EPA-600/R-97-153
December 1997
HYNOL PROCESS EVALUATION
By
Robert H. Borgwardt
Atmospheric Protection Branch
Air Pollution Prevention and Control Division
U.S. Environmental Protection Agency
Research Triangle Park, NC 27711
Prepared for:
U.S. Environmental Protection Agency
Office of Research and Development
Washington, DC 20460
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NOTICE'
This document has been reviewed in accordance with
U.S. Environmental Protection Agency policy and
approved for publication. Mention of trade names
or commercial products does not constitute endorse-
ment or recommendation for use.
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FOREWORD
The U. S. Environmental Protection Agency is charged by Congress with pro-
tecting the Nation's land, air, and water resources. Under a mandate of national
environmental laws, the Agency strives to formulate and implement actions lead-
ing to a compatible balance between human activities and the ability of natural
systems to support and nurture life. To meet this mandate, EPA's research
program is providing data and technical support for solving environmental pro-
blems today and building a science knowledge base necessary to manage our eco-
logical resources wisely, understand how pollutants affect our health, and pre-
vent or reduce environmental risks in the future.
The National Risk Management Research Laboratory is the Agency's center for
investigation of technological and management approaches for reducing risks
from threats to human health and the environment. The focus of the Laboratory's
research program is on methods for the prevention and control of pollution to air,
land, water, and subsurface resources; protection of water quality in public water
systems; remediation of contaminated sites and groundwater; and prevention and
control of indoor air pollution. The goal of this research effort is to catalyze
development and implementation of innovative, cost-effective environmental
technologies; develop scientific and engineering information needed by EPA to
support regulatory and policy decisions; and provide technical support and infor-
mation transfer to ensure effective implementation of environmental regulations
and strategies.
This publication has been produced as part of the Laboratory's strategic long-
term research plan. It is published and made available by EPA's Office of Re-
search and Development to assist the user community and to link researchers
with their clients.
E. Timothy Cppelt, Director
National Risk Management Research Laboratory
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ABSTRACT
Fuel-cell vehicles (FCVs) have the potential to offer a major improvement in efficiency relative to
current motor vehicles. The potential of FCVs to nearly eliminate pollutant emissions and reduce the economic
pressures of petroleum imports will be major factors contributing to the sustainability of the current system of
highway transport. Greenhouse gas emissions, of which the transportation sector is a major contributor, are
another part of the sustainability issue. FCVs and the fuels used in them offer an economical and effective
technological option for mitigation of anthropogenic emissions of carbon dioxide (C02), the predominant
greenhouse gas. The hydrogen that is required for fuel cells can be produced from natural gas (which
contributes to the buildup of C02 in the atmosphere) and biomass (which does not). This report examines
process alternatives for the optimal use of these resources for production of FCV fuel, emphasizing maximum
displacement of petroleum and maximum reduction of overall fuel-cycle C02 emissions at least cost. Three
routes are evaluated: (i) production of methanol from biomass and from natural gas by independent processes,
(ii) production of methanol or hydrogen by hydrogasification of biomass using natural gas as co-feedstock
supplemented with, and without, the use of carbonaceous municipal wastes as co-feedstocks, and (iii)
production of methanol or hydrogen by addition of natural gas to a biomass-to-methanol process originally
designed for biomass only. The results show that the combined use of natural gas and biomass in a single
process can reduce net fuel-cycle C02 emissions by 20% relative to separate systems and reduce the cost
of fuel production to a range competitive with the current cost of gasoline. A plant optimized for efficiency and
size, with 25% of the feedstock energy consisting of biomass, should be able to produce methanol at a cost
of $0.42/gal ($6.09/GJ), or hydrogen at $5.98/GJ. This technology represent a "no regrets" approach to C02
mitigation and the most cost-effective use of biomass as a source of fuel energy.
This report covers the period June 1995 to May 1997 and was completed June 3,1997.
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CONTENTS
Page
Abstract iv
List of Figures viii
List of Tables ix
Abbreviations and Symbols x±±
Metric Equivalents
1. Introduction 1-1
Sustainable T ransport 1-1
Background 1-2
National Goals for Alternative Fuels 1-3
The Biomass Supply Issue 1-4
Biomass as an Independent Source of Transportation Fuel 1-5
Impact of Fuel-Cell Vehicle Technology 1-6
Methodology and Assumptions of this Analysis 1-6
2. Conclusions 2-1
General 2-1
Conclusions Specific to Hynol 2-2
3. Recommendations 3-1
4. The BCL Process 4-1
Background 4-1
The Biomass Gasifier 4-2
Process Block Flow Diagram 4-2
Steam Reforming Block 4-2
Heat-Recovery Steam Generator Block 4-2
Distillation Block 4-2
Methanol Synthesis Block 4-10
Biomass Drier 4-10
BCL Process Performance Evaluation 4-10
5. The Hynol Process 5-1
General Description 5-1
Biomass Gasification Block 5-4
Reformer Unit and Power Block 5-4
Distillation Block 5-9
Methanol Synthesis Block 5-9
Biomass Block 5-16
Integrated Process Flowsheet 5-19
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CONTENTS (Continued)
6. Performance Comparisons 6-1
Objectives 6-1
Methanol Production from Natural Gas by Conventional Steam Reforming 6-1
Vehicle C02 Emission Reduction Potential of Methanol Produced by the BCL
Process and Conventional Steam Reforming 6-2
C02 Emission Reduction by the Hynol Process 6-3
Fuel-Cycle COz Emissions 6-4
7. Hynol Sensitivity Analysis 7-1
Introduction 7-1
Procedure 7-1
Ratio of Natural Gas Feed to Biomass Feed 7-3
Purge Location 7-8
Effect of Reformer Temperature 7-11
Effect of Biomass Conversion 7-14
Use of Unreacted Carbon as Process Fuel 7-14
Effect of Reformer Steam/Carbon Ratio 7-19
Approach to Equilibrium 7-19
8. Hynol Process Optimization 8-1
Background 8-1
Initial Conditions for Process Optimization 8-3
Additional Assumptions from Sensitivity Analysis 8-3
Assumptions Regarding Tar Formation 8-4
Pressure and Pressure Drops 8-4
Allocation of Pressure Drops in the Gasifier Loop 8-5
Allocation of Pressure Drops in Methanol Synthesis Loop 8-5
Process Optimization for 50-atm Methanol Converter 8-5
Effect of Methanol Converter Pressure >50 atm 8-10
Process Optimization with 90-atm Methanol Converter 8-10
Summary of Optimized Hynol Process Performance 8-30
9. Hydrogen Production 9-1
Introduction 9-1
Process Modifications for Hydrogen Production 9-1
Evaluation of Process Performance 9-15
10. Cost Estimates 10-1
C02 Mitigation Cost 10-4
Stationary Source Mitigation 10-4
Costs Compared to Gasoline 10-6
11. Waste Methane as Co-Feedstock 11-1
Introduction 11-1
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CONTENTS (Continued)
Landfill Gas as Supplemental Process Feed for Methanol Production 11 -3
Hydrogen Production Using Waste Gas as Supplemental Feed 11-3
12. Sewage Sludge 12-1
Introduction 12-1
Case 1: Total Displacement of Woody Biomass by Sludge 12-2
Re-Optimization of Operating Conditions for Sludge Addition 12-2
Case 2: Intermediate Displacement of Woody Biomass with Sludge 12-4
Case 3: Low Displacement of Woody Biomass with Sludge 12-4
Case 4: Sludge Without Digester Gas 12-4
Performance Comparisons 12-4
13. Natural Gas as Co-Feedstock for the BCL Methanol Process 13-1
Introduction 13-1
BCL Process Modification for Natural Gas Addition 13-1
Performance Evaluation 13-6
14. References 14-1
Appendix A. Modification of the BCL Process Proposed by Katofsky A-1
Appendix B. Energy Requirements for Drying Biomass B-1
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LIST OF FIGURES
Figure Page
4-1 Gasifier / combustor system of the BCL process 4-3
4-2 Overall block flow diagram for methanol production by the BCL process 4-4,4-5
4-3 Details of steam reforming block of the BCL process 4-8
4-4 Details of the heat recovery steam generator and power block of the BCL process 4-9
4-5 BCL process, details of the methanol distillation block 4-11
4-6 BCL process, details of the methanol synthesis block 4-12
4-7 Biomass drying operation using heat recovered from the BCL process 4-13
5-1 Block diagram of Hynol process 5-2
5-2 Biomass gasification block for Hynol process 5-5
5-3 Equilibrium conditions for the reaction: H2S + ZnO = ZnS + H20 5-7
5-4 Hynol reformer unit and power block 5-8
5-5 Hynol distillation block 5-12
5-6 Methanol synthesis block 5-14
5-7 Hynol biomass block 5-17
5-8 Integrated Hynol process flowsheet 5-20
7-1 Hynol process block flow diagram 7-2
' 7-2 Total fuel-cycle C02-equivalent emission reduction for gasoline ICEVs displaced by
methanol FCVs with 2.5 factor improvement of fuel economy 7-5
7-3 Effect of Hynol purge location on total natural gas requirement and
fuel-cycle C02 emission reduction 7-11
8-1 Conceptual Hynol flowsheet of Steinberg and Dong 8-2
8-2 Pressure drop allocation in the gasifier loop. From entry of recycle stream from
methanol converter to reformer exit 8-6
8-3 Pressure drop allocation in gasifier loop. From reformer exit to entry at
gasifier loop compressor 8-7
8-4 Block diagram of Hynol process optimized for 50-atm methanol synthesis 8-8
8-5 Details of gasifier block for 50-atm optimization 8-11
8-6 Details of reformer block for 50-atm optimization 8-13
8-7 Details of distillation block for 50-atm optimization 8-16
8-8 Details of methanol synthesis block, 50-atm optimization 8-18
8-9 Block diagram of Hynol process optimized for 90-atm methanol synthesis 8-21
8-10 Reformer block for 90-atm Hynol optimization 8-24
9-1 Block diagram for Hynol process configured for hydrogen production 9-2
9-2 Details of gasifier block for hydrogen production 9-4
9-3 Reformer block for hydrogen production 9-7
9-4 Shift reactor block for hydrogen production 9-10
9-5 Pressure swing adsorber block for hydrogen production 9-13
10-1 Comparison of fuel production costs with gasoline 10-7
11-1 Block flow diagram for optimized Hynol process using landfill gas as co-feedstock 11-4
11 -2 Reformer operation with landfill gas for partial displacement of natural gas as fuel 11 -7
11 -3 Block flow diagram for hydrogen production using waste gas as co-feedstock 11 -8
13-1 Block diagram for BCL process modified for natural gas addition 13-2
13-2 Details of the gas preparation block for BCL process modified for natural gas addition 13-7
13-3 Details of the reformer and power block for modified BCL process 13-9
13*4 Details of the distillation block for modified BCL process 13-11
13-5 Methanol synthesis block for modified BCL process 13-13
13-6 Modification of BCL char combustor for steam recovery 13-18
A-1 BCL process flowsheet as established for this report A-2, A-3
A-2 Energy balance on BCL reformer based on Katofsky flowsheet data A-4
A-3 BCL gasifier material balance A-5
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LIST OF TABLES
Table Page
1-1 COMPOSITION OF BIOMASS FEED ASSUMED FOR THIS STUDY 1-7
1-2 COMPOSITION OF NATURAL GAS ASSUMED FOR THIS STUDY 1-7
4-1 BCL PROCESS STREAM COMPOSITIONS 4-6, 4-7
5-1 PRINCIPAL STREAM DATA FOR HYNOL PROCESS 5-3
5-2 STREAM DATA FOR FIGURE 5-2 5-6
5-3 STREAM DATA FOR REFORMER AND POWER BLOCK 5-10
5-4 HEAT DUTIES FOR REFORMER AND POWER BLOCK 5-11
5-5 STREAM DATA FOR DISTILLATION BLOCK 5-13
5-6 STREAM DATA FOR METHANOL SYNTHESIS BLOCK 5-15
5-7 STREAM DATA FOR BIOMASS BLOCK 5-18
5-8 ELECTRIC POWER CONSUMPTION 5-19
5-9 ELECTRIC POWER AVAILABLE 5-19
6-1 SUMMARY OF TAILPIPE COa EMISSION REDUCTIONS 6-4
6-2 GRAMS OF C02-EQUIVALENT EMISSIONS PER 106 Btu OF FUEL AVAILABLE TO USER .. 6-5
7-1 EFFECT OF NATURAL GAS/BIOMASS RATIO 7-6
7-2 SUMMARY OF EFFECTS OF NATURAL GAS / BIOMASS FEED RATIO 7-7
7-3 EFFECT OF PURGE LOCATION ON SIMULATIONS WITH REDUCED
NATURAL GAS / BIOMASS FEED RATIO 7-9
7-4 SUMMARY OF EFFECTS OF PURGE LOCATION ON HYNOL PROCESS
PERFORMANCE WITH REDUCED NATURAL GAS / BIOMASS FEED RATIO 7-10
7-5 SIMULATION RESULTS FOR VARIABLE REFORMER TEMPERATURE 7-12
7-6 SUMMARY OF REFORMER TEMPERATURE EFFECT ON HYNOL
SYSTEM PERFORMANCE 7-13
7-7 SIMULATION RESULTS FOR VARIABLE BIOMASS CONVERSION 7-15
7-8 SUMMARY OF EFFECTS OF CARBON CONVERSION ON HYNOL
PROCESS PERFORMANCE 7-16
7-9 SIMULATION RESULTS ON EFFECT OF RECOVERING ENERGY
CONTENT OF UNREACTED GASIFIER CARBON 7-17
7-10 SUMMARY OF EFFECTS OF WASTE CARBON UTILIZATION ON HYNOL
PROCESS PERFORMANCE 7-18
7-11 SIMULATION RESULTS FOR VARIABLE STEAM/CARBON RATIO 7-20
7-12 SUMMARY OF EFFECTS OF STEAM/CARBON RATIO ON HYNOL
PROCESS PERFORMANCE 7-21
7-13 SIMULATION RESULTS FOR VARIABLE APPROACH TO EQUILIBRIUM
IN METHANOL SYNTHESIS REACTOR 7-22
7-14 SUMMARY OF EFFECTS OF MSR APPROACH TO EQUILIBRIUM ON
HYNOL PROCESS PERFORMANCE 7-23
7-15 SIMULATION RESULTS FOR VARIABLE APPROACH TO EQUILIBRIUM IN REFORMER 7-25
7-16 SUMMARY OF EFFECTS OF REFORMER APPROACH TO EQUILIBRIUM
ON HYNOL PROCESS PERFORMANCE 7-26
8-1 STREAM DATA FOR 50-atm OPTIMIZATION 8-9
8-2 STREAM DATA FOR GASIFICATION BLOCK 8-12
8-3 STREAM DATA FOR REFORMER AND POWER BLOCK 8-14
8-4 HEAT DUTIES FOR REFORMER BLOCK, 50-atm OPTIMIZATION 8-15
8-5 STREAM DATA FOR DISTILLATION BLOCK 8-17
8-6 STREAM DATA FOR METHANOL SYNTHESIS BLOCK 8-19
8-7 EFFECT OF METHANOL SYNTHESIS REACTOR PRESSURE ON THE MSR RECYCLE
RATIO, HEAT EXCHANGER REQUIREMENTS, AND STEAM PRODUCTION 8-20
LX
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LIST OF TABLES (Continued)
Table Page
8-8 STREAM DATA FOR OPTIMIZED HYNOL PROCESS WITH 90-atm
METHANOL CONVERTER 8-22
8-9 STREAM DATA FOR GASIFICATION BLOCK, 90-atm OPTIMIZATION 8-23
8-10 STREAM DATA FOR REFORMER AND POWER BLOCK 8-25
8-11 STREAM DATA FOR DISTILLATION BLOCK, 90-atm OPTIMIZATION 8-26
8-12 STREAM DATA FOR METHANOL SYNTHESIS BLOCK, 90-atm OPTIMIZATION 8-27
8-13 HEAT DUTIES FOR GASIFIER BLOCK, 90-atm OPTIMIZATION 8-28
8-14 HEAT DUTIES FOR REFORMER BLOCK, 90-atm OPTIMIZATION 8-28
8-15 HEAT DUTIES FOR DISTILLATION BLOCK, 90-atm OPTIMIZATION 8-29
8-16 HEAT DUTIES FOR METHANOL SYNTHESIS BLOCK, 90-atm OPTIMIZATION 8-29
8-17 HEAT DUTIES FOR BIOMASS BLOCK, 90-atm OPTIMIZATION 8-29
8-18 ELECTRIC POWER CONSUMPTION: 50-atm HYNOL OPTIMIZATION 8-30
8-19 ELECTRIC POWER PRODUCTION: 50-atm HYNOL OPTIMIZATION 8-31
8-20 ELECTRIC POWER CONSUMPTION: 90-atm HYNOL OPTIMIZATION 8-31
8-21 ELECTRIC POWER PRODUCTION: 90-atm HYNOL OPTIMIZATION 8-32
8-22 SUMMARY OF PERFORMANCE PARAMETERS FOR 50- AND 90-atm
HYNOL OPTIMIZATIONS 8-32
9-1 STREAM DATA FOR HYNOL CONFIGURED FOR HYDROGEN PRODUCTION 9-3
9-2 STREAM DATA FOR GASIFICATION BLOCK, HYDROGEN PRODUCTION 9-5
9-3 STREAM DATA FOR REFORMER BLOCK FOR HYDROGEN PRODUCTION 9-8
9-4 HEAT DUTIES FOR REFORMER BLOCK, HYDROGEN PRODUCTION 9-9
9-5 STREAM DATA FOR SHIFT REACTOR BLOCK FOR HYDROGEN PRODUCTION 9-11
9-6 HEAT DUTIES FOR SHIFT REACTOR BLOCK, HYDROGEN PRODUCTION 9-12
9-7 STREAM DATA FOR PRESSURE SWING ADSORBER BLOCK 9-14
9-8 HEAT DUTIES FOR BIOMASS BLOCK, HYDROGEN PRODUCTION 9-14
9-9 ELECTRIC POWER CONSUMPTION FOR HYDROGEN PRODUCTION BY HYNOL 9-16
9-10 ELECTRIC POWER AVAILABLE FOR HYDROGEN PRODUCTION 9-16
10-1 COST ESTIMATE FOR METHANOL PRODUCTION BY THE HYNOL
PROCESS: BASE CASE 10-2
10-2 COST ESTIMATE FOR METHANOL PRODUCTION BY THE HYNOL
PROCESS: 90-atm OPTIMIZED CASE 10-3
10-3 COST ESTIMATE FOR HYDROGEN PRODUCTION BY THE HYNOL PROCESS 10-5
11-1 WASTE METHANE CHARACTERISTICS AND POTENTIAL MARKET 11-2
11 -2 STREAM DATA FOR OPTIMIZED HYNOL PROCESS WITH 90-atm METHANOL
CONVERTER AND 1.62 mols LANDFILL GAS ADDED TO GASIFIER EFFLUENT AS
SUPPLEMENTAL FEEDSTOCK 11-5
11-3 COMPARISON OF BASE-CASE OPTIMIZED HYNOL SYSTEM AND SYSTEMS USING
LANDFILL GAS AS SUPPLEMENTAL FEEDSTOCK FOR METHANOL PRODUCTION 11-6
11-4 STREAM DATA FOR HYNOL CONFIGURED FOR HYDROGEN PRODUCTION
USING LFG AS CO-FEEDSTOCK 11-9
11-5 RESULTS OF SIMULATIONS OF HYDROGEN PRODUCTION USING WASTE GAS AS
SUPPLEMENTAL FEEDSTOCK FOR NATURAL GAS DISPLACEMENT 11-11
11-6 HEAT DUTIES AND HEAT TRANSFER AREAS FOR WASTE GAS UTILIZATION AS
CO-FEEDSTOCKS 11-12
11-7 C02 EMISSION REDUCTION POTENTIAL OF WASTE GAS USED AS HYNOL
CO-FEEDSTOCK 11-13
12-1 STREAM DATA FOR HYNOL PROCESS USING SEWAGE SLUDGE FOR
TOTAL DISPLACEMENT OF WOODY BIOMASS, 90-atm, 1.2 mols purge 12-3
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LIST OF TABLES (Continued)
Table Page
12-2 STREAM DATA FOR HYNOL PROCESS UTILIZING SEWAGE SLUDGE
FOR TOTAL DISPLACEMENT OF WOODY BIOMASS, 90-atm, 2.4 mols purge 12-5
12-3 STREAM DATA FOR HYNOL PROCESS UTILIZING SEWAGE SLUDGE
FOR TOTAL DISPLACEMENT OF WOODY BIOMASS , 70-atm, 2.4 mols purge 12-6
12-4 STREAM DATA FOR HYNOL PROCESS UTILIZING SEWAGE SLUDGE
FOR TOTAL DISPLACEMENT OF WOODY BIOMASS, 70-atm, 3.6 mols purge 12-7
12-5 STREAM DATA FOR HYNOL PROCESS UTILIZING AN INTERMEDIATE AMOUNT
OF SEWAGE SLUDGE FOR PARTIAL DISPLACEMENT OF WOODY BIOMASS 12-8
12-6 STREAM DATA FOR HYNOL PROCESS UTILIZING A SMALL AMOUNT OF
SEWAGE SLUDGE FOR PARTIAL DISPLACEMENT OF WOODY BIOMASS 12-10
12-7 STREAM DATA FOR HYNOL PROCESS UTILIZING SEWAGE SLUDGE FOR FULL
DISPLACEMENT OF WOODY BIOMASS WITHOUT DIGESTER GAS ADDITION 12-11
12-8 PERFORMANCE SUMMARY OF SEWAGE SLUDGE SIMULATIONS 12-12
12-9 SUMMARY OF ENERGY REQUIREMENTS FOR SEWAGE SLUDGE UTILIZATION 12-13
12-10 HEAT EXCHANGER DUTIES AND AREAS FOR HYNOL PROCESS UTILIZING
SEWAGE TREATMENT SLUDGE, DIGESTER GAS, AND WOODY BIOMASS
AS CO-FEEDSTOCKS 12-14
13-1 PRINCIPAL STREAM COMPOSITIONS FOR BCL PROCESS MODIFIED FOR
NATURAL GAS ADDITION, OPTION 1 13-4
13-2 PRINCIPAL STREAM COMPOSITIONS FOR BCL PROCESS MODIFIED FOR
NATURAL GAS ADDITION, OPTION 2 13-5
13-3 STREAM DATA FOR BCL PROCESS MODIFIED FOR NATURAL GAS ADDITION.
GAS PREPARATION BLOCK 13-8
13-4 STREAM DATA FOR BCL PROCESS MODIFIED FOR NATURAL GAS ADDITION.
REFORMER AND POWER BLOCK 13-10
13-5 STREAM DATA FOR MODIFIED BCL PROCESS. DISTILLATION BLOCK 13-12
13-6 STREAM DATA FOR MODIFIED BCL PROCESS. METHANOL SYNTHESIS BLOCK 13-14
13-7 HEAT EXCHANGER DUTIES AND AREAS FOR MODIFIED BCL PROCESS (OPTION 2:
NATURAL GAS USED AS BOTH PROCESS FEED AND AS REFORMER FUEL) 13-15
13-8 ELECTRIC POWER UTILIZATION FOR BCL PROCESS MODIFIED FOR
NATURAL GAS ADDITION 13-16
13-9 COMPARISON OF PERFORMANCE PARAMETERS FOR MODIFIED BCL PROCESS
AND OPTIMIZED HYNOL PROCESS USING THE SAME AMOUNT OF BIOMASS AND
NATURAL GAS FEEDSTOCKS 13-19
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ABBREVIATIONS AND SYMBOLS
ADG Anaerobic digester gas
APPCD Ai'r Pollution Prevention and Control Division
atm Pressure in atmospheres (absolute)
BCL Battelle Columbus Laboratory
BFW Boiler feed water
Btu British thermal unit
°C Degrees centigrade
CCAP Climate Change Action Plan
CE-CERT College of Engineering, Center for Environmental Research and Technology
DOE U.S. Department of Energy
EJ Exajoule (10*8 joules)
EPA U.S. Environmental Protection Agency
EPACT Energy Policy Act
FCV Fuel cell vehicle
FFVs Flexible fuel vehicles
GJ Gigajoule (109 joules)
h Hour
HE Heat exchanger
HHV Higher heating value
HP High pressure
HRSG Heat recovery steam generator
IC Internal combustion
ICEVs Internal combustion engine vehicles
ICI Imperial Chemical Industries
IFC International Fuel Cells, Inc.
IP Intermediate pressure
IPCC Intergovernmental Panel Climate Change
K Degrees kelvin
kW Kilowatt
LFG Landfill gas
LHV Lower heating value
m Meter
MC Methanol condenser
Mg Metric tonne (megagram)
MSR Methanol synthesis reactor
MTBE Methyl f-butyl ether
MTCI Manufacturing and Technology Conversion International, Inc.
NG Natural gas
NMOC Non-methane organic compounds
AP Pressure drop
PGPH Process-gas preheater
PNGV Partnership for a New Generation of Vehicles
POX Partial oxidation
PSA Pressure swing adsorber
psi Pounds per square inch
RH Reheater
SH Superheater
SPR Steam pyrolysis reactor
T Tank
TG Turbine-generator
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METRIC EQUIVALENTS
It is EPA's policy to use metric units in its publications. However, certain nonmetric units are used
in this document for the reader's convenience. Readers more familiar with metric units may use the
following factors to convert to that system.
Yields
Nonmetric Multiply bv Metric
atm 101.3 kPa
Btu 1.054 kJ
°F 5/9 (°F - 32) °C
ft3 0.0283 m3
gal 3.79 L
0.00379 m3
in. 0.0254 m
mi 1.609 km
mpg 0.425 km/L
psi 6.89 kPa
ton 907.1 kg
0.907 tonne
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SECTION 1
INTRODUCTION
Sustainable Transport
It is generally agreed that the system of road transport, which currently meets 98 percent of U.S.
transportation needs, is not sustainable. The impacts of an expanding vehicle population on pollution, health,
energy security, and economic drain of due to imported petroleum are some of the most immediate problems.
Long term effects such as greenhouse gas emissions-carbon dioxide (C02) in this case-are also important.
Highway vehicles account for 79 % of the current U.S. petroleum use and about 23 % of the total U.S. C02
emissions. Globally, transport is the fastest growing contributor to the buildup of greenhoue gas concentration
in the atmosphere.
Methods being discussed to cope with these problems (D. Gordon, 1995) include: (1) reducing
demand and incentives for personal transport by including the external costs of petroleum use in its market
price, with reallocation of revenues to more optimal uses of national resources that can reduce the soical costs,
(2) changing land-use patterns to reduce the need for personal motorized transport, (3) developing and
bringing to market technologies that improve efficiency and reduce transportation impacts. Recent experience
with attempts to introduce a nominal "energy tax" on motor fuel revealed strong opposition to any approach
aimed at reducing demand, regardless of its perceived merits for conservation of energy resources or pollution
reduction. The second approach requires changes in personal habits and preferences in addition to changes
in the way urban development now occurs; either change could occur only slowly and, on a national scale,
would necessarily be a long term approach. This report is concerned with the third method.
New technologies that can effectively deal with many of the problems of road transport are nearing the
commercialization stage, especially under the auspices of the Partnership for a New Generation of Vehicles
(PNGV), a combined effort established in 1993 between the federal government and the U.S. automotive
industry. Fuel cell vehicles (FCVs), which comprise one part of the PNGV research and development
enterprise, are expected to be effective in reducing impacts of transportation on the environment, including
greenhouse gas emissions. One aspect of the sustainability issue not addressed by PNGV is alternative fuels
that are suited for use by the new vehicles, and that aspect is to be considered here.
An alternative to petroleum fuels must be competitive in selling price if it is to gain acceptance in the
marketplace. It should also be compatible with the existing refuelling infrastructure, provide a driving range
comparable to gasoline, and require no greater time for vehicle refuelling. Most important of all, it should be
compatible with the fuel cell vehicles when they begin commercialization. As an alternative fuel, natural gas
is competitive in price and can greatly reduce emissions of criteria pollutants from conventional vehicles, but
is lacking in the other requirements. FCVs, furthermore, require hydrogen-either as compressed gas or as
a liquid hydrogen-carrier that can be reformed on board to produce hydrogen. Compressed hydrogen is the
most efficient FCV fuel, but also lacks compatibility with the refuelling infrastructure now in place for petroleum
fuels; the distribution problem is a major barrier to the use of hydrogen gas as a transportation fuel. Natural
gas can be converted to methanol (CH3OH) which is a liquid hydrogen- carrier of much greater energy density
than hydrogen gas and can be relatively easily and efficiently reformed to hydrogen on board the vehicle. It
is compatible with the existing refueling infrastructure, is already in distribution for flexible fuel vehicles, and
is a prime candidate as FCV fuel.
1-1
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U.S. reserves of natural gas, though large (resource base = 4 x 1014 m3), are finite arid have many uses
other than transportation fuel. Methanol, and/or hydrogen, can also be produced from renewable biomass.
Biomass, produced as dedicated energy crops such as perennial grasses or short-rotation woody crops, could
supplement natural gas as a source of FCV fuel energy. Fuels produced from biomass have the important
advantage that their use contributes little to net greenhouse gas emissions. In FCVs, methanol produced from
biomass could reduce the total fuel cycle C02 emission by 86 % relative to conventional gasoline vehicles.
Because of their potential for reducing greenhouse gas emissions, large scale production of biomass in energy
plantations is under consideration for fossil fuel displacement. The main barrier to utilization of biomass, either
directly to displace coal for electric power production, or for conversion to liquid transportation fuel, is cost.
The object of this evaluation is to examine the optimal use of the two resources, natural gas and biomass to
achieve minimum impacts of road transport on the environment and least cost from the perspective of long-
term national interests.
Background
Several technologies, in varying stages of development, are under evaluation for production of
alternative transportation fuels such as alcohols or hydrogen from biomass. The main reasons for considering
biomass as an energy source for transport are to enhance national energy security and to reduce greenhouse
gas emissions (COz in this case). This report summarizes an EPA evaluation of one of those technologies that
appears to have potential for displacement of petroleum imports and reduction of C02 emissions from U.S.
highway vehicles. In making that evaluation, additional factors to be discussed in this section were considered
in terms of relative prospects for achieving national goals for alternative transportation fuels. Those goals
include benefits to the national economy and the reduction of environmental impacts of the petroleum fuels now
used almost exclusively. Environmental benefits in general, and greenhouse gas emissions in particular, are
the main focus of these process evaluations. Because of the magnitude of the fuel requirements for the U.S.
transportation sector, the way in which petroleum displacement is effected can have a large impact on each
of those national goals; a careful examination of options is therefore needed~not only with regard to alternative
transportation fuel, but also for the optimum use of available domestic energy resources for maximum overall
benefit.
The EPA Science Advisory Board (U.S. EPA, 1990) identified the risk of global climate change due to
emissions of greenhouse gases as one of the highest-priority environmental problems facing the Agency. Of
the U.S. anthropogenic C02 emissions-the predominant greenhouse gas-30 % are generated by the
transportation sector. The issue of climate change and mitigation of greenhouse gas emissions from mobile
sources favors an alternative fuel derived from renewable energy such as biomass because such fuels
contribute no net C02 to the atmosphere: C02 emitted by their combustion is withdrawn from the atmosphere
when the biomass is regrown. The other dominant source of anthropogenic carbon emissions is power plants,
for which the main technological approach for mitigation is separation and recovery of COa followed by
sequestration in the ocean or in depleted oil and gas wells. An international effort to evaluate the engineering
feasibility has been underway since 1992 (Blok et al., 1992).
Among the renewables, biomass is the only energy source that can be converted to liquid fuels
compatible with the existing vehicle refueling infrastructure and can provide a driving range comparable to
gasoline. Biofuels are cleaner than petroleum fuels: in addition to the potential for reducing greenhouse gas
emissions from mobile sources-a goal that has very limited technological options-alcohol fuels can reduce
toxic emissions by as much as 90 % per vehicle mile relative to gasoline. Because biomass suitable for fuel
production can be grown on land unsuited or poorly suited for food crops and on excess farmland, biofuels
have additional potential for creation of jobs in the domestic agricultural and industrial sectors in accordance
with other national goals.
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National Goals for Alternative Fuels
Although the magnitude and effects of future climate change are debated, the costs of current
petroleum dependence are already demonstrably real and severe (DeLuchi et al., 1987). Approximately half
of the 120 billion gallons of transportation fuel currently consumed annually in the U.S. is produced from
imported petroleum. As domestic supplies continue to decline, and consumption continues to increase (to an
estimated 160 billion gallons in the year 2020), imported petroleum and its cost, already $50 billion per year,
will impose a growing drain on the national economy. That drain now accounts for 75% of the trade deficit.
Global production of petroleum is also expected to start to peak as early as year 2000, followed by an annual
decline of 2.7%. Consumption of petroleum by Asian countries is expected to increase over 100 % over the
next 15-20 years, further adding to cost and supply pressures. In addition to the direct cost of petroleum fuels
which is reflected in the selling price, their indirect costs-which include environmental and health impacts-are
estimated at $30 billion per year. In recognition of these prospects, the DOE National Energy Strategy of 1991 -
92 called for the development of alternative fuels derived from domestic resources, including renewable
biomass, to begin displacing petroleum. If alternative fuels derived from biomass could relieve some of these
problems of petroleum dependence, they would provide a strong "no regrets" approach to the climate change
issue also (i.e., the results would be beneficial and desirable independent of the climate change issue). As
stated by the EPA Science Advisory Board: "Some risks are potentially so serious, and the time for recovery
so long, that risk reduction actions should be viewed as a kind of insurance premium and initiated in the face
of incomplete and uncertain data.... Preemptive actions are especially justifiable if they lead to unrelated but
immediate and substantial benefits, such as improved ambient air quality and reduced U.S. dependence on
imported oil." [emphasis added]
The Clean Air Act Amendments of 1990 addressed another problem associated with petroleum fuels--
the fact that they are primarily responsible for noncompliance with EPA air pollution standards for ozone and
carbon monoxide (CO). Motor vehicles using petroleum fuels are responsible for 80% of the total U.S.
emissions of CO and 45% of the anthropogenic emissions of nitrogen oxides (N0„) and volatile organic
compounds (VOCs). Reformulation of gasoline and addition of oxygenates, primarily methyl f-butyl ether
(MTBE) produced from methanol, is a first step that helps reduce the environmental problems of conventional
fuel, but does not address the long-term problems of national dependence on petroleum, including greenhouse
gas emissions. Car and truck fleets in severe ozone nonattainment areas are required to utilize clean fuels
than include alcohols and natural gas as well as reformulated gasoline.
The Energy Policy Act of 1992 established a goal of 30 % displacement of petroleum with alternative
fuels by the year 2010. Half of that fuel is to be derived from domestic resources. Although the Act identifies
no specific alternative fuel as the preferred choice, three objectives are given as primary considerations: (1)
it should achieve maximum displacement of imported petroleum, (2) it should have maximum benefit to the
national economy by minimizing the total cost to the Nation, and (3) it should reduce environmental impacts
of fuel production and use, including greenhouse gas emissions. The third criterion is a new factor that
responds to growing global concerns regarding climate change as reflected by the comments cited above by
the EPA Science Advisory Board and the initiatives developed by the Climate Change Action Plan of 1993.
