733 $41,776,129 $40,817,733 $40,817,733 $40,817,733Annual DepreciationGeneral PlantDDB$38,064,676 $27,189,054 $19,420,753 $13,871,967 $9,908,548 $7,077,534 $5,055,381SL$19,032,338 $15,860,282 $13,594,527 $12,137,971 $11,559,972 $11,559,972 $11,559,972Remaining Value$95,161,690 $67,972,636 $48,551,883 $34,679,916 $24,771,369 $17,693,835 $12,638,453Actual$38,064,676 $27,189,054 $19,420,753 $13,871,967 $11,559,972 $11,559,972 $11,559,972Steam PlantDDB$4,321,018 $3,996,942 $3,697,171 $3,419,883 $3,163,392 $2,926,138 $2,706,677 $2,503,677 $2,315,901SL$2,880,679 $2,804,872 $2,738,645 $2,682,262 $2,636,160 $2,601,011 $2,577,788 $2,567,873 $2,567,873Remaining Value$53,292,560 $49,295,618 $45,598,446 $42,178,563 $39,015,171 $36,089,033 $33,382,355 $30,878,679 $28,562,778Actual$4,321,018 $3,996,942 $3,697,171 $3,419,883 $3,163,392 $2,926,138 $2,706,677 $2,567,873 $2,567,873Net Revenue($24,819,978) $3,424,773 $11,492,845 $17,318,919 $19,887,405 $19,166,264 $20,344,120 $32,042,896 $32,042,896Losses Forward($24,819,978) ($21,395,205) ($9,902,360) $0 $0 $0 $0 $0Taxable Income($24,819,978) ($21,395,205) ($9,902,360) $7,416,559 $19,887,405 $19,166,264 $20,344,120 $32,042,896 $32,042,896Income Tax$0 $0 $0 $2,892,458 $7,756,088 $7,474,843 $7,934,207 $12,496,729 $12,496,729Annual Cash Income$17,565,716 $34,610,769 $34,610,769 $31,718,311 $26,854,681 $26,177,531 $26,676,563 $22,114,040 $22,114,040Discount Factor 1.21 1.1 1 0.909090909 0.826446281 0.751314801 0.683013455 0.620921323 0.56447393 0.513158118 0.46650738 0.424097618Annual Present Value $217,657,607$15,968,833 $28,603,942 $26,003,583 $21,664,033 $16,674,644 $14,776,534 $13,689,295 $10,316,363 $9,378,512Total Capital Investment + Interest $22,510,821.23 $125,954,363.41 $70,610,779.49Net Present Worth$0DCFROR WorksheetYearFixed Capital InvestmentWorking CapitalLoan PaymentLoan Interest PaymentLoan PrincipalEthanol SalesBy-Product CreditTotal Annual SalesAnnual Manufacturing CostRaw MaterialsTar reforming catalystsSteam reforming catalystsZnOMixed Alcohol catalystsBaghouse BagsOther Variable CostsFixed Operating CostsTotal Product CostAnnual DepreciationGeneral PlantDDBSLRemaining ValueActualSteam PlantDDBSLRemaining ValueActualNet RevenueLosses ForwardTaxable IncomeIncome TaxAnnual Cash IncomeDiscount FactorAnnual Present ValueTotal Capital Investment + InterestNet Present Worth10 11 12 13 14 15 16 17 18 19 20($9,541,997)$0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0$0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0$0 $0 $0 $0 $0 $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 $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 $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 $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 $27,025,885 $27,025,885 $27,025,885$0$0$0$0$0 $542,966 $0 $0 $0 $0 $542,966 $0 $0 $0 $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 $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 $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 $40,817,733 $40,817,733 $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 $1,148,142 $1,062,031 $982,379$2,567,873 $2,567,873 $2,567,873 $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 $14,160,417 $13,098,386 $12,116,007$2,567,873 $2,567,873 $2,567,873 $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 $32,042,896 $32,042,896 $32,042,896$0 $0 $0 $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 $32,042,896 $32,042,896 $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 $12,496,729 $12,496,729 $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 $22,114,040 $22,114,040 $22,114,0400.385543289 0.350493899 0.318630818 0.28966438 0.263331254 0.239392049 0.217629136 0.197844669 0.17985879 0.163507991 0.148643628$8,525,920 $7,545,930 $7,046,215 $6,405,650 $5,823,318 $5,293,925 $4,685,429 $4,375,145 $3,977,404 $3,615,822 $3,287,111($1,418,357.09)Minimum Ethanol Selling Price ($/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