The Climate Change Action Plan commits the U.S. to a goal of reducing emissions of greenhouse
gases to their 1990 levels by the year 2000. Two actions relevant to the transportation sector are outlined: a
natural gas strategy and a renewable energy strategy. Because domestic natural gas is abundant (160 x 101?
cubic feet of proven reserves, 1500 x 101z cubic feet of resource base) and emits less C02 per unit of energy
than other fossil fuels (14 kg carbon (C)/GJ, compared to oil with about 20 kg C/GJ and coal with about 25 kg
C/GJ), the EPA was directed to encourage the use of natural gas as a vehicle fuel to meet the pollution control
objectives of the Clean Air Act, especially compliance with urban ozone standards. Incentives to encourage
the use of renewable energy sources including biomass such as wood, carbonaceous wastes, and energy
crops, are outlined by the Energy Policy Act. These two strategies, use of natural gas and renewable energy
to displace petroleum for transport, are consistent with the recommendations of the Intergovernmental Panel
on Climate Change (IPCC, 1996).
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Given the objective of utilizing natural gas and renewable biomass to displace petroleum fuels, a
crucial problem is raised: how to achieve the greatest overall benefit from those resources in meeting the
national goals; not only greenhouse gas mitigation, but also energy security, economic stability, job creation,
and reduction of local air pollution. The possible solutions include the direct use of compressed natural gas
to displace gasoline and the conversion of biomass independently to liquid fuel. This, basically, is the current
concept, or leading option, for existing technology. Its implications will be examined and then compared with
the different option of using both resources to produce a single fuel. First, the effect of natural gas alone and
biomass alone.
Conventional vehicles powered by compressed natural gas could reduce the greenhouse effects of
their emissions by about 16% compared to vehicles using petroleum fuels (DeLuchi et al., 1987). That estimate
includes emissions from the production, distribution, and compression of the gas and also the emissions of
unburned methane, which is a much more efficient absorber of solar radiation, per unit of emission, than C02.
On the other hand, a fuel derived solely from biomass would have very little net C02 emission (which occurs
only from fossil fuels used for fertilizer production, biomass cultivation, or biomass transport). Comparisons
of the complete life-cycle greenhouse gas emissions (Michaelis, 1995), that include vehicle manufacturing and
fuel production, as well as vehicle emissions, indicate a potential reduction of C02-equivalent emissions per
vehicle mile of 13% for compressed natural gas and 74% for methanol derived from woody biomass. Such
comparisons suggest a strong advantage of biofuels for reduction of C02 emissions from a given vehicle of
conventional design, but provide no guidance as to the total effect that might be achieved with the highway fleet
as a whole, which would depend on the amount of biomass that could be sustainably produced for conversion
to transportation fuel.
The Biomass SudpIv Issue
As indicated above, comparisons of greenhouse gas emissions per vehicle mile are informative only
if the amount of biomass available for production of vehicle fuel is unlimited; otherwise, the amount of petroleum
that can be displaced, and therefore the total C02 emission reduction potential, will be constrained by the
biomass supply. Complicating this issue is the fact that biomass production is currently viewed primarily as
a coal replacement for electric power production. If biomass is to be considered a practical alternative energy
source, one must ask how much biomass might be produced as energy crops in the U.S., and for what
purpose: be it coal displacement or petroleum displacement, and at what cost for those options.
An assessment of the potential for energy crop production (Graham, 1994) indicates that (I) the total
U.S. land area that is suitable for perennial energy crops such as short-rotation woody crops, but not suitable
for conventional crops, is 15.4 million ha, (ii) the total excess cropland that is suitable for energy crops and
could be converted to that use without affecting domestic and export food production is about 16.2 million ha,
and (iii) on the total 31.6 million ha of suitable lank, about 13.5 Mg of dry biomass could be produced per ha-yr.
Assuming 10% loss in harvest, 7.4 EJ/yr (7 quads) of biomass energy could thus be produced. A more recent
assessment (Graham et al., 1995) estimates the total production of wood and grass crops at 5.8 EJ/yr (5.5
quads) for the target year 2020. The extent to which either of these projections might be realized depends on
the price that could be obtained per tonne of biomass produced and the crop yield per ha-yr because the
amount of suitable land that will be dedicated to energy crop production depends on the economic return to
the producer.
To achieve 5.8 EJ/yr of biomass energy production in the year 2020 will require an improvement of
current wood crop yield by a factor of 2.3. If that improvement is realized as anticipated (Graham et al., 1995),
the projected cost of woody biomass production will be $1,66/GJ ($1.75/million Btu). The most promising
herbaceous energy crop, switchgrass, is estimated to cost $1.7 to $2.8/GJ to produce, excluding transport.
If used for electric power production to displace coal, biomass would have to compete with coal on the basis
of energy content. The delivered cost of contract coal is about $1.14 to $1.5/GJ. Added to the constraint
imposed by this cost differential is the fact that a major portion of the land suitable for energy crop production
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will be too distant from any potential energy conversion plant site for that area to be utilizable from a logistical
standpoint if the usual method of transport by truck is the only option. Although coal is transported long
distances to power plants, it is generally loaded on unit trains at a single source; whereas, the sources of
biomass are diffuse, and biomass has a lower bulk and energy density, requiring greater volumes be
transported. Use of truck transport to railheads where it could be temporarily stored for long-distance shipment
by rail to conversion plants would be the most practical solution, but would require establishment of storage
and transshipment facilities at many locations. If economic factors are the only incentive for energy crop
production, it is unlikely that the displacement of coal by biomass could approach the 5.8EJ/yr potential. If the
target fuel price for electric power production is $1.33/GJ, the projected displacement of coal by biomass is only
about 0.53 EJ/yr (0.5 quad). In order to obtain 5.8 EJ/yr of fossil fuel displacement, a premium of S0.42/GJ
($0.44/million Btu) is required or about $4.4/tonne of C02 emission reduction from stationary sources.
A total of 5.8 EJ/yr of biomass is not sufficient to significantly displace both coal and petroleum as a
U.S. energy source. If biomass is to be used to displace fossil fuel for the purpose of reducing greenhouse
gas emissions, the option of displacing coal for electric power production needs to be compared with the
alternative option of its conversion to a fuel that can be used to displace petroleum for transport. Because the
Climate Change Action Plan identifies natural gas as part of the domestic resource base to be used for that
purpose in the transportation sector, it must be included in the comparison. Factors to be considered include
not only cost per unit of C02 emission reduction, but also the economic incentive for production of the maximum
biomass supply, ancillary environmental and economic benefits beyond reduction of greenhouse gas
emissions reduction, and impacts on energy security and domestic job creation. Finally, one must consider
a range of technological options that may be more appropriate to future vehicle designs.
Biomass as an Independent Source of Transportation Fuel
Given that the biomass supply will be a limiting factor for direct displacement of fossil fuel, how can
it be utilized to attain greenhouse gas reduction and other national goals most effectively? The best current
technologies for production of an alternative fuel from biomass are-in terms of production cost-the Battelle
Columbus Laboratory (BCL) process for methanol production and the enzymatic hydrolysis process being
developed by the National Renewable Energy Laboratory for ethanol production. Of the biomass energy
content, 50-60% is lost in conversion to ethanol and 39% is lost when converted to methanol by the BCL
process. One might therefore expect to displace less than 3.5 EJ/yr of petroleum with biogenic alcohol.
Marrison and Larson (1995) estimate the cost of alcohol production from biomass in the year 2020 to be $12.4
and $11.2/GJ for the minimum cost of methanol and ethanol, respectively, in the North Central area of the U.S.
where the major portion of land suitable for energy crop production is located.
Excluding natural gas for the moment, and assuming that biomass alone is converted to methanol by
the existing BCL process, the biomass cost accounts for about 38 % of the methanol production cost (Williams
et al., 1994). In terms of farmgate price, the $12.4/GJ production cost would sustain an average biomass cost
of $61/tonne or $3.36/million Btu. This biomass price would justify the full 5.5 quads of biomass production and
displace up to 3 quads of petroleum. The first requirement for establishment of a viable biofuel industry would
therefore be satisfied with enough price elasticity to provide a strong incentive for the fanner. This factor
therefore favors petroleum displacement. The current production cost of gasoline, $0.60/gal ($4.5/GJ, HHV),
is projected to increase to $0.90/gal by the year 2010. The cost differential for displacing 3 quads of petroleum
by the BCL process would be about $3.00/GJ, taking into account the relative efficiencies of methanol and
gasoline in internal combustion (IC) engines, or $0.40/gal of gasoline displaced, or $45/tonne of C02 avoided.
Thus, if natural gas is not considered as part of the strategy, if the analysis is confined to conventional vehicles
with IC engines, and if only the BCL process is considered for conversion of biomass to methanol, biofuel
cannot compete with petroleum fuels and the result would favor coal displacement. The situation is much
different when one takes into consideration new transportation technology currently under development based
on fuel cells. Fuel cell vehicles (FCVs) are expected to become commercialized within the next decade and
are now undergoing demonstration (as transit buses) in Chicago, Washington, D.C., and Vancouver, B.C.
Light duty vehicles should begin demonstrations within the next 2 years. These fuel cells require hydrogen,
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obtained either from compressed gas cylinders or from a liquid hydrogen-carrier that can be reformed onboard
to produce hydrogen, methanol being the most readily adaptable and lowest-cost liquid for this purpose. Using
hydrogen, these vehicles will be capable of achieving fuel economies about 3 times greater than the gasoline
vehicles now in use. Using methanol as fuel, the fuel economies are expected to be about 2.5 times that of
conventional vehicles (Williams, et al., 1994). The potential for reduction of petroleum requirements and
greenhouse gas emissions is directly related to their increased efficiency.
In thermochemical processes, natural gas and biomass are compatible with the production of either
methanol or hydrogen. Moreover, the production of these fuels by either feedstock is leveraged by the other
if used as co-feedstock in a single process, thus creating the potential for greater petroleum displacement than
either resource could achieve alone. Most important, perhaps, is the compatibility of these alternative fuels
with fuel cells which can further leverage petroleum displacement due to their increased thermal efficiency
relative to gasoline.
Impact of Fuel-Cell Vehicle Technology
As indicated by the preceding discussions, a sustainable road transport system must achieve greater
reductions of emissions, including C02, and greater reduction of petroleum dependence than will be possible
by incremental improvements in internal combustion engines. A 40-mpg fuel economy goal for such vehicles,
using petroleum fuels, would result in only modest improvements in the status quo and would not diminish
dependence on foreign oil producers, urban air pollution, or the threat of global warming (Kelly and Williams,
1992). Compressed natural gas is an alternative that is price-competitive with liquid fuels, but constrained in
driving range and refueling options; when used directly in conventional vehicles it could reduce CO emissions
by half, and greenhouse gas by about 17%, but would not significantly affect NOx.
Fuel-cell vehicles operating on methanol or hydrogen derived from natural gas would reduce
greenhouse gas emissions by about half and would effectively eliminate all criteria pollutant emissions
including CO, VOC, NO*, and particulates. If the methanol or hydrogen is derived from biomass, greenhouse
gas emissions per vehicle would be reduced by about 90%. As already discussed, however, the amount of
biomass available will limit the size of the vehicle fleet that could be fueled by that source of energy, therefore
restricting the total reduction of greenhouse gas emissions from road transport as a whole. The cost of that
fuel, $12.4/GJ, will also be non-competitive in a free market when used in conventional vehicles such as
flexible fuel vehicles (FFVs) and would therefore be unable to become established as a viable option. It would
also have difficulty competing with FCVs equipped with partial-oxidation reformers that will use gasoline as the
hydrogen source. It will be shown here that use of natural gas as co-feedstock for production of methanol or
hydrogen from biomass will reduce the production cost to the point where these fuels will cost substantially less
than gasoline used in conventional vehicles and about the same as gasoline used in FCVs. The two existing
barriers to use of alternative transportation fuel from biomass-cost and supply-will no longer exist when FCVs
displace internal combustion engine vehicles (ICEVs)--as they must, if the system of road transport is to be
sustainable-and the advantages of FCVs with regard to reduction of greenhouse gas emissions and reduced
petroleum dependence in particular, will be further enhanced.
When these factors are taken into account, it will be shown that the result not only favors petroleum
displacement for the purpose of reducing greenhouse gas emissions, but also displaces gasoline at no cost
while achieving greater C02 emission reduction from the vehicle fleet than can be expected from biomass or
natural gas alone. Consequently, the option to be considered will also be expanded to include the likelihood
that FCVs will be successfully developed and become commercially available in the same time frame as the
alternative fuel technology.
Methodology and Assumptions of this Analysis
Given the assumption that reduction of greenhouse gas emissions is a continuing national objective
and that natural gas and biomass are to be utilized for the purpose of achieving that objective, the purpose of
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this evaluation is to assess the most effective way in which those two resources can be used to maximum
effect and minimum cost. Three basic approaches will be considered in detail: first, the direct conversion of
biomass alone to liquid transportation fuel by the best current technology; second, the production of
transportation fuel from natural gas alone; and third, the production of liquid fuel from both biomass and natural
gas in a single process. The existing process options will thus be broadened to include the use of both natural
gas and biomass as feedstocks for production of a single transportation fuel that is compatible with FCVs. As
already indicated, this alternative option has been determined to be the most effective means of displacing
petroleum and reducing overall greenhouse gas emissions, and this report is intended to document the basis
for that conclusion.
A process specifically designed to utilize natural gas and biomass as co-feedstocks is the principal
focus of this assessment. That process, called Hynol, was conceived at Brookhaven National Laboratory
(Steinberg and Dong, 1994) and is the main focus of this study. The Hynol process has been under evaluation
by APPCD, using the process simulation software Aspen Plus (1995), In many configurations using broad
ranges of operating assumptions. This report will document those simulations and the data that support the
conclusions regarding thermal efficiency, energy ratio, net C02 emission reduction potential, and cost. Those
data will then be used to compare the Hynol process with other options that can utilize both biomass and
natural gas, either in separate processes, or as co-feedstocks for a single process, to produce methanol or
hydrogen as transportation fuel. The results of these comparisons are expected to provide a starting point for
. future engineering design studies of the Hynol process that will more fully assess its relative cost compared
to other alcohol production technologies, that should include the enzymatic hydrolysis process for ethanol
production.
Another approach is to modify a thermochemical process that was designed to produce methanol
or hydrogen from biomass alone so that natural gas may be added to increase the output of transportation fuel.
Both of these options will be evaluated and compared with the alternative options of using the same natural
gas and biomass in separate processes.
Because the Hynol process has not been demonstrated in hardware at any scale to provide operating
data, a number of assumptions are made regarding the effects of such variables as carbon conversion during
gasification, steam/carbon ratio required for steam reforming of the gasification products without carbon fouling,
pressure drops in the various units, and approach to equilibrium in the reforming and methanol synthesis steps.
In order for comparisons between different processes to be meaningful, the same assumptions were made
regarding these variables wherever possible. Performance estimates are made as a function of the expected
range of the critical operating variables.
The biomass feed was assumed to be a short-rotation woody crop produced specifically for alcohol
production and dried to a moisture content of 10 wt% by a circulating steam drier that has no direct atmospheric
emission. The compositions of the biomass and the natural gas assumed for all cases are given in Tables 1-1
and 1-2, respectively.
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TABLE 1-1. COMPOSITION OF BIOMASS FEED ASSUMED FOR THIS STUDY
Basis: 100 kg dry biomass containing 10 wt% moisture
Component
kg
kg-mols
C
51.07
4.255
H
7.48
3.74
0
51.52
1.61
N
0.15
0.00536
S
0.08
0.0025
Ash
0.79
Total
111.09
TABLE 1-2. COMPOSITION OF NATURAL GAS ASSUMED FOR THIS STUDY
Component
Mol fraction
I
o
0.947
C2H8
0.028
co2
0.002
N,
0.023
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SECTION 2
CONCLUSIONS
General
A sustainable system of road transport for the future will require a large reduction of the environmental and
economic impacts of the current vehicle and fuel technologies.
Displacement of imported petroleum with natural gas as transportation fuel can reduce the economic burden
and environmental impact of the current transportation system, but its domestic supply is finite and its
marketability is constrained by limited range and refueling problems.
- Any alternative transportation technology for the future must take into account greenhouse gas emissions;
emissions from road transport are currently a major part of the total anthropogenic C02 emissions and are the
fastest growing source category.
Clean alternative fuels can be produced from renewable biomass produced on dedicated energy plantations,
and such fuels could eliminate most of the net overall fuel-cycle C02 emissions from vehicles currently using
petroleum based fuels.
Liquid alternative fuels produced from biomass can provide the driving range and compatibility with the existing
refueling infrastructure that are needed to compete with existing fuels, but cannot be produced at a cost that
would allow wide acceptance in a free market if produced from biomass alone.
Even if marketable, the amount of bio-fuei that could be produced on a sustainable basis from energy crops
alone would not be sufficient to displace enough petroleum to appreciably affect greenhouse gas emissions
from U.S. road transport.
The limitations of biomass and natural gas as individual sources of alternative fuels could be overcome if those
resources are used together in a process to convert them to a liquid fuel such as methanol.
Any alternative fuel considered for a sustainable future transportation technology should be compatible with
fuel-cell vehicles. More than any other factor, vehicles powered by fuel cells offer the greatest potential for
increased fuel economy and reduction of vehicle emissions, including greenhouse gases.
Fuel-cell vehicles will require hydrogen fuel, either as compressed gas or as a liquid hydrogen-carrier.
Methanol is the liquid hydrogen-carrier most easily and efficiently converted to hydrogen on board a vehicle
and can be used directly as the hydrogen source in some types of fuel ceils now being developed. It can be
produced from biomass, from natural gas, or from a combination of both resources.
Methanol produced from biomass and natural gas as combined feedstock will be competitive with gasoline
costs and provide major reduction of net C02 emissions as well as elimination of particulates and other criteria
pollutant (CO, VOC, NOx) emissions when used in fuel cell vehicles.
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Using natural gas as the primary energy source, and biomass providing 25% of the energy input for methanol
production, the amount of petroleum displaced would be leveraged by a factor of 4.8 relative to the use of
biomass alone, and the production cost would be reduced by 40%. Overall C02 emission reduction would be
twice as great as the potential for C02 reduction using fuels produced from biomass alone.
Thermochemical processes can produce either methanol or hydrogen fuel from biomass for use in FCVs; the
yield of either fuel can be leveraged and its cost reduced by using natural gas as co-feedstock.
Municipal carbonaceous wastes such as sewage sludge, landfill gas, digester gas, and solid wastes normally
sent to landfills should be acceptable co-feedstocks with energy crops for production of methanol or hydrogen.
Landfill gas or digester gas can displace part of the natural gas used by the process, limited only by the amount
of such waste gas available at a given location.
The BCL process, using only biomass and no natural gas, will yield 14.77 kg-mols of methanol per tonne of
biomass at a cost of $12.4/GJ. Used in FCVs, that methanol will displace 194 gallons of gasoline with a net
overall fuel-cycle C02 emission reduction of 2.04 tonnes.
The thermal efficiency of the BCL process is 51.1% when its full electric power requirements are supplied by
combustion of biomass; i.e., no electric power is imported.
Natural gas can be used as co-feedstock to enhance the methanol yield of the BCL process by adding it, with
steam, to the gasified biomass entering the reforming step. The result will eliminate the need for a shift reactor
and Selexol unit normally required by a BCL system that uses biomass only. Import of electric power, or use
of biomass to produce electric power, for the process will also be no longer required. Overall efficiency will be
marginally less than the Hynol process, but should not require development or demonstration of new
technology.
Conclusions Specific to Hvnol
The Hynol process, using 1 tonne of biomass and 6.276 kg-mols of natural gas, will yield 72 kg-mols of
methanol at a cost of $6.09/GJ ($0.42/gal) and displace 944 gallons of gasoline with a net overall fuel-cycle
COz emission reduction of 7.24 tonnes.
The thermal efficiency is estimated at 68.4% when the process is configured to require no electric power
import.
Optimum performance with respect to overall net fuel-cycle C02 emission reduction from FCVs using methanol
will be obtained with natural gas comprising 75% of the Hynol feedstock energy.
Biomass carbon conversion as low as 82% can be accepted if un-gasified carbon is utilized as part of the fuel
for the reformer furnace. This can be accomplished by using it in a separate furnace to preheat combustion
air.
Desulfurization must take place prior to reforming to protect the reformer tubes and catalyst. Zinc oxide at
268°C will reduce the equilibrium hydrogen sulfide concentration entering the reformer to the required level of
0.02 ppm.
Maximum thermal efficiency will require recovery of heat as high-pressure steam between the gasifier at 800°C
and the desulfurizer at 268°C. This can be accomplished only if tars are not formed during biomass
gasification.
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Acceptable performance will be obtained with a reformer operating within -30K temperature-equivalent
approach to equilibrium.
Acceptable performance will require that the methanol converter operate at a temperature-equivalent approach
to equilibrium not exceeding +18K. This is a critical variable, and the design value should be +12K or less.
The Lurgi type methanol reactor is superior to the ICI design for this system because the medium-pressure
steam recovered provides a significant amount of the electric power required by the process.
The net overall C02-equivalentfuel cycle emission reduction would be 20% more than could be achieved if the
biomass were converted to methanol by the BCL process and the natural gas were converted to methanol by
conventional steam reforming.
Compared to gasoline used in current ICE vehicles, methanol produced by the Hynol process and used in
conventional flexible-fuel vehicles would cost $91 per tonne of C02 emission avoided. Used in dedicated
methanol vehicles, the mitigation cost would be $52/tonne of C02 emission avoided. When fuel-cell vehicles
become long term replacements for ICE vehicles, total fuel-cycle C02 emissions can be reduced 65 % at no
cost; a cost saving of $34/tonne C02 could actually be realized from the transportation sector for trading against
other emissions from stationary sources, the expected mitigation costs for which are expected to be at least
$72/tonne.
Compared to fuel-cell vehicles using gasoline and partial oxidation reformers, methanol FCVs will achieve 39
% greater C02 emission reduction at no greater fuel cost per mile traveled.
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SECTION 3
RECOMMENDATIONS
This analysis indicates that biomass, as a limited energy resource, would be most effectively utilized for
production of transportation fuel to displace petroleum rather than used as a boiler fuel for coal displacement
The sustainability of road transport, in terms of minimization of its environmental impacts by means of
technological improvements, will be most effectively realized by the introduction of fuel-cell vehicles. The
greatest environmental benefits will accrue if those vehicles use non-petroleum fuels, especially those derived
at least in part from biomass. Of all the federal agencies involved in transportation R&D, none has a greater
stake (in terms of vital interest in the outcome) than the EPA in the type of vehicle and fuel technologies that
- eventually replace those now in use for road transport.
From the standpoint of greenhouse gas emission reduction, the analysis of technologies undertaken here
indicates that processes that can utilize both natural gas and biomass for production of fuel for fuel-cell vehicles
would be substantially superior to processes that would utilize biomass and/or natural gas as separate sources
of fuel energy. Consequently, any mitigation strategy for greenhouse gases should be focused on alternative
fuels (methanol or hydrogen) produced by the dual feedstock.
The Hynol process appears to be superior to other technologies for production of fuel from natural gas and
biomass. Its advantage is not large, however, and may disappear upon further optimization of other processes
that could use biomass gasifiers that are already demonstrated (or upon field assessment of Hynol
performance in actual hardware).
This report is based on computer simulations and laboratory data. Although this procedure is well accepted
engineering practice, it does not substitute for pilot testing. Technical assumptions made for this assessment
concerning the performance of the Hynol process that need to be verified by actual tests include the ability to
operate a reformer at 32 atm pressure and 950°C, attainment of a carbon conversion of 87% in the gasifier,
tie feasibility of operating a ceramic (or other type of heat exchanger) at inlet temperature of 950°C and 32 atm
on the hot side and 110 to 900°C and 42 atm on the cold side. Most important is the assumption that no tars
will be formed during biomass gasification. If tars and/or corrosive alkali deposits occur in the heat-recovery
steam generator between the gasifier and desulfurizer, the thermal efficiency will be reduced and the system
may require import of electric power.
The option by which natural gas might be utilized in existing biomass-to-methanol and biomass-to-hydrogen
processes needs to be further examined by process simulation to establish the potential for C02 emission
reduction and cost reduction relative to the Hynol process. Effects of the operating and design variables should
be further assessed, including: the reforming temperature, pressure drop between the reforming step and the
methanol converter, the methanol synthesis pressure, and the ratio of natural gas fed to the process to that
used as reformer fuel. Optimization of these variables could yield substantial improvement.
The conclusions of this study are potentially significant with regard to national policy objectives for
environmentally sustainable road transport and the potential for future compliance with EPA's air standards
for ozone and particulates in urban areas. Relative to other options that require changes in public preferences
or behavior, or changes in land use policy, or economic disincentives for demand, the technologies discussed
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represent an assured means of attaining a sustainable road transport system for the foreseeable future with
optimal utilization of resources that will achieve maximum environmental benefits at least economic cost.
Consequently, the preliminary cost estimate and process evaluations made here need to be carried out in
further detail by an architectural and engineering firm to confirm those conclusions independently.
Use of biomass to any significant degree as a source of transportation fuel will require transport of biomass
over much longer distances than are practical or economical by truck. Transport by train, similar to the
methods used for coal, must be considered for biomass if a significant portion of the land that is suitable for
production of energy crops is to be utilized for that purpose. It is therefore recommended that an independent
study be performed to assess the modifications required of the current rail and agricultural transport system,
if any, and the methods of storage and cost for delivery of biomass to large, centrally located plants in various
regions of the U.S. where natural gas is also available.
3-2
-------
SECTION 4
THE BCL PROCESS
Background
This section discusses what is considered to be the best current example of technology for production
of methanol or hydrogen from biomass without natural gas addition. The "BCL Process" utilizes a biomass
gasifier developed at the Battelle Columbus Laboratory as documented in the report 'Conversion of Forest
Residues to a Methane-Rich Gas in a High-Throughput Gasifier" { Feldmann, et al., 1988). That gasifier has
been tested successfully at the 10-ton/day level in a 10-in diameter gasifier at the Battelle Laboratory and is
currently under evaluation at the 200-ton/day level at the McNeil Generating Station in Burlington, VT,to
produce medium Btu fuel for gas turbines. A detailed analysis and cost evaluation of the Battelle gasifier
incorporated into a process to produce methanol or hydrogen was carried out by the Princeton Center for
Energy and Environmental Studies (Katofsky, 1993). The integrated system proposed by Katofsky, which
- includes the Battelle gasifier and the additional process steps for production of methanol, is discussed in
Appendix A.
Although Katofskys report and subsequent paper (1994) provide most of the technical detail needed
to assess the performance of the BCL process with regard to thermal efficiency-given the assumptions on
which it is based-several important factors are not addressed in sufficient detail to make meaningful
comparisons with other processes or other configurations of the BCL system. Not included, for example, are
the configuration of, energy balance on, and heat recovery from, the biomass gasifier. Also not identified are
the energy requirements for biomass drying, methanol/water separation by distillation, or the method of
integration of the overall process for maximum thermal efficiency. The description of the BCL system proposed
by Katofsky has been recalculated in Appendix A to identify and quantify the energy requirements of each step
of the process. It shows that the operating conditions specified by Katofsky did not provide adequate fuel to
satisfy the energy requirement of the reforming step: a 19.1% shortage. The BCL process discussed in this
section of the report has therefore recalculated the energy balance to fully assess the duty of each component,
including biomass gasification, biomass drying, steam reforming, and electric power inputs and outputs. The
basis for these calculations is explicitly defined as 100 kg of dry biomass feed to the gasifier, including that
biomass required as combustor fuel, but not including any biomass that may be necessary as additional fuel
to generate electric power if the power produced within the process is insufficient to meet process needs (this
must necessarily be calculated after the fully integrated process has been evaluated and will be the final step
of this analysis).
Some other factors that should be noted are: (1) The steam/carbon ratio in the gas stream fed to
the reformer remains at the same value as the original Katofsky version: 1.13 mols steam/kg-atom carbon.
This ratio may not be sufficient to avoid carbon deposition on the reforming catalyst. Conventional natural gas
reformers require steam/carbon ratios in the range 2.5 to 3.5; and (2) In order to provide sufficient thermal
energy to dry the biomass (considered here to initially contain 48.4-wt % moisture) to 10 % moisture, all of the
heating value of the combustor flue gas that is not required for air preheat is used for biomass drying, thus
preventing recovery of electric power from that high temperature stream. In addition, all available low
temperature heat must be recovered for the biomass drier. This energy is recovered with heat exchangers that
heat air for partial drying of the biomass at 35°C prior to final drying with flue gas. Recovery of this heat with
air imposes a significant drain in electric power for air blowers. The thermal energy required to dry 100 kg
biomass containing an initial moisture content of 48.4% to a final moisture content of 10% is 56525 kcal, as
shown in Appendix B. The other details of the BCL process steps are based on the same assumptions as
Katofsky's regarding temperatures, pressures, inputs, and outputs.
4-1
-------
Th? pipmass Softer
TTie BCL gasifier consists of an entrained flow reactor into which dried biomass and steam are injected.
It operates at atmospheric pressure and 927°C. The heat required for gasification is provided indirectly by an
external combustor that is fueled with unreacted char from the gasifier and additional fresh biomass. When
burned with air, these fuels heat sand to 1094°C which is circulated through the gasifier to provide the thermal
energy required for endothermic gasification reactions. The sand is then returned to the combustor. Figure
4-1 shows the material balance on the gasifier-combustor system based on a total biomass feed of 111.1 kg,
including 11.1 kg of biomass moisture (10 wt %). Details of the calculations of the energy balance are given
in Appendix A. Of the 100 kg of dry biomass used, 84.48 kg is fed to the gasifier and 15.52 kg is fed to the
combustor as fuel. Additional combustor fuel enters from the gasifier as unreacted char (carbon conversion
in the gasifier Is assumed to be 88 % as specified by Katofsky). Sand is circulated between the combustor and
gasifier units at a rate of 1420 kg per 100 kg of dry biomass fed to the gasifier. Flue gas leaving the combustor
at 623°C is fed to the biomass drier.
Proves? PIqqH Flpyv piagrgm
Figure 4-2a, -2b shows each of the process steps involved in conversion of the gasified biomass to
methanol based on 100 kg of total dry biomass fed to the gasifier per hour and 5.394 kgmols of gasified
' biomass delivered as process feed. Temperatures, pressures, and enthalpy changes for each step are noted.
Table 4-1 shows the composition of the principal process streams.
Steam Reforming Block
The product stream from the gasifier is cooled from 927 to 400° C in a heat-recovery steam generator
(HRSG-1) producing high pressure (HP) steam from boiler feed water (BFW) preheated to 140°C in the
methanol synthesis block. The effluent gas is quenched, at conditions specified by Katofsky, in a scrubber that
removes particulates, tar, and sulfur. A three-stage compressor raises the pressure from atmospheric to 2.4,
5.7, and 15.3 atm with interstage cooling to 40°C. The third stage effluent, at 148°C, is fed to the reformer with
added steam.
Figure 4-3 gives details of the steam reformer block. Thermal energy required for the endothermic
reforming reactions is obtained from two sources, 0.307 mol of purge gas from the methanol unit and 0.603
mol of fuel gas extracted from the gasifier after heat recovery. These fuel gases are preheated to 290°C and
burned with preheated air (15% excess) to provide the thermal input required by the steam reforming reactor.
Heat-Recoverv Steam Generator Block
Process gas leaves the steam reformer at 867°C and 14.3 atm, and passes through a series of heat
exchangers, Figure 4-4, in which the gas is cooled to 337°C. These heat exchangers generate additional
steam which, after addition of the steam from HRSG-1 and superheating, passes through turbines to produce
some of the electric power required by the process. Part of the turbine exhaust is reheated and fed to the
reformer and shift reactor as process steam. Preheated boiler feed water is obtained from the heat exchanger
HE-2 on the effluent stream of the methanol reactor.
Distillation Block
The temperature at which the process gas leaves the heat-recovery steam generator block is
determined by the enthalpy required of that stream to satisfy the thermal inputs needed for separation of
methanol and water by distillation at a reboiler temperature of 103°C and delivery to the shift reactor at the
required temperature of 400°C. Using a 15-stage distillation column operating at atmospheric pressure with
a recycle ratio of 0.22, and a distillate-to-feed ratio of 0.94, the reboiler requires 18,090 kcal, or 5025 cal/sec.
4-2
-------
Process gas to
reformer
927°C. 1 atm
Kamois
M
-------
5.7 atm
HzO 0.158 mol
2.95 kW
Air
3.43 mols
H,0 2.39 mols
3.42 kW
40°C
148°C
15.5 atm
2.4 atm
Purge
r h20
0.0625 mol
2.88 kW
HjO, 20°C
2.07 mols
82°C
1 atm
Steam
3.232 mols
527°C
Quench
HzO
0.792 mol
Air
6.20 mols
Shift
Steam
0.667 mol
387°C
ca l/sec
Com bust or
Biomass
84.48 kg
Biomass
15.52 kg
Steam
QE)—
Reformer
+13074
ca l/sec
HRSG-1
-8078
ca l/sec
HRSG-2
-10271
ca l/sec
Steam
1.478 kgmols
Figure 4-2a. Overall block flow diagram for methanol production by the BCL process.
Basis: 100 kg total dry biomass fed to gasifier per Figure 4-1.(continued).
4-4
-------
2.06 kW
Methanol
synthesis
HE-A3
-4762
cal/sec
10t
HE-MC
-4190
cal/sec
HRSG-3
-4633
cal/sec
27°C
4761
cal/sec
\^J Condenser
HE-BF
-2063
cal/sec
CH3OH
1.601 mols
HE-A2
910
cal/sec
104.5 atm
4.1 kW
Distillation
H20
0.101 mol
3.26 kW
-851
cal/sec
24 atm
-883
cal/sec
48 atm
Selexol
3.56 kW
HE-A1
1913
cal/sec
Reboiler
5025
cat/sec
C02 1.0816 mois
~ H20 1.44 mols
HjO
0.9 mol
Figure 4-2b. Continuation of Figure 4-2a.
4-5
-------
TABLE 4-1. BCL PROCESS STREAM COMPOSITIONS OF FIGURE 4-2, kg mols
Stream Number
Component
1
2
3
4
5
6
H,0
1.647
0.2644
0.2665
3.277
1.931
2.34
H,
0.777
0.0702
0.707
0.707
3.22
3.48
CO
1.728
0.1562
1.572
1.77
1.515
1.515
o
a
0.416
0.0376
0.378
0.952
1.208
0.126
X
o
0.589
0.0532
0.536
0.536
0.165
0.165
C,H4
0.194
0.0175
0.176
0
0
0
C,H,
0.027
0.0024
0.0246
0.0246
0
0
N,
0.016
0.0014
0.0146
0.0146
0.0146
0.0146
CH,OH
0
0
0
0
0
0
SUM
5.394
0.603
3.675
6.686
8.05
8.72
°C
927
82
40
585
867
373
|_ atm
1.0
1.0
1.0
14.8
14.3
13.8
(continued)
4-6
-------
TABLE 4-1 (Continued). BCL PROCESS STREAM COMPOSITIONS OF FIGURE4-2,kg mols
Stream Number
Component
7
8
9
10
1'
H,0
0
0.00135
0.108
0
0.00135
H,
3.476
15.32
12.0
0.1462
11.85
CO
1.515
2.36
0.859
0.0105
0.848
CO,
0.126
0.542
0.436
0.0051
0.416
X
o
0.165
10.80
10.81
0.131
10.64
C,H,
0
0
0
0
0
C,H.
0
0
0
0
0
N,
0.0146
1.05
1.059
0.0128
1.035
CH,OH
0
0.0866
1.697
0.00107
0.0866
SUM
5.30
30.2
26.95
0.307
24.9
°C
40
83
260
27
27
atm
12.8
104.5
96.0
94.9
94.9
4-7
-------
C02 0.441 mols
HjO 0.919
N2 2.795
O, 0.100
Flue gas
Process gas
4.255
Fan
1.24 kW
69°C
Air
3.52 mols
APH
Fuel gas
235°C
Purge
FPH
325'C
PPH
Steam
3.23 mols
527°C
1956°C
Reformer
To
HRSG-2
Stream composition, kq-mols
Stream
1
Stream
2
Stream
3
Stream
4
Stream
5
H,0
0.2665
3.270
0.2644
0
1.931
H,
0.707
0.707
0.0702
0.167
3.22
CO
1.572
1.572
0.1562
0.0128
1.77
CO.
0.378
0.378
0.0376
0.0069
0.952
CH,
0.536
0.536
0.0532
0 133
0.165
C,H.
0.176
0.176
0.0175
0
0
C.H.
0 0246
0.0246
0.0024
0
0
N,
0.0146
0.0146
0.0014
0.0129
0.0146
CH,OH
0
0
0
0.0011
0
SUM
3.675
6.686
0.603
0.334
8.05
Deq. C
313
585
82
27
1140
Atm
1.0
14.8
1.0
94.9
14.3
HEAT DUTIES, BCL STEAM REFORMING BLOCK
ux
Description
Duty, caVsec
APH
Air preheater
1042
FPH
Fuel qas preheater
835
PPH
Process stream preheater
4415
SPR
Steam reformer
13074 i
Skin loss
444 |
Figure 4-3. Details of steam reforming block of the BCL process.
4-8
-------
Steam 3.16 mols
310°C
BFW 140°C
400°C, 1 atm
HRSG-1
Steam
to shift reactor
0.667 mol
Steam
310°C
97.3 atm
BFW 1403C
1.72 mols
769°C
HRSG-2
15.5
Atm
303°C
Steam 527°C
3.23 mols
to reformer
-©
0.05 atm
7.15 kW
4.25 kW
HEAT DUTIES OF REFORMER HEAT-RECOVERY
AND POWER-BLOCK UNITS
Unit
Duty, cal/sec
Heat recovery steam qenerator, HRSG-1
8076
Heat recovery steam qenerator, HRSG-2
4831
Reformer steam feed superheater
1167
Turbine steam reheater
829
Turbine steam superheater
3453
Figure 4-4. Details of the heat recovery steam generator arid power block of the BCL process.
Feed received from the reformer block. Figure 4-3.
4-9
-------
The hot gas leaving the reboiler passes through a heat exchanger that preheats the crude methanol/water
mixture to 65°C before it enters the distillation column.
The process gas stream leaving the distillation unit at 137°C is cooled to 127°C in a condensing heat
exchanger, HE-D1, before entering a Selexol scrubber which removes 90% of the C02 and all of the remaining
water vapor at 12.8 atm pressure and 127°C. The spent scrubber solution is regenerated by depressurization.
The process gas stream is finally cooled to 40°C in another heat exchanger, HE-D2, prior to compression for
methanol synthesis. Both HE-D1 and HE-D2 are air cooled, the heated air being used for biomass drying.
Figure 4-5 summarizes the operating data for the distillation block.
Methanol Synthesis Block
The synthesis gas from the distillation block is compressed in three stages to 24, 48 , and 104.5 atm,
with interstage cooling to 40°C, and enters the methanol synthesis recycle loop at 130°C. The methanol
synthesis block is shown in Figure 4-6. The reactor is thermally neutral, cooled by the recycle stream. Part
of the reactor feed is preheated to 250 °C as required to initiate the reaction at the entrance. Heat exchanger
HRSG-M1 recovers energy from the hot reactor effluent to produce low pressure steam for the gasifier. Heat
exchanger HE-M1 preheats boiler feed water to 140°C for the steam generators; and heat exchanger HE-M2
heats air for the biomass drier. The methanol condenser, HE-MC, is water cooled; part of the cooling water
constitutes the feed for HE-M1 and -M2. The purge stream is fed as part of the fuel, to the combustor in the
reformer block.
Biomass Drier
Woody biomass normally has an initial moisture content in the range 45 to 50 wt%. As shown in
Appendix B, drying 100 kg of biomass from an initial moisture content of 48.4 wt% to a final moisture content
of 10 wt% requires 56537 kcal. For the BCL process, all available low-temperature heat must be recovered
for the drying step. From HE-M2,40 mols of air at 95°C is obtained as waste heat from the methanol synthesis
block. An additional 17.7 mols of air at 117 °C is obtained from heat exchanger HE-A2 as waste heat from the
distillation block. These combined air streams can dry the biomass to a moisture content of 46.4%, discharging
moist air at 41 °C. To achieve additional drying, flue gas from the gasifier/combustor must be used. Because
this flue gas must be discharged at a higher temperature (120 °C), a second stage of drying is necessary as
indicated in Figure 4-7. The second stage cannot discharge wet flue gas at 120aC and dry the biomass to a
final moisture content of 10 % if the original biomass moisture exceeds 46.9 %, which is the value assumed
for the results shown in Figure 4-7. The dried biomass, discharged at 100 °C, is assumed to be fed to the
gasifier without cooling. Given the recoverable heat sources indicated, a maximum initial moisture content of
the biomass is limited to 46.9%.
BCL Process Performance Evaluation
Based on the preceding material and energy balances on each operating block of the BCL process,
both the total electric power required to operate the system and the amount of electric power that it can
produce can be determined. The thermal efficiency of the process is defined by these values; if the electric
power required is greater than that produced, then an import of electric power is required and an additional
amount of biomass must be used as fuel for production of that power-assuming that the process is to remain
neutral with regard to C02 emissions. It is this final value of total biomass (biomass gasified + biomass burned
as fuel for the gasifier combustor + biomass burned as fuel to generate electric power) that determines the
methanol yield of significance: the mols of methanol per kg of total biomass used by the process. The following
procedure is followed to obtain that yield.
Basis: One hour operation, 100 kg dry biomass fed to gasifier and gasifier-combustor.
4-10
-------
CHjOH
1.609 mols
Distillation
From HRSG-2
337*0, 13.8 atm
6.51 mols
~ H20 0.106 mol
Steam, 387°C
0.667 mol
Shift
From crude
methanol condenser
27°C
Reboiler
HE-A2
Air to drier
117-0
4C°C
HE-A1
Selexol
0.38 kW
0.80 kW
Air 25'C
5.7 mols
127°C
12.8 atm Vj
Air to drier
117°C
C02 1.0816 mols
HjO 1.44 mols
H20
* 0.90 mol
HEAT DUTIES OF METHANOL DISTILLATION-BLOCK UNITS
| Unit
Duty, cal/sec
Distillation column reboiler
5025
Column crude methanol feed preheater
341
Column methanol condenser
5013
Drier air heater, HE-A1
1913
Drier air heater, HE-A2
910
Figure 4-5. BCL process; details of the methanol distillation block.
4-11
-------
30.2 mols. 50°C
Feed from compressor
104.5 atm, 130°C
BFW 145°C
BFW 40°C
5 mols
Methanol
synthesis
2.06 kW
4.93 mols
94.9 atm
HE-M2
95.5 atm
70CC
HRSG-3
HE-A3
Condenser
HE-MC
10.0 mols
250°C
22.47 mols
Steam to
gasifier
1.428 mols
157CC
CH3OH 1.609 mols
HjO 0.106 mot
2.65 kW
Crude methanol
to distillation
Air 40 mols
25'C
BFW 40°C
Air to drier 95°C
HEAT DUTIES OF METHANOL BLOCK UNITS OF BCl PROCESS
Unit
Duty, cal/sec
Heat recovery steam generator, HRSG-3
4633
Drier air heater. HE-A3
4762
Boiler feed water heater, HE-M2
2063 |
Crude methanol condenser, HE-MC
4190 I
Methanol converter feed heater
4761 I
Figure 4-6. BCL process; details of the methanol synthesis block.
4-12
-------
Biomass 100 kg
88.5 kg moisture (46.9 wt%)
25°C
Air 40 mols, 9S°C
from methanol synthesis
HE-A3
Air 17.7 mols, 117°C
from distillation block
HE-A1, HE-A2
_~ Air, 41 °C, 57.7 mols
1.753 mols H20
Flue gas 6.766 mols, 623°C
from gasifier/combustor
Dner, stags 2
Rue gas, 120°C
CO,
HjO
Biomass 100 kg, 100°C
Moisture 11.1 kg
1.083 mols
0.194
4.881
3.151
Figure 4-7. Biomass drying operation using heat recovered from the BCL process.
4-13
-------
Electric Power Required by the Process:
Feed compressor (3-stage, 40°C intercool) 9.25 kW
Main (methanol synthesis) compressor (3-stage, 40°C intercool) 10.92
Recycle compressor 2.0
Reformer combustor air blower 1.21
Gasifier-combustor air blower 2.15
Boiler feed water pump 0.8
Drier air blowers 3.83
Condenser pumps (methanol unit, distillation unit) 0.2
Total electric power needed 30.36 kW
Electric Power Available:
Heat recovery steam generators, HRSG-1 and -2 7.15 kW
4.25
Total electric power produced 11.40 kW
Imported electric power required = 30.36 -11.4 = 18.96 kW, or 62.5% of the total electric power used by the
process.
Energy ratio (HHV of methanol produced/HHV of biomass fed to the gasifier):
r . 1.609 x 172000 AAni
Energy ratio - = 0.601
100 x 4604
Thermal efficiency:
The thermal efficiency discounts from the energy ratio that amount of energy that must be imported
as electric power. If the process is to operate with no net C02 emission, this power must be generated on site
from biomass. It is assumed, in accordance with Katofsky, that this power is generated by a gas turbine
combined cycle using gasified biomass with an efficiency of 39.58%.
Thermal efficien = (16Q9 * 172000) - [(30.36-11.4)860.5/0.3958]
erma efficiency 100x 4604
= 0.5116
If the imported energy (41220 kcal) is to be obtained from biomass combustion, the additional biomass required
to produce that power is:
41220 kcal x —^^ biomass _ g ^ ^ biomass
4604 heal
and the final yield of methanol per 100 kg of total biomass utilized by the process is:
1 509 x ^ { 477 ^ mo Is MeOH
108.95 100 kg dry biomass
4-14
-------
SECTION 5
THE HYNOL PROCESS
General Description
The biomass gasification that is the main subject of this report proposes the use of natural gas as an
integrated co-feedstock. It was conceived at the Brookhaven National Laboratory (Steinberg and Dong,
1994a&b) and has been under evaluation by EPA/APPCD as a potential source of clean transportation fuel
which could reduce greenhouse gas emissions to maximum effect and least cost. This section summarizes
the APPCD evaluations of that process, called Hynol, utilizing the BCL process discussed in Section 4 as a
basis for comparison. Because natural gas is used in the Hynol process, but not in the BCL process, the
comparison must include the option of using that natural gas for direct production of transportation fuel as well
as the use of biomass for production of that fuel. The latter case will be treated in Section 6 of this report; here,
the Hynol process will be described in sufficient detail to identify and quantify all energy requirements and
materials throughputs. The detail provided is also intended to establish a basis for cost comparisons to be
undertaken in later sections of the report For these reasons, all assumptions regarding feedstock composition,
conversion, reaction equilibria, and other operating criteria that affect performance are taken to be the same
as those given by Katofsky for the BCL process-everywhere that such assumptions are applicable to the
Hynol system.
The basis chosen for this analysis is, like the preceding discussion of the BCL system, 100 kg of dry
biomass feed per hour. This basis is intended to limit the magnitude of the numbers carried through the report,
beyond those significant figures required to specify accurate material and energy balances. It is nevertheless
understood that the actual plant size would be several orders of magnitude larger than 100 kg/hr if appreciable
impact is to be achieved on petroleum displacement or greenhouse gas reduction in the transportation sector.
Due to this latter point, the configuration of each part of the Hynol process discussed here includes all heat
exchangers, steam recovery systems, and hardware arrangements that could be utilized to maximize thermal
efficiency and minimize net C02 emission. It also includes estimates of kinetic effects on approach to
equilibrium for the reforming and methanol synthesis steps, distillation of the methanol product, pressure
drops, and ancillary equipment needed to dry biomass and utilize its unreacted residue. The final configuration
of the process is therefore considerably more complex than the version originally proposed. The overriding
criterion for the base case configuration to be discussed is that no electric power is to be imported to satisfy
the energy required by the process.
Figure 5-1 is a simplified Hynol block flow diagram. This thermochemical process utilizes the extra
hydrogen provided by natural gas addition to increase the efficiency of biomass conversion in three ways: (1)
at equilibrium conditions, it would allow complete conversion to synthesis gas of all carbon in the biomass, (2)
the gasifier design would not require an external biomass combustor or internal partial oxidation of the biomass
in order to generate the energy needed for gasification: excess hydrogen from the methanol reactor is recycled
to the gasifier through a heat exchanger that returns high-temperature sensible heat from the reformer effluent
which, combined with the exothermic heat of reaction of the hydrogen with biomass, is sufficient to satisfy the
gasification enthalpy requirements, and (3) because of the extra hydrogen and the recycle of unconverted CO
and C02 to the gasifier, no shift reactor is required for methanol production and a lower methanol synthesis
pressure is possible. Table 5-1 gives the compositions of the principal streams identified in Figure 5-1.
Because Hynol is a thermochemical system, it can be configured to produce either methanol or
hydrogen as transportation fuel. The production of methanol is considered first in this report. The process
configuration for methanol production shown in Figure 5-1 is the base case against which all other
5-1
-------
Biomass
51.06 kg C
7.49 kg H
51.53 kg O
0.15 kg N
0.08 kg S
0.792 kg ash
Natural gas
3.79 mols fTl
Steam 19.79 mols
r5"@
HRSG
Desulfurizer
Filter
Air
31.5 mols
Reformer
Gasifier
Combustor
Unreacted
carbon
0.549 mol
Ceramic
heat
exhanger
Natural gas (fuel)
2.60 mols
Steam 1.4 mols
HjO condensate
14.9 mols
HRSG
«—[D
Purge 1.2 mols
(to reformer fuel)
Methanol
converter
CHjOH 7.293 mols
HjO 1.85 mols
Figure 5-1. Block diagram of Hynol process.
5-2
-------
TABLE 5-1. PRINCIPAL STREAM DATA FOR HYNOL PROCESS, FIGURE 5-1
Stream composition, kg-mols
Stream
1
Stream
2
Stream
3
Stream
4
Stream
5
Stream
6
Stream
7
Stream
8
H,0
2.997
0.168
0.163
0.265
0.0008
1.407
0
0
6.10
26.23
193.93
220.16
0.9361
8.714
0
0
CO
1.698
6.540
17.63
24.17
0.0851
0.7923
0
0
O
O
0.9006
2.335
10.77
13.10
0.0520
0.4838
0.00758
0.0052
CH.
2.713
0.247
4.92
5.165
0.0237
0.2210
3.589
2.462
C,Hfi
0
0
0
0
0
0
0.1061
0.0728
N,
0.847
0.935
18.74
19.68
0.0905
0.8421
0.08717
0.0598
CH,OH
0
0
2.44
2.44
0.0118
0.1096
0
0
SUM
15.26
36.45
248.6
285
1.20
12.57
3.79
2.60
Deq. C
800
51
40
51
40
950
90
90
Atm
29
22.3
30.0
36.0
30
29.5
28.0
1.5
5-3
-------
configurations of the Hynol process and other process options will be compared; it is therefore discussed In
detail in this section. The following discussion addresses each of the individual operating blocks that comprise
Figure 5-1.
Biomass Gasification Block
Figure 5-2 and Table 5-2 show the configuration, operating conditions, and heat duties of each step
of the process between the gasification unit and the reforming unit. Like the BCL system discussed in Section
4, the data are based on 100 kg of dry biomass and 11.1 kg of biomass moisture (10 wt%) fed to the gasifier
and 87% conversion of the carbon in the biomass. The elemental composition of this feed, shown in Figure
5-1, includes both biomass and moisture. The design of the fluidized bed gasifier for the Hynol process is
described in the report "Hynol Process Engineering: Process Configuration, Site Plan, and Equipment Design"
(Unnasch, 1996). Gasified products leave the gasifier at 800°C and entrained particulates are removed in
ceramic candle filters. The particulate-free gas enters the first heat recovery system where process steam for
the gasifier is produced together with part of the steam needed for the reforming step. Low pressure steam
from the methanol synthesis block (stream 19), after superheating in heat exchanger SH-G1, is used as
process steam for the gasifier (stream 20); 0.47 kW of electric power is extracted in turbine TG-7 to equalize
the steam pressure with the gasifier pressure. The remainder of low-pressure superheated steam is used to
extract 36.8 kW of electric power in condensing turbine TG-6. Steam generator HRSG-1 and superheater SH-
G2 produce high pressure steam from which an additional 3.72 kW of power is extracted prior to use as
reformer feed.
The steam added to the gasifier feed (recycle) stream is sufficient to eliminate equilibrium constraints
on the biomass conversion obtainable in the gasification step. The biomass conversion, assumed to be 87%
of the carbon, will thus be determined by reaction kinetics only.
Following heat recovery, the process gas (stream 8) has a temperature of 324°C. Addition of the
natural gas co-feedstock at 90°C reduces the temperature to 268°C prior to entry into the desulfurizing unit.
Catalysts used for reforming and methanol synthesis are poisoned by sulfur compounds exceeding 0.1 ppm
in the gas feed, and design conditions are usually established for 0.02 ppm sulfur. Desulfurization is carried
out by reaction with ZnO. Assuming a biomass sulfur content of 0.05 wt%, the gas entering the desulfurizer will
contain 82 ppm H2S and 15.7 mol% HzO. At equilibrium, the gas leaving the desulfurizer will contain 0.02 ppm
H^. Equilibrium data for the Zn0-H2S-H20 system are given in Figure 5-3. Because natural gas contains a
few ppm of sulfur compounds in addition to H2S, the desulfurization unit also contains a hydrogenation catalyst
to convert all of the sulfur to H2S which can be removed by ZnO. A small stream of hydrogen-rich gas from the
purge stream of the methanol unit is used for the hydrogenation reaction.
Superheated steam from SH-G3 (2.79 mols) and from SH-S1 of the power block (19.79 mols) is added
after desulfurization to provide a steam/carbon ratio of 2.5:1 in the process gas stream entering the steam
reformer. Laboratory tests conducted by EPA/RTP, at the operating pressure and gas feed composition
corresponding to the Hynol reforming step, verified that steam/carbon ratios of 2.0 or greater are sufficient to
prevent carbon deposition on a nickel reforming catalyst at 900-1000°C (Dong and Karwowski, 1997).
Reformer Unit and Power Block
The reformer unit shown in Figure 5-4 includes the steam pyrolysis reactor, a process gas preheater,
and other preheaters that recover energy from the flue gases leaving the reformer's combustor. The
combustor is fueled with natural gas and with purge gas from the methanol reactor. Because both of these fuel
streams are under pressure of approximately 28 atm, they are expanded through turbines TG-4 and TG-5 to
recover 5.9 kW of electric power. The fuel gas is then preheated to 388°C before entering the combustion
furnace at 1 atm pressure. The combustion air (15% excess) is preheated to 325°C by the reformer flue gas
and then to 535°C in a heat exchanger in the char combustor (of the biomass block, Figure 5-7) that utilizes
unreacted biomass from the gasifier as fuel. The amount of natural gas required as reformer fuel is
5-4
-------
Desulfurtzatfon
8FW
Natural gas
SH-G2
TG-l
1161
SH-Gl
TG-7
TG-6
120'
HE-1
1221
HEAT DUTIES FOR GASIFIER BLOCK
I.D.
Description
Duty, cal/sec
Surface area, m2
SH-Gl
Steam superheater
7758
0.1204
SH-G2
Steam superheater
2380
0.0756
SH-G3
Steam reheater
1182
0.0742
HRSG-1
Heat recovery steam generator
7820
0.2457
Figure 5-2. Biomass gasification block for Hynol process
5-5
-------
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
19
20
22
24
25
83
TABLE 5-2. STREAM DATA FOR FIGURE 5-2
I.D.
Flow,
kg-mol
Temp.,
°C
Dried biomass, 10 wt% moisture
111.1 (kg)
25
Heated recycle from methanol block
12.57
950
Gasifier effluent
15.26
800
Unreacted carbon
0.549
800
Steam superheater SH-G1 effluent
15.26
616
Steam superheater SH-G2 effluent
15.26
558
Steam superheater SH-G3 effluent
15.26
528
Heat recovery steam generator HRSG-1 effluent
15.26
324
Natural gas feed to process
3.79
90
Superheated I.P. process steam to reformer
2.79
527
Boiler feed water
2.79
90
Saturated HP steam
2.79
310
Superheated HP steam
2.79
527
Process steam
2.79
367
Natural gas
6.39
90
Superheated medium pressure steam
9.19
527
Turbine condensate
5.82
30
Sat steam from methanol block, HRSG-4
9.19
250
Superheated process steam for gasifier
1.4
481
Heat exchanger HE-1 cold side feed
12.57
97
Desulfurized gasifer effluent: reformer feed
19.05
268
Natural gas fuel to TG-4 and reformer furnace
2.60
90
Recycle stream from methanol block
11.17
40
5-6
-------
500°C
8.0
268°C
203°C
w
0Y
I
o
I*
CD
o
4.0
0.0013
0.0015
0.0017
0.0019
0.0021
,K
Figure 5-3. Equilibrium conditions for the reaction: H2S + ZnO = ZnS + H20
5-7
-------
133
Purge
26
NGPH
Natural gas [£5
APH-1
29i Air
TG-4
TG-5
FPH-1
110°C
30°C
FPH-2
379°C
J24j—~
From gasifier block
APH-2
28
PGPH-2
809° C
Reformer
HE-1
Cond-1
TG-3
TG-2
44 43
To distillation
• block
/——* ,u L
140| ¦' »
42
[5T] ~ To drier
SH-S2
SH-S1
39!
HRSG-2
HRSG-3
BFW
¦95
46
|45
Figure 5-4. Hynol reformer unit and power block
5-8
-------
determined by the heat duties of the process-gas preheaters, PGPH-1 (13700 cal/sec), PGPH-2 (18980
cal/sec), that recover energy from the flue gas, the duty of the external air preheater that recovers energy
(13720 cal/sec) from the char combustor flue gas, and the reformer SPR (127300 cal/sec) with an assumed
skin loss of 2.3% of total thermal input. Flue gas is discharged at 126 °C. Given the amount of purge gas, its
composition, and the amount of residual carbon available from the gasifier for preheating the reformer
combustion air, the base case requires 2.60 mols of natural gas as supplementary reformer fuel. Alternative
configurations, in which all of the natural gas is fed to the process and fuel for the reformer is extracted from
the gasifier effluent and purge, were also examined and found less efficient than the configuration of Figure
5-1.
Process gas exiting the SPR at 1000°C passes through a ceramic heat exchanger, HE-1, that heats
the gasifier feed stream to 950°C. The process gas stream enters the main power block at 813°C where the
second heat-recovery steam generator and superheater run turbine-generators TG-2 and -3 producing a total
of 29.7 kW of electric power. Exhaust steam from TG-2 at 28 atm pressure is reheated in heat exchanger SH-
S1 to provide 17.03 mols of superheated steam for the process feed stream to the reformer; the remaining 2.76
mols of steam required for the reformer to bring the steam/carbon ratio to 2.5:1 is obtained from the gasifier
block, stream No. 10. The final unit of the reformer block is steam generator HRSG-3 from which the 15337
cal/sec of low pressure steam required for the biomass drier is delivered at 160°C. The temperature of the
process stream leaving the power block (153°C) is determined by the enthalpy required to operate the reboiler
for distillation of the crude methanol to remove its water content.
Operating conditions of the streams identified in Figure 5-4 are shown in Table 5-3; heat duties are
given in Table 5-4.
Distillation Block
Crude methanol condensed in the methanol synthesis block contains 20.2 mol% water. Water is
separated from the crude methanol in a 15-stage distillation column operating at 1 atm pressure, 0.22 recycle
ratio, and a distillate/feed ratio of 0.81. This separation requires a major energy input of 23320 cal/sec for the
reboiler. The methanol product has a purity of 98.5 mol% (99.1 wt%). Referring to Figure 5-5 and Table 5-5,
the process gas leaving the reformer block enters the reboiler of the distillation column to which It delivers the
required 23320 cal/sec and then heats the column feed to 65°C in heat exchanger HE-2. Accumulation tanks
T1 and T2 separate water condensate (to be reused as boiler feed water) from the process gas prior to its
compression and entry into the methanol synthesis block. Heat exchangers HE-3 and HE-5 recover energy
for boiler feed water and preheat the natural gas. Process gas enters the compressor Comp-1 at 22.3 atm and
exits the first compressor stage at 28 atm. After cooling to 40°C, it is again compressed to 36 atm and enters
the methanol synthesis block at 67°C.
Methanol Synthesis Block
CO and H2 combine in an exothermic reaction to produce methanol in a catalytic converter at 260°C.
In order to maintain isothermal operation, the catalyst tubes within the reactor are surrounded by boiling water
at a pressure of 39.2 atm which yields steam at 250°C. This Lurgi-type reactor was compared with the
alternative ICI-type which uses cooled process gas, fed between stages of reaction, to maintain constant
temperature. The Lurgi system was found to result in slightly greater efficiency due to the higher temperature
of the steam recovered. Figure 5-6 and Table 5-6 show the configuration and operating conditions of the
methanol synthesis block.
Recovery of senstole heat from the reactor effluent is essential for maximum thermal efficiency of the
process and is achieved with heat exchangers HE-4 and HE-6 which preheat the boiler feed water and process
gas feed for the reactor, respectively. HRSG-4 recovers 9.19 mols of medium pressure steam from the
exothermic methanol synthesis reaction and delivers that steam to the gasification block where it is
superheated in SH-G1 to supply the gasifier (1.4 mols) and part of the reformer steam (1.97 mols), while the
5-9
-------
TABLE 5-3. STREAM DATA FOR REFORMER AND POWER BLOCK, FIGURE 5-4
Stream
No.
I.D.
Flow,
kg-mol
Temp.,
°C
Pressure,
atm
10
Process steam from superheater SH-G3
2.79
527
28.0
23
Turbine condensate
1.67
30
0.05
24
Process feed stream from gasifier
19.05
268
28.0
25
Natural gas reformer fuel
2.60
90
28.5
26
Purge gas from methanol synthesis block
1.2
40
30.0
27
Natural gas and purge fuel for SPR furnace
3.80
30
1.5
28
Fuel to furnace
3.80
388
1.0
29
Air inlet
31.5
25
1.0
30
Air to first preheater
31.5
69
1.5
31
Air to second preheater
31.5
325
1.25
32
Air to furnace
31.5
535
1.0
33
Flue gas
34.8
126
1.0
34
Superheated steam to reformer
17.03
527
28.0
35
Preheated process gas to SPR
38.84
697
27.5
36
Reactor effluent
51.4
1000
24.8
37
Heat exchanger HE-1 cold-side effluent
51.4
812
24.3
38
Superheater SH-S1 hot-side effluent
51.4
754
24.0
39
Superheater SH-S2 hot-side effluent
51.4
620
23.7
40
Heat recovery steam generator HRSG-3 effluent
51.4
153
23.2
41
Saturated HP steam
18.67
310
97.3
42
Superheated HP steam from char combustor
0
527
97.3
43
Superheated HP steam to turbine/generator TG-3
1.67
527
97.3
44
HP steam to turbine/generator TG-2
17.03
527
97.3
45
Boiler feed water from heat exchanger HE-3
6.05
90
97.3
46
Boiler feed water from reboiler/HE-2 condensate
12.10
115
97.3
91
Steam to drier
5.59
145
2.0
95
Condensate water from drier
5.59
110
2.0
5-10
-------
TABLE 5-4. HEAT DUTIES FOR REFORMER AND POWER BLOCK
I.D.
Description
Duty, cal/sec
Surface area, mz
HE-1
Ceramic heat exchanger
23450
0.4634
SH-S1
IP steam reheater
7220
0.1065
SH-S2
HP steam superheater
15928
0.2947
HRSG-2
Heat recovery steam generator
50760
3.257
NGPH
Natural gas preheater
181
0.026
APH-1
First air preheater
15950
2.198
APH-2
Second air preheater (in biomass block)
13720
0.354
FPH-1
Furnace fuel gas preheater
943
0.0160
FPH-2
Furnace fuel gas reheater
3620
0.1534
PGPH-1
First process gas preheater
13700
2.116
PGPH-2
Second process gas preheater
18980
1.231
COND-1
Steam condenser, TG-3 exhaust
4617
1.92
HRSG-3
Drier steam generator
15337
2.00
SPR
Steam pyroivsis reactor (reformer)
127300
Skin loss
4550
5-11
-------
COND-2
|6lj-» CHjON
40
80
Purgs
§§
58
H3—~
P3
COMP-1
methanol
tyntnesis block
HE-5
HE-3
41
BFW ~
HEAT DUTIES FOR DISTILLATION BLOCK
I.D.
Description
Duty, cal/sec
Surface area, mJ
Reboiler
Distillation column reboiler
23315
1.851
HE-2
Column feed preheater
1800
0.145
HE-3
Heat exctianqer, boiler feed water
9950
2.539
HE-5
Natural qas preheater
1090
0.123
COND-2
Condenser, column overheads
23060
Figure 5-5. Hynol distillation block.
5-12
-------
TABLE 5-5. STREAM DATA FOR DISTILLATION BLOCK, FIGURE 5-5
Stream
No.
I.D.
Flow,
kg-mol
Temp.,
°C
Pressure,
atm
23
Turbine TG-3 condensate from COND-1
3.837
30
1.0
26
Purge to reformer fuel, SPR block
1.2
40
30.0
40
Effluent from reformer block, HRSG-2
51.4
153
23.2
41
Natural gas
6.39
25
30
42
Preheated natural gas
6.39
90
29.5
45
Boiler feed water to HRSG-1 and HRSG-2
9.07
90
97.3
46
Boiler feed water to HRSG-2
12.62
115
97.3
48
Reboiler effluent
51.4
115
22.9
49
Column feed preheater, HE-2 effluent
51.4
110
22.8
50
Water condensate
12.10
115
22.8
51
Condenser feed
38.76
110
22.8
52
Condensate water
2.25
50
22.3
53
Condenser effluent to gasifier loop compressor
36.5
50
22.3
54
Feed to methanol synthesis reactor block
36.5
67
36.0
55
Effluent from methanol condenser
260.9
40
30.0
56
Recycle to gasifier
11.17
40
30.0
57
Recycle to methanol converter
248.6
40
30.0
58
Crude methanol to distillation column
9.143
40
30.0
59
Distillation column feed
9.143
65
1.0
60
Reboiler effluent
9.143
103
1.0
61
Methanol product
7.293
30
1.0
62
Column bottoms
1.737
103
1.0
63
Cooling water
26.7
25
1.0
64
Water to pond
15.2
90
1.0
67
Boiler feed water to MSR block HRSG-4
9.19
90
39.2
5-13
-------
HRSG-4
Methanol
synthesis reactor
HE-6
73,
HE-4
69,
74,
Comp-2
Cond-3
57
¦E]—*
To gasifier
55
T3
HEAT DUTIES, METHANOL SYNTHESIS BLOCK
I.D.
Description
Duty, cal/sec
Surface area, m2
HE-4
Heat exchanaer, reactor feed oreheater
96579
17.17
HE-6
Heat exchanaer, reactor boiler feed water
7530
0.627
Cond-3
Crude methanol condenser
45000
16.22
MSR
Methanol synthesis reactor
18970
HRSG-4
Steam generator
16970
Figure 5-6. Methanol synthesis block.
5-14
-------
19
26
54
55
56
57
67
68
69
70
71
72
73
74
75
77
78
TABLE 5-6. STREAM DATA FOR METHANOL SYNTHESIS BLOCK, FIGURE 5-6
I.D.
Flow,
kg-mol
Temp.,
GC
Pressure,
atm
Saturated steam to gasifier, HE-1
9.19
250
39.2
Purge to reformer furnace (fuel)
1.20
40
30.0
Feed from gasifier/reformer loop
36.45
67
36.0
Crude methanol condenser effluent
270
40
30.0
Recycle to gasifier
11.17
40
30.0
Recycle from methanol loop compressor
248.6
40
30.0
Boiler feed water from HE-3, distillation block
9.19
90
39.2
Boiler feed water
9.19
246
39.2
MSR recycle compressor effluent
248.6
48
36.0
Methanol synthesis reactor feed
285
51
36.0
Heat exchanger HE-4 effluent/connverter feed
285
217
35.5
MSR reactor effluent
270
260
32.5
Heat exchanger HE-6 effluent
270
247
32.0
Heat exchanger HE-4 effluent
270
78
32.0
Condenser cooling water
299
25
1.0
Water to distillation condenser, and compressor
interstage coolers
73
25
3.0
Water to pond
226
60
1.0
5-15
-------
remainder (5.82 mols) is used to produce 27.4 kW of electric power in condensing turbine TG-6 of the gasifier
block. Condenser Cond-3 separates crude methanol from the process gas stream at 40°C which is then sent
to the distillation block for separation of the water. A gas purge (stream 26, which controls the accumulation
of nitrogen in the system) is sent to the reformer furnace as part of its fuel makeup. It should be noted that both
biomass and natural gas contain nitrogen in varying amounts and the purge rate will therefore depend on those
feedstock compositions. Another part of the process gas leaving the methanol condenser is recycled to heat
exchanger HE-1 for preheating prior to return to the gasifier. Most of the process gas is recycled to the
methanol reactor via compressor Comp-2.
Compressor Comp-2 moves the recycle stream through the methanol synthesis loop and overcomes
the pressure drop across the converter and associated heat exchangers of that part of the Hynol system. The
methanol converter effluent is at a pressure of 32.5 atm and exits Cond-3 at 30 atm and 40°C; it is compressed
to the reactor feed pressure of 36 atm in two stages, initially to 33 atm, with 40°C intercooling. This compressor
is the largest consumer of electric power in the process and determines the ratio of natural-gas-to-biomass
that may be fed to the process without importing or exporting electric power. Because the flow rate of the
recycle stream increases with that ratio, there exists a maximum recycle flow rate beyond which the electric
power consumption exceeds that which can be generated by the process. At the conditions of Figure 5-1,
the total electric power consumed matches that produced, and no imported power is required.