## Appendix G

Process Parameters & Operation SummaryEnergy Efficiencies Tar Reformer Alcohol Synthesis Alcohol SynthesisGasifier Efficiency - HHV % 76.6 Inlet Molar Flow (MMscf/hr) 6.70Syngas from Conditioning 279,888 Relative Alcohol Distribution After ReactorGasifier Efficiency - LHV % 76.1Space Velocity (hr-1)2,476 Recycled from initial flash tank 215 Methanol 8.5%Overall Plant Efficiency - HHV % 47.4 Reactor Volume (ft3) 2,705 Recycled from MolSieve Flush 5,026 Ethanol 81.7%Overall Plant Efficiency - LHV % 45.8 Inlet: Total 285,128 Propanol 8.8%Temperature (°F) 1,086 Butanol 0.9%DryerCarbon as CO (mol%) 49.0% Pentanol + 0.1%Inlet: Carbon as tar (ppmv) 5,758 Conditioned Syngas H2:CO Ratio 1.00Temperature (°F) 60.0 H2:CO Ratio (mole) 0.87 Recycled Gas H2:CO Ratio 1.06 Flash SeparatorMoisture Content (wt%) 50.0%Temperature (°F) 110Outlet: Reformer Conversions:At Reactor InletPressure (psia) 970Temperature (°F) 219.7 CO2 --> CO 32.5% Temperature (°F) 570Moisture Content (wt%) 5.0% Methane --> CO 80.0% Pressure (psia) 991 Relative Alcohol Distribution After Flash TankInlet Flue Gas (°F) 1,206 Ethane --> CO 99.0% H2:CO Molar Ratio 1.00 Methanol 8.3%Outlet Flue Gas (°F) 235.9 Ethylene --> CO 90.0% CO2 (mol %) 5.0% Ethanol 81.7%Dew Point Flue Gas (°F) 173.9 Benzene --> CO 99.0% Methane (mol%) 1.5% Propanol 8.9%Difference 62.0 Tar --> CO 99.9% H2O (wt%) 0.64% Butanol 0.9%Ammonia --> CO 90.0% Pentanol + 0.1%GasifierInlet Molar Flow (MMscf/hr) 6.5Temperature (°F) 1,633 Outlet:Space Velocity (hr-1)4,000 Vapor Losses From Flash TankPressure (psia) 23.0 Temperature (°F) 1,633 Reactor Volume (ft3) 1,616 Methanol 3.5%H2:CO Molar Ratio After Gasifier 0.60 Carbon as CO (mol%) 75.7% Ethanol 1.8%Methane (vol%) 9.0% Carbon as tar (ppmv) 43 CO Conversion - Overall 59.5% Propanol 0.8%Benzene (vol%) 0.07% H2:CO Ratio (mole)1.00CO Conversion - Singlepass 59.4% Butanol 0.3%Tar (wt%) 0.91% Methane (vol%) 1.2% Conversion To: Pentanol + 0.1%Tar (g/Nm3) 9.5 Benzene (ppmv) 2.7 CO2 21.9%Char (wt%) 12.7% Tars (ppmv) 3 Methane 3.4% Cleaned Gas Recycled to Reactor 0.1%H2S (ppm) 413 Tars (g/Nm3) 0.01 Ethane 0.3%Residual Heat (MBtu/hr) 0 H2S (ppm) 205 Methanol 0.3%Residual SyngasNH3 (ppm) 80 Ethanol 28.2% Recycled to synthesis reactors (lb/hr) 215Char CombustorPropanol 4.6% To Tar Reformers (lb/hr) 214,787Temperature (°F) 1,823QuenchButanol 0.6% To Fuel System (lb/hr) 10,739Pressure (psia) 21.4 Benzene (ppmv) 3.1 Pentanol + 0.1% To Reformer for Process (lb/hr) 204,048Ratio Actual:Minimum air for combustion 1.20 Tars (ppmv) 4 Total 59.4%Residual Heat (MBtu/hr) 0.0 Tars (g/Nm3) 0.01Overall Water DemandH2S (ppm) 235 Selectivity (CO2 Free) gal/gal etoh 1.94Syngas UsageNH3 (ppm) 83 Alcohols 90.1% gal/gal total alcohols 1.65To Reformer (lb/hr) 168,120 Hydrocarbons 9.9%To Char Combustor (lb/hr) 0Acid Gas RemovalTo Fuel System (lb/hr) 64,703 Inlet:At Reactor OutletCO2 (mol/hr) 2,041 Temperature (°F) 570Fuel SystemCO2 (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) 846Into 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) 0.488H2S (mol/hr) 1H2S (ppmv) 49Fraction H2S removed 99.6%

## Appendix H

Process Flow Diagrams (PFDs)

## Appendix I

Syngas and Char CorrelationsThe 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 expe rimental 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 pr oduction 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 elem ental 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 wood a CO 133.46 -0.1029 0.000028792 mol% dry gas

CO 2 -9.5251 0.037889 -0.000014927 mol% dry gas

CH 4 -13.82 0.044179 -0.000016167 mol% dry gas

C 2 H 4 -38.258 0.058435 -0.000019868 mol% dry gas

C 2 H 6 11.114 -0.011667 0.000003064 mol% dry gas

H 2 17.996 -0.026448 0.00001893 mol% dry gas

C 2 H 2 -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. Residua l 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 biom ass 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 H 2 S in the syngas.

a Scf = standard cubic feet. The standard conditons are 1 atm pressure and 60°F temperature.• A set amount of nitrogen is parsed to the char (6.6%). All remaining nitrogen is set as

NH 3 in the syngas.

• The amount of hydrogen in the syngas (including tar, H 2 S, NH 3 , 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 th ere is no net heat for an adiabatic reaction.

The syngas amount and composition will be de pendent upon the biomass composition and the gasifier temperature. As an example, the re sulting 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

H2O Gasifier Temperature (1,633 °F)

Figure 2. Syngas Composition for Woody Biomass Used in Design Report

References

Bain, R. (January 14, 1992). Material and Energy Balances fo r 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-Ric h Gas in a High-Throughput Gasifier

.

## Appendix J

Alcohol Synthesis Catalyst ReferencesA literature search was conducted to review ex isting 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 descri ptions 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 de rivation. Common to all of these catalysts is the addition of alkali metals which shift the pr oduct slate towards alcohol producti on. 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 (m odified methanation).

• Other, which includes alternate catalysts, such as Rhodium based catalysts, which are not specifically used for mixed al cohols but have been develope d 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 sulfidecatalysts 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 pr oduce mixtures of metha nol 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 certa in reaction conditions were met (temperature, pressure) a mixture of methanol and higher alc ohols resulted. At the same time, Fischer and Tropsch observed that hydrocarbon sy nthesis catalysts produced linear alcohols as byproducts. From this they were able to de velop the “Synthol” process for producing higher alcohols. Some development c ontinued, but it wasn’t until the 1970s oil embargo that significant interest re-appeared, and resear chers renewed efforts to produce higher alcohols for liquid fuels applications. As petroleum pr ices 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, me thyl-tertiary-butyl ether (MTBE) became the oxygenate of choice because refi ners 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 tofind alternatives for petroleum based fuels, a nd 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 mix ed alcohol production can range from below 150 to near 400 g/kg/hr. In comparison, me thanol synthesis produc tivities can often be over 1000 g/kg/hr, or 1 kg/kg/hr. Less commonl y, 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: (CO initial – CO final ) / CO initial Sometimes researchers present conversions exclusive of CO 2 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 pr oduct. 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%. Se lectivities are sometimes

shown on a CO 2 -exclusive basis.

Yield – see Productivity.

The overall stoichiometric reaction for higher alcohol synthesis (HAS) can be summarized as:



## ( )

2 21 2 CO 2 H C H OH 1 H O

n n nn n + + → +−

The value of “ n ” typically ranges from 1 to 8. The stoichiometry suggests an optimum H 2 /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 (CO 2 ). The overall reaction is exothermic; therefore, maintaining constant reaction temperature is an important desi gn 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 sh ift the product slate from methanol to higherbranched 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, ke tones, and ethers are also form ed, along with large amounts of

CO 2

.

Typical high pressure and low pressure reaction conditions (as provided by Nexant) are listed in Table 4. In general there is a trade-off betw een maximizing CO conversion and maximizing the higher alcohol selectivity and yield.

Table 4. “Typical” Modified Methanol Cataly st Conditions

H 2 /CO

ratio Temperature (°F) Pressure (psia) CO

conversion (per pass) Total

Alcohol

Yield (g/kg/hr) C 2 +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 S uperiori - methanol pl us 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 devel oper still involved in this area is the Standard Alcohol Company of America. They have a bench-scale pro cess to produce a mixed alcohols product known as

Envirolene TM . Envirolene is composed of methanol through octanol, with approximately 50% of the product as ethanol. The process uses a modifi ed 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. Ca rbon chain growth schemes have been developed that describe the product distri bution relatively accurately. This is shown in Figure 3.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 additi on of the alkali promoter helps to shift the product slate from hydrocarbons to linear alcohols, a lthough hydrocarbons remain a significant byproduct. Typical reaction conditions ar e 220 – 350°C (430 – 660°F) and 5-20 MPa (725 – 2,900 psia). One commonly-researched catalyst system is a MoS 2 -based system that is alkali and/or Cobalt-promoted. This has the tendency to increase etha nol and other higher alcohols selectivity. CO 2 is still a substantial byproduct due to water-gas shift (WGS) activity of the catalysts. Other potential byprodu cts 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 U CC jointly developed a sulfided mixed alcohol catalyst based on molybdenum (MoS 2 ). Sulfided catalysts have the advantage of being sulfurtolerant (up to 100 ppm) which has the potential to reduce upstream cleanup costs. IFP, inconjunction with Idemitsu Kosan (Japan), de veloped a process base d on Cu-Co and Cu-Ni catalyst systems. Dow built a 2-ton-per-day (T PD) 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 th is area are Power Energy Fuels Inc. (PEFI), Western Research Institute (WRI), and Pearson Technologies. PEFI c ontinues to develop the