The pressure drop across the methanol synthesis reactor and its associated heat exchangers and
condenser is assumed to be 6.0 atm for the base case. This variable also strongly affects the electric power
consumed by the recycle compressor and therefore affects the yield of methanol under the constraint that
electric power is not imported. The sensitivity of process performance to this pressure drop and other major
process variables will be examined in Sections 7 and 8.
Biomass Block
The biomass block, Figure 5-7 and Table 5-7, Includes the biomass drying step and a char combustor.
The char combustor bums unreacted carbon from the gasifier (assumed for the base case to be 13 mol% of
tie carbon fed to the gasifier as biomass) to produce HP steam in HRSG-5 for power generation in turbine TG-
3 and/or to preheat air (in APH-2) for the reformer combustor. By utilizing the gasifier char in this manner, its
heating value is thereby recovered and its disposal problems are avoided. Unlike the BCL system, the flue
gas from this combustor is exhausted to the atmosphere and is not used for biomass drying.
Biomass is dried with LP steam obtained from the final heat recovery step of the reformer/power block,
HRSG-3 (Figure 5-4). This drier is indirectly heated with a steam-jacketed vessel under partial vacuum (0.5
atm). The biomass is fed at a moisture content of 50 wt% and dried to 10% in a fluidized bed. Recycled steam
is the fluidizing medium. The general steam drying concept is that of Modo Chmetics and Stork Friesland as
described by Liinanki et al. (1994), and demonstrated at the 80-tonne/day level in Finland.
The amount of air preheat that can be provided by the char combustor for the reformer combustion
air (and therefore the amount of natural gas required as supplementary reformer fuel) is determined by the
amount of steam that is required to produce sufficient electric power to balance the overall power requirements
of the Hynol process. Because the electric power required for the base case is produced entirely within the
process, all of the heat output of the char combustor can be used for air preheat, and no duties are shown for
the steam generator in Figure 5-7. Other cases require that part of that energy be used to produce additional
electric power, in which case the air preheat is reduced and more natural gas may be required for reformer fuel;
Figure 5-7 therefore identifies those streams although they do not apply to the base case. The specific
electric power requirements of all process equipment are itemized in Table 5-8; their total is shown to be 76.4
kWh per 100 kg of dry biomass fed to the gasifier (or 7.293 kg-mols of methanol produced). Electric power
cogenerated by the process is similarly shown h Table 5-9, which totals 76.5 kWh. The adjustable parameters
of the base case (primarily the ratio of natural gas to biomass fed) were selected to achieve this match, given
5-16
-------
COND-4
HRSG-3
HRSG-5
To r«form«r
furnace
B3
HEAT DUTIES, BIOMASS BLOCK
I.D.
Description
Dutv. cal/sec
Surface area, m2
CAP
Air preheater for char combustor
1450
0.0834
CSG
Char combustor steam generator
0
0
CSH
Steam superheater
0
0
APH-2
Second-staqe SPR furnace air preheater
13720
0.221
HRSG-3
Drier steam qenerator (In reformer block)
15337
1.762
COND-4
Drier overhead steam condenser
12000
1.918
Figure 5-7. Hynol biomass block.
5-17
-------
No
1
4
31
32
40
42
90
91
93
94
95
97
98
99
100
102
TABLE 5-7. STREAM DATA FOR THE BIOMASS BLOCK, FIGURE 5-7
I.D.
Flow,
kg-mol
Temp.,
°C
Pressure,
atm
Dried biomass, 10 wt% moisture
111.1 (kg)
105
29
Unreacted char from gasifier
0.549
800
1.0
Air from reformer air preheater APH-1
31.5
325
1.25
Air to reformer furnace
31.5
535
1.0
Hot side effluent from reformer block HRSG-3
51.4
153
23.2
HP superheated steam to turbine TG-3
527
97.3
Wet biomass feed, 45 wt% moisture
181-8 (kg)
25
1.0
LP saturated steam
5.59
145
2.0
Steam recycle
1.54
105
0.5
Steam recycle
1.54
187
1.0
Condensate/boiler feed water
5.59
110
2.0
Water condensate
3.93
25
0.05
Air to preheater, char combustor
3.14
25
1.0
Air to furnace
3.14
300
1.5
Boiler feed water from distillation block, HE-3
90
97.3
Flue gas to stack
3.14
126
1.0
5-18
-------
the conditions of pressure drop, approach to equilibrium, and other operating variables assumed for the base
case. Later sections of the report will discuss the sensitivity to changes in these assumptions.
Integrated Process Flowsheet
The thermally integrated Hynol process, including all steps discussed above, is shown in Figure 5-8.
TABLE 5-8. ELECTRIC POWER CONSUMPTION
I.D.
Unit Description
Power required, kWh
COMP-2
Methanol loop compressor
39.75
COMP-1
Gasifier loop compressor
15.88
B1
Reformer furnace air FD fan
11.1
P3
HP boiler feed water pump for HRSG-4
0.6
P5
HP boiler feed water pump for HRSG-2
1.5
P1
MP boiler feed water pump for HRSG-2
1.76
LH
Lock hopper
1.7
B3
Char combustor air FD fan
1.1
B2
Drier steam recycle fan
1.2
P4
Low pressure cooling water pump, Cond-2,-3,
and compressor cooler
1.7
P2
Low pressure cooling water pump for HE-3
0.11
Total power requirement
76.4
TABLE 5-9. ELECTRIC POWER AVAILABLE
I.D.
Source
Power, kWh
HRSG-1
Heat recovery steam generator, TG-1
3.68
HRSG-2
Heat recovery steam generator, TG-2
22.73
HRSG-2
Heat recovery steam generator, TG-3
6.97
HRSG-4
Heat recovery steam generator, TG-6
36.77
HRSG-4
Heat recovery steam generator, TG-7
0.47
TG-4,5
Natural gas and purge expander
5.88
Total power available
76.50
5-19
-------
Btomass
DOfiUff
Nfttura! g*s
NG1
Rcfomw
BFW2
BFW2
DMDtetton
Methanol
oonvwtef
Ntturalgts
Figure 5-8. Integrated Hynol process flowsheet.
5-20
-------
SECTION 6
PERFORMANCE COMPARISONS
Objectives
The expected advantages of the Hynol process are: (1) to achieve greater reduction of C02 emissions
from mobile sources with alternative fuels derived from a given amount of biomass and natural gas than could
be obtained from those feedstocks if they were used in separate processes, (2) to achieve greater
displacement of petroleum fuel than could be obtained by conversion of biomass or natural gas individually
to liquid fuel, and (3) to minimize the cost of petroleum displacement while maximizing the overall net reduction
of C02 emissions from mobile sources. This section will compare the C02 emissions and gasoline
displacement of Hynol and the BCL process based on the data developed in Sections 4 and 5. As discussed
in Section 1, the greatest potential for reduction of the environmental impacts of road transport, including
emissions of criteria pollutants and greenhouse gases, is fuel-cell vehicles which will become commercialized
in the next decade. These comparisons will therefore assume that each process will produce methanol for use
in fuel cell vehicles. Because Hynol uses natural gas and BCL does not, the analysis will first consider the
methanol yield and CO, emissions that would be obtained if the natural gas were converted to methanol directly
by the conventional steam reforming process. The sum of that methanol and the methanol produced from
biomass by the BCL process will then be compared with that obtained from Hynol using the same total amount
of natural gas and biomass. The basis for these comparisons will be, like the preceding section, 100 kg of dry
biomass and 6.39 kg-mols of natural gas
Methanol Production from Natural Gas bv Conventional Steam Reforming
The raw material and fuel requirements for production of methanol from natural gas by conventional
technology (steam reforming, Lurgi process) are given by the California Fuel Methanol Cost Study (Bechtel,
Inc., 1988) for a plant producing 10,000 metric tonnes/day of fuel grade methanol. The facility consists of four
trains of 2,500 tonnes/day each. Natural gas is used as both feed and supplemental fuel for the facility; 300.4
x 106 scfd of gas is required for feed and 15.6 x 10s scfd for fuel. The natural gas contains 95.8 mol% CH4,
0.77% C02,0.33% N2, 0.03% H* 0.1% C?, 0.62% C3,0.33% C* and 0.05% C6; MW = 16.87, LHV = 930 Btu/scf.
From these data the methanol yield is:
Methanol {MeOH) yield per mol of natural gas (NG) feedstock =
iniwi . u rtu '000 kg MeOH kg mol MeOH „
10,000 tonnes MeOH - - x
tonne MeOH 32 kg MeOH
1 359 scf NG lb mol NG = Q ?g2 kg mols MeOH
316 x 106 scf NG & mol NG 0.454 kg mol NG kg mol NG
and the total amount of natural gas required to produce 1 kg-mol of methanol is therefore:
6-1
-------
mols NG
0.782 mol CHflH
= 1.279
mols NG
mol CH^OH
and the total C02 emission, per mol of methanol produced and utilized is:
1.279 mol NG 44 kgC02 kg C02
= 36.2/
mol MeOH mol NG mol MeOH
Vehicle CO. Emission Reduction Potential of Methanol Produced bv the BCL Process and Conventional Steam
Reforming
Section 4 showed that 100 kg of biomass will produce 1.477 kg-mols of methanol when utilized in the
BCL process. If that methanol is used in fuel cell vehicles having 2.5 times the fuel efficiency of gasoline-ICE
vehicles (LHV basis) then it will displace:
1.477 mol MeOH x 32 kg Me0H Uters MeOH x
mot MeOH 0.796 kg MeOH
*al Me0H 56800 Btu 8al gasoline ^ = ]9 3? . lm£
3.785 Uters MeOH gal MeOH 115000 Btu
As discussed in Section 5, the Hynol process uses 3.79 mols of natural gas as process co-feedstock and 2.60
mols as reformer fuel, or 6.39 total mols of natural gas in addition to 100 kg of dry biomass to produce 7.293
kg mols of methanol. If that natural gas were used to produce methanol by the conventional steam reforming
route discussed above, it would yield 6.39 x 0.782 = 5.00 mols of methanol which, by the same calculation
indicated above, would displace 65.55 gallons of gasoline and:
Total gasoline displaced - 65.55 + 19.37 - 84.92 gallons
Gasoline, with an average carbon content of 86.5 wt%, will emit the following amount of C02 per gallon of
gasoline burned:
CO2 emission per gallon of gasoline combustion =
3.785 liters gasoline 0.75 kg gasoline ^
gallon gasoline liters gasoline
0.865 kg carbon mol carbon 44 kg C02 _ ^ kg CP2
kg gasoline 12 kg carbon mol carbon gallon gasoline
and the net vehicle C02 emission avoided by displacement of the gasoline by methanol in fuel cell vehicles
is:
6-2
-------
84.92 gallons gasoline
9.0 kg CO.
2
gallon gasoline
56.27 kg CO, ]
1 x 5.00 mob MeOHl
mol MeOH
482.9 kg C02 emission avoided
CO- Emission Reduction bv the Hvnol Process
The Hynol process requires 6.39 total mols of natural gas and 100 kg of biomass to produce 7.293 kg-mols
of methanol. The net C02 emission from vehicles powered with methanol produced by the Hynol process is
therefore 6.39 x 44 = 281.2 kg.
The gasoline displaced by 7.293 kg-mols of methanol used in fuel cell vehicles is:
- -A, , .. 32 kg MeOH liters MeOH
7.293 mols MeOH x
mol MeOH 0.796 kg MeOH
Me0H 56800 B,u 2.5 . 95.62 gal gasalm,
3.785 liters MeOH gal MeOH 115000 Btu
so the net COs emission reduction from fuel-cell vehicles using Hynol methanol is:
(95.62 x 9.0) - 281.2 = 579.4 kg
and the ratio of Hynol C02 emission reduction to that of the combined BCL process and conventional steam
reforming process is:
579 4
Hynol CO, emission reduction advantage - — = 1.20
2 482.9
Thus, the net reduction of tailpipe C02 emission from fuel-cell vehicles using methanol produced by the Hynol
process would be 20 % greater than the emission reduction of the same vehicles using methanol produced
from the identical feedstocks by the best biomass gasification technology and the conventional natural gas
conversion technology.
As indicated by the data in Table 6-1, the net tailpipe C02.emission reduction from a given vehicle
would be 67.3% for the Hynol case and 63.1% for the BCL + steam reforming case; Hynol, however will also
displace 13% more gasoline from the vehicle fleet. The combination of these effects is a 20% advantage for
Hynol in terms of total C02emission reduction that can be obtained from a given amount of the two feedstocks.
This improvement in utilization efficiency of the two resources-biomass and natural gas~for the displacement
of petroleum and reduction of greenhouse gas emission is attributed to the effect of natural gas on the yield
of methanol obtainable from the biomass, as discussed in Section 5. The magnitude of this enhancement can
be deduced from the above data: since 1 mol of natural gas alone would produce 0.782 mol of methanol, the
6.39 mols of natural gas used for Hynol production would account for 5.00 mols of the methanol product. The
other 2293 mob of methanol is therefore derived from the 100 kg of biomass. Section 3 showed that the BCL
process can produce 1.477 mols of methanol from 100 kg of biomass; therefore, the enhancement of biomass
yield due to synergistic effects of natural gas is 2.293/1.477 = 1.55, or 55 %.
6-3
-------
TABLE 6-1. SUMMARY OF TAILPIPE C02 EMISSION REDUCTIONS
Hynol
BCL
Steam
reforming
BCL+steam
reforming
Dry biomass feed, kg
100
100
0
100
Natural gas feed, kg-mols
6.39
0
6.39
6.39
Methanol produced, kg-mols
7.293
1.477
5.00
6.477
Gasoline displaced, gallons
95.62
19.37
65.55
84.92
Net C02 emission, kg
281.2
0
281.2
281.2
C02 emission reduction, kg
579.4
174
308.9
482.9
More important pertiaps than the Improvement in biomass utilization and 20 % improvement in tailpipe
C02 reduction is the fact discussed in Section 1 that the biomass supply will be the limiting factor that
determines how much petroleum displacement and greenhouse gas emissions are possible by its conversion
to alternative fuel for use in the transportation sector. If these dual objectives are to be accomplished on a
national scale sufficient to affect either appreciably, leveraging of the biomass supply will be necessary,
especially as the vehicle population increases. Fuel production from biomass can be leveraged efficiently only
if a feedstock capable of producing the same fuel is also utilized. Natural gas is probably the only other
resource that meets these requirements and, although the potential displacement of gasoline by both resources
is just 13 % better for the Hynol route, it is also 4.9 times better than the petroleum displacement that could be
expected for biomass alone.
Fuel-Cvcle CO. Emissions
Although tailpipe C02.emission reduction discussed above is the most accurately quantifiable effect
of displacing gasoline with methanol derived from biomass and natural gas, other emissions associated with
feedstock production and fuel use also contribute to the overall effect on greenhouse gas emissions. The
production of biomass, for example, involves emissions from fertilizer manufacture, N20 and NOx emission from
fertilizer use, and diesel fuel used in planting, harvest, and transport. Natural gas production involves the
release of C02 extracted from the well with the natural gas~which may, or may not, be removed as an
impurity-methane leakage during extraction and transport, fuels used in drilling and operation of natural gas
wells, fuels used in purification of the natural gas, and energy needed for pipeline transport. Because some
of these greenhouse emissions are not C02, they must be accounted for according to their C02 equivalents.
An accounting of these total fuel-cycle emissions will now be addressed.
The process energy consumed and C02-equivalent emissions at each stage of the fuel cycle for
production of gasoline, natural gas, and biomass are given by DeLuchi (1993) in Table 6-2. From the data of
Table 6-2, the production of 6.39 kg-mols of natural gas for the Hynol process will have an equivalent C02
emission of:
C02 equivalent emission -
coi 212800 kcal 3.968 Btu ,AA .
11.115 6.39 mols NG = 60.0 kg CO-,
106 Btu mol NG kcal 2
6-4
-------
TABLE 6-2. GRAMS OF COz-EQUIVALENT EMISSIONS PER 108 Btu OF FUEL AVAILABLE TO USER
Fuel cycle stage
Standard gasoline
from crude
Methanol production
from wood
Natural
gas
C02 from natural gas wells
0
0
1163
Gas leaks/flares
1318
0
2265
Fertilizer manufacture
0
2180
0
NsO, NOx from fertilizer
0
1979
0
Feedstock recovery
3051
8488
1792
Feedstock transport
2723
2892
0
Fuel production
12705
0
1468
Fuel distribution
1442
5195
4427
Total C02 equivalent, g/106 Btu
21238
20734
11115
All values in Table 6-2 under "Methanol production from wood" are associated with biomass cultivation and
harvest, with the exception of "fuel distribution" which concerns transport by truck to refueling stations (using
diesel fuel, not methanol)- The data further assume that the methanol is produced by the BCL process, which
yields 1.477 mol of MeOH per 100 kg of biomass feedstock. The C02 equivalent emissions associated with
the cultivation, harvest, and transport of that biomass is therefore:
Equivalent C02 emission per 100 kg of biomass production =
(20.734 - 5.195)** C°2 1.477 mob MeOH 172000 3 968 Btu
106 Btu mol MeOH kcal
kg C02
15.5
100 kg biomass
The emissions resulting from distribution of the methanol fuel (assuming distribution by trucks using diesel fuel)
from the Hynol plant gate to vehicle refueling stations are:
C02 emission by distribution of 7.293 mo Is of Hynol methanol -
c tnc^ /iu 172000 kcal 3.968 Btu oc ,
5.195 7.293 mols MeOH , :—:— = 25.85 kg C02
106 Btu mol MeOH kcal 2
and the C02-equivalent emissions that would have occurred from the production of the gasoline that was
displaced by methanol are:
-------
CO2 equivalent emission by production of one gallon of gasoline
kg C02 Btu kg CO.
21.238 — 2- 125000 — = 2.65 2
106 Btu gallon gasoline gallon gasoline
and the total C02 emission due to both production and combustion of gasoline is therefore: 2.65 + 9.0 = 11.65
kg C02 per gallon of gasoline displaced by methanol used in fuel cell vehicles.
And,
Net overall fuel-cycle C02 emission reduction for Hynol base case ~
95.62(11.65) - 6.39(44) - 60.0 - 15.5 - 25.85 - 731.4 kg
so the comparative fuei-cycle C02 emission reduction that could be obtained by methanol produced in
separate processes by the BCL and steam reforming routes is:
BCL + steam - reforming C02 emission reduction =
(65.55 + 19.37)11.65 - 44(6.39) - 60.0 - 15.5 - 25.85f 5 00'1477) = 609.7 kg CO,
\ 7.293
and the relative advantage of Hynol to the alternative option of producing methanol from biomass and natural
gas by the BCL and steam reforming route using the same biomass and natural gas feedstocks is:
Hynol fuel cycle C02 emission reduction
BCL + steam-reforming fuel cycle C02 emission reduction
731.4 kg C02
609.7 kg C02
1.20
Again, there is a 20 % improvement in overall reduction of fuel-cyle greenhouse gas emissions for Hynol
relative to existing technology.
6-6
-------
SECTION 7
HYNOL SENSITIVITY ANALYSIS
Introduction
This section examines a number of process variables that might be expected to affect the performance
of the Hynol process. Since Hynol has not been demonstrated at any level in actual hardware, these
simulations must assume reasonable values of biomass conversion, reactor temperatures, pressures, and
configurations of the basic process components. The base case discussed in Section 5 represents the best
current estimate of those operating conditions. The effect of changing those values to account for the
possibility that they may not be achieved because, for example, of lower biomass conversions, or inability to
operate the reformer at maximum temperature, is predicted by the additional simulations discussed here.
The Hynol process involves two basic loops, each containing a compressor to overcome the pressure
drop imposed by the reactors, heat exchangers, and other process equipment within that loop. As indicated
by Figure 7-1, the gasifier loop includes the gasifier, particulate filter, heat recovery steam-generator HRSG-1,
a desulfurization unit, the reformer, the ceramic heat exchanger, HRSG-2, the distillation reboiler, and a
condenser. The methanol recycle loop includes the methanol converter, heat exchanger HE-4, and the crude
methanol condenser. Each variable that is changed as process Input alters the relative flows between these
two loops and thus affects the energy consumption by the compressors, which are the primary consumers of
electric power. Changes in relative flow rates in the two loops also affect equipment size and capital cost. If,
for example, the natural gas feed is increased with all other input variables constant, the flow rate of the
methanol recycle loop is also increased (as well as the methanol output). If no electric power is to be imported,
this recycle rate and power consumption determines the maximum amount of natural gas that can be fed per
unit of biomass feed. As pointed out in Section 5, it is that condition that defines the natural gas/biomass feed
ratio of the base case. As another example, reduction of the reformer temperature, all other input conditions
unchanged, reduces the recycle rate in the methanol recycle loop, as does also an increase in the purge rate.
Because the purge stream is used as part of the fuel for the reformer, the amount and location of its extraction
(Figure 7-1, location G or M) affects not only the flow rate in both loops, but also the amount of natural gas
required as supplementary reformer fuel. Each of these effects will be examined in this section.
Procedure
The overall energy balance on the process must provide:
1. Energy for biomass gasification
2. Energy for biomass drying
3. Energy for separation of water from methanol by distillation
4. Energy for steam reforming, provided as:
> Natural gas
> Purge gas taken from the process
> Un-gasified biomass (carbon)
7-1
-------
Purge Natural gas Steam
location G
HRSG-1
Filter
GASIFIER LOOP
Biomass
Reformer
Gasifier
Ceramic
heat exchanger
Carbon
+ ash
Steam
HRSG-2 Drier Distillation
steam reboiler
* H20
Gasifier loop
compressor
Recycle loop
compressor
Purge
location M
4
METHANOL RECYCLE LOOP
Methanol
converter
HE-4
Methanol
H,0
Figure 7-1. Hynol process block flow diagram.
7-2
-------
5. Electric power required by the system:
> To operate compressors to overcome pressure drops in reactors and
heat exchangers
> To operate air blowers for the reformer combustor, char combustor, and
biomass drier
> To operate high pressure (HP) pumps for boiler feed water (BFW)
> To operate condenser pumps and lock hoppers
The base case of the Hynol process is configured to satisfy the above requirements with energy recovered
within the process so that no imported electric power is required while also accomplishing the following
objectives:
1. Maximizing fuel-cycle C02 emission reduction from the total FCV fleet using the methanol fuel produced
by the process.
2. Maximum improvement over the fuel-cycle C02 emission reduction that could be achieved with the same
biomass and natural gas feedstocks if they were converted to methanol by conventional steam reforming
of the natural gas and by conversion of biomass by the BCL process.
3. Maximum thermal efficiency in conversion of the feedstocks while achieving the above two goals.
4. Minimum fuel production cost per unit of net C02 emission reduction.
This section will examine the effects of process variables on the first three of the above criteria while
including relevant comparisons that will affect process cost, which will be addressed in Section 10. In order
to assess the effect of each process variable as its value is changed, the following procedure was followed:
1. Adjust the steam feed to the gasifier to eliminate any equilibrium constraint on the biomass conversion.
2. Adjust the recycle rate to the gasifier to match enthalpies of the gasifier input and output streams, given the
assumed carbon conversion in the gasifier.
3. Adjust heat recovery from the gasifier effluent to provide 0.02 ppm equilibrium H2S concentration in
the desulfurizer effluent.
4. Adjust the steam feed to the reformer to provide the required steam/carbon ratio, normally 2.5 kg mols
steam/kg atom carbon
5. Evaluate each additional process block (HRSG-1, methanol converter, methanol distillation, HRSG-2,
biomass drying, and steam reforming) to determine material and energy inputs and outputs.
6. Determine from the above results the amount of electric power required from the char combustor and the
amount of air preheat available for the reformer.
7. If the power required exceeds that available, reduce the natural gas feed to the process and repeat the
above steps until energy balance is achieved without electric power import.
Ratio of Natural Gas Feed to Biomass Feed
The base-case Hynol system, Figure 7-1, utilizes 3.79 kg-mols of natural gas as process feed and 2.60
kg-mols of natural gas as reformer fuel per 100 kg of dry biomass feed. The total natural gas used therefore
constitutes 75 % of the energy input to the process. The principal objectives of the Hynol process are to (1)
achieve maximum overall C02 emission reduction and (2) to achieve maximum petroleum displacement at
7-3
-------
minimum cost. It is therefore reasonable to question whether more biomass and less natural gas might be
more consistent with these objectives. As discussed in the preceding section, the criterion for assessing these
objectives is a comparison between Hynol and the production of methanol by the BCL process in combination
with the production of methanol by conventional steam reforming of natural gas when those two processes use
the same total amount of those feedstocks as the Hynol system.
It was shown in Section 5 that the base-case Hynol process can achieve 20 % greater net C02
emission reduction from the vehicle fleet than might be obtained with the BCL and steam reforming combination
of processes. Figure 7-2 shows the results of other simulations with lower ratio of natural gas to biomass feed
(NG/biomass). Like the base case, these simulations assume 87 % biomass carbon conversion, 1000°C
reformer temperature, and 2.5 steam/carbon ratio in the reformer feed stream. It also assumes that the
unreacted char from the gasifier is burned in a fired heat exchanger to preheat the combustion air to the
reformer. (This avoids the need to dispose of that char and recovers its heating value; a separate combustor
is used to avoid ash deposition on the reformer tubes which are at higher temperature and therefore subject
to fouling by the char ash.) The following additional parameters are constant for all simulations shown in
Figure 7-2:
950° C primary (ceramic) heat exchanger exit temperature; i.e., 50 °C approach to the reformer's exit
temperature of 1000°C
-10K approach to equilibrium in reformer; +12K approach to equilibrium in methanol converter
Char combustor fan power = 1.1 kW
Drier recycle fan power = 1.2 kW
Lock hopper power = 1.7 kW
Condenser pumps = 1.8 kW total
Biomass is dried from 50 to 10% moisture using process energy
Enthalpy of the biomass feed (and moisture) is -176800 kcal
Enthalpy of the unreacted carbon and ash leaving the gasifier is +2536 kcal
7.7 atm total pressure drop in gasifier loop
6.0 atm total pressure drop in methanol recycle loop
40° C methanol condenser
Distillation at atmospheric pressure in single, 15-stage column
The methanol product is used in fuel cell vehicles having fuel economy 2.5 times that of gasoline
spark-ignition vehicles.
The abscissa of Figure 7-2 is the total net COs-equivaIent emission reduction for the full fuel cycle that
was defined in Section 6. It is evident from Figure 7-2 that Hynol exceeds the C02 reduction capability of the
BCL + steam reforming option for all ratios greater than 19 kg-mols NG/tonne biomass (at which natural gas
represents 47 % of the total energy input). For the range of natural gas utilizations evaluated, the base-case
conditions (63.9 kg-mols NG/tonne biomass) yield greatest net C02 emission avoided.
Table 7-1 summarizes the data resulting from the six process simulations denoted by the symbols in
Figure 7-2, including the base case. Table 7-2 gives the pertinent performance parameters calculated from
those data. The following comments are noted regarding these results: The total natural gas feed includes
both process feed and natural gas used as fuel for the reformer (in addition to the process purge). As the ratio
of total natural gas feed to biomass feed was reduced, the purge was taken from the gasifier loop instead of
the methanol recycle loop; this change was necessary to maintain a reasonable gas throughput in the gasifier
loop, as explained in the next section. No imported electric power is required for any of the six Hynol cases;
the electric power produced matches the power requirements with a small reserve for the two highest
NG/biomass ratios, but the lower ratios produce significant excess power. The abscissa of Figure 7-2 does
not credit this exportable power to the net C02 emission reduction because the objective is to produce
transportation fuel, not electric power. Reduction of the NG/biomass ratio also required that the amount of
steam fed to the gasifier be increased in order to avoid equilibrium limitation of the biomass conversion. At the
lowest NG/biomass ratio (6.8 mols/tonne), the steam rate rate was insufficient to allow more than 89%
7-4
-------
70
60
50
¦s 40
30
20
10
BASE CASE
—
BCL +
-
STEAM REFORMING /
—
HYNOL
—
l/ I
»
1
8
C02 Emission Avoided, tonnes
Figure 7-2. Total fuel-cycle COz-equivalent emission reduction for gasoline ICEVs displaced by methanol
FCVs with 2.5 factor improvement of fuel economy.
7-5
-------
TABLE 7-1. EFFECT OF NATURAL GAS / BIOMASS RATIO
Basis: 1 hour, 100 kg dry biomass, 1000° C reformer, 2.5 steam / carbon ratio
Total natural gas feed, kg-mols
Process variable
0.68
1.62
2.775
4.19
5.42
6.39*
Gasifier throughput, mots
22.86
20.75
17.86
16.61
14.6
15.26
Purge rate, mols
4.71
4.71
4.71
3.60
3.80
1.20
Purge location
G
G
G
G
G
M
Reformer steam feed, mols
19.0
17.0
16.1
17.1
17.56
19.79
Gasifier steam feed, mols
3.3
3.3
2.5
2.1
1.4
1.4
Reformer throughput, mols
41.8
40.2
39.4
42.9
44.2
51.4
Gasifer loop compressor feed, mols
22.2
23.5
25.4
29.1
31.8
36.45
Recycle loop compressor feed, mols
23.9
44.16
70.0
113.3
219.4
248.6
Methanol converter throughput, mols
42.5
. 61.9
87.4
132
238
270
Methanol product, mols
1.623
2.681
3.885
5.259
6.466
7.293
Methanol loop compressor power, kW
-3.77
-6.99
-11.15
-18.1
-35.1
-39.75
SPR furnace air blower power, kW
-5.63
-6.33
-7.22
-8.63
-10.0
-11.1
Gasifier loop compressor power, kW
-9.58
-10.18
-11.0
-12.66
-13.9
-15.88
High pressure BFW pumps, kW
-2.3
-2.8
-2.9
-3.3
-3.1
-3.9
HRSG-2 power, kW
+9.91
+17.8
+21.22
+25.94
+20.2
+30.96
HRSG-1 power, kW
+11.1
+18.6
+30.96
+40.0
+38.66
+40.92
Reformer fuel expander, kW
+10.81
+11.5
+11.9
+10.1
+11.2
+5.9
Exportable power, kW
+4.2
+15.8
+25.9
+27.7
+2.2
+1.36
Power from char combustor, kW
0
0
0
0
0
0
Distillation reboiler duty, kcal
18840
30340
44265
59100
73890
83930
Ceramic heat exchanger duty, kcal
143560
124190
102870
93560
79020
84410
NG fuel for reformer, mols
0.38
0.46
0.575
1.19
1.32
2.60
NG feed to process, mols
0.30
1.160
2.20
3.00
4.10
3.79
Reformer duty, kcal
226040
263340
310490
366840
420030
458300
Reformer combustion air, mols
16.0
18.0
20.5
24.5
28.5
31.5
*Base case
7-6
-------
TABLE 7-2. SUMMARY OF EFFECTS OF NATURAL GAS / BIOMASS FEED RATIO
ON HYNOL PROCESS PERFORMANCE
Parameters calculated from data of Table 71
Total natural gas feed, kg-mols/100 kg biomass
Process variable
0.68
1.62
2.755
4.19
5.42
6.39*
Thermal efficiency, %
42.4
54.3
61.3
65.4
67.6
68.4
C02 emission from natural gas, kg
29.9
71.3
122
184
238
281
Gasoline displaced, gallons
21.3
35.1
50.9
68.95
84.8
95.6
Net tailpipe C02 emission avoided, kg
162
245
336
437
525
579
Tailpipe C02 reduction per vehicle, %
84.3
77.4
73.3
70.3
68.7
67.3
Methanol produced, mols
1.623
2.681
3.885
5.259
6.466
7.293
Total fuel cycle C02-equivaIent
emission reduction, kg
190
298
416
545
660
731
Ratio, Hynol fuel cycle emission
reduction / BCL + steam reforming
emission reduction
0.767
1.00
1.093
1.160
1.204
1.200
Ratio, sum of gasifier + reformer +
methanol converter throughputs /
methanol produced, mols/mol
66.0
45.8
37.2
36.4
47.1
46.1
Sum above per tonne of fuel-cycle
C02 emission avoided, mols/tonne
563
412
348
351
450
465
*Base case
7-7
-------
conversion even though a 136% increase in steam over the base case was allowed. The last two columns
of Table 7-2 compare the sum of the throughputs of the gasifier, reformer, and methanol converter per unit of
methanol produced by the Hynol system and the C02 reduction potential. Because the throughputs of these
reactors are a measure of their size, this ratio is a rough measure of the relative capita! cost of the six
operating modes. It is seen to have a minimum value between 2.755 and 4.19 mols natural gas/100 kg
biomass fed.
Purge Location
Because the Hynol process is a closed-loop recycle system, a purge stream is necessary to control
the accumulation of inerts (e.g., nitrogen) that would otherwise become the dominant constituent in the process
streams. The amount of purge needed depends on the nitrogen content of the feedstocks, which is relatively
low in biomass but can be a significant component of natural gas and varies substantially with its source. The
Hynol base case assumes a purge rate of 1.20 mols per 100 kg of dry biomass fed which is sufficient to
maintain the nitrogen concentration below 7 mol % in the recycle stream entering the gasifier (for a biomass
containing 0.15 wt% nitrogen and natural gas containing 2.3 mo!% nitrogen). The base case also assumes
that the purge is withdrawn from the methanol recycle loop.
As indicated in Tables 7-1 and 7-2, the purge location is changed from the methanol loop to the gasifier
loop when the NG/biomass ratio is reduced below the base case conditions. There are two reasons for this
change: (1) convergence of the simulation model could not be obtained when the same purge rate (1.2 mols)
was taken from the gasifier loop with reduced NG/biomass ratios; and (2) the simulation results shown in
Tables 7-3 and 7-4 for the case where the purge is taken form the methanol loop while reducing the natural
gas feed to the process to a ratio of 2.20 mols/100 kg biomass (from the base-case ratio of 3.79 mols of
natural gas/100 kg biomass). When the purge is maintained at 1.2 mols, ttiis simulation shows that the
gasifier throughput and reformer throughput are increased substantially. Because of the increased throughput,
the steam rate and heat duty of the reformer are also increased, thus requiring a higher natural gas
consumption as fuel for the reformer furnace. The result of a reduction of natural gas feed to the process under
these conditions is an increase in the amount of natural gas needed as fuel for the reformer furnace which
offsets the reduction to the process as feedstock. Another result is that the capital cost of the gasifier,
reformer, ceramic heat exchanger, heat recovery systems, and gas cleanup units is increased due to the
higher throughput.
To avoid this problem, the second simulation shown in Tables 7-3 and 7-4 shows the effect of
withdrawing the purge from the gasifier loop while maintaining the same 2.2 mols of natural gas feed to the
process as feedstock. The purge rate is increased to avoid the convergence problem indicated above and
to supply the reformer with fuel derived partially from the biomass. The reformer throughput, the methanol
converter throughput, and the recycle loop compressor throughput-all of which affect capital cost-are also
reduced substantially. Figure 7-3 shows the net effect of attempting to reduce the NG/biomass ratio according
to the two methods discussed above. By withdrawing the purge from the gasifier loop at an increased rate,
a final ratio of 27.75 moles/tonne biomass is achieved, whereas the withdrawal from the methanol loop results
in a final ratio of 47.2 mots/tonne. For the latter case, the total fuel-cycle C02 emission reduction is only 2.5%
better than that of the alternative BCL + steam reforming route, compared to a 16% improvement for the former
case.