Ecalene TM technology and process, which is a modi fication of Dow’s Sygmal process using polysulfite catalyst. According to Nexant, prog ress has not moved beyond the bench scale and a planned 500 gallon/day pilot plant is no longer be ing pursued. However, 2-3 other pilot plants are under funding consideration using various bioma ss resources. WRI had worked with PEFI in the past, however currently they are not, but ar e conducting their own benc h-scale experiments, particularly reactor and catalyst testing. Pearson Technologie s has developed a 30-ton-per-day biomass gasification and alcohol s conversion facility in Aber deen, 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 usin g the Pearson technology on the island of Kauai. Sasol (South Africa) is a worl d leader in FT fuels and chemicals production as well as technology development. They curren tly 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 alc ohols, acetic acid, and ketones.

Due to the severe process conditions of higher alcohol synthesis, it w ould be expected that catalyst life would not be significantly long b (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 perfor mance 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.Figure 4. Kinetic Analysis of HAS with Fischer-Tropsch Mechanism

Other Catalysts

Some research has been conducte d 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 superc ritical conditions. Rh-based catalysts have been primarily developed for selective ethanol sy nthesis 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 potentia l 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 thermochem ical production of ethanol, a moly-sulfide-based (MoS 2 ) 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.Doc # Title DateAuthor(s) Company/University Affiliation Journal / Conference Synopsis9542Synthesis of Higher Alcohols fromSyngas over Cu-Co2O3/ZnO, Al2O3Catalyst2005Mahdavi, V.; M.H. Peyrovi,et.al.Shahid Beheshti University, Tehran Iran;National Iranian Oil Com pany,Petroleum Research Institute Applied Catalysis, ElsevierFixed bed flow reactor; changes inactivity and selectivity resultingfrom feed and operatingconditions; alcohols selectivitygreater than 80%Mixed Alcohols from Syngas: State ofTechnology Final Report2005Nexant Nexant Subcontract to NREL/DOEState of technology of m ixedalcohols and recommendations forareas of focus for researchGridley Ethanol Dem onstration ProjectUtilizing Biomass GasificationTechnology: Pilot Plant Gasifier andSyngas Conversion Testing2005TSS Consultants City of Gridley, CANREL Subcontract Report SR510-37581Pearson Technologies PilotGasifier; Rice Straw into Ethanolthermochem ically; use PSA toremove CO2 from system; Notm uch data, this is a publiclyavailable report, supposedly thereis a more detailed proprietaryreportN/AProduction of Mixed Alcohol Fuels2004Lucero, Andrew J. W estern Research Institute DOE workshop on BLGUpdate on efforts to use PEFI'smodified MoS2 catalyst;com pression costs dominate opcosts; temp selection is critical;target alcohol com position given;Single-pass CO conversions up to20% to alcohols11631Solvent Effects on Higher AlcoholsSynthesis Under SupercriticalConditions: a Therm odynam icConsideration2004Qin, Zhangfeng; et.al.Institute of Coal Chemistry, ChineseAcademy of SciencesFuel Processing Technology,ElsevierSupercritical HAS, used redlichkwong soave; used varioussolvents; low ethanol; nice chart ofequilibrium constants, standardethalpies, entropies, and gibbsfree energies9543Higher-Alcohols Biorefinery:Improvement of Catalyst for EthanolConversion2004Olson, Edwin S.; et.al.EERC, North DakotaApplied Biochem & BiotechReact methanol (from gasification)with ethanol (from biochem) toyield primarily C4 (C3+) alcohols(Guerbet rxn)Preliminary Screening - Technical andEconom ic Assessm ent of SynthesisGas to Fuels and Chemicals