7-8
-------
TABLE 7-3. EFFECT OF PURGE LOCATION ON SIMULATIONS WITH REDUCED
NATURAL GAS / BIOMASS FEED RATIO
Basis: 100 kg biomass, 87% carbon conversion, 1000°C reformer, 2.5 steam / carbon ratio
Purge location
Process variable
Gasifier loop
Methanol loop
Gasifier throughput, mols
17.86
23.61
Purge rate, mols
4.71
1.20
Steam to reformer, mols
16.1
26.15
Steam to gasifier, mols
2.5
3.8
Reformer throughput, mols
39.4
61.0
Gasifier loop compressor feed, mols
25.4
35.6
Recycle loop compressor feed, mols
70.06
123.7
Methanol converter throughput, mols
87.4
148
Methanol product, mols
3.885
5.255
Methanol loop compressor power, kW
-11.2
-19.5
Gasifier loop compressor power, kW
-11.0
-15.4
SPR furnace air blower power, kW
-7.22
-10.5
High pressure BFW pumps, kW
-2.9
-4.4
HRSG-2 power, kW
+21.2
+36.2
HRSG-1 power, kW
+31.0
+37.0
Reformer fuel expander, kW
+11.9
+5.8
Imported (exported) power, kW
(25.9)
(23.3)
Power from char combustor, kW
0
0
Distillation reboiler duty, kcal
44270
59600
Ceramic heat exchanger duty, kcal
102870
145520
NG fuel for reformer, mols
0.575
2.52
NG feed to process, mols
2.20
2.20
Reformer duty, kcal
310500
399900
Reformer combustion air, mols
20.5
30.0
7-9
-------
TABLE 7-4. SUMMARY OF EFFECTS OF PURGE LOCATION ON HYNOL PROCESS
PERFORMANCE WITH REDUCED NATURAL GAS / BIOMASS FEED RATIO
Parameters calculated from data of Table 7-3
Purge location
Process variable
Gasfier loop
Methanol loop
Total natural gas fed, mols
2.775
4.720
C02 emission from natural gas, kg
122.1
207.7
Gasoline displaced, gallons
50.9
68.9
Net tailpipe C02 emission avoided, kg
336
412
Tailpipe C02 reduction per vehicle, %
73.3
66.5
Methanol produced, mois
3.885
5.255
Energy ratio
0.637
0.619
Thermal efficiency, %
61.3
61.0
Total fuel cycle C02-equivalent
emission reduction, kg
416
516
Ratio, Hynol fuel cycle emission
reduction / BCL + steam reforming
emission reduction
1.093
1.025
Ratio, sum of gasifier + reformer +
methanol converter throughputs /
methanol produced
37.2
44.3
Sum above per tonne of fuel-cycle
C02 emission avoided
348
451
7-10
-------
70
/ HYNOL
Purge location M
(methanol loop)
BCL +
STEAM REFORMING
HYNOL
O Purge location G
(gasifier loop)
C02 Emission Avoided, tonnes
Figure 7-3. Effect of Hynol purge location ori total natural gas requirement and fuel-cycle C02 emission
reduction. Natural gas feed to process = 22 mob per tonne of biomass in both cases.
Effect of Reformer Temperature
In Table 7-5 are the results of Hynol process simulations that assume different temperatures of the
reformer, 900 and 950°C, in addition to the 1000°C base case. All other conditions are the same as the base
case, except that one run was made with a lower value of carbon conversion, 78 instead of 87 mol %. In all
cases the system is independent of electric power import. The simulation results are interpreted by the
calculated performance parameters shown in Table 7-6. The base case (1000°C reformer) yields superior
performance by all measures. However, at 950°C, the performance is only marginally less than the base case
in terms of comparisons with the BCL option, including the overall net C02 emission reduction, the throughput
per unit of methanol product, and thermal efficiency. Since 950°C is within the upper limits of conventional
reformer technology, and would be expected to provide longer life of the reformer tubes, and longer catalyst
life than the higher temperature, it may be considered the more practical choice for initial Hynol process
application.
Three runs were made at 90CC, one of which assumes that the natural gas feed is maintained at the
same value as the 950°C case; it shows that this natural gas feed cannot be accommodated at the lower
reforming temperature without import of electric power. For the next case shown at 900°, the natural gas feed
was therefore reduced to bring the energy requirements of the process in balance with power available and
illustrates the fact discussed above regarding the effect of the NG/biomass feed ratio on electric power
requirements-as the ratio is reduced, the power available increases. The final case evaluated at 900°C
7-11
-------
TABLE 7-5. SIMULATION RESULTS FOR VARIABLE REFORMER TEMPERATURE
Steam/carbon ratio = 2.5, 24.8 atm, -10K approach to equilibrium, purge location M
Reformer temperature, °C
Process variable
1000*
950
900"
900
900
Biomass carbon conversion, %
87
87
87
87
78
Gasifier throughput, mols
15.26
16.93
19.93
21.4
18.92
Purge rate, mols
1.20
1.20
1.20
1.20
1.20
Steam feed to reformer, mols
19.79
20.9
23.3
23.8
20.75
Steam feed to gasifier, mols
1.2
1.5
1.6
2.1
1.3
Reformer throughput, mols
51.4
54.2
59.7
61.1
55.1
Gasifier loop compressor feed, mols
36.5
38.2
41.5
41.6
38.9
Recycle loop compressor feed, mols
248.6
260.7
296
236.6
317
Methanol converter throughput, mols
270
283
323
264
343
Methanol product, mols
7.293
7.225
7.178
6.907
6.613
Methanol loop compressor power, kW
-39.75
-41.6
-47.0
-37.6
-50.7
Gasifier loop compressor power, kW
-15.9
-16.6
-18.0
-18.0
-16.9
SPR furnace air blower, kW
-11.1
-11.2
-11.2
-12.0
-10.7
High pressure BFW pumps, kW
-3.9
-3.7
-4.0
-4.5
-4.2
HRSG-2 power, kW
+31.0
+29.8
+24.7
+20.9
+22.2
HRSG-1 power, kW
+40.9
+36.6
+26.7
+33.4
+38.1
Reformer fuel expander, kW
+5.9
+5.9
+6.1
+6.1
+5.7
Imported power requirement, kW
0
0
5.5"
0
0
Power from char combustor, kW
0
+6.6
+18.0
+17.4
+23.0
Distillation reboiler duty, kcal
83930
82710
84298
80930
77570
Ceramic heat exchanger duty, kcal
84410
91920
109300
117900
102900
NG fuel for reformer, mols
2.60
2.79
2.69
2.79
2.43
NG feed to process, mols
3.79
3.76
3.76
3.54
3.54
Reformer duty, kcal
458300
442000
429950
422180
398970
Reformer combustion air, mols
31.5
32.0
32.5
34.0
30.5
•Base case "Electric power import required
7-12
-------
TABLE 7-6. SUMMARY OF REFORMER TEMPERATURE EFFECT ON
HYNOL SYSTEM PERFORMANCE
Parameters calculated from data of Table 7-5
Reformer temperature, °C
Process variable
1000
(base case)
950
900
900
Biomass carbon conversion (assumed), %
87
87
87
78
Total natural gas fed, kg-mols
6.39
6.40
6.33
5.97
C02 emission from natural gas, kg
281
281.6
278.5
262.7
Gasoline displaced, gallons
95.6
94.7
90.55
86.69
Net tailpipe C02 emission avoided, kg
579
570.7
536.5
517.5
Tailpipe C02 reduction per vehicle, %
67.3
67.0
65.8
66.3
Carbon used for reformer fuel, mol
0.549
0.549
0.549
0.936
Methanol produced, mols
7.293
7.225
6.907
6.613
Energy ratio
0.690
0.684
0.659
0.659
Thermal efficiency, %
68.4
67.7
65.2
65.2
Total fuel cycle C02-equivalent emission
reduction, kg
731
720.7
677
652.3
Ratio, Hynol fuel cycle emission reduction /
BCL+steam reforming emission reduction
1.20
1.182
1.118
1.119
Ratio, sum of gasifier + reformer + methanol
converter throughputs / methanol produced
46.1
49.0
50.1
62.9
Sum above per tonne of fuel-cycle C02
emission avoided
465
491
518
637
7-13
-------
reforming temperature assumes that the carbon conversion in the gasifier is at the low end of the anticipated
range, 78 instead of the 87 mol % assumed for the other cases. As indicated by the performance parameters
in Table 7-6, a 900°C reforming temperature will yield only about 12 % improvement over the BCL option in
terms of C02 emission reduction as opposed to 18-20 % improvement for the higher reforming temperatures,
it is therefore concluded that temperatures less than 950°C should not be considered.
Effect of Biomass Conversion
For the reforming temperature of 1000°C, two additional levels of biomass carbon conversion, 82.5
and 78%,were examined and are compared with the base case of 87% conversion in Tables 7-7 and 7-8. The
range of conversions assumed for these simulations is consistent with the values determined by in-house
hydrogasification kinetics studies completed in 1996 using a thermobalance reactor, poplar wood, a reactor
feed-gas stream representative of the Hynol recycle stream with regard to composition, and reaction
temperature and pressure (Dong and Cole, 1996).
The simulation results in Table 7-7 show a reduction in methanol yield as biomass conversion
decreases, as would be expected. The difference between 87 and 78% conversion is a reduction of methanol
yield by 8.07%. The lower methanol yield is due mainly to the reduction of natural gas feed to the process (0.5
mol, or 7.8%) which is necessary to avoid electric power import. In each case, sufficient electric power is
generated within the process that no import or export is required. The principal performance data calculated
from these simulations are summarized in Table 7-8. The thermal efficiency of the process is not significantly
affected by reduced biomass conversion from 87 to 78 %. This is explained by the fact that unreacted carbon
from the gasifier is utilized as fuel for preheating the reformer combustion air and/or for generating electric
power for use in the process; its energy value is therefore retained. The main effect of reduced biomass
carbon conversion is the lower C02 emission reduction which results from the reduced methanol yield. At 78%
conversion, the C02 reduction potential is nevertheless significantly better than the BCL option. As indicated
by the last two rows of Table 7-8, the ratio of the sum of reactor throughputs to methanol production increases
as carbon conversion is reduced, signaling higher cost per unit of methanol produced and CO? emission
avoided.
Use of Unreacted Carbon as Process Fuel
All data shown for the Hynol process assume that the unreacted carbon from incomplete gasification
of biomass is utilized in a separate combustor/HRSG to provide part of the energy needed by the process. The
energy produced by burning this carbon is used for (a) generating some of the electric power required by
compressors and pumps, and/or (b) preheating the air fed to the reformer furnace. To assess the extent to
which this use affects process performance, a comparison is made in Tables 7-9 and 7-10 with and without
utilization of that carbon. The case chosen for this comparison is the 82.5% conversion simulation of Table
7-7 in which 0.7446 mol of carbon (8.935 kg) is discharged from the gasifier per 100 kg of dry biomass fed.
Normally, this carbon would be burned to produce 8.3 kWh of electric energy to balance the process energy
requirements and 12850 kcal/sec thermal preheat for the reformer combustion air, thus partially displacing
some of the natural gas required to fuel that furnace. These data are reproduced in Tables 7-9 and 7-10 under
"100% carbon utilized as fuel."
If all of the unreacted, waste carbon discharged from the gasifier is disposed of rather than burned for
energy recovery, the effects on process performance are shown by the simulation results shown in Tables 7-9
and 7-10 under "0% carbon utilized as fuel." In this case, the 8.3 kWh of electric power and air preheat are
not available, requiring two compensating changes: (1) natural gas fed to the process must be reduced from
3.629 to 3.463 mols in order to balance the electric power produced with that required by the methanol loop
compressor, and (2) the natural gas used as reformer fuel must be increased from 2.39 to 2.62 mols to
compensate for the reduced thermal input from char combustion. The result of these effects is compared in
Table 7-10 which shows about 3% reduction of thermal efficiency and about 3% lower fuel-cycle C02 emission
reduction relative to the "100% carbon utilization" case.
7-14
-------
TABLE 7-7. SIMULATION RESULTS FOR VARIABLE BIOMASS CONVERSION
Basis: 100 kg biomass, 1000°C reformer, 2.5 steam/carbon ratio, 1.20 mol purge rate
Biomass carbon conversion, mol %
Process variable
87
(base case)
82.5
78
Gasifier throughput, mols
15.26
15.05
14.66
Purge location
M
M
M
Steam feed to reformer, mols
19.79
18.70
17.50
Steam to gasifier, mols
1.4
1.35
1.2
Reformer throughput, mols
51.4
49.4
47.4
Gasifier loop compressor feed, mols
36.5
35.2
34.0
Recycle loop compressor feed, mols
248.6
246
258.5
Methanol converter throughput, mols
270
267
229
Methanol product, mols
7.293
6.970
6.704
Methanol loop compressor power, kW
-39.75
-39.4
-41.4
Gasifier loop compressor power, kW
-15.9
-15.3
-14.9
SPR furnace air blower, kW
-11.1
-10.4
-10.2
High pressure BFW pumps, kW
-3.9
-3.5
-3.6
HRSG-2 power, kW
+31.0
+30.3
+28.3
HRSG-1 power, kW
+40.9
+38.5
+33.1
Reformer fuel expander, kW
+5.9
+5.6
+5.5
Imported power, kW
0
0
0
Power from char combustor, kW
0
0
+9.6
Distillation reboiler duty, kcal
83934
80100
76824
Ceramic heat exchanger duty, kcal
84409
82590
79920
NG fuel for reformer, mols
2.60
2.38
2.34
NG feed to process, mols
3.79
3.639
3.55
Reformer duty, kcal
458300
439900
424430
Reformer combustion air, mols
31.5
29.5
29.0
7-15
-------
TABLE 7-8 SUMMARY OF EFFECTS OF CARBON CONVERSION ON
HYNOLPROCESS PERFORMANCE
Parameters calculated from data of Table 7-7
Biomass carbon conversion, mol %
Process variable
87
(Base case)
82.5
78
Total natural gas fed, mols
6.39
6.030
5.89
C02 emission from natural gas, kg
281
265.3
259.2
Gasoline displaced, gallons
95.6
91.4
87.9
Net tailpipe C02 emission avoided, kg
579
557
532
Tailpipe C02 reduction per vehicle, %
67.3
67.7
67.2
Methanol produced, mols
7.293
6.974
6.704
Energy ratio
0.69
0.689
0.675
Thermal efficiency, %
68.4
68.2
66.8
Total fuel cycle COs-equivalent
emission reduction, kg
731
702
670
Ratio, Hynol fuel cycle emission
reduction / BCL + steam reforming
emission reduction
1.200
1.197
1.160
Ratio, sum of gasifier + reformer +
methanol converter throughputs /
methanol produced
46.1
47.6
50.9
Sum above per tonne of fuel-cycle
C02 emission avoided
465
472
509
7-16
-------
TABLE 7-9. SIMULATION RESULTS ON EFFECT OF RECOVERING ENERGY
CONTENT OF UNREACTED GASIFIER CARBON
Basis: 100 kg biomass, 82.5% carbon conversion, 1000°C reformer,
2.5 steam/carbon ratio
Carbon utilized as fuel, %
Process variable
100
0
Gaslfier throughput, mols
15.05
15.08
Purge rate, mols
1.20
1.20
Purge location
M
M
Steam to reformer, mols
18.7
18.7
Steam to gasifier, mols
1.35
1.35
Reformer throughput, mols
49.4
49.4
Gasifier loop compressor feed, mols
35.2
35.2
Recycle loop compressor feed, mols
246
239
Methanol converter throughput, mols
267
260
Methanol product, mols
6.970
6.953
Methanol loop compressor power, kW
-39.4
-38.2
Gasifier loop compressor power, kW
-15.3
-15.3
SPR furnace air blower power, kW
-10.4
-11.8
High pressure BFW pumps, kW
-3.5
-3.5
HRSG-2 power, kW
+30.3
+29.9
HRSG-1 power, kW
+38.5
+39.1
Reformer fuel expander, kW
+5.6
+6.1
Imported power, kW
0
0
Power from char combustor, kW
0
0
Distillation reboiler duty, kcal
80100
79650
Ceramic heat exchanger duty, kcal
82590
83050
NG fuel for reformer, mols
2.38
2.74
NG feed to process, mols
3.639
3.639
Reformer duty, kcal
439900
439900
Reformer combustion air, mols
29.5
33.5
7-17
-------
TABLE 7-10. SUMMARY OF EFFECTS OF WASTE CARBON
UTILIZATION ON HYNOL PROCESS PERFORMANCE
Parameters calculated from data of Table 7-9
Carbon utilized as fuel, %
Process variable
100
0
Waste carbon from gasifier, kg
8.935
8.935
Total natural gas fed, mols
6.030
6.379
C02 emission from natural gas, kg
265.3
280.6
Gasoline displaced, gallons
91.4
91.2
Net tailpipe COz emission avoided, kg
557
540
Tailpipe C02 reduction per vehicle, %
67.7
65.8
Methanol produced, mols
6.970
6.953
Energy ratio
0.689
0.664
Thermal efficiency, %
68.3
65.6
Total fuel cycle COz-equivaIent
emission reduction, kg
702
681
Ratio, Hynol fuel cycle emission
reduction / BCL + steam reforming
emission reduction
1.197
1.119
Ratio, sum of gasifier + reformer +
methanol converter throughputs /
methanol produced
47.6
46.7
Sum above per tonne of fuel-cycle
C02 emission avoided
472
476
7-18
-------
Sequester of the unreacted carbon, rather than its use as reformer fuel, is another option although less
desirable from a practical standpoint. In that case, however, the overall fuel cycle C02 emission reduction
would be 714 kg, or 1.7% better than the "100% utilization" case.
Effect of Reformer Steam/Carbon Ratio
Conventional steam reformers, using natural gas as process feed, normally operate with steam/carbon
ratios ranging from 2 to 3 in the reformer feed stream. This steam ratio is necessary to prevent carbon
formation on the reforming catalyst and the resulting deactivation of the catalyst. Because Hynol contains a
substantial amount of hydrogen in the feed stream (about 40 mol%), which would be expected to have a
similar effect as HzO in suppressing carbon formation, a lower ratio may be sufficient for the Hynol process.
A steam/carbon ratio of 2.5 is assumed for the base case and other sensitivity simulations discussed here and
suitability of that ratio has been verified by in-house experimental studies using a fixed-bed catalytic reactor
with conventional nickel catalyst at Hynol operating pressure, and reformer feed gas composition at
temperatures of 900, 950, and 1000° C.
Tables 7-11 and 7-12 compile the results of process simulations in which the steam/carbon ratio is 2.0
or 3.0 and compares them with the base case which assumes a ratio of 2.5—all other inputs remaining
unchanged. The comparisons show that the operating steam/carbon ratio has little effect on actual process
performance. The most important effect of higher steam/carbon ratio is the increased throughputs of the
reformer and methanol converter, which are directly related to the required size of those units. The effect will
therefore be on capital cost, and the lowest operable steam/carbon ratio is desirable for that reason. A ratio
of 2.5 is consistent with good catalyst protection without significant sacrifice of thermal efficiency or C02
reduction potential.
Approach to Equilibrium
As indicated earlier, all simulations assume that sufficient steam is added to the gasifier feed stream
to ensure that biomass conversion in the gasifier is not limited by reaction equilibrium. The other reactors, for
reforming and methanol synthesis, will not attain equilibrium due to kinetic limitations, equipment design,
operating conditions, and catalyst aging. These factors are accounted for in the simulations by a temperature
input to the reactor model that artificially transposes the equilibrium temperature calculations to a slightly
different value-a value estimated by comparison with operating data-that yields the appropriate conversion.
In the case of conventional steam reformers, a temperature compensation of -1 OK is normally used and a +12K
temperature for methanol converters for systems of good design and fresh catalyst. Those values were used
by Katofsky (1993) for the BCL process simulations and have been used also for the simulations reported here
for Hynol. The effect of less favorable assumptions regarding approach to equilibrium in those reactors will
now be considered.
Tables 7-13 and 7-14 give the results of simulations in which values of +12,+18, and +24K were
assumed for the approach to equilibrium in the methanol synthesis reactor. Equilibrium is assumed at 260°C,
and the temperature approach is positive because the reaction is exothermic. It is evident from these
comparisons that this variable has a strong effect on process performance. The ideal approach to equilibrium
would be zero, but that is not attainable in practice; as the real approach becomes larger, throughputs increase
in the gasifier and reformer as well as the methanol synthesis reactor. Attempts to counter this effect by
reducing the natural gas/biomass feed ratio were only partially effective and resulted in reduced methanol yield,
as would be expected. At +24 K, it was no longer possible to maintain balance between electric power
production and demand due to the large compressor throughput. That balance could be restored by reduction
of the assumed carbon conversion in the gasifier (which provides more carbon for reformer fuel), but only with
sacrifice of methanol yield. As shown in the summary of Table 7-14, the thermal efficiency and C02 emission
reduction potential drop sharply as the approach to equilibrium becomes greater, and the performance is
unacceptable at +24 K. Clearly, the design of the methanol synthesis reactor to achieve the closest approach
7-19
-------
TABLE 7-11. SIMULATION RESULTS FOR VARIABLE STEAM/CARBON RATIO
Basis: 100 kg biomass, 1000°C reformer, 87% carbon conversion
Reformer steam/carbon ratio
Process variable
2.5
(base case)
2.0
3.0
Gasifier throughput, mols
15.26
15.82
14.65
Purge rate, mols
1.20
1.20
1.20
Purge location
M
M
M
Steam teed rate, mots
19.79
15.67
24.1
Reformer throughput, mols
51.4
47.8
55.4
Gasifier loop compressor feed, mols
36.5
36.4
36.8
Recycle loop compressor feed, mols
248.6
237
288
Methanol converter throughput, mols
270
259
309
Methanol product, mols
7.293
7.277
7.404
Methanol loop compressor power, kW
-39.75
-37.9
-46.1
Gasifier loop compressor power, kW
-15.9
-15.8
-16.0
SPR furnace air blower power, kW
-11.1
-10.9
-12.0
High pressure BFW pumps, kW
-3.9
-3.1
-4.3
HRSG-2 power, kW
+31.0
+22.0
+37.0
HRSG-1 power, kW
+40.9
+43.7
+38.1
Reformer fuel expander, kW
+5.9
+5.8
+6.1
Imported power requirement, kW
0
0
0
Power from char combustor, kW
0
+2.0
+3.0
Distillation reboiler duty, kcal
83930
82120
85450
Ceramic heat exchanger duty, kcal
84410
88200
80130
NG fuel for reformer, mols
2.60
2.54
2.75
NG feed to process, mols
3.79
3.79
3.87
Reformer duty, kcal
458300
448400
472800
Reformer combustion air, mols
31.5
31.0
34.0
7-20
-------
TABLE 7-12. SUMMARY OF EFFECTS OF STEAM/CARBON RATIO ON
HYNOL PROCESS PERFORMANCE
Parameters calculated from data of Table 7-11
Reformer steam/carbon ratio
Process variable
2.5
(Base case)
2.0
3.0
Total natural gas fed, mols
6..39
6.33
6.62
C02 emission from natural gas, kg
281.0
278.5
291.3
Gasoline displaced, gallons
95.6
95.4
97.1
Net tailpipe C02 emission avoided, kg
579
580
582
Tailpipe C02 reduction per vehicle, %
67.3
67.5
66.7
Methanol produced, mols
7.293
7.277
7.404
Energy ratio
0.690
0.695
0.682
Thermal efficiency, %
68.4
68.8
67.5
Total fuel cycle C02-equivalent
emission reduction, kg
731
732
735.7
Ratio, Hynol fuel cycle emission
reduction / BCL + steam reforming
emission reduction
1.200
1.209
1.179
Ratio, sum of gasifier + reformer +
methanol converter throughputs /
methanol produced
46.1
44.3
51.2
Sum above per tonne of fuel-cycle
C02 emission avoided
465
441
515
7-21
-------
TABLE 7-13. SIMULATION RESULTS FOR VARIABLE APPROACH TO EQUILIBRIUM
IN METHANOL SYNTHESIS REACTOR (MSR)
8.0 atm MSR pressure drop, 32.5 atm MSR pressure, 260° C MSR temperature,
100 kg biomass, 1000°C reformer, 2.5 steam/carbon ratio, 1.20 mol purge
MSR approach to equilibrium, K
Process variable
12
18
24
24
Carbon conversion in gasifier, %
87
87
87
78
Gasifier throughput, mols
17.0
17.69
21.5
18.2
Steam feed to reformer, mols
20.3
21.1
25.6
19.1
Steam feed to gasifier, mols
2.1
2.4
3.4
2.9
Reformer throughput, mols
52.6
53.2
58.8
49.8
Gasifier loop compressor feed, mols
36.1
35.4
34.9
32.6
Recycle loop compressor feed, mols
182.5
189.9
20.1
190.7
Methanol converter throughput, mols
204
212
225
212
Methanol product, mols
6.876
6.52
5.443
5.654
Methanol loop compressor power, kW
-48.4
-50.3
-53.0
-50.5
Gasifier loop compressor power, kW
-19.8
-19.4
-19.0
-17.8
SPR furnace air blower power, kW
-10.8
-11.1
-11.1
-9.9
High pressure BFW pumps, kW
-4.0
-4.3
-4.8
-4.3
HRSG-2 power, kW
+37.6
+39.5
+39.4
+41.4
HRSG-1 power, kW
445.9
+38.1
+20.4
+25.5
Reformer fuel expander, kW
+5.8
+5.9
+5.9
+4.6
Exportable (imported) power, kW
0
0
(10.0)
0
Power from char combustor, kW
0
+7.4
+18.0
+17.6
Distillation reboiler duty, kcal
78880
74100
64300
64740
Ceramic heat exchanger duty, kcal
95900
100400
128800
102200
NG fuel for reformer, mols
2.57
2.63
2.65
2.31
NG feed to process, mols
3.47
3.19
2.35
2.73
Reformer duty, kcal
444580
430770
400100
386200
Reformer combustion air, mols
30.8
31.5
31.5
28.2
7-22
-------
TABLE 7-14 SUMMARY OF EFFECTS OF MSR APPROACH TO EQUILIBRIUM ON
HYNOL PROCESS PERFORMANCE
Parameters calculated from data of Table 7-13
MSR approach to equilibrium, K
Process variable
12
18
24
24
Carbon conversion, mol %
87
87
87*
78
Total natural gas fed, mols
6.04
5.82
5.00
5.04
C02 emission from natural gas, kg
265.8
256.1
220
221.8
Gasoline displaced, gallons
90.1
85.5
71.4
74.1
Net tailpipe C02 emission avoided, kg
546
513
422
445
Tailpipe C02 reduction per vehicle, %
67.2
66.7
65.7
67.1
Methanol produced, mols
6.876
6.52
5.443
5.654
Energy ratio
0.679
0.662
0.616
0.636
Thermal efficiency, %
67.2
65.3
59.3
63.0
Total fuel cycle C02-equivaient
emission reduction, kg
688
646.5
530
559
Ratio, Hynol fuel cycle emission
reduction / BCL + steam reforming
emission reduction
1.171
1.128
1.011
1.067
Ratio, sum of gasifier + reformer +
methanol converter throughputs /
methanol produced
39.8
43.4
56.1
49.5
Sum above per tonne of fuel-cycle
C02 emission avoided
398
438
576
500
•System cannot achieve 87% carbon conversion with 24 K MSR approach to equilbibrium
and 8 atm pressure drop without import of electric power (10 kW).
7-23
-------
possible is crucial to performance of the process. The value should not exceed +18 K under any operating
conditions.
The effect of approach to equilibrium in the reformer was also examined, with the results shown in
Tables 7-15 and 7-16. In this case the reactions are endothermic and the approach is therefore negative.
Values of -20 and -40 K were assumed and compared to the -10 K value assumed for the Hynol base case.
The results show that the effect is much smaller for the reformer than it is for the methanol reactor. For the
reformer operating at 950°C, a -20 K approach to equilbrium will not adversely affect performance, but the
value should probably not exceed -30 K, since a decline in performance at -40 K is evident in Table 7-16.
7-24
-------
TABLE 7-15. SIMULATION RESULTS FOR VARIABLE APPROACH TO EQUILIBRIUM
IN REFORMER (SPR)
2.7 atm SPR pressure drop, 24.8 atm SPR pressure, 950° C SPR temperature,
2.5 steam/carbon ratio, 1.20 mol purge, 100 kg biomass
SPR approach to equilibrium, K
Process variable
10
20
40
Carbon conversion in gasifier, %
87
87
78
Gasifier throughput, mols
16.93
17.25
18.36
Steam feed to reformer, mols
20.9
21.2
21.8
Steam feed to gasifier, mols
1.5
1.55
1.75
Reformer throughput, mols
54.2
54.8
56.4
Gasifier loop compressor feed, mols
38.2
38.5
39.5
Recycle loop compressor feed, mols
261
265
279
Methanol converter throughput, mols
283
289
304
Methanol product, mols
7.225
7.218
7.198
Methanol loop compressor power, kW
-41.6
-42.3
-44.6
Gasifier loop compressor power, kW
-16.6
-16.8
-17.2
SPR furnace air blower power, kW
-11.2
-11.2
-11.6
High pressure BFW pumps, kW
-3.7
-3.8
-4.2
HRSG-2 power, kW
+29.8
+30.3
+31.2
HRSG-1 power, kW
+36.6
+37.1
+34.2
Reformer fuel expander, kW
+5.9
+5.9
+6.1
Exportable (imported) power, kW
0
0
0
Power from char combustor, kW
+6.6
+6.6
+11.9
Distillation reboiler duty, kcal
82710
82550
82540
Ceramic heat exchanger duty, kcal
91920
94930
102900
NG fuel for reformer, mols
2.64
2.64
2.74
NG feed to process, mols
3.76
3.76
3.76
Reformer duty, kcal
460080
442410
444100
Reformer combustion air, mols
32.0
32.0
33.0
7-25
-------
TABLE 7-16 SUMMARY OF EFFECTS OF REFORMER (SPR) APPROACH TO
EQUILIBRIUM ON HYNOL PROCESS PERFORMANCE
Parameters calculated from data of Table 7-15
SPR approach to equilibrium, K
Process variable
10
20
40
Carbon conversion, mol %
87
87
87
Total natural gas fed, mols
6.40
6.40
6.50
C02 emission from natural gas, kg
281.6
281.6
286
Gasoline displaced, gallons
94.7
94.6
94.4
Net tailpipe C02 emission avoided, kg
571
570
564
Tailpipe C02 reduction per vehicle, %
66.9
66.9
66.3
Methanol produced, mols
7.225
7.218
7.198
Energy ratio
0.684
0.683
0.674
Thermal efficiency, %
67.7
67.6
66.6
Total fuel cycle COz-equivalent
emission reduction, kg
720.7
719.7
711.3
Ratio, Hynol fuel cycle emission
reduction / BCL + steam reforming
emission reduction
1.182
1.180
1.154
Ratio, sum of gasifier + reformer +
methanol converter throughputs /
methanol produced
49.0
50.0
52.6
Sum above per tonne of fuel-cycle
C02 emission avoided
491
502
532
7-26
-------
SECTION 8
HYNOL PROCESS OPTIMIZATION
Background
The Hyriol process evaluations discussed up to this point have established the essential features of
a base-case operating system that can achieve maximum performance with regard to C02 emission reduction.
The focus of those evaluations has been on the balance between energy consumption and energy production
without specific regard to secondary factors that affect equipment size and cost. One factor that has been
shown to be important is the throughput of each of the two process loops, as reflected by comparisons of the
ratios of the sum of throughputs to methanol product. Those comparisons indicated significant differences
between process configurations that otherwise achieve equal C02 reduction goals. The optimization of that
base case, to be discussed in this section, will address modifications that are aimed at reducing capital cost
while maintaining the C02 emission-reduction goal already defined.
The following discussion will briefly review the features of the base-case Hynol process and how it
differs from the original Hynol concept. The reasons for the modifications made will be discussed in terms of
their effects on the performance goals and practical operational requirements. Because these changes result
in a process considerably more complex than the original Hynol concept, the reasons for their inclusion in the
final optimized configuration must be made clear. Only after the full range of operating equipment has been
included in the design can final modifications to minimize cost be undertaken in a meaningful way.
The Hynol process proposed to the EPA by Steinberg and Dong (1994a&b) was configured for a
uniform pressure of 30 atm throughout the system. The conceptual flow sheet developed by Steinberg and
Dong, shown in Figure 8-1, considered only the basic reaction steps and associated enthalpy changes
necessary to convert biomass and natural gas to methanol. As indicated in Section 5, much more equipment
must be added to the conceptual flowsheet to obtain a system from which reasonable estimates of operating
efficiency can be based. These additions include, for example: heat exchangers, desulfurization equipment,
condensers, and steam generators. Each of these items imposes a pressure drop which, when added to the
pressure drops of the reformer and methanol converter {normally 3.2 atm and 8 atm, respectively), will require
a compressor in each of the two process loops to maintain gas flow through that loop. The conceptual flow
sheet also lacks thermal integration: it doesnt include a drying step for biomass or a distillation step to separate
water from the methanol product. A thermally integrated process must account for the energy consumption
of those steps as well as the two compressors needed to overcome pressure drop. Additional energy must
be accounted for to operate the air blowers required by the reformer furnace, for pumps delivering high-
pressure boiler feed water and condenser cooling water. The base case configuration of Hynol discussed in
Section 5 modified the conceptual flowsheet to include this ancillary equipment in a manner that recovers
sufficient energy to overcome pressure drops, dry the biomass, separate water from methanol, and operate
the electrical pumps and fans.
8-1
-------
Steam
20 kg
Biomass
100 kg
HYDRO-
GASIFIER
30atm
800° C
Steam for
reforming
15.8 kgmols
Water
CO
C02
ch4
HjO
H2
N,
5.62%
13.07
7.90
0.20
69.56
2.50
MeOH 1.16
METHANOL
CONVERTER
30 aim
260°C
5
Methanol
201.4 kg
T
Water
75.3 kg
19kmols
CO
13.42 %
C02
7.72
CH4
19.46
HjO
19.76
h2
37.44
n2
2.16
HjS
0.03
39.9 kmols
CO
C02
CH4
HjO
Hz
N,
21.09%
2.87
3.26
13.10
58.65
1.03
Purge gas
0.6 mol
Steam for dryer
Water
Methane
50 kg
CLEAN-UP
K, Na, CI
STEAM
REFORMER
30 atm
1000°C
FIRING
, Steam
120 kg
¦ Flue gas
Methane
fuel
36.8 kg
Air
Figure 8-1. Conceptual Hynol flowsheet of Steinberg and Dong (1994b).