withEmphasis on the Potential for Biomassderived Syngas2003Spath, P.L.; Dayton, D.C.National Renewable energy Laboratory(NREL) Report TP-510-34929overview of potential for syngasconversion to fuels and chem icals;includes hydrogen, ammonia,m ethanol, DME, mixed alcohols,oxosynthesis, MTG, isosynthesis,ethanol fermentation, etc. Verygood report11635Palladium-based Catalysts for theSynthesis of Alcohols2003M. Josefina Perez-Zurita;Cifarelli, M.; et.al.Universidad Central de VenezuelaJournal of Molecular Catalysis,ElsevierCatalyst analysis from XRD; highmethane and methanolselectivities9527Synthesis of Higher Alcohols in a SlurryReactor with CS-promoted ZincChromite Catalyst inDecahydronaphthalene2003Xiaolei Sun, G.W . RobertsNorth Carolina StateApplied Catalysis, ElsevierContinuous Slurry Reactor; Csshifts products away from MeOH,towards Higher Alcohols9535Fuel from the Synthesis Gas - the Roleof Process Engineering2003Stelm achowski, Marek, et.al.Technical University of Lodz, PolandApplied Energy, ElsevierMeOH, FT, and HAS, synthesis,slurry phase reactor; simulationN/AAlcohol Synthesis over Pre-ReducedActivated Carbon-supportedMolybdenum-Based Catalysts2003Dadyburjor, D.B., Li, Xianguo;et.al.W est Virginia University; OceanUniversity of China Moleculesformation of alcohols at expenseof CO conversion; prereducedrather than sulfided; getting lots ofHCsN/ACo-Production of Fuel Alcohols &Electricity via Refinery Coke Gasification2003Ravikumar, Ravi; andShepard, PaulFluorGasification Technologies ConfUse of Power Energy Fuels Inc'sEcalene process for m ixedalcohols synthesis in refineryapplication (Pet Coke); modelingand economic analysis9519The Influence of Clay on K2C O3/CoMoS2Catalyst in the Production ofHigher Alcohol Fuel2002Iranm ahboob, Jam shid; Hill,D onaldUconn, M ississippi StateFuel Processing Technology,ElsevierClay had significant im pact onHAS when K2CO 3 present;oxygenates selectivity of 70% ;testing of feed conditions9536Alcohol Synthesis from Syngas overK2C O3/CoS/MoS2on Activated Carbon2002Iranm ahboob, Jam shid; Hill,D onaldUconn, M ississippi StateCatalysis LettersExperim entation;Effects ofTem p,H2:CO, G HSV tested; fixed bed;active chem icals on carbondecreased surface areadram atically9528Advances in Catalytic Synthesis andUtilization of H igher Alcohols2000Herm an, R.G.Lehigh Univ.Catalysis Today, Elsevierfocus on oxygenates, particularlyisobutanol; slurry phase bed andDouble Bed designs; som eliterature results and op conditionsreported; injection of EtO H m ayprom ote HASThe Econom ical Production of AlcoholFuels from Coal-derived Synthesis Gas:Final Report, O ctober 1991 - M arch19981999 W est Virginia UniversitySubm itted to National EnergyTechnology Lab (NETL),Contract #DE-AC22-91PC91034-257 years of research on m ixedalcohols from M oS2 catalystsl;extensive work; higher alcoholselectivity > 40% ; productivity at200 g/kg/h; hydrocarbon selectivity< 20% ; Use Coal feedstock9545Synthesis of Higher Alcohols -Enhancem ent by the Addition ofM ethanol or Ethanol to the Syngas1999Lachowska, M.Polish Academ y of SciencesReact.Kinet.Catal.Lett ElsevierIncrease in C3+alcohols usingm odified m ethanol catalystN/AForm ation of Ethanol and HigherAlcohols from Syngas1999Iranm ahboob, Jam shid;D onald HillM ississippi State UniversityM ississippi State University,Thesis ReportTalk of biom ass; M oS2; Ph. D.Thesis; unsupported (K2CO 3/CoM oS2), and supported on Clay.M ention of m ethanol rem oval andadsorption for kinetic driving force9515Selective Synthesis of M ixed AlcoholsCatalyzed by Dissolved Base-activatedHighly Dispersed Slurried Iron1999M ahajan, D. et.al.Brookhaven National LaboratoryFuel (Elsevier)Ultrafine catalyst system , lowertem ps (250C) and pres (<6MPa);up to 95% selectivity (c1-c4alcohols); and 80% COConversion; nanom eter iron oxides"NANOC AT"; synthesis rem ains