8-2
-------
Initial Conditions for Process Optimization
The stream compositions shown in Figure 8-1 assume that the chemical reactions occurring in the
reformer and methanol synthesis reactor reach the equilibrium limit--a condition not feasible in practice. The
operating conditions shown for the gasifier on the other hand impose an equilibrium limit of 87 % conversion
of the carbon in the biomass fed to the gasifier. Rate processes occurring in the gasifier, including gas/solid
mass and heat transfer in the fluidized bed, together with the kinetics of the desired chemical reactions, would
therefore result in carbon conversions less than 87 %. Consequently, the modifications discussed here include
the addition of steam in sufficient amount to the gasifier to eliminate this equilibrium limitation on biomass
conversion and methanol yield. This steam, added to the recycle stream entering the ceramic heat exchanger
(HE-1), allows as much of the biomass to gasify as will be permitted by the rate processes alone.
Rate processes are also taken into account for the simulations of the reforming step and methanol
synthesis step discussed in the previous section by introduction of a "temperature approach to equilibrium" for
those reactors. Values of -10 and +12K, respectively, are known from industrial experience with those
operations to represent actual conversions achievable in operating equipment of good design, and those
values are used here. The composition of the process streams given in Section 5 for the modified Hynol
process include these adjustments for rate processes and are therefore expected to reflect achievable values
more accurately than the theoretical limits suggested by Figure 8-1.
The ratio of steam to carbon entering the reformer shown in Figure 8-1 is 0.96. In actual industrial
practice, ratios of 2.5-3.5 are required to avoid carbon deposition and catalyst deactivation during the reforming
of natural gas. EPA laboratory studies of the steam reforming of a reactor feed stream representing the
composition, temperature, and pressure of the Hynol gasifier effluent (Dong and Karwowski, 1997) have shown
that a steam/carbon ratio of 2.0 is sufficient to avoid carbon deposition on a conventional nickel reforming
catalyst. The presence of hydrogen as a major constituent of that stream assists in reducing this steam
requirement for Hynol. A ratio of 2.5 is assumed for the modified Hynol system described in Section 5 to
provide a safety factor that ensures maximum catalyst activity, and that ratio is adopted here as well.
Additional Assumptions from Sensitivity Analysis
The Hynol configuration described in Section 5 and evaluated in Section 6, which includes all of the
modifications discussed above, Is therefore the first step in optimization-establishing a workable system from
the basic concept-and represents the starting point for the refinements that will now be discussed. As
demonstrated in Section 6, the base-case Hynol configuration should achieve 68-69 % thermal efficiency and
20 % greater C02 emission reduction than the best alternative technologies for production of FCV fuel from
biomass and natural gas. The degree to which these performance projections will be affected by changes in
the principal process operating variables, as explored in the previous section, point to the following conclusions
which will be taken into account for the optimization to be made in this section of the report.
1) The approach to equilibrium in the methanol converter is critical to achieving maximum
performance. Optimum performance requires that the approach to equilibrium not exceed
appreciably the +12K assumed for the base case; e.g., 18K is too much. The equilibrium
approach in the reformer is much less important than that of the methanol reactor.
2) A reformer temperature of 950° C (reduced from the original 1000° C) will not significantly
sacrifice overall performance and will improve reformer tube life and catalyst life. A 950° C
reformer temperature is therefore considered optimum and will be assumed for all simulations
further considered here.
3) A steam/carbon ratio of 2.5 can be sustained without sacrifice of Hynol performance and will be
considered optimum for all simulations discussed in this section.
8-3
-------
4) Steam will be added in sufficient amount to the recycle stream to eliminate equilibrium constraint
on carbon conversion. An 87 % conversion of the biomass carbon is assumed, limited by mass
transfer and chemical kinetics during gasification.
5) In order to achieve maximum C02 emission reduction from the vehicle fleet using methanol
produced from biomass and natural gas, the optimum ratio of natural gas to biomass is 3.79 mols
per 100 kg. An additional 2.60 mols of natural gas is required for reformer fuel for the base case.
The natural gas feed rate (as process feedstock) will thus be maintained at 3.79 mols/100 kg of
biomass feed for these optimization simulations while minimizing the amount of natural gas
required as reformer fuel. For a given level of biomass conversion in the gasifier, this means that
all of the ungasified biomass char is used as supplementary reformer fuel (with natural gas)
instead of the alternative use for electric power production.
Assumptions Regarding Tar Formation
A crucial assumption throughout these Hynol process evaluations is that tars will not be formed during
hydrogasification of biomass at the specified conditions. Tar production is a major problem for biomass
gasifiers that utilize oxygen or air to provide the energy for gasification by in-situ partial oxidation of the
biomass. The presence of high molecular weight, condensable organic compounds in the gasification product
stream requires low temperature scrubbing or high temperature catalytic cracking for their removal before
further use of the product stream can be considered. Steinberg and Dong (1994a & b) propose that tars will
not be formed in the Hynol gasifier because their unsaturated carbon bonds will react rapidly with hydrogen.
That notion tends to be supported by laboratory tests at EPA/APPCD which have not produced visible
condensation products during biomass hydrogasification experiments at the temperature and pressure of the
Hynol gasification step (Dong and Cole, 1996). It must be noted that laboratory tests of biomass
hydrogasification conducted elsewhere (Garg et al.,1988) have resulted in substantial tar formation when
conducted at atmospheric pressure and at temperatures (550 C°) below that of Hynol. These laboratory tests
are far from definitive, however, and the question will remain open until the pilot gasification tests are
completed at CE-CERT using realistic operating equipment. This issue is one of the most important ones that
must be answered by those pilot tests.
The significance of the tar issue is that heat from the gasifier effluent can be recovered efficiently only
in the absence of condensable tars. The configuration assumed here includes heat exchangers that (1)
superheat the steam from the methanol converter so that the steam can be used to generate most of the
electric power production, (2) generate steam for the gasifier, and (3) generate part of the steam required for
the reforming step. The four heat exchangers used for these purposes are located between the gasifier and
the desulfurizer and cool the gas stream to the 268°C required for H2S elimination. If these heat exchangers
cannot be utilized due to tar condensation, a major reconfiguration of the process flow sheet will be necessary
with significant negative effects on thermal efficiency and performance projections.
Pressure and Pressure Drops
Within the limits allowed by the necessary process modifications outlined above, the basic concept
of a uniform system pressure, averaging about 30 atm, was not altered substantially in the configurations
considered so far: this pressure represents a compromise, as determined by Steinberg and Dong (1994a),
between a higher pressure which favors the methanol synthesis equilibrium and a lower pressure which favors
the equilibrium of the reforming step (and the gasification step). The pressure drops added in Section 5
resulted in a reduction of reformer pressure to 24.8 atm and an increase of methanol converter pressure to
32.5 atm. Because the pressure drops in each of the two process loops will necessarily be substantial,
requiring large compressors, and because the pressure drops assumed for heat exchangers are dependent
on their design (high pressure drops increase heat transfer coefficients which reduce area, size, and cost),
optimization must seek the combination of pressures in those two reactors that will minimize equipment size
and cost without sacrificing the performance goals established by the base case analysis.
8-4
-------
There are two primary considerations in allocating pressure drops between the two process loops.
The first consideration is the difference in pressure between the final component of the gasifier loop and the
feed pressure of the methanol converter. This pressure differential affects the energy requirements for
compression and will thus be the principal determinant of the energy demand at which the overall system can
be operated without import of electric power. The second consideration is the pressure assumed for the
methanol converter which affects the converter throughput and therefore the process cost- the converter
represents the largest single cost component of the system. A third factor must also be considered indirectly
when adjusting reactor pressure and the cumulative pressure drops in the two process loops and that factor
is the areas of the heat exchangers and condensers. A tradeoff exists between the value of energy that could
be recovered by adding heat exchanger area (and increased pressure drop) and the capital cost of adding that
area. In this analysis, the amount of heat exchange area used is determined by the condition that the
recovered energy must be sufficient to generate all of the electric power and thermal energy required to
operate the process.
Allocation of Pressure Drops in the Gasifier Loop
Figure 8-2 identifies the equipment items that impose pressure losses in the gas stream from the point
where the recycle exits the methanol synthesis loop to the point where it exits the reformer and ceramic heat
exchanger. They include the ceramic heat exchanger, the gasifier, steam superheaters, a HRSG, the
desulfurization reactor (and particulate filter), preheaters located in the flue gas exiting the reformer, and the
catalytic reforming reactor. The area of each heat exchanger is shown as determined by Aspen simulations
for 100 kg of dry biomass fed to the gasifier. A maximum pressure of 33 atm is assumed for the reformer
operating at 950°C. The pressure drop of the reformer is 3.2 atm, and the pressure drop of the gasifier is taken
to be 0.5 atm in accordance with the assumption of Katofsky (1993) for the BCL system. The pressure drop
of the desulfurization reactor/particulate filter is taken to be 2.6 atm. The pressure drop for each heat
exchanger is allocated in proportion to its area. Given these conditions, the required pressure entering the
gasifier loop as recycle from the methanol synthesis loop is 42 atm.
Figure 8-3 indicates the process equipment located after the reformer, leading to the compressor that
feeds the methanol synthesis loop. Here again, the pressure drop of each heat exchanger is allocated In
proportion to its area, using the same factor, 0.755 atm/m2, that is used in Figure 8-2. Given the 32.9 atm
pressure leaving the reformer and the 10.77 m2 total heat exchanger surface yields a final pressure of 24.8 atm
at the point where the stream leaves the gasifier loop to enter the compressor that feeds the methanol
synthesis loop.
Allocation of Pressure Drops in Methanol Synthesis Loop
Given a maximum specified pressure of 33 atm in the reformer, the pressure of the feed stream
entering the compressor feeding the methanol synthesis loop will be, as indicated above, 24.8 atm. The overall
pressure drop in the methanol converter/heat exchanger system that was given by Katofsky (1993) for an ICI-
type converter is 8 atm. For consistency in comparison, the same total pressure drop of 8 atm is assumed here
although the Lurgi converter imposes only 0.1 atm pressure drop. The remaining 7.9 atm pressure drop is
allocated to the heat exchanger that heats the converter feed and the condenser that separates the crude
methanol/water from the converter effluent. These conditions result in a compressor outlet pressure of 50.6
atm, a converter pressure of 50 atm, and a recycle pressure of 42 atm returning to the gasifier and to the
methanol loop compressor. The block diagram, Figure 8-4, shows the basic features of this configuration and
Table 8-1 gives the composition of the principal streams.
Process Optimization for 50-atm Methanol Converter
As was done in Section 5 for the base case, the following data summarize the detailed operating
stream conditions in each of the operating blocks of the Hynoi process when the methanol converter pressure
8-5
-------
Natural gas Steam -
Desutfurization
40.79 atm
38.19 atm
PGPM-1
1.737 m*
HRSG-1
0.272 m*
36.88 atm
41.00 atm
SH-G3
0.060 m'
36.12 atm
SH-G2
0.0435 m:
32.92 atm
SH-G1
0.0467 m*
41,12 atm
41.62 atm
From methanol
synthesis loop
Gasifier
To HRSG-2
Figure 8-2. Pressure drop allocation in the gasifier loop. From entry of recycle stream
from methanol converter to reformer exit.
B-6
-------
29.47 atm
30.69 atm
HE-6
0.688 rrv
28.37 atm
31.41 atm
HE-5
0.0798 m!
32.28 atm
28.23 atm
32.47 atm
SH-S1
0.0879 m1
24.79 atm
32.54 atm
To methanol
synthesis loop
HE-1
0.500 rrv
32.92 atm
HE-3
4.56 m5
HRSG-2
1.154 m'
Figure 8-3. Pressure drop allocation in gasifier loop. From reformer exit to entry at
gasifier loop compressor.
8-7
-------
Biomass
51.06 kg C
7.49 kg H
51.53 kg O
0.15 kg N
0.08 kg S
0.792 kg ash
Natural gas
3.79 mols
Steam 20.45 mols
Desuifurizer
HRSG
Air
31.0 mols
Filter
Reformer
32.9 atm
Gasifler
41.1 atm
Combusior
Unreacted carbon
0.549 mol
Natural gas (fuel)
2.50 mols
Steam 1.6 mols
HjO condensate
15.71 mols
Ceramic
heat
exhanger
HRSG
4 [_5j
Purge 1.2 mols
(to reformer fuel)
50 atm
Methanol
converter
CH3OH 7.224 mols
H,0 1.947 mols
Figure 8-4. Block diagram of Hynol process optimized for 50-atm methanol synthesis.
8-8
-------
TABLE 8-1. STREAM DATA FOR 50-atm OPTIMIZATION (FIGURE 8-4)
Stream composition, kg-mols
Stream
1
Stream
2
Stream
3
Stream
4
Stream
5
Stream
6
Stream
7
Stream
8
H,0
3.364
0.1541
0.0725
0.153
0.0006
1.606
0
0
H,
5.782
26.184
106.06
132.25
0.9040
8.761
0
0
CO
1.443
6.050
6.619
12.67
0.0564
0.5468
0
0
CO,
0.9037
2.582
6.997
9.58
0.0596
0.5780
0.00758
0.0050
I
o
3.365
0.8890
9.650
10.540
0.0823
0.7972
3.589
2.367
C,HS
0
0
0
0
0
0
0.1061
0.0700
N,
0.8602
0.9474
10.347
11.29
0.0882
0.8547
0.08717
0.0575
CH,OH
0
0
1.038
1.038
0.0088
0.0857
0
0
SUM
15.72
36.81
140.8
177.5
1.20
13.23
3.79
2.50
Deg. C
800
51
40
51
40
900
56
90
Aim
41.12
24.8
42.0
50.6
42.0
41.62
40.79
1.5
8-9
-------
is increased from the base case of 32.5 atm to the optimized pressure of 50 atm: Figure 8-5 and Table 8-2
show details of stream operating conditions and heat exchange duties for the gasification block, Figure 8-6 and
Tables 8-3 and 8-4 show details of the streams and heat exchangers comprising the reforming block, Figure
8-7 and Table 8-5 show the distillation block, and the methanol synthesis block is given in Figure 8-8 and Table
8-6. The biomass block is the same as the base case.
It is evident from these data (which will be compared more explicitly later) that a substantial reduction
of throughput is achieved in the methanol synthesis loop by raising the converter pressure from 32.5 to 50 atm.
This corresponds to a 40% reduction in throughput (and a similar reduction in size and cost). Although the
pressure in the gasifer loop has been increased relative to the base case (40.79 atm entering
thedesutfurization unit, compared to 28 atm in the base case), the throughputs of the gasifier and reformer are
not significantly altered by the change in converter pressure. The only other equipment change is the addition
of a compressor for the natural gas feed to the process. This compressor, which raises the pressure from 28.5
to 40.79 atm, draws 1.15 kW per 3.79 mols of natural gas fed.
Effect of Methanol Converter Pressure >50 atm
The optimization data for 50-atm converter pressure discussed to this point utilized the pressure drop
imposed by equipment in the gasifier loop to establish a maximum pressure differential between the reformer
and methanol synthesis reactor. The result is a 44% reduction of the recycle ratio in the methanol loop and
a 40% reduction of converter throughput. These effects result without imposing significant change in the
operating condition of the gasifier loop.
This section examines the effect of further increase of converter pressure. Table 8-7 summarizes
significant results of simulations carried out with assumed pressures up to 90 atm. When the converter
pressure is raised above 50 atm, the pressure differential between it and the reformer cannot be
accommodated by the cumulative pressure drops of equipment in the gasifier loop; an expander must be
added between the two loops. Addition of the expander allows the pressure to be reduced to the 42-atm
maximum needed for the gasification loop and recovers additional electric power from the clean gas stream
entering the gasifier loop from the methanol loop. It is evident from the results shown in Table 8-7 that the
recycle ratio and methanol converter throughput continue to decrease as converter pressure increases. It is
also significant to note that the heat exchanger area and condenser area decrease (offset to some degree by
higher heat transfer area in the converter for steam production). Throughputs and operating conditions of the
gasification loop are not significantly affected.
Process Optimization with 90-atm Methanol Converter
Figure 8-9 is a block diagram for the 90-atm case summarized in Table 8-7. It shows the expander
located between the two process loops and the main inputs to the process. Compositions of the principal
process stream are given in Table 8-8. As done for the base case and the 50-atm optimization, detailed data
on stream operating conditions for the gasification block are given in Table 8-9; the reformer block, shown in
Figure 8-10, includes changes in the reformer of Figure 8-6 that include a new heat exchanger to preheat boiler
feed water for the methanol converter/steam generator and deletes electric power recovery due to that
increased duty. Also shown for the 90-atm case is a new stream added to the reformer fuel. This stream
consists of non-condensabies (CO, C02, CH4, H2, and N2) from the distillation step. These components are
slightly soluble in methanol at 90 atm pressure (the amount is not significant at 50 atm, and is not considered
in that simulation) and are responsible for the slight reduction in methanol yield as pressure is increased. They
are sfripped from the methanol and water during distillation and can be either returned to the system to
increase methanol yield or purged as fuel to the reformer. Simulations of the two approaches showed the latter
to be more effective. Other details of the reformer block are given in Table 8-10. Table 8-11 summarizes
details for the distillation block and Table 8-12 shows the stream data for the methanol synthesis block. Heat
exchanger duties for the 90-atm optimized system are shown in Tables 8-13 through 8-17.
8-10
-------
Desulfurization
HRSG-1
BFW
Comp-3
Natural Gas
SH-G3
TG-2
SH-G2
TG-1
•119
SH-G1
Gasifer
Biomass
HE-1
HEAT DUTIES FOR GASIFIER BLOCK. 50-atm OPTIMIZATION
I.D.
Description
Duty, cal/sec
Surface area. m*
SH-G1
Steam superheater
4698
0.0546
SH-G2
Steam superheater
2380
0.0438
SH-G3
Steam superheater
2290
0.0600
HRSG-1
Heat recovery steam generator
10890
0.2555
Figure 8-5. Details of gasifier block for 50-atm optimization.
8-11
-------
1
2
3
4
5
6
7
B
9
10
11
12
13
14
15
16
17
18
19
20
22
24
25
B3
TABLE 8-2. STREAM DATA FOR GASIFICATION BLOCK (FIGURE 8-5)
I.D.
Flow,
kg-mol
Temp.,
°C
Pressure,
atm
Dried biomass, 10 wt% moisture
111-1 (kg)
25
1
Heated recycle from methanol block
13.23
900
41.62
Gasifier effluent
15.72
800
41.12
Urireacted carbon
0.549
800
41.12
Steam superheater SH-G1 effluent
15.72
697
41.08
Steam superheater SH-G2 effluent
15.72
645
41.05
Steam superheater SH-G3 effluent
15.72
592
41.00
Heat recovery steam generator HRSG-1 effluent
15.72
329
40.79
Natural gas feed to process
3.79
56
40.79
Superheated I.P. process steam to reformer
5.19
527
38.19
Boiler feed water
3.66
116
42.5
Sat. I. P. steam
3.66
310
42.25
Superheated I.P. steam for SPR
2.06
500
38.19
Process steam for gasifier
1.60
527
42.0
Natural gas
6.29
25
28.5
Saturated I.P. steam from methanol block
14.63
250
39.2
Feed to ZnO desulfurization unit
19.51
268
40.79
Superheated steam for reformer
3.13
527
38.19
Superheated steam
11.50
399
38.8
Turbine condensate
11.50
30
0.05
Heat exchanger HE-1 cold-side feed
13.23
108
42.0
Desulfurized gaslfer effluent: reformer feed
19.51
268
38.2
Natural gas fuel to HE-5 and reformer furnace
2.50
25
28.5
Recycle stream from methanol block
11.63
40
42.0
8-12
-------
Air
30|
APH-t
TG-4
TG-3
271
From gasitiar block
404 "C
591 *C
28
:32
HE-1
37!
SH-SZ
91
SH-Sl
HRSG-2
4af
Condensate (rem driw
44
Figure 8-6. Details of reformer block for 50-atm optimization.
8-13
-------
TABLE 8-3. STREAM DATA FOR REFORMER AND POWER BLOCK (FIGURE 8-6)
Stream
No.
I.D.
Flow,
kg-mol
Temp.,
°C
Pressure,
atm
10
Process steam from superheater SH-G3
5.19
527
36.9
24
Process feed stream from gasifier
19.51
268
38.2
25
Natural gas reformer fuel
2.50
90
28.5
26
Purge gas from methanol synthesis block
1.2
40
42.0
27
Natural gas and purge fuel for SPR furnace
3.70
30
1.5
28
Fuel to furnace
3.70
377
1.0
29
Air inlet
31.0
25
1.0
30
Air to first preheater
31.0
69
1.5
31
Air to second preheater
31.0
320
1.25
32
Air to furnace
31.0
533
1.0
33
Flue gas
34.3
125
1.0
34
Superheated steam to reformer
15.26
527
38.19
35
Preheated process gas to SPR
39.96
697
36.1
36
Reactor effluent
52.52
950
32.92
37
Heat exchanger HE-1 hot-side effluent
52.52
765
32.54
38
Superheater SH-S1 hot-side effluent
52.52
719
32.47
39
Superheater SH-S2 hot-side effluent
52.52
621
32.28
40
Heat recovery steam generator HRSG-3 effluent
52.52
164
31.41
41
Saturated HP steam
14.90
310
97.3
42
Superheated HP steam to turbine/generator TG-2
14.90
527
97.3
43
Turbine/generator TG-2 exhaust steam
14.90
405
38.2
44
HRSG-2 hot-side effluent
52.52
285
31.41
45
Heat exchanger HE-6 hot-eide effluent
52.52
199
30.89
46
Boiler feed water from reboiler/HE-2 condensate
12.48
129
97.3
67
Boiler feed water from HE-3
14.63
116
40.0
68
Boiler feed water to methanol converter
14.63
246
39.5
91
Steam to drier
5.59
149
3.0
8-14
-------
TABLE 8-4. HEAT DUTIES FOR REFORMER BLOCK, 50-atm OPTIMIZATION
I.D.
Description
Duty, cal/sec
Surface area, m2
HE-1
Ceramic heal exchanger
23840
0.500
SH-S1
IP steam reheater
5750
0.100
SH-S2
HP steam superheater
12240
0.235
HRSG-2
Heat recovery steam generator
39900
1.256
HE-6
Boiler feed water heater for MSR
9910
0.8456
NGPH
Natural gas preheater
181
0.026
APH-1
First air preheater
15380
1.968
APH-2
Second air preheater (in biomass block)
13720
0.153
FPH-1
Furnace fuel gas preheater
943
0.022
FPH-2
Furnace fuel gas reheater
3690
0.154
PGPH-1
First process gas preheater
14480
1.737
PGPH-2
Second process gas preheater
19860
1.000
HRSG-3
Drier steam generator
15337
1.878
SPR
Steam pyrolysis reactor (reformer)
122070
Skin loss
4500
8-15
-------
COND-2
Prom
fwfBrmer btock
[«Ol
-59'
HE-2
methanol
*ym*»sis otoc»c
M
ME-5
w
461
HEAT DUTIES FOR DISTILLATION BLOCK: 50-atm OPTIMIZATION
I.D.
Description
Dutv, cal/sec
Surface area. mJ
Reboiler
Distillation column reboiler
22810
1.467
HE-2
Column feed prehaater
1820
0.117
HE-3
Condenser, boiler feed water heater
14680
4.789
HE-5
Natural qas preheater
425
0.033
COND-2
Condenser, column overheads
22600
Figure 8-7. Details of distillation block for 50-atm optimization.
8-16
-------
TABLE 8-5. STREAM DATA FOR DISTILLATION BLOCK (FIGURE 8-7)
Stream
No.
I.D.
Flow,
kg-mol
Temp.,
°C
Pressure,
atm
26
Purge to reformer fuel, SPR block
1.2
40
42.0
40
Effluent from reformer block, HRSG-2
52.52
164
31.41
21
Natural pas
2.50
25
28.5
23
Preheated natural gas
2.50
90
28.5
46
Boiler feed water to HRSG-2
12.48
129
97.3
47
Reboiler effluent
52.52
130
28.37
49
Column feed preheater, HE-2 effluent
52.52
125
28.29
50
Water condensate
12.48
129
28.29
51
Condenser feed
40.33
125
28.29
52
Condensate water
3.19
50
24.8
53
Condenser effluent to gasifier loop compressor
36.85
50
24.8
54
Feed to methanol synthesis reactor block
36.85
53
50.6
55
Effluent from methanol condenser
153.7
40
42.0
83
Recycle to gasifier
11.63
40
42.0
57
Recycle to methanol converter
140.8
40
42.0
58
Crude methanol to distillation column
9.171
40
42.0
59
Distillation column feed
9.171
65
1.0
60
Reboiler effluent
9.171
103
1.0
61
Methanol product
7.224
30
1.0
62
Column bottoms
1.926
103
1.0
63
Cooling water
32.3
25
1.0
64
Water to pond
14.9
116
1.0
65
Boiler feed water for HRSG-2
2.78
116
97.3
11
Boiler feed water for HRSG-1
3.62
116
42.5
67
Boiler feed water to HE-4 and MSR block
14.63
116
39.2
8-17
-------
HRSG-4
Methanol
synthesis
reactor
HE-4
Coinp-2
«—©
To gasifier
-n—~
Cond-3
T3
HEAT DUTIES, METHANOL SYNTHESIS BLOCK: 50-atm OPTIMIZATION
l.D.
Description
Duty, cal/sec
Surface area, m2
HE-4
Heat exchanqer, reactor feed preheater
64060
9.559
HE-6
Heat exchanqer, reactor boiler feed water
9910
0.8456
Cond-3
Crude methanol condenser
39100
5.81
MSR
Methanol synihesis reactor
30180
HRSG-4
Steam generator
30180
Figure 8-8. Details of methanol synthesis block, 50-atm optimization.
8-18
-------
16
26
54
55
56
57
68
69
70
71
72
74
75
77
78
TABLE 8-6. STREAM DATA FOR METHANOL SYNTHESIS BLOCK (FIGURE 8-8)
I.D.
Flow,
kg-mol
Temp.,
°C
Saturated steam to gasifier, HE-1
14.63
250
Purge to reformer furnace (fuel)
1.20
40
Feed from gasifier/reformer loop
36.81
67
Crude methanol condenser effluent
162.9
40
Recycle to gasilier
11.63
40
Recycle to methanol loop compressor
140.8
40
Boiler feed water from HE-6, distillation block
14.63
246
MSR recycle compressor effluent
140.8
48
Methanol synthesis reactor feed
177.5
51
Heat exchanger HE-4 effluent/converter feed
177.5
225
MSR reactor effluent
162.9
260
Heat exchanger HE-4 effluent
162.9
95
Condenser cooling water
800
25
Water to distillation condenser, and compressor
interstage coolers
73
25
Water to pond
727
35
8-19
-------
-r C -W 1 - I (
TABLE 8-7. EFFECT OF METHANOL SYNTHESIS REACTOR (MSR) PRESSURE ON THE
MSR RECYCLE RATIO, HEAT EXCHANGER REQUIREMENTS, AND STEAM PRODUCTION
Basis: 100 kg dry biomass, 3.79 mols natural gas feed to process
MSR pressure, atm
32.5
(base case)
50
60
75
90
MSR recycle rate, mols
249
140.8
100
80.3
63.0
Recyle ratio
6.8
3.8
2.7
2.2
1.74
MSR throughput, mols
220
163
123
102
84.6
MSR steam production, mols
9.19
14.6
16.3
16.98
17.70
Converter feed temperature, °c
217
225
225
225
225
MSR heat exchanger area, m2
17.1
9.56
7.46
6.09
4.86
MSR condenser area, m2
16.2
5.81
5.1
4.36
3.79
Total heat exchange area, m5
33.3
15.37
12.56
10.45
8.65
Gasifier throughput, mols
15.26
15.72
15.87
15.15
14.84
Reformer throughput, mols
51.4
52.52
52.4
51.61
51.00
Reformer duty, kal/sec
127.3
122.0
122.3
121.5
121.0
Methanol produced, mols
7.293
7.224
7.254
7.210
7.202
8-20
-------
Biomass
51.06 kg C
7.49 kg H
51.53 kg O
0.15 kg N
0.08 kg S
0.792 kg ash
Steam 19.82 mols
Natural gas
3.79 mols
0-0
HRSG
Desulfurizer
Air
31.3 mols
Filter
Reformer
32.4 atm
Gasifier
41.1 atm
Combustor
Unreacted carbon
0.549 mol
Distillation purge Natural gas (fuel)
0.1627 mol 2.486 mols
Steam 1.15 mols
HjO condensate
14.77 mols
Ceramic
heat
exhanger
HRSG
Purge 1.2 mols
(to reformer fuel)
90 atm
Methanol
converter
CH3OH 7.202 mols
H20 1.763 mols
Figure 8-9. Block diagram of Hynol process optimized for 90-atm methanol synthesis.
8-21
-------
TABLE 8-8. STREAM DATA FOR OPTIMIZED HYNOL PROCESS WITH
90- atm METHANOL CONVERTER (FIGURE 8-9)
Stream composition, kg-mols
Stream
1
Stream
2
Stream
3
Stream
4
Stream
5
Stream
6
Stream
7
Stream
8
H,0
3.026
0.1515
0.0173
0.0735
0.0003
1.153
0
0
H,
5.679
26.051
49.41
75.46
0.9418
8.869
0
0
CO
1.237
5.863
1.589
7.452
0.0303
0.2852
0
0
o
o
o
0.7090
2.366
2.709
5.074
0.0516
0.4862
0.00758
0.0050
ch4
3.398
0.9241
4.557
5.481
0.0869
0.8179
3.589
2.354
c,hr
0
0
0
0
0
0
0.1061
0.0696
N,
0.7944
0.8815
4.40
5.276
0.0838
0.7890
0.08717
0.0572
CH,OH
0
0
0.279
0.276
0.0053
0.0495
0
0
SUM
14.84
36.24
62.95
99.1
1.20
12.45
3.79
2.486
Deq. C
800
51
40
51
40
900
56
90
Atm
41.16
24.8
82.0
90.6
82.0
41.66
40.79
1.5
8-22
-------
TABLE 8-9. STREAM DATA FOR GASIFICATION BLOCK, 90-atm OPTIMIZATION {FIGURE 8-5)
Stream
No.
I.D.
Flow,
kg-mol
Temp.,
°C
Pressure,
atm
1
Dried biomass, 10 wt% moisture
111.1 (kg)
25
1
2
Heated recycle from methanol block
12.45
900
41.66
3
Gasifier effluent
14.84
800
41.16
4
Unreacted carbon
0.549
800
41.16
5
Steam superheater SH-G1 effluent
14.84
664
41.11
6
Steam superheater SH-G2 effluent
14.84
596
41.07
7
Steam superheater SH-G3 effluent
14.84
552
41.03
8
Heat recovery steam generator HRSG-1 effluent
14.84
333
40.85
9
Natural gas feed to process
3.79
56
40.85
10
Superheated I.P. process steam to reformer
5.60
527
38.25
11
Boiler feed water
2.87
116
42.5
12
Sat. I. P. steam
2.87
310
42.25
13
Superheated I.P. steam for SPR
1.72
500
38.25
14
Process steam for gasifier
1.15
527
42.0
15
Natural gas
6.30
25
28.5
16
Saturated I.P. steam from methanol block
17.70
250
39.2
17
Feed to ZnO desulfurization unit
18.63
268
40.85
18
Superheated steam for reformer
3.88
527
38.25
19
Superheated steam
13.82
402
38.8
20
Turbine condensate
13.82
30
0.05
22
Heat exchanger HE-1 cold^side feed
12.45
51
42.0
24
Desulfurized gasifer effluent: reformer feed
18.63
268
38.25
25
Natural gas fuel to HE-5 and reformer furnace
2.51
25
28.5
83
Recycle stream from expander and methanol block
11.30
40
42.0
8-23
-------
12TC
NGPH
130°C
338 C
FPH-1
348 C
PPH-2
From gassier btocfc
24
398°C
APH-2
P GPH-2
812°C
Reformer
To distillation
j9l]—Todriar
SH«S1
HRSG-3
HRSG-2
Condensate from aher
Figure 8-10. Reformer block for 90-atm Hynol optimization.
8-24
-------
TABLE 8-10. STREAM DATA FOR REFORMER AND POWER BLOCK (FIGURE 8-10)
Stream
No.
I.D.
Flow,
kg-mol
Temp.,
°C
Pressure,
atm
10
Process steam from superheater SH-G3
5.45
527
38.25
24
Process feed stream from gasifier
18.63
268
38.25
25
Natural gas reformer fuel
2.51
90
28.5
26
Purge gas from methanol synthesis block
1.2
40
42.0
27
Natural gas and purge fuel for SPR furnace
3.71
30
1.5
28
Fuel to furnace
3.71
377
1.0
29
Air inlet
31.3
25
1.0
30
Air to first preheater
31.3
69
1.5
31
Air to second preheater
31.3
315
1.25
32
Air to furnace
31.3
477
1.0
33
Flue gas
34.6
127
1.0
34
Superheated steam to reformer
14.22
527
38.25
35
Preheated process gas to SPR
38.45
697
35.57
36
Reactor effluent
51.00
950
32.37
37
Heat exchanger HE-1 hot-side effluent
51.00
760
32.03
39
Superheater SH-S1 hot-side effluent
51.00
688
31.86
40
Heat recovery steam generator HRSG-3 effluent
51.00
165
29.46
41
Saturated HP steam
14.22
310
39.5
42
Purge from distillation block (non-condensibles)
0.1627
30
1.0
44
HRSG-2 hot-side effluent
51.00
330
31.20
45
Heat exchanger HE-6 hot-side effluent
51.00
222
30.76
46
Boiler feed water from reboiler/HE-2 condensate
11.37
129
40.0
65
Boiler feed water from condenser HE-3
3.36
116
40.0
67
Boiler feed water from HE-3
17.70
116
40.0
68
Boiler feed water to methanol converter
17.70
246
39.5
91
Steam to drier
5.59
149
3.0
95
Condensate water from drier
5.59
110
3.0
8-25
-------
TABLE 8-11. STREAM DATA FOR DISTILLATION BLOCK 90-atm OPTIMIZATION
Stream
No.
I.D.
Flow,
kg-mol
Temp.,
°C
Pressure,
atm
21
Natural gas
2.51
25
28.5
23
Preheated natural gas
2.51
90
28.5
26
Purge gas to reformer fuel
1.2
40
42.0
40
Effluent from reformer block, HRSG-3
51.00
164
29.46
46
Boiler feed water to HRSG-2
11.37
129
40.0
47
Reboiler effluent
51.00
132
28.39
49
Column feed preheater, HE-2 effluent
51.00
127
28.31
50
Water condensate
11.37
131
28.31
51
Condenser feed
39.45
127
28.29
52
Condensate water
3.36
50
24.8
53
Condenser effluent to gasifier loop compressor
36.24
50
24.8
54
Feed to methanol synthesis reactor block
36.24
53
90.6
55
Effluent from methanol condenser
84.58
40
82.0
56
Recycle to gasifier
11.30
40
82.0
57
Recycle to methanol converter
62.95
40
82.0
58
Crude methanol to distillation column
8.965
40
82.0
59
Distillation column feed
8.965
65
1.0
60
Reboiler effluent
8.965
103
1.0
61
Methanol product
7.202
30
1.0
62
Column bottoms
1.781
103
1.0
63
Cooling water
33.3
25
1.0
64
Water to pond
9.88
116
1.0
65
Boiler feed water for HRSG-2
2.85
116
40.0
66
Boiler feed water for HRSG-1
2.87
116
43.0
67
Boiler feed water to HE-6 and MSR block
17.70
116
39.2
8-26
-------
TABLE 8-12. STREAM DATA FOR METHANOL SYNTHESIS BLOCK, 90- atm OPTIMIZATION
Stream
No.