aproblem9538Synthesis of Short Chain Alcohols overa Cs-prom oted Cu/ZnO/Cr2O3catalyst 1998M ajocchi, L.; Forzatti, Pio;et.al.Politecnico di M ilano; Snam progettiSpA Applied Catalysis, ElsevierEffect of Tem p, C O:H2 ratio,GHSV. Tem p is key to AlcoholProductivity, 290 - 315 C are bestfor C2-C3; HC's m inim ized9511Higher Alcohols from Synthesis GasUsing Carbon-Supported DopedM olybdenum -Based Catalysts1998Xianguo Li; Dadyburjor, DadyB. et.al.W est Virginia UniversityInd. Eng. Chem . Res. (ACS)Studied catalyst prep param etersand reaction conditions; catalystprom oted w/ K and/or Co9537AKinetic M odel for the Synthesis ofHigh-M olecular-W eight Alcohols over aSulfided Co-K-M o/C Catalyst1998G unturu, A.K.; Dadyburjor,D.B., et.alW est Virginia UniversityInd. Eng. Chem . Res. (ACS)Kinetics, recycle reactor, tem psand feed variables explored;includes rate expressions9529Developm ent of a Process for HigherAlcohol Production via Synthesis Gas1998Beretta, A.; Tronconi, E. et.al. Politecnico di M ilano; Snam progetti SpA Ind. Eng. Chem . Res. (ACS)High-tem p catalyst; single-stageadiabatic pilot scale reactor,Kinetics developed; targetingoxygenates, ethanol low;sim ulation of m ulti-stage adiabaticreactor9530High Tem perature Calcined KMoO3/gam m a-Al2O3Catalysts for M ixedAlcohols Synthesis from Syngas: Effectsof M o Loadings1998Bian, G uo-zhu, et.al. (U niv. ofS&T, China)University of Sci & Tech, China; Univ ofTokyo Applied Catalysis, Elseviertotal yields of m ixed alcoholsdecreased but selectivity increasedfrom 3-50% ; m ainly linear C1-C4,sm all am t iso-C49514Kinetic Analysis of M ixed AlcoholSynthesis from Syngas over K/M oS2Catalyst1997Park, T.Y.; Nam , In-sik, et.al.Pohang Univ. of Sci. & Tech (Postech),Korea Ind. Eng. Chem Res. (ACS)K2CO3-prom oted-MoS2;m echanistic kinetics m odels;m ixed alcohol form ationm axim ized around 320C9518Synthesis of M ixed Alcohols from CO2contained syngas on supported M oS2catalysts1997G ang, Lu, et.al.East China University of Sci & TechApplied Catalysis, ElsevierCO2 in feed inhibits higheralcohols; product distributionstrongly influenced by C O2 andH2O in feed; high H2:CO ratioused; 0-100 ppm H2S9524Development of a Mechanistic KineticModel of the Higher Alcohol Synthesisover a Cs-doped Zn/Cr/O Catalyst - 1)Model Derivation and Data Fitting1996Beretta, A.; Forzatti, Pio, et.al. Politecnico di Milano; Snamprogetti SpA Ind. Eng. Chem. Res. (ACS)Kinetics model; goal is branchedalcohols, not linear because ofbetter octane properties9525A Kinetic Model for the Methanol-HigherAlcohol Synthesis from CO/CO2/H2overCu/ZnO-based Catalysts IncludingSimultaneous Formation of MethylEsters and Hydrocarbons1994Breman, B.B.; Beenackers,A.C.C.M.; et.al.University of Groningen, TheNetherlandsChemical Engineering Science,PergamonOverview of past work on modifiedmethanol catalysts and Kinetics;methanol primary product; claimsto be most accurate kineticsmodels to date; substantial HC'sproduced as unwanted byprod.9526Thermodynamics of Higher AlcoholSynthesis1993Mawson, S.; Roberts, G.W .;et.al. North Carolina State Energy & FuelsAspen Plus simulation usingRGIBBS; C4 alcohols aredominant product, especially asisobutyl; conversions > 90%;Thermodynamics11636SiO2-supported Biometallic RutheniumCluster-derived Catalysts for Selectivityand Activity Control in COHydrogenation Toward C1-C5 Alcohols1992Xiao, F.; Ichikawa, M.; et.al. Hokkaido UniversityJournal of Molecular Catalysis,Elsevier Ruthenium Catalyst;9517The Thermodynamics of Higher AlcoholSynthesis1992Roberts, G.W ., et.al.North Carolina StateSymposium on NG upgrading II,Division of Petroleum Chem,Inc., ACSBasically same info as presentedin 1993,9523An Overview of the Higher AlcoholSynthesis1992Smith, K.; Klier, K. Univ. British Columbia; LehighSymposium on NG upgrading II,Division of Petroleum Chem,inc., ACSGood brief review on catalystchoices and Operating Conditionsat that time, including Rh/SiO2;target for octane seemed to be70:30 c2+/c1 (required for gasolineblending?)9521Non-Stoichiometric Spinel-typeCatalysts for the Synthesis of Methanoland Higher Alcohols1992Trifiro, F., et.al.Dipartim ento di Chimica Industriale e deimateriali, Bologna, ItalySymposium on Octane andCetane EnhancementProcesses for Reducedemissions Motor Fuels, Divisionof Petroleum Chemistry, ACSNon-stoichiometric spinel-typesolid oxides; again showspropensity of K to shift towardshigher alcohols11632The Conversion of Synthesis Gas toHigher Oxygenated Fuel on Rh-basedCatalysts: Effect of Chemical Additives1991Balakos, M.W ., Chaung, S. S.C, University of AkronFuel Science and TechnologyInternationalRhodium catalyst, high selectivitiesto CH4 & C2+ Oxygenates9522Higher Alcohol Synthesis1991Forzatti, Pio, et.al.Politecnico di Milano; SnamprogettiSpA Catal. Rev. - Sci. EngExtensive writeup (book chapter?)on status of HAS catalysts;includes kinetics andthermodynamics; and commercialdesigns from Lurgi, IFP,SnamprogettiProduction of Oxygenates fromSynthesis Gas: A Technology Reviewfor US Dept of Energy for Design ofGeneric Coal conversion Facilities1991Pittsburgh Energy TechnologyCenter M.W . Kellogg Company Contract #DE-AC22-91PC89854Oxygenates overview, includingmixed alcohols; discuss dowtechnology as 85% selectivity andCO perpass conversion of 20-25%; effect of catalyst additivesClassical and Non-classical Routes forAlcohol Synthesis1991Herman, R.G. Lehigh UniversityBook Chapter, New Trends inCO Activation, L. Guczi, Ed.ElsevierImpressive summary of state ofmixed alcohols technology.Includes detailed kinetics work andexperimental results to date9539Catalytic Conversion of Synthesis Gasto Methanol and Other OxygenatedProducts1991Stiles, Alvin, et.al.University of DelawareInd. Eng. Chem. Res.; ACSExperiments, bench-scale, highaldehyde production based onconditions; benchmark establishedon pure methanol; alkaline earth &cobalt additives9531Kinetic Modelling of Higher AlcoholSynthesis over Alkali-promoted Cu/ZnOand MoS2Catalysts1990Smith, K. Herman, R.; Klier, K.Lehigh UniversityChemical Engineering ScienceKinetics of both modified FT &modified MeOH; different reactionpathways for each9541Production of Mixed Alcohols fromSynthesis Gas1990Uchiyama, S., et.al.Idemitsu Kosan Co.Sekiyu GakkaishiComparison of modified FT &modified MeOH Catalysts;includes summary of operatingconditions for industry: HalderTropsche/Snamprogetti/Enichem,Dow-Union Carbide, IFP/IdemitsuKosan; Lurgi; some economics9568Production of Methanol-higher AlcoholMixtures from Natural Gas via SyngasChemistry1990Coutry, Chaumette,Raimbault, TraversFrench Petrol Institute (IFP)European Applied ResearchConf. On Natural GasEnichem; Snam Progetti; HaldorTopsoeProduction of Methanol-higher AlcoholMixtures from Natural Gas via SyngasChemistry1990LurgiProduction of Methanol-higher AlcoholMixtures from Natural Gas via SyngasChemistry1990IFP & Idemitsu-KosanProduction of Methanol-higher AlcoholMixtures from Natural Gas via SyngasChemistry1990Dow, Union CarbideProduction of Methanol-higher AlcoholMixtures from Natural Gas via SyngasChemistry1990C1 Chem Group (Japan), Texaco,others9532Synthesis of Alcohols from CarbonOxides and Hydrogen. 4. - LumpedKinetics for the Higher AlcoholSynthesis over a Zn-Cr-K OxideCatalyst1987Tronconi, E.; Forzatti, P; et.al.Politecnico di Milano; SnamprogettiSpA Ind. Eng. Chem. Res. (ACS)Kinetics; discussions of CO2recycle9512Mixed Alcohols from Synthesis Gas1986Quarderer, GeorgeDow ChemicalAIChE Spring National MeetingVery low water, 50:50 MeOH/HASsplit; economics provided; reactordesigns considered as well; lots ofgood dataSRI PEP#85-1-4 Mixed Alcohols from Syngas1986Nirula, SatishSRIPEP Review #85-1-4Dow, Union Carbide; includeseconomics; also includes state ofart discussion