I.D.
Flow,
kg-mol
Temp.,
°C
Pressure,
atm
19
Saturated steam to qasifier, HE-1
17.70
250
39.2
54
Feed from gasifier/reformer loop
36.24
67
90.6
55
Crude methanol condenser effluent
84.58
40
82.0
56
Recycle to qasifier
11.30
40
82.0
57
Recycle to methanol loop compressor
62.95
40
82.0
67
Boiler feed water from HE-3, distillation block
17.70
246
39.5
69
MSR recycle compressor effluent
62.95
48
90.6
70
Methanol synthesis reactor feed
99.1
51
90.6
71
Heat exchanger HE-4 effluent/converter feed
99.1
225
90.1
72
MSR reactor effluent
84.58
260
90.0
74
Heat exchanger HE-4 effluent
84.58
119
86.0
75
Condenser cooling water
800
25
1.0
77
Water to distillation condenser, and compressor
interstage coolers
73
25
3.0
78
Water to pond
727
33
1.0
8-27
-------
TABLE 8-13. HEAT DUTIES FOR GASIFIER BLOCK, 90-atm OPTIMIZATION
I.D.
Description
Duty, cal/sec
Surface area, m2
SH-G1
Steam superheater
5760
0.0697
SH-G2
Steam superheater
2945
0.0636
SH-G3
Steam superheater
1790
0.0637
HRSG-1
Heat recovery steam generator
8544
0.2150
TABLE 8-14. HEAT DUTIES FOR REFORMER BLOCK, 90- iatm OPTIMIZATION
I.D.
Description
Duty, cal/sec
Surface area, m2
HE-1
Ceramic heat exchanger
23660
0.4725
SH-S1
IP steam reheater
8774
0.1440
HRSG-2
Heat recovery steam generator
41610
0.9186
HE-6
Boiler feed water heater for MSR
12020
0.627
NGPH
Natural gas preheater
181
0.026
APH-1
First air preheater
15200
1.981
APH-2
Second air preheater (in biomass block)
10450
0.057
FPH-1
Furnace fuel gas preheater
943
0.023
FPH-2
Furnace fuel gas reheater
3620
0.160
PGPH-1
First process gas preheater
13900
2.180
PGPH-2
Second process gas preheater
19190
1.275
HRSG-3
Drier steam generator
15337
1.806
SPR
Steam pyrolysis reactor (reformer)
121020
Skin loss
4350
8-28
-------
TABLE 8-15. HEAT DUTIES FOR DISTILLATION BLOCK, 90-atm OPTIMIZATION
I.D.
Description
Duty, cal/sec
Surface area, m2
Reboiler
Distillation column reboiler
22700
1.470
HE-2
Column feed preheater
1760
0.112
HE-3
Condenser, boiler feed water heater
15240
4.685
HE-5
Natural gas preheater
477
0.0361
COND-2
Condenser, column overheads
22460
TABLE 8-16. HEAT DUTIES FOR METHANOL SYNTHESIS BLOCK, 90-atm OPTIMIZATION
I.D.
Description
Duty, cal/sec
Surface area, m2
HE-4
Heat exchanger, reactor feed preheater
35700
4.862
HE-6
Heat exchanger, reactor boiler feed water
12020
0.627
Cond-3
Crude methanol condenser
30660
3.79
MSR
Methanol synthesis reactor
36050
HRSG-4
Steam generator
36050
TABLE 8-17. HEAT DUTIES FOR BIOMASS BLOCK, 90- atm OPTIMIZATION
I.D.
Description
Duty, cal/sec
Surface area, m2
CAP
Air preheater for char combustor
1380
0.0799
CSG
Char combustor steam generator
2640
0.0560
CSH
Steam superheater
632
0.0101
APH-2
Second-stage SPR furnace air preheater
10450
0.0570
HRSG-3
Drier steam generator (in reformer block)
15337
1.806
COND-4
Drier overhead steam condenser
12000
1.918
8-29
-------
Unlike the base case and 50-atm optimization, in which all of the electric power requirements are
obtained from heat recovered within the process, the 90-atm case needs 4.6 kW of power from the char
combustor; this reduces the reformer combustor air preheat available from the char combustor by about 24%,
but does not result in any imported power. Electric power consumption and production for the 50- and 90-atm
optimizations are itemized in Tables 8-18 through 8-21 which show that the total power required is not
significantly increased by raising the converter pressure to 90 atm and in both cases are less than the power
required for the base case (76 kWh per 100 kg of biomass fed).
Summary of Optimized Hvnol Process Performance
Table 8-22 compares performance of the two optimized systems with the performance of the base
case previously established. These comparisons include parameters calculated from the tabulated data that
indicate C02 emission reduction potential. It Is clear that the performance of the base case in terms of gasoline
displacement, thermal efficiency, and improvement of C02 reduction relative to the BCL plus conventional
steam reforming are maintained while the overall process equipment size (and cost) are greatly reduced. The
conclusion from these comparisons is that the process operating with a 90-atm methanol converter represents
the best overall performance, and cost estimates that will be discussed in Section 10 will therefore be made
on that basis.
TABLE 8-18. ELECTRIC POWER CONSUMPTION: 50-atm HYNOL OPTIMIZATION
I.D.
Unit Description
Power required, kWh
COMP-2
Methanol loop compressor
23.15
COMP-1
Gaslfier loop compressor
24.39
B1
Reformer furnace air FD fan
10.9
P3
HP boiler feed water pump for HRSG-4
1.0
P5
HP boiler feed water pump for HRSG-2
1.56
P1
MP boiler feed water pump for HRSG-2
0.83
LH
Lock hopper
1.7
COMP-3
Natural qas compressor for process feed
1.15
B3
Char combustor air FD fan
1.1
B2
Drier steam recycle fan
12.
P4
Low pressure cooling water pump, Cond-2,-3,
and compressor cooler
1.95
P2
Low pressure coolinq water pump for HE-3
0.11
Total power requirement
69.0
8-30
-------
TABLE 8-19. ELECTRIC POWER PRODUCTION: 50-atm HYNOL OPTIMIZATION
I.D.
Source
Power, kWh
HRSG-1
Heat recovery steam generator, TG-1
47.31
HRSG-2
Heat recovery steam generator, TG-2
15.18
TG-3,4
Natural gas and purge expander
6.41
Total power available
68.90
TABLE 8-20. ELECTRIC POWER CONSUMPTION: 90-atm HYNOL OPTIMIZATION
I.D.
Unit Description
Power reguired, kWh
COMP-2
Methanol loop compressor
5.63
COMP-1
Gasifier loop compressor
45.89
B1
Reformer furnace air FD fan
11.0
P3
IP boiler feed water pump for HRSG-4
1.23
P5
IP boiler feed water pump for HRSG-2
0.65
P1
MP boiler feed water pump for HRSG-1
0.21
LH
Lock hopper
1.7
COMP-3
Natural gas compressor for process feed
1.15
B3
Char combustor air FD fan
1.1
B2
Drier steam recycle fan
1.2
P4
Low pressure cooling water pump, Cond-2,-3,
and compressor cooler
1.95
P2
Low pressure cooling water pump for HE-3
0.11
Total power requirement
71.82
8-31
-------
TABLE 8-21. ELECTRIC POWER PRODUCTION: 90-atm HYNOL OPTIMIZATION
l.D.
Source
Power, kWh
HRSG-1
Gasifer heat recovery steam qenerator, TG-1
57.04
HRSG-2
Reformer heat recovery steam qenerator
0
HRSG-3
Char combustor heat recovery steam generator
4.6
TG-5
Recycle expander
3.70
TG-3,4
Natural qas and purge expander
6.41
Total power available
71.75
TABLE 8-22. SUMMARY OF PERFORMANCE PARAMETERS FOR
50- AND 90-atm HYNOL OPTIMIZATIONS
Process parameter
MSR Pressure, atm
32.5 (base case)
50
90
Natural gas used as process feed, mols
3.79
3.79
3.79
Natural gas used as reformer fuel, mols
2.60
2.50
2.486
COz emission from natural gas, kg
282.5
278.1
277.5
Methanol produced, mols
7.293
7.224
7.202
Gasoline displaced, gallons
95.6
94.7
94.4
Tailpipe C02 reduction per vehicle, %
67.3
67.5
67.5
Energy ratio
0.69
0.69
0.69
Thermal efficiency, %
68.4
68.5
68.4
Total fuel-cycle C02 emission reduction, kg
731
727
724
Ratio, Hynol fuel-cycle emissions reduction/
BCL + steam reforming reduction
1.200
1.205
1.202
Ratio, sum of gasifier, reformer, and
methanol converter throughputs/methanol
product
46.1
32.0
20.9
Sum above per tonne of fuel cycle COz
emission reduction
465
318
208
8-32
-------
SECTION 9
HYDROGEN PRODUCTION
Introduction
Fuel cells operate by chemical reduction of hydrogen at the cell's anode (which yields the electrical
energy) and protons that react with oxygen ions at the cathode. The hydrogen can be obtained directly from
cylinders of compressed gas, from liquid hydrogen, or indirectly from a liquid hydrogen-carrier that can be
reformed onboard a vehicle to produce the required hydrogen. Methanol is the most easily reformed hydrogen
carrier, is easily transported and stored, and provides for rapid refueling and longer range than compressed
hydrogen. Development of fuel cells that can utilize methanol directly without reforming is also progressing
steadily, but fuel cell vehicles will achieve highest efficiency when using hydrogen fuel. Hydrogen does not
require an onboard reformer, and the energy loss associated with the reforming step is thus avoided. It does,
however, require compression to pressures of at least 340 atm (5000 psi) in order to provide the minimum
driving range comparable to conventional gasoline vehicles with an acceptable volume of hydrogen tanks. The
energy required for compression therefore detracts somewhat from the energy otherwise saved by conversion
of syngas to methanol and subsequent reforming of methanol to hydrogen.
This section examines the reconfiguration of the Hynol process to produce hydrogen instead of
methanol. Because natural gas is the normal feedstock for conventional technologies that produce either
methanol or hydrogen, its use as Hynol co-feedstock makes possible the leveraging of hydrogen yield from
biomass in the same manner as for methanol production, previously discussed. Katofsky (1993) has shown
that biomass can be converted to hydrogen by any of five gasification processes, of which the BCL system was
found to be most efficient and lowest cost. The reconfiguration requires only replacement of the methanol
converter and distillation system with shift converters and pressure swing adsorbers (PSAs). Two such
reconfigurations of the Hynol process were considered here: the first assumed that the reformer operates at
the 1000° C temperature specified for the original Hynol concept, and the second assumes that the reformer
is operated at 950°C, as shown in the previous section to provide equal performance for methanol production.
Because the results of the simulations show that the lower reforming temperature will yield similar performance
also for hydrogen production, the detailed summaries of operating conditions that follow will be based on a
reformer temperature of 950° C. As for the methanol simulations, the feedstocks are assumed to consist of
100 kg of dry woody biomass containing 10 wt% moisture and 3.79 kg mols of natural gas, and a steam/carbon
ratio of 2.5 for the reforming step.
Process Modifications for Hvdroaen Production
Figure 9-1 shows the block diagram for hydrogen production. Table 9-1 gives the composition of the
principal process streams identified in Figure 9-1. The gasification block consists of the biomass gasifier
operating at 800° C and 30.5 atm pressure, and a heat recovery steam generator (HRSG-1) and superheater
that provide steam for the gasifier and part of the steam for the reforming step. Natural gas preheated to 130°
C is added to the process stream leaving the HRSG, giving a temperature of 268° C entering the ZnO
desulfurizer. Details of the gasification block are shown in Figure 9-2, and conditions of each process stream
identified are indicated in Table 9-2.
9-1
-------
Natural gas
3.79 mols
Desulfurizer
Biomass (10% moisture)
51.06 kg C
7.49 kg H
51.52 kg O
0.151 kg N
Filter
HRSG-1
Natural gas (fuel)
2.633 mols
Steam
Reformer
Gasifier
Unreacted
carbon
0.549 mol
Air
32.5 mols
Furnace
Heat
exchanger
Steam 0.57 mol
0-
Shift
Purge 1.2 mols 4.
(to reformer fuel)
Shift
PSA
H, 21.74 mols
75atm
BFW
PSA
85 mols
25° C
C02 7.42 mols
H5O 0.156 mol
Condensate
7.48 mols
Figure 9-1. Block diagram for Hynol process configured for hydrogen production.
9-2
-------
TABLE 9-1. STREAM DATA FOR HYNOL CONFIGURED FOR HYDROGEN PRODUCTION (FIGURE 9-1)
Stream composition, mols
Stream
1
Stream
2
Stream
3
Stream
4
Stream
5
Stream
6
Stream
7
Stream
8
Stream
9
Stream
10
Stream
11
H,0
2.240
13.057
10.062
7.632
0.1562
0
0
0
0.570
0
0
H,
6.214
26.593
29.588
32.018
32.45
1.0122
9.697
0.453
9.262
0
0
CO
1.042
5.761
2.766
0.3361
0.350
0.0331
0.3173
0.0142
0.3030
0
0
CO,
0.4055
1.995
4.990
7.420
7.420
0
0
0
0
0.00758
0.00526
CH,
3.127
0.6284
0.6283
0.6284
0.655
0.0619
0.5930
0.0266
0.5665
3.589
2.493
c,hr
0
0
0
0
0
0
0
0
0
0.1061
0.0737
N,
0.8535
0.9407
0.9407
0.9407
0.9806
0.0927
0.8878
0.0398
0.848
0.08717
0.0606
SUM
13.88
48.98
48.98
48.98
42.02
1.20
11.50
0.515
11.55
3.79
2.633
Deg. C
800
407
167
227
40
40
40
40
900
150
150
Atm
30.5
22.9
21.9
22.3
20.8
1.5
20.8
20.8
31.0
30.3
1.5
-------
(7]{ Desulfurization }¦ ¦ -fTs] »
HRSG-1
To reformer block
s—
SH-1
HE-1
~'
Gasrfier
Comp-1
Prom PSA block
HEAT DUTIES FOR GASIFIER BLOCK, HYDROGEN PRODUCTION
I.D.
Description
Duty, cal/sec
Surface area, m2
SH-1
Steam supemeater
2997
0.0436
HRSG-1
Heat recovery steam generator
14910
0.2909
Figure 9-2. Details of gasifier block for hydrogen production.
9-4
-------
TABLE 9-2. STREAM DATA FOR GASIFICATION BLOCK, HYDROGEN PRODUCTION (FIGURE 9-2)
Stream
I.D.
Flow,
Temp.,
Pressure,
No.
kg-mol
°C
atm
1
Dried biomass, 10 wt% moisture
1111 (kg)
25
1
2
Heated recycle from PSA block
11.55
900
31.04
3
Unreacted carbon
0.549
800
30.54
4
Gasifier effluent
13.88
800
30.54
5
Steam superheater SH-G1 hot-side effluent
13.88
725
30.52
6
Heat recovery steam generator HRSG-1 effluent
13.88
317
30.40
7
Feed to ZnO desulfurization unit
17.67
268
30.40
8
Preheated natural gas from HE-5, PSA block
6.423
130
28.5
9
Natural gas feed to process
3.79
150
28.5
10
Natural gas to reformer as fuel
2.633
150
28.5
11
Boiler feed water from shift reactor block
5.19
148
32.5
12
Saturated steam
5.19
310
32.25
13
Superheated steam for gasifier
0.57
527
31.9
14
Superheated steam to reformer
4.62
527
31.9
15
Process gas to reformer
17.67
268
29.40
16
Recycle stream from PSA block
10.98
40
20.80
17
Recycle compressor effluent
10.98
85
31.9
18
Reformer effluent
48.98
950
23.50
9-5
-------
The process gas leaving the desulfurizer will have an equilibrium HZS concentration of 0.02 ppm.
Added to that stream prior to the first stage of preheat to 527° C in the reformer is 14.08 mols of superheated
steam. The 4.62 mols of steam produced in HRSG-1 is then added, followed by final preheat to 697° C before
entering the catalytic reforming reactor. Details of the reforming block are shown in Figure 9-3 and Tables 9-3
and 9 -4. A skin loss of 2.3% of heat input is assumed, as in the methanol reformer. The heat duties of the
reforming reactor, the process gas preheaters, the fuel preheaters, and air preheater require 2.633 mols of
natural gas fuel in addition to the 1.2 mols of process gas purged from the PSA section. The natural gas,
received at a pressure of 28.5 atm, is expanded through a turbine-generator to reduce it to the 1.5 atm required
for use as furnace fuel. This power is added to that generated by the steam recovered within the process, but
is not included in the thermal efficiency calculation and C02 emission calculations. Process gas leaving the
reformer at 950° C enters the ceramic heat exchanger where it heats the recycle stream entering the gasifier
to 900° C, thus providing the total enthalpy required for energy balance of the gasifier (assuming 87% carbon
conversion).
The primary changes in process configuration required for hydrogen production occur in the next two
blocks, the first of which (shift reactor block, Figure 9-4, and Tables 9-5 and 9-6) involves two catalytic shift
reactors in which CO is reacted with steam to produce hydrogen according to:
CO + H20 = C02 + H2
The first shift reactor operates at 475° C and 22.4 atm. The gas entering the reactor is cooled to 407° C by heat
exchanger HE-2 and steam superheaters SH-2 and -3 which provide the 14.08 mols of steam needed for the
reformer and extract electric power for operation of the compressors. Due to the heat of reaction, there is no
heat duty for either of the shift reactors. Heat exchangers HE-2 and -3 generate high-pressure saturated
steam for superheater SH-3. Heat exchanger HE-4 produces 5.6 mols of low pressure steam for the biomass
drier, cooling the process stream to 167° C before entering the second shift reactor operating at 21.3 atm and
227° C.
Adsorption efficiency in the PSA section increases with pressure and decreases with temperature.
Therefore, the last step in the shift reactor block is heat exchanger HE-5 which cools the process stream to
40° C and condenses out most of the moisture which is separated in tank T-1. HE-5 also acts as preheater
for the boiler feed water (BFW) required by HE-2, HE-3, and HRSG-1. As was discussed for the methanol
converter, a trade-off exists between the size and area of this condenser and the temperature to which the
BFW can be heated: the closer the BFW temperature approaches the gas temperature entering the condenser,
the greater the heat-exchange area required. A maximum limit of 12 m2 is assumed in this case and, given an
entering temperature on the hot side of 227° C, an exit temperature of 40° C, and a water temperature entering
the cold side at 25° C, the heat duty is 42600 cal/sec and the BFW temperature is 148° C, requiring a heat
exchange area of 11.39 m2. After extracting BFW for HRSG-1 and -2, and heat exchangers HE-2 and -3, 43.9
mols of water remains. This hot water enters heat exchanger HE-6 to preheat the natural gas for the process
and for the reformer to 130° C. Water is discharged from this final stage of heat recovery at 121° C.
Turbine-generator TG-3 extracts 17.87 kW of electric power from the high pressure steam superheated
in SH-3 while reducing it to the pressure required for the reformer. Another 10.51 kW is recovered in
condensing turbineTG-4 from the remaining steam from SH-3.
9-6
-------
125" C
NGPH \
Natural gas
Purge
128° C
Air
25j
24,
TG-1
345° C
FPH-1
TG-2
FPH-2
PGPH-1
From gasiffer block
556° C
32
PGPH-2
Reformer
HE-1
SH-2
^ To sNft reactor biock
Figure 9-3. Reformer block for hydrogen production.
9-7
-------
TABLE 9-3. STREAM DATA FOR REFORMER BLOCK FOR HYDROGEN PRODUCTION (FIGURE 9-3)
Stream
No.
I.D.
Flow,
kg-mol
Temp.,
°C
Pressure,
atm
10
Preheated natural gas from HE-6, shift reactor block
2.633
130
28.5
14
Superheated steam from HRSG-1, gasifer block
4.62
527
29.40
15
Process stream from gasifier block
17.67
268
29.40
18
Reformer effluent
48.98
950
23.50
19
Heat exchanger HE-1 hot-side effluent
48.98
786
23.33
20
Steam reheater SH-2 hot-side effluent
48.98
736
23.29
24
Air inlet
31.5
25
1.0
25
Air to preheater
31.5
69
1.5
26
Air to furnace
31.5
325
1.0
27
Flue gas
35.85
125
1.0
28
Purge from PSA block
1.20
40
1.5
29
Preheated purge to reformer fuel
1.20
117
1.5
30
Tubine-generatorTG-1 effluent
2.633
39
6.0
31
Preheated feed to turbine-generator TG-2
2.633
172
6.0
32
Preheated fuel to furnace
3.833
377
1.0
33
Superheated steam from SH-2
14.08
527
29.40
34
Process gas feed to first peheater PPH-1
31.75
377
29.40
35
Effluent from preheater PPH-1
31.75
527
27.83
36
Process gas feed to reformer from second preheater
36.37
697
26.70
9-8
-------
TABLE 9-4. HEAT DUTIES FOR REFORMER BLOCK, HYDROGEN PRODUCTION (FIGURE 9-3)
I.D.
Description
Duty, cal/sec
Surface area, m2
HE-1
Ceramic heat exchanqer
19320
0.3965
SH-2
IP steam reheater
5714
0.0915
NGPH
Natural qas preheater
181
0.026
APH
Air preheater
16450
2.238
FPH-1
Furnace fuel pas preheater
943
0.019
FPH-2
Furnace fuel gas reheater
3010
0.165
PGPH-1
First process qas preheater
13090
3.322
PGPH-2
Second process qas preheater
18020
1.911
HRSG-3
Drier steam generator
15337
0.671
SPR
Steam pyrolysis reactor (reformer)
119560
Skin loss
4100
9-9
-------
From reformer block
SH-2
TG-3
SH-3
HE-2
43
Shift
Natural gas
HE-6
231
H6-3
42
To PSA block
HE-4
€1—~
©—
Shift
T-1
Figure 9-4. Shift reactor block for hydrogen production.
9-10
-------
TABLE 9-5. STREAM DATA FOR SHIFT REACTOR BLOCK FOR HYDROGEN PRODUCTION
(FIGURE 9-4)
Stream
No.
I.D.
Flow,
kg-mol
Temp.,
°C
Pressure,
atm
11
Heated boiler feed water for HRSG-1, gasifier block
5.19
148
32.5
19
Heat exchanger HE-1 hot-side effluent
48.98
786
23.33
20
Steam reheater SH-2 hot-side effluent
48.98
736
23.29
21
Superheater SH-3 hot-side effluent
48.98
614
23.20
22
Heat exchanger HE-2 hot-side effluent
48.98
407
22.94
23
First shift reator effluent
48.98
475
22.44
33
Superheated steam to reformer
14.08
527
29.4
37
Heat exchanger HE-3 hot-side effluent
48.98
311
22.16
38
Drier steam generator, HE-4 hot-side effluent
48.98
168
21.84
39
Second shift reactor effluent
48.98
227
21.34
40
Cold water condenser feed
69.2
25
3.0
41
Condenser cold-side effluent
69.2
148
2.0
42
High pressue boiler feed water
16.17
148
97.3
43
Water to pond
47.84
121
1.0
45
Steam from HE-3 cold side
7.12
310
97.3
46
Steam from HE-2 cold side
9.05
310
97.3
47
Superheated steam
16.17
527
97.3
48
Steam to turbine-generator TG-4
2.09
527
97.3
49
Steam condensate
2.09
30
0.05
50
Steam turbine-generator TG-3 effluent
14.08
375
30.0
51
Natural gas
6.423
25
28.5
52
Preheated natural gas to process and reformer
6.423
130
28.5
53
Condensate from drier, biomass block
5.6
110
1.5
54
Steam to drier, biomass block
5.6
141
1.5
55
Condenser effluent
48.98
40
20.8
56
Water condensate
7.48
40
20.8
57
Feed to PSA block
41.50
40
20.8
9-11
-------
TABLE 9-6. HEAT DUTIES FOR SHIFT REACTOR BLOCK, HYDROGEN PRODUCTION
I.D.
Description
Duty, cal/sec
Surface area, m2
SH-2
Steam reheater for reformer feed
5710
0.0915
SH-3
Steam superheater
13800
0.2669
HE-2
Steam generator
22725
0.4960
HE-3
Steam generator
17880
0.8388
HE-4
Steam generator for biomass drier
15537
0.671
HE-5
Condenser and BFW preheater
42580
11.39
HE-6
Natural gas preheater
1800
0.193
Process gas enters the pressure swing adsorber block (Figure 9-5 and Table 9-7) at a pressure of 20.8
atm and 40° C where C02 and H20 are removed by adsorption on activated carbon. Two adsorber trains are
shown; when the first adsorber is loaded, the process stream is switched to the second vessel and the
pressure is released to expel the adsorbed gases from the first vessel. The next PSA stage adsorbs all of the
remaining CO, CH4, and N2, and 33% of the hydrogen on zeolite. The unadsorbed hydrogen {21.74 mols) is
extracted at a pressure of 20.1 atm. The pressure on the adsorber is then reduced to 1.5 atm to expel the
adsorbate. Part of these desorbed gases are extracted as purge (which controls accumulation of inerts) and
the remainder is recompressed to 20.8 atm. Compression occurs in three stages (Comp-2) with interstage
cooling to 40° C; 95 % of the compressed gas is returned to the gasifier after further compression in Comp-1
to the 31.2 atm pressure required for entry to the gasfication block. A nominal recycle of 0.515 mol is returned
to the PSA unit for control purposes. Product hydrogen is assumed to be delivered at 75 atm pressure as was
the case for the Katofsky (1993) simulation of hydrogen production from biomass by the BCL route.
The hydrogen adsorption efficiency of 33% in PSA-2 is much lower than the 86% specified by
Katofsky. The difference is partly due to the presence of nitrogen which reduces adsorption efficiency, but is
mainly due to the amount of hydrogen needed for return to the gasifier in order to provide the enthalpy required
by the gasifier energy balance. If a 34% adsorption efficiency is assumed, the recycle stream to the PSA is
eliminated, and at 35% efficiency there is insufficient stream flow to return to the gasification block. The
hydrogen adsorption efficiency in PSA-2 is therefore an important variable that must be controlled close to 33%
for proper operation of the system.
Compressors Comp-2, Comp-3, and Comp-4 are each three-stage compressors with interstage
cooling to 40°C. The three water-cooled heat exchangers required for each of those compressors are rated
as follows:
Duty, cal/sec Area, m2
Comp-2 2388 0.320
Comp-3 2100 0.550
Comp-4 2449 0.642
9-12
-------
To gasrfier block
HE-7
Purge «
Comp-3
Comp-4
PSA
PSA
From shift block
57
HEAT DUTIES FOR PSA BLOCK, HYDROGEN PRODUCTION
I.D.
Description
Dutv. cal/sec
Surface area, m*
HE-6
Heat exctianoer, compressor effluent
2440
0.321 1
1
Figure 9-5. Pressure swing adsorber (PSA) block for hydrogen production.
9-13
-------
TABLE 9-7. STREAM DATA FOR PRESSURE SWING ADSORBER BLOCK, FIGURE 9-5
Stream
No.
I.D.
Flow,
kq-mol
Temp.,
°C
Pressure,
atm
16
Recycle stream to gasifier block
10.98
40
20.8
57
Feed from shift reactor block
41.50
40
20.8
58
Feed to first PSA unit
42.02
40
20.8
59
Adsorbate purge from first PSA unit
7.576
40
20.8
60
Feed to second PSA unit
34.44
40
20.8
61
Hydrogen product
21.74
40
20.1
62
Compressed hydrogen
21.74
25
75.0
63
Recycle teed
12.70
40
1.5
64
Purge to reformer fuel
1.20
40
1.5
65
Recycle compressor Comp-2 effluent
11.50
148
20.8
66
Recycle to PSA system
0.515
40
20.8
67
Compressed hydrogen product (5000 psi)
21.74
40
340
The biomass block for hydrogen production is the same as that for methanol, Figure 5-7. In this case,
however, the char combustor does not provide partial preheat of the combustion air for the reformer; instead,
all of the combustion energy is utilized for production of 19.8 kW of electric power for operation of the
compressors and displace as much imported electric power as possible. Table 9-8 gives the heat duties
required for the char combustor and the areas of the heat exchangers needed for the biomass block.
TABLE 9-8. HEAT DUTIES FOR BIOMASS BLOCK, HYDROGEN PRODUCTION
,0.
Description
Duty, cal/sec
Surface area, m2
CAP
Air preheater for char combustor
1699
0.0978
CSG
Char combustor steam generator
10910
0.0860
CSH
Steam superheater
2790
0.0092
HRSG-3
Drier steam generator (in reformer block)
15337
0.671
COND-4
Drier overhead steam condenser
12000
1.918
9-14
-------
Evaluation of Process Performance
The overall material balance shows 21.74 mols of hydrogen can be produced from 100 kg of dry
woody biomass and 3.79 mols of natural gas fed to the process. An additional 2.633 mols of natural gas is
required as reformer fuel.
It is assumed that the product hydrogen will be utilized as vehicle fuel, compressed to a final pressure
of 340 atm {5000 psi) for storage on-board the vehicle. This pressure is currently considered to be within the
capability of existing carbon filament cylinders of sufficient capacity to fit in conventional passenger cars and
provide a driving range when used in FCVs that is equivalent to a tank of gasoline used in ICEVs (Thomas
et al., 1996). Tables 9-9 and 9-10 show the electric power requirements and power available. The energy
required to compress the hydrogen product to that pressure accounts for most of the energy consumption.
As the data show, the total power requirement exceeds the power that can be generated within the process;
therefore, electric power must be imported: 58.1 kWh per 21.74 mols of hydrogen delivered. This imported
power must be deducted when calculating thermal efficiency, and its equivalent C02 emission must also be
accounted for. Assuming that the imported electric power is generated from natural gas with an energy
conversion efficiency of 35%, the C02 emission from that fuel will be:
C02 emitted by production of imported power =
krnl 1 mols CH, 1
58.1 kWh x 860.5 — ! - — = 0.670 mol CH,
kWh 212800 kcal 0.35 4
Total COz emission = (3.79 + 2.633 + 0.67) x 44 = 312 kg
Section 6 discussed the displacement of gasoline by methanol which has an energy efficiency in FCVs that
is 2.5 times that of gasoline ICEVs. The energy efficiency of hydrogen in FCVs is expected to be 3 times
greater than gasoline ICEVs. The gasoline displacement potential of hydrogen is therefore:
Gasoline displaced - 21.74 mols H2 x 57798 ^ ^ ^so^ne
1 Btu
""'7a mol H2 115000 Btu *
x 3 = 130 gallons
0.252 kcal
Then, by the procedure shown in Section 4 for evaluating the total equivalent fuel-cycle COz emission avoided:
Total C02 emission from gasoline production and use,
including oil extraction, oil transport, refining to gasoline,
and transport to refueling stations =130x11.65= 1515 kg
C02 emissions from biomass production including cultivation,
harvest, transport, and fertilizer production = 15.5 kg
C02 emission from natural gas extraction, cleanup, compression,
and transmission, including fuel for imported power = 66.6 kg
COa emission from distribution of product hydrogen to user
(like methanol, assumed to be proportional to energy delivered) = 29.3 kg
9-15
-------
TABLE 9-9. ELECTRIC POWER CONSUMPTION FOR
HYDROGEN PRODUCTION BY HYNOL
I.D.
Unit Description
Power required, kWh
COMP-2
PSA effluent recycle compressor
29.98
COMP-2
Product H, compressor (20.1 to 75 atm)
26.79
COMP-4
High pressure H, compressor (75 to 340 atm)
32.88
COMP-1
Gasifier loop compressor
4.10
B1
Reformer furnace air FD fan
11.44
P3
HP boiler feed water pumps
1.24
P5
IP boiler feed water pump
0.29
P1
LP boiler feed water pumps
0.18
LH
Lock hopper
1.7
B3
Char combustor air FD fan
1.1
B2
Drier steam recycle fan
1.2
Total power requirement
110.9
TABLE 9-10. ELECTRIC POWER AVAILABLE FOR HYDROGEN PRODUCTION
I.D.
Source
Power, kWh
TG-5
Turbine-generator, char combustor (biomass block)
19.80
TG-3
Turbine-generator, shift reactor block
17.87
TG-4
Condensing turbine-generator, shift reactor block
10.51
TG-1,2
Natural gas and purge expanders
4.67
Total power available
52.85
9-16
-------
COj emission from all natural gas used =
312 kg
and the total fuel-cycle C02 emission reduction is:
1515 -15.5 - 66.6 - 29.3 - 312 = 1092 kg
This reduction compares to 731 kg for the base case methanol system and 724 kg for the optimized
(for cost) methanol system. The hydrogen case is therefore capable of greater gasoline displacement and
more C02 emission reduction than can be achieved by production of methanol from the same feedstocks.
The thermal efficiency of hydrogen production {HHV basis) is:
Thermal efficiency -
21.74x63317 + [(52.85-110.92-4.67) x (860.5/0.3958)] - 68 1 %
212800[(3.79+2.633) 0.947] + 372800[(3.79 +2.633) 0.028] + [100x4604]
which compares to 68.4 % for methanol production.
9-17
-------
SECTION 10
COST ESTIMATES
This section develops preliminary cost estimates for the production of methanol by the Hynol process
configured as the base case discussed in Section 5 and as the 90-atm optimized case discussed in Section
8. The cost of hydrogen production by the system discussed in the preceding section is then addressed.
The plant size assumed for these estimates is 7870 tonnes/day of biomass which is the optimum size
determined by Marrison and Larson (1995) from a balance between the economy of scale afforded by large
plant size and tie delivered cost of biomass which increases with plant size (due to the cost of transport from
increasingly remote supply regions). The corresponding delivered biomass cost for that optimum plant size,
assuming woody biomass produced in the North Central U.S., was determined by Marrison and Larson to be
$61 /tonne, and that value is used for all costing reported here. The cost of natural gas is taken to be $2.50/106
Btu (S2.37/GJ).
The gasilier cost is based on the FIuor/EPRI estimate for the Texaco gasifier (Buckingham et al.,
1981), re-sized on the basis of relative gas throughputs and indexed to 1994 dollars. Other equipment costs
are based on data published by Princeton University (Williams, et al., 1994) which evaluated five biomass
gasification processes for methanol or hydrogen production in plants of 1650 tonne/day size. The cost of
components used for those assessments (reformers, compressors, shift reactors, etc.) were scaled up from
1650 to 7870 tonnes/day using a scaling exponent of 0.70 for all plant facilities except the methanol converter,
which uses an exponent of 0.66.