## Appendix K

Alcohol Synthesis KineticsA more rigorous method for modeling the producti on 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 ki netics-based model can more accu rately predict the effect of varying inlet conditions to the synthesis reactor, especially when unconverted syngas or unwanted co-products (i.e., methanol, hydrocar bons) are recycled to the reactor inlet.

Two major barriers to using kine tics-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 eq uations – 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 co nditions explored. Using the model outside the range of operating conditions introduces increase d uncertainty the more removed the estimate is from the experimental conditions used to devel op the model. Catalyst performance is very sensitive to many factors that can arise durin g their production. Two seemingly identical catalysts based on chemical formulation can vary greatly in performance because of preparation techniques and catalyst support characteris tics (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 ty pes of added promoters or inhibitors that help tune the catalyst to promote desired reac tions while inhibiting unde sirable reactions. An “optimum” catalyst is usually a compromise of co mpeting and interacting promoters, inhibitors, and supports to give a catalyst with the “best” cumulative properties. Wh en precious metals are used, economics also become a major factor in determining the catalyst formulation.

Very little published data exist for the MoS 2 -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 La ngmuir-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 slig htly different equations and parameters. The procedure for estimating the para meters is not discusse d here, but can be found in detail in Larson, et al. (Larson, et al ., 2006) and are presented here in Table 1.Table 1. Kinetic parameter estimates from Larson, et al. Blank cells do not have values for the corresponding rate equati on parameter and rate equation.



## Parameter Methanol Ethanol Propanol Hydrocarbons

A

m , A

e , A p , A h 14.6233 3.0518 0.2148 9.3856

E

m , E
e , E p , E h 143.472 24.986 89.3328 95.416

n m , n e , n p , n h 3 1 1 1

K 1 7.6393E-9

K 2 0.6785

K 3 0.9987

K
e 0.7367

K

p 0.6086

K

h 1.2472

K

z 0.8359

The rate equations were created using the typi cal Langmuir-Hinshelwood kinetic approach found in most catalyst kinetics textbooks. Equation 1 is th e gross rate of methanol production. It is the only rate equation in the set that has a reverse reaction component and th erefore an equilibrium value. Chemical equilibrium is reached when th e partial pressure of me thanol is sufficiently high to cause the reverse reaction rate, th e dissociation of methanol to CO and H 2 , plus the consumption rate of meth anol to other products, to equal the forward reaction rate of CO and H 2 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 effe cts of methanol being

consumed to make ethanol or methane as de scribed 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 K K = (3)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 irreve rsible reactions of meth anol to ethanol, ethanol to propanol, and methanol to methane. Ethanol can react further with H 2 and CO to make propanol. Implicit in this set of equations is the simplifying assumption that any products generated other than methanol, ethanol or pr opanol produce methane even though in higher molecular weight hydrocarbons are observed expe rimentally. Small amounts of butanol and pentanol, both expected and experi mentally observed, have been ignored in the analysis since

CO conversions were low resu lting in minimal production of in creasingly higher alcohols.

Using the rate equations above, plus a rate equati on for the water-gas-shift reaction , a system of differential equations an d initial conditions can be easily written for a plug flow reactor, as in this report, or for a Continuous Stirred Tank R eactor (CSTR). The solution of the system of ordinary differential equations was programmed into the commercially available equation solver, PolyMath, to evaluate the kine tics over a range of conditions.

Examples of the results given by the model are sh own 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 “adde d” to the reactor in let to make 1 mol%

concentration. CO and H 2 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 consta nt 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.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 ConversionFlow 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 ConversionFlow 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 -1For 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 sustaina ble in a process without modifying the process to produce more methanol.

Values of CO conversion and pr oduct rates at the reactor outlet for a plug flow reactor at 68,

with and without methanol adde d 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

CH 4 = 3237

EtOH = 9861

PrOH = 3851 XCO = 6.38

MeOH = 1974

CH 4 = 5770

EtOH = 19600

PrOH = 6543

610K XCO = 27.08

MeOH = 4282

CH 4 = 26400

EtOH = 25600

PrOH = 21400 XCO = 25.94

MeOH = 4197

CH 4 = 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 eff ect of adding methanol at 610K is less pronounced than at 570K. The CO conversion decreases slig htly 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 me thanol 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 dete rmine when it is infeasible to increase the capital and catalyst costs to increase production.

The effect of adding methanol on the ethanol yi eld suggests that other configurations may be more effectively used to maximize ethanol pr oduction. Since the methanol production at thereactor exit is limited by chemical equilibrium, th e amount of methanol that can be recycled to the reactor is also limited. Instead of recy cling only methanol produced from a mixed alcohol catalyst, it may be possible to economically split th e 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 me thanol, contrary to the way these reactors are typically operated when methanol is the desi red end product. These alternate designs would require other factors to be considered, such as lower sulfur and CO 2 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 CO 2 that can be fed to the reactor without adversely affe cting product selectivit y 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 Fuel s from Coal-Derived Synthesis Gas”, Final Report, Contract # DE-AC22-91P C91034-25, 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 Pu lp and Paper Industry, Volume 2, Detailed Biorefinery Design and Performance Simula tion”, Final Report unde r contract DE-FC26- 04NT42260 with the U.S. Department of Energy a nd with cost-sharing by the American Forest and Paper Association, December 2006.F1147-E(12/2004) REPORT DOCUMENTATION PAGE Form Approved
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National Technical Information Service U.S. Department of Commerce 5285 Port Royal Road Springfield, VA 22161

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14. ABSTRACT (Maximum 200 Words) 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 level by 2012.

15. SUBJECT TERMS process design; thermochemical; pilot; cellulosic; ethanol

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