Projected production costs are based on a discounted cash flow rate-of-retum of 10 % after taxes and
after adjusting for inflation. Other economic factors assumed are:
15-year depreciation period with constant annual expenses
Capital recovery factor of 15.45%
13% after-tax rate of return, including 2.7% inflation rate
26% corporate income tax rate
Table 10-1 shows the estimated cost of methanol production for the base-case Hynol configuration.
The total capital investment is taken as 125% of the total cost of the itemized equipment that comprise the plant
facilities. The operating cost and production cost assume 329 days of total operation per year. The indicated
production cost of methanol for these conditions is $0.464/gal or $6.75/GJ. In accordance with the data
reported by Princeton (Katofsky, 1993),the higher heating value of methanol is the basis for the expressing
the energy content
Table 10-2 shows costs for the optimized Hynol system. As discussed in Section 9, this system utilizes
higher pressure in the methanol converter (90 vs 32.5 atm for the base case). As indicated in Table 10-1, the
methanol synthesis unit represents the largest single cost component of the Hynol process and, as discussed
in Section 8, the methanol converter throughput and size are reduced by increased pressure. This cost
reduction of the methanol converter is the main reason for the lower production cost relative to the base case.
As indicated by Table 10-2, the production cost of methanol by the optimized system is $0.418/ga! or $6.09/GJ.
This is 9.9 % lower than the base case.
10-1
-------
TABLE 10-1. COST ESTIMATE FOR METHANOL PRODUCTION
BY THE HYNOL PROCESS: BASE CASE
Basis: 7870 tonne/day (biomass) plant, 4.129 x 105 GJ/day methanol product,
biomass delivered @ $61/tonne, natural gas @ $2.50/10® Btu ($2.37/GJ)
Component
Cost, Smillion
Plant facilities investment
Feed preparation
39.3
Gasifier
300.7
Ceramic heat exchanger
100.2
Gasifier loop compressor
115.0
Reformer
174.2
Methanol synthesis
521.7
Utilities and auxiliaries
81.4
Total plant facilities investment (PFI)
1333
Total capital investment @ 125% of PFI
1666
Operating cost per day
Biomass
0.4801
Natural gas
0.9930
Operation and maintenance
0.2434
Catalysts
0.00304
Purchased energy
0
Capital charge
0.3042
Total operating cost
2.024
Methanol production cost = $0.464/gal ($6.75/GJ)
10-2
-------
TABLE 10-2. COST ESTIMATE FOR METHANOL PRODUCTION BY THE
HYNOL PROCESS: 90-atm OPTIMIZED CASE
Basis: 7870 tonne/day (biomass) plant, 4.078 x 105 GJ/day (HHV) methanol product,
biomass delivered @ $61/tonne, natural gas @ $2.50/10fl Btu ($2.37/GJ)
Component
Cost, $million
Plant facilities investment
Feed preparation
39.3
Gasifier
292.5
Ceramic heat exchanger
97.5
Gasifier loop compressor
114.5
Reformer
173.2
Methanol synthesis
242.5
Utilities and auxiliaries
81.4
Total plant facilities investment (PFI)
1041
Total capital investment @ 125% of PFI
1301
Operating cost per day
Biomass
0.4801
Natural gas
0.9752
Operation and maintenance
0.1901
Catalysts
0.00102
Purchased energy
0
Capital charge
0.2376
Total operating cost
1.884
Methanol production cost = $0.418/gal ($6.09/GJ)
10-3
-------
Table 10-3 is a cost estimate for hydrogen production. In this case, as discussed in the previous
section, the components of the gasifier loop are the same as for methanol production, but the methanol
synthesis bop is replaced by two shift reactors and pressure swing adsorbers. Additional compressor capacity
is necessary to deliver the product hydrogen at the 5000 psi pressure required for on-board use as vehicle fuel.
The estimated production cost of hydrogen, $5.98/GJ, is slightly lower than the cost of methanol production.
As shown in Section 9, the C02 emission reduction achievable by hydrogen (assuming use in fuel cell vehicles)
is significantly greater than for methanol: 1092 vs 724 kg total fuel-cycle C02 emission reduction for the same
amount of process feedstocks.
CO- Mitigation Cost
The relative costs of C02 emission mitigation will now be examined on the basis of the methanol
production costs shown above and will then be compared with the published cost estimates of COs mitigation
from stationary sources. Use of methanol as transportation fuel would occur in three stages. First, as is
already the case, it would be used in flexible fuel vehicles as M-85 or M-100. Next, it would be used in
dedicated vehicles designed specifically for methanol. These uses represent the transition period between
conventional vehicles and fuel cell vehicles and would establish the refueling network necessary for large
displacement of ICE vehicles and their petroleum fuels In the third stage. The cost of C02 emission reduction
in each case can be estimated on the basis of the relative production costs of methanol and gasoline and the
efficiencies of each in the respective application.
The production cost of gasoline is approximately $0.60/gal., including crude production, transport, and
refining. The cost of methanol production by the optimized Hynol process is expected to be $0.42/gal. Based
on the relative energy content and efficiency of IC engines, 1 gallon of methanol will displace 0.519 gallon of
gasoline (relative efficiency = 1.05) and it will therefore cost $0.80 to displace 1 gallon of gasoline costing
$0.60. Using the procedure discussed in Section 6, the total fuel-cycle equivalent C02 emissions from the
production, refining, and distribution of gasoline is 11.65 kg per gallon of gasoline used, and the C02 emissions
from the production, distribution, and use of methanol by the Hynol process are 4.91 kg per gallon of methanol.
In flexible fuel vehicles, an overall C02 emission reduction of 2.19 kg would result from the displacement of 1
gallon of gasoline by 1.93 gallons of methanol at a cost differential of $0.20. The mitigation cost per tonne of
C02 emission avoided is therefore $91.
Used in dedicated methanol vehicles with IC engines, 1 gallon of methanol is expected to displace
0.627 gallon of gasoline (relative efficiency = 1.27). In that case, the fuel-cycle C02 emission reduction will be
3.82 kg per gallon of gasoline displaced, and the net mitigation cost will be $52.3 per tonne of C02 emission
avoided.
When used in fuel cell vehicles, 1 gallon of methanol will displace 1.23 gallons of gasoline that would
have been used in conventional vehicles with a net reduction of 9.42 kg of fuel-cycle C02 emissions. In this
case, the value of gasoline displaced by 1 gallon of methanol is $0,738, and the cost of C02 mitigation is -$33.7
per tonne. Thus, mitigation of C02 emissions relative to current conditions could be achieved at no cost. If
emission trading of greenhouse gas credits were to become a future reality (as some predict), the negative cost
of reduction in the transportation sector might be traded against the cost of stationary source emissions.
Stationary Source Mitigation
As indicated in Section 1, a considerable amount of international R&D has been underway to assess
the feasibility and cost of C02 mitigation from coal-burning power plants (Blok et al., 1992; Herzog, 1997).
Those technologies involve the recovery of C02 from the combustion gases, usually by amine scrubbing, and
sequestration in the ocean or in underground repositories. Feasibility has been established for a number of
technologies, and preliminary costs have been estimated. Application to a U.S. power plant has been
examined in detail (Doctor et al.t 1996) by Argonne National Laboratory for a case in which C02 would be
10-4
-------
TABLE 10-3. COST ESTIMATE FOR HYDROGEN PRODUCTION BY THE HYNOL PROCESS
Basis: 7870 tonne/day (biomass) plant, 4.532 x 10s GJ/day (HHV) methanol product,
biomass delivered @ $61/tonne, natural gas @ $2.50/10® Btu ($2.37/GJ)
Component
Cost, Smillion
Plant facilities investment
Feed preparation
39.3
Gasifier
272.9
Ceramic heat exchanger
84.9
Gasifier loop compressor
47.8
Reformer
181.6
Shift reactors
36.9
Pressure swing adsorbers
118.9
PSA recycle compressor
49.9
Hydrogen product compressor (340 atm)
91.0
Utilities and auxiliaries
63.5
Total plant facilities investment (PFl)
986.7
Total capital investment @ 125% of PFl
1233
Operating cost per day
Biomass
0.4801
Natural gas
0.9981
Operation and maintenance
0.1802
Catalysts
0.0325
Purchased energy
0.2285
Capital charge
0.2253
Total operating cost
2.145
Hydrogen production cost = $5.98/GJ
10-5
-------
recovered from an oxygen-blown, coal gasification combined-cycle power plant having a net energy output of
411 MW for the complete energy cycle. The base-case plant without C02 recovery has a thermal efficiency
(HHV) of 36.6%, a total fuel-cycle C02 emission of 329.4 tonnes/hr, and a power generating cost of
$0.0586/kWh.
Addition of C02 recovery reduces the thermal efficiency to 33.6% and increases the power generating
cost to $0.0946/kWh while reducing emissions to 113.8 tonnes COVhr (after accounting for emissions resulting
from generation of the makeup power required to maintain 411 MW output). The total cost of C02 recovery,
purchased makeup power, transporting 254.8 tonnes/hr of recovered C02 by a 500-mile pipeline, and its
sequestration in depleted oil or gas wells, was estimated at $61 per tonne of C02 sequestered. The net
reduction of C02 emissions compared to the base case is 215.6 tonnes/hr and the net cost per tonne of COa
emission avoided is $72.7/tonne.
Clearly, the mitigation of C02 emissions from stationary sources such as power plants would cost
substantially more than mitigation from mobile sources if the latter were to use methanol produced from
biomass and natural gas instead of gasoline. In addition to the real cost reduction represented by these
comparisons, the displacement of imported oil has external cost reductions and environmental benefits that
would not accrue from C02 mitigation from power plants.
Costs Compared to Gasoline
Assuming a natural gas price of $2.37/GJ ($2.50/106Btu), the production cost of methanol is estimated
here to be $6.09/GJ (HHV), and hydrogen is estimated to cost $5.98/GJ (HHV). The production cost of
gasoline is about $0.60/gal or $4.51/GJ (HHV). If the fuel economy of gasoline used in current vehicles with
internal combustion engines (ICEVs) is 27 miles per gallon, it takes 1.33 x 10® joules to travel 27 miles, and
the cost per mile is 2.2 cents.
Figure 10-1 compares the costs of methanol and hydrogen with gasoline as a function of the price of
natural gas used in the Hynol process. The cost basis for this comparison is the fuel production cost per
vehicle-mile traveled which avoids assumptions about relative fuel tax rates, distribution costs, and markup
that cannot be forecast. The fuel economy of methanol used in ICEVs as M85 in flexible fuel vehicles or M100
in dedicated methanol vehicles during the period of transition to fuel-cell fuel vehicles is 1.05 and 1.27,
respectively, relative to the energy required to travel a given distance by gasoline ICEV. The comparison is
made for gasoline at the current production cost of $4.5/GJ for oil priced at $20/bbl and for DOE's projected
cost for the year 2010 at an oil price of $36/bbl. The comparisons show that dedicated methanol vehicles
would be in a competitive cost range with gasoline at current prices and FFVs would be competitive at future
gasoline prices.
When fuel-cell vehicles are introduced commercially, probably within the next decade, the cost per
vehicle mile will be about 45% less than current gasoline cost and will remain less than gasoline for the
foreseeable future even if the price of natural gas should escalate faster than the cost of gasoline. The relative
fuel efficiency of methanol FCVs assumed in Figure 10-1 is 2.5. With hydrogen FCVs, having a relative fuel
efficiency of 3.0, the cost per vehicle mile will be about 55% less than the current cost of gasoline, assuming
the current price of natural gas.
The initial commercialization of FCVs in the U.S. will probably occur with vehicles that utilize gasoline
reformed on board to hydrogen via partial oxidation (POX) reformers. Chrysler is expected to demonstrate a
POX/FCV in 1999. The efficiency of gasoline-FCV systems is reported (Espino and Robbins, 1997) to be 75-
83 % for conversion of gasoline to hydrogen by the fuel processor and 45-50 % for the fuel cell, or 38.8-41.5
% efficient for conversion of gasoline (LHV) to electric energy. A hydrogen FCV, according to the same
source, will convert 55-60 % of hydrogen's LHV to electric energy (and requires no fuel processor). Taking
the average of these ranges as a basis for comparison, 57 % of hydrogen's energy is converted to electric
energy while the POX converts 38 % of the gasoline energy.
10-6
-------
Gasoline in conventional vehicles
(Year 2010)
Gasoline in conventional vehicles
(Year 1996)
Natural Gas Price, $/GJ
Figure 10-1. Comparison of fuel production costs with gasoline. Basiis: foL hllT
per gallon used in an internal combustion engine vehicle (ICEV). - _ uvdrogen
ICEV, DMV = M100 in dedicated ICEV, FCV = methanol in fuel-cell vehicle, H2 FCV y
In fuel-cell vehicle.
10-7
-------
Directed Technologies (1997) gives the fuel economy of a hydrogen FCV as 275 miles from 3.58 kg
of compressed hydrogen. The energy content of hydrogen (LHV) is 57800 kcal/mol, and the energy
consumption is therefore:
mol //, ]rral VJU 1 Wf,
3.58 kg H, ^ 57800 1.163— x 0.57 - 249 —
" 2 kg H2 mol H2 kcal 275 miles mile
The LHV of gasoline is 115000 Btu/gal. Assuming that the POX vehicle will have the same efficiencies
(auxiliary batter, electric motor, and transmission), and operates over the same driving cycle as the above
hydrogen vehicle, the gasoline consumption of the POX vehicle will be:
X gal x 115000— (0.38) 0.2928— - 249 Wh
gal Btu
or X = 0.0195 gal per mile, or 5.36 gal per 275 miles or 51.3 miles per gallon.
Because the POX vehicle requires a fuel reforming reactor, a sulfur removal reactor, a shift reactor, and a CO
removal reactor that are not required by the hydrogen vehicle, a POX FCV will be heavier, which will reduce
its fuel economy. Based on other data, these components are expected to reduce efficiency by about 10 %;
the fuel economy of the POX FCV will therefore be:
46.2 miles/gallon of gasoline
and the cost per vehicle mile will be about 60/46.2 = 1.3 cents per mile. Comparing this with the methanol
curve of Figure 10-1, the fuel cost is seen to be nearly equal for FCVs fueled by either gasoline or methanol
based on current prices of natural gas and gasoline.
Similar calculations show that the total fuel-cycle C02 emission-reduction potential per vehicle mile will
be 39 % greater for a methanol FCV than for a gasoline POX vehicle.
10-8
-------
SECTION 11
WASTE METHANE AS CO-FEEDSTOCK
Introduction
Landfill gas (LFG) from municipal landfills and anaerobic digester gas (ADG) from municipal
wastewater treatment plants offer potential sources of methane that could displace part of the natural gas used
by the Hynol process. Such use would reduce the fossil C02 emitted by the Hynol process and assist the
reduction of methane emissions from landfills and wastewater treatment plants. A double effect on reduction
of greenhouse gas emissions would thus be achieved, together with a possible reduction of feedstock cost.
The simulations discussed in this section examine the feasibility of this approach both for production of
methanol and for production of hydrogen in a situation where a Hynol plant could be located near either of
these waste methane sources.
Table 11-1 indicates the amount and quality of landfill gas and disgester gas available in the U.S.
(Thorneloe, 1992). Although the total amount is substantial, the quantity that could be secured from a single
source is not sufficient to completely displace the natural gas needed for a large Hynol plant-the objective of
which is, as was assumed for the previous process simulations, to maximize displacement of conventional
liquid transportation fuel at least cost. A large Hynol plant size is therefore a requirement. In the case of landfill
gas, considerations of the long term availability and constancy of rate require flexibility in adapting the Hynol
process to utilize it as one of the raw materials. These simulations assume a maximum of 13 % displacement
of the natural gas required by Hynol with waste methane as supplementary feedstock. For the optimum Hynol
plant size of 7870 tonnes (biomass) per day, this amounts to about:
-,g?Qtonnes biomass^Q^ kg biomass ^ y^kg mol methane
day tonne biomass 100 kg biomass
lb mol scf CH. „ scf CH,
2.2-——-*359 - i- = 3.91x10s — ±
kg mol lb-mol CH4 day
or 7.76 x 10® scf/day of landfill gas. As in the other simulations, woody biomass and natural gas are
considered as the primary feedstocks.
Several options for utilization of waste methane in the Hynol system are possible: (1) direct
replacement of part of the natural gas fed to the process as feedstock for methanol or hydrogen production,
{2) replacement of part of the natural gas used as reformer fuel, and (3) in the case of hydrogen production,
feeding with the natural gas as process feedstock, or feeding separately to the process through the PSA unit,
or feeding to the reformer as fuel. Each of these options will be examined here using as a base case for
comparison the 90-atm Hynol optimizations discussed in Section 9, with a reformer temperature of 950°C.
11-1
-------
TABLE 11-1. WASTE METHANE CHARACTERISTICS AND POTENTIAL MARKET
WASTE METHANE CHARACTERISTICS
Characteristic
Landfill Gas
Anaerobic Digester Gas
Average
Range
Average
Range
Heating Value
HHV-Btu/lt3
500
350-650
650
600-700
Major Constituents
(%)
CH,
CO,
N,
o,
HjO
50
45
5
<1
Sat.
35-58
40-55
0-15
0-15
90-120°F
65
35
Sat.
90-95°F
Contaminants
Sulfur (ppmV H,S)
Halides (ppmV)
NMOC (ppmV)
21
132
2,700
1-700
N/A
237-1400
100
<0.01
<100
10-2,000
POTENTIAL MARKET
Source
filORAL
U.S.
Methane
(Tg/yr)
Potential
Electric Power
(MW)
Methane
(Tg/yr)
Potential
Electric Power
(MW)
Landfills
35
11,000
9
2.700
Digesters
- Water Treatments
- Animal Waste
35
15
13.000
5,300
5
3
1,800
1,000
Coal Mines
35
13,000
4
1,400
11-2
-------
As for all prior simulations, natural gas is assumed to have the following composition: 94.7% CH4,
2.8% G,^ 2.3% N2, and 0.2% C02. Landfill gas is assumed to contain 50% CH4, 45% C02, 4.8% N2, and 0.2%
02. Anaerobic digester gas from water treatment plants is assumed to contain 65% CH4 and 35% C02.
Because the Hynol catalysts are subject to poisoning and reformer tubes are subject to corrosion by the
impurities in these wastes, particularly sulfur and chlorides, the waste gas must be cleaned thoroughly before
use as co-feedstock. This applies especially to landfill gas which will require cleaning by the multi-step method
developed by IFC under EPA sponsorship for utilization of landfill gas for production of electric power in fuel
cells. Consequently, landfill gas cannot be considered "free," but its use must include the cost of cleanup.
Current estimates of that cost, based on an 800 kW LFG fuel cell (Sandelli, 1992), is about $1.1 per million
Btu for the LFG pretreatment system.
Landfill Gas as Supplemental Process Feed for Methanol Production
Figure 11-1 is a block diagram for the case in which landfill gas, in the amount of 1.62 mols, is used
as process co-feedstock for methanol production. Table 11-2 gives the composition of the principal process
streams for that case. Comparison with the base case, in which only natural gas is used as co-feedstock with
biomass, is made in Table 11-3. The following results are noted: (1) the throughputs of the gasifier and
reformer are increased when LFG is added, (2) the methanol converter throughput is reduced, (3) the methanol
yield is slightly increased, and (4) the natural gas required to fuel the reformer (together with the unconverted
carbon from the gasifier and the purge) is increased. The amount of increase in NG fuel required by the
reformer relative to the base case, due to the higher throughput and consequent increased head duty, is 0.494
mol. The net reduction of the natural gas requirement is therefore 0.81-0.494 = 0.316 mol.
Because C02 is, like H20, an oxidant of carbon at the reformer temperature, one would expect that the
large amount of CO, fed as a component of LFG would reduce the steam/carbon ratio required to avoid carbon
deposition in the reforming catalyst. Another simulation is therefore shown in Table 11-3 for the same case
as above, except that the steam/carbon ratio is reduced from 2.5 to 2.0. Reforming experiments at RTP with
nickel catalysts at Hynol base-case conditions (Dong and Karwowski, 1997) show that a steam/carbon ratio
of 2.0 is probably feasible without the extra C02. As indicated by the results summarized in Table 11-3, the
reformer heat duty and throughput are reduced, and the NG required as reformer fuel is reduced to 2.86 mols
or 0.374 mol greater than that required for the base case reformer. The net reduction of natural gas required
by the Hynol process, due to the use of LFG as a supplemental methane source in this case, is therefore 0.81 -
0.374 = 0.436 mol. In either case, no imported energy is required and, in the second case, 12.35 kW of power
is exportable. Thus, the natural gas component could be increased (and the methanol yield improved) to some
extent for that case without violating the condition for no electric power import be required.
An alternative to the use of LFG as process feed is to use it to displace part of the natural gas required
as fuel for the reformer furnace. Figure 11-2 shows the operating conditions of the reformer for that case. (The
rest of the Hynol system is the same as the base case.) This simulation shows that the natural gas required
for reformer fuel is reduced from 2.486 to 1.688 mols or 0.798 mol, about the same as the methane content
of the LFG fed. The 8.36 kW of power required to compress the LFG for feeding to the process as feedstock
is avoided.
Hydrogen Production Using Waste Gas as Supplemental Feed
For a system configured to produce hydrogen, the same two options discussed above for utilizing LFG
or ADG are applicable with the additional possibility of adding them to the pressure swing adsorber (PSA)
instead of the gasifier effluent. In the latter case, the C02 component of the waste gas would be removed by
the PSA unit before entering the process and less compression of the waste gas would be required (to 20.8
instead of 40.9 atm). The three options for feeding the waste gas for hydrogen production are indicated in the
block diagram, Figure 11-3. The primary stream compositions corresponding to that diagram are given in
Table 11-4 for the case in which LFG is fed to the gasifier effluent (feed location A). Because the CO?
component of the waste gas helps reduce the steam requirement for reforming in this case, a steam/carbon
11-3
-------
Natural gas
2.98 mols
Biomass
51.06 kg C
7.49 kg H
51.53 kg O
0.15 kg N
0.08 kg S
0.792 kg ash
Steam 26.85 mols
CD-
Air
35 mete
Filter
Reformer
S50T
Gasifier
800*C
Combustor
Unreacted carbon
0.549 mot
Natural gas
(fuel) 2.98 mols
Steam 2.5 mols
Heat
exchanger
HjO condensate
23.7 mols
Purge 1.2 mols
(to reformer fuel)
Methanol
converter
250°C
CHaOH 7.338 mols
WjO 1.25 mols
Figure 11-1. Block flow diagram for optimized Hynoi process using landfill gas as co-feedstock.
11-4
-------
TABLE11-2. STREAM DATA FOR OPTIMIZED HVNOL PROCESS WITH 90-atm METHANOL CONVERTER
AND 1.62 mols LANDFILL GAS ADDED TO GASIFIER EFFLUENT AS SUPPLEMENTAL FEEDSTOCK
Stream composition, kfl-mols
Stream
1
Stream
2
Stream
3
Stream
4
Stream
5
Stream
6
Stream
7
Stream
8
Stream
9
a,o
4.988
0.1768
0.0070
0.086
0.0003
2.504
0
0
0
H,
6.201
26.695
16.76
43.46
0.7320
9.913
0
0
0
CO
2.793
7.530
1.990
9.52
0.0870
1.178
0
0
0
o
p
2.420
4.747
4.760
9.51
0.208
2.816
0.0060
0.0060
0.729
ch4
3.018
0.4876
0.753
1.24
0.0329
0.4454
2.822
2.822
0.810
C,Hc
0
0
0
0
0
0
0.0834
0.0834
0
N,
1.812
1.958
3.05
5.01
0.133
1.806
0.0685
0.0685
0.0778
CH,OH
0
0
0.147
0.147
0.0064
0.0870
0
0
0
SUM
21.23
41.60
27.48
68.98
1.20
18.75
2.98
2.98
1.62
Deg.C
800
51
40
51
40
900
56
56
56
Atm
41.16
24.8
82.0
90.6
82.0
41.66
40.79
40.79
40.79
11-5
-------
TABLE 11-3. COMPARISON OF BASE-CASE OPTIMIZED HYNOL SYSTEM AND SYSTEMS UTILIZING
LANDFILL GAS {LFG) AS SUPPLEMENTAL FEEDSTOCK FOR METHANOL PRODUCTION
Basis: 100 kg dry biomass feed, 87% carbon conversion in gasifier, 950°C reformer
Simulation
Base
case
Run
1
Run
2
Run
3
Run
4
NG feed to gasifier effluent, mols
3.79
3.385
2.98
2.98
3.79
NG feed to reformer furnace, mols
2.60
2.082
2.98
2.86
1.687
LFG feed to gasifier effluent, mols
0
0.81
1.62
1.62
0
LFG feed to reformer furnace, mols
0
0
0
0
1.62
Gasifier throughput, mols
14.84
17.72
21.23
22.07
14.84
Steam feed to reformer, mols
19.82
22.70
26.85
21.45
19.82
Steam/carbon ratio
2.50
2.50
2.50
2.00
2.50
Reformer throughput, mols
51.0
57.2
65.3
60.7
51.0
Reformer duty, cal/sec
121020
126100
132070
129140
121020
Methanol converter throughput, mols
84.6
59.65
54.1
48.6
84.6
Methanol loop recycle, mols
63.0
35.9
27.5
22.0
63.0
Methanol product, mols
7.202
7.277
7.338
7.312
7.202
Total electric power consumed, kWh
71.7
73.9
86.6
85.5
71.7
Total electric power produced, kWh
71.6
88.7
89.3
97/8
70.9
MSR duty, cal/sec
36050
39670
41470
42840
36050
11-6
-------
Purge
1.2 mols
26 r
NGPH
Air
31.7 mols
30
TG-4
TG-3
352°C
Natural gas
1.687 mols
FPH-1
27F
FPH-2
From gasifier block
APH-2
790°C
LFG
1.62 mols
Reformer \
Figure 11-2. Reformer operation with landfill gas for partial displacement of natural gas
as fuel. Stream data are given in Table 8-1 a
11-7
-------
Natural gas
Waste gas
addition pont A
Biomass {10% moisture)
51.06 kg C
7.49 kg H
51.52 kg O
0.151 kg N
Filter
HRSG-1
Natural gas (fuel)
Steam
Reformer
Gaslfier
Unreacted
carbon
0.549 mol
Air
Furnace
Heat
exchanger
Waste gas
addition point C
Steam
0-
Shift
Purge 12 mols
(to reformer fuel)
Shift
PSA
BFW
PSA
RO
Waste gas
addition point B
Condensate
Figure 11 -3. Block flow diagram for hydrogen production using waste gas as co-feedstock.
11-8
-------
TABLE 11-4. STREAM DATA FOR HYNOL CONFIGURED FOR HYDROGEN PRODUCTION USING
LFG AS CO-FEEDSTOCK (FIGURE 11-3, FEED LOCATION A)
Stream composition, mols
Stream
1
Stream
2
Stream
3
Stream
4
Stream
5
Stream
6
Stream
7
Stream
8
Stream
9
Stream
10
Stream
11
H,0
2.213
11.65
8.484
5.60
0.163
0
0
0
0.600
0
0
H,
6.284
25.91
29.078
31.96
32.52
0.9334
9.798
0.5573
9.241
0
0
CO
1.170
6.548
3.383
0.4989
0.526
0.0457
0.4805
0.0273
0.4532
0
0
CO,
0.4451
2.076
5.241
8.125
8.125
0
0
0
0
0.0060
0.729
X
o
3.226
0.7517
0.7518
0.7518
0.793
0.0690
0.7240
0.04118
0.6828
2.822
0.810
C,H„
0
0
0
0
0
0
0
0
0
0.0834
0
N,
1.5085
1.655
1.655
1.6557
1.746
0.1518
1.594
0.0907
1.503
0.0685
0.0778
SUM
14.85
48.59
48.59
48.59
43.87
1.20
12.60
0.7165
12.48
2.98
1.62
Deg. C
800
402
155
227
40
40
40
40
900
150
150
Attn
30.5
22.9
21.9
21.3
20.8
1.5
20.8
20.8
31.0
30.3
1.5
-------
ratio of 2.0 is assumed; for addition to the PSA unit, the C02 does not enter the reformer and in that case, a
reforming steam/carbon ratio of 2.5 is assumed.
Table 11-5 summarizes the results of simulations of hydrogen production using LFG or ADG as Hynol
co-feedstock, added either to the gasifier effluent (feed point A), or to the PSA unit (feed point B), or added to
the reformer as fuel (feed point C). Table 11-6 gives the heat exchanger duties and areas corresponding to
the ieed point A" options for the two waste gas co-feedstocks, corresponding to the units identified in Sections
8 and 9.
Table 11-7 summarizes the results of the above simulations on C02 emission reduction (and methane
emission reduction). One may conclude from these results that the use of LFG to displace natural gas in the
Hynol process would be most effective if utilized as fuel for the reformer rather than as direct feed to the
process.
Finally, it should be noted that, if a Hynol plant is located near a landfill for utilization of LFG as co-
feedstock for either methanol or hydrogen production, solid carbonaceous wastes as well as waste gases can
be considered as co-feedstocks to displace part of the high-cost woody biomass. Such carbonaceous wastes
could Include paper, fabrics, rubber, and woodwaste. It is necessary that those solid wastes be free of
extraneous materials that could interfere with the high-pressure feed system used for the Hynol gasifier.
11-10
-------
TABLE 11-5. RESULTS OF SIMULATIONS OF HYDROGEN PRODUCTION USING
WASTE GAS (WG) AS SUPPLEMENTAL FEEDSTOCK FOR NATURAL GAS (NG) DISPLACEMENT
Basis: 100 kg dry biomass, 87% carbon conversion in gasifier, 950°C reformer,
H2 product at 340 atm (5000 psi)
Waste gas addition point
Base case
A
B
C
A
Waste gas type
NA
LFG
LFG
LFG
ADG
NG feed to gasifier effluent, mols
3.79
2.98
2.98
3.79
2.678
NG feed to reformer furnace, mols
2.633
2.68
2.63
1.823
2.60
WG feed to gasifier effluent, mols
0
1.62
0
0
1.62
WG feed to PSA unit, mols
0
0
1.62
0
0
WG feed to reformer furnace, mols
0
0
0
0
0
Gasifier throughput, mols
13.88
14.85
15.68
13.88
14.37
Steam feed to reformer, mols
18.70
16.60
18.28
18.70
16.05
Steam/carbon ratio
2.50
2.00
2.50
2.50
2.00
Reformer throughput, mols
48.98
48.59
49.27
48.98
47.13
Reformer duty, cal/sec
119560
120440
120250
119560
118560
Shift reactor throughput, mols
48.98
48.59
49.27
48.98
47.13
PSA-1 unit throughput, mols
42.0
43.87
44.0
42.0
42.09
Hydrogen product, mols
21.74
21.79
21.67
21.74
21.45
Total electric power consumed, kWh
110.9
120.0
120.0
110.9
115.3
Total electric power produced, kWh
52.9
61.8
58.1
52.1
59.5
Steam fed to gasifier, mols
0.57
0.60
0.55
0.57
0.85
Air fed to reformer furnace, mols
32.5
33.0
33.0
31.5
32.5
Condenser HzO condensate, mols
7.48
5.44
7.14
7.48
5.28
Boiler feed water (BFW), mols
69.2
59.0
67.0
69.2
57.0
PSA-1 C02 extraction, mols
7.42
8.13
8.10
7.42
7.89
PSA-1 HzO extraction, mols
0.156
0.163
0.159
0.156
0.158
11-11
-------
TABLE 11-6. HEAT DUTIES AND HEAT TRANSFER AREAS FOR WASTE GAS
UTILIZATION AS CO-FEEDSTOCKS
Basis: 1.62 mols waste gas, feed location A; 950°C reformer, 2.00 steam/carbon ratio.
Methanol units identified in Figures 8-5,6,7,8; hydrogen units identified in Figures 9-2,3,4,5
Methanol production
Hydrogen production
Unit ID
Duty, cal/sec
Area, m2
Unit ID
Duty, cal/sec
Area, m2
SH-G1
5760
0.0664
SH-1
3158
0.0459
SH-G2
5275
0.0980
HRSG-1
15720
0.3028
SH-G3
3125
0.0941
HE-1
21070
0.440
HRSG-1
15060
0.3975
SH-2
4760
0.0798
HE-1
37940
0.832
NGPH
181
0.026
SH-S1
10520
0.1996
APH
16040
2.128
HRSG-2
45540
1.684
FPH-1
943
0.0199
HE-6
14830
3.07
FPH-2
2880
0.149
NGPH
181
0.026
PGPH-1
17920
1.76
APH-1
14810
1.984
PGPH-2
14150
2.55
APH-2
13720
0.153
SH-3
13820
0.2783
FPH-1
943
0.0174
HE-2
21860
0.495
FPH-2
3410
0.1899
HE-3
18840
0.9392
PGPH-1
19150
2.763
HE-4
15337
0.764
PGPH-2
23600
1.210
HE-5
36230
9.745
HRSG-3
15337
2.104
HE-6
1240
0.105
Reboiler
22850
1.44
HE-7
7850
HE-2
1607
0.0975
CAP
1699
0.0978
HE-3
19770
4.25
CSG
10910
0.0860
HE-4
23560
3.045
CSH
2790
0.097
HE-5
980
0.0655
Cond-4
1200
1.918
HE-6
14830
3.07
HRSG-3
15337
0.766
COND-3
24588
3.72
CAP
0.0834
APH-2
13720
0.153
Cond-4
12000
1.918
11-12
-------
TABLE 11-7. CO, EMISSION REDUCTION POTENTIAL OF WASTE GAS
USED AS HYNOL CO-FEEDSTOCK
Basis: 100 kg dry blomass, 87% carbon conversion in gasifier, 950°C reformer
Waste gas utilization option
LFG
LFG
LFG
ADG
Fuel product
Methanol
Methanol
Hydrogen
Hydrogen
Waste gas addition point
A
C
A
A
Waste gas fed, mols
1.62
1.62
1.62
1.62
Total NG used, mols
6.39
5.477
5.66
5.278
C02 emission from NG, kg
262
241
249
232
COz emission from imported electric power, kg
0
0
29.5
28.2
C02 emission from biomass production, kg
15.5
15.5
15.5
15.5
C02 emission from NG production, kg
55.9
51.4
59.4
55.5
C02 emission from fuel product distribution, kg
26.0
25.5
29.4
28.9
Gasoline displaced, gal
96.2
94.4
130.4
128.3
Full fuel-cycle C02 emission reduction, kg
760
724
1136
1134
C02 equivalent of methane emission avoided, kg
256
256
256
370.6
Total C02 equivalent emission reduction, kg
1016
980
1392
1505
/
11-13
-------
SECTION 12
SEWAGE SLUDGE
Introduction
The previous section analyzed the effects of substituting waste methane for part of the natural gas
used by the Hynol process for production of methanol or hydrogen. This section will discuss the prospects for
displacing woody biomass with sewage sludge as an alternative source of the solid feedstock. The
composition of sewage sludge, as obtained from analyses of samples from the Riverside County, CA, plant,
is taken to be for these simulations as follows: 31.66 wt % C, 4.53% H, 4.02% N, 1.51% S, 17.78% O, and
40.5% ash. The sludge is assumed to be dried completely before use and to have a higher heating value of
3054 kcal/kg (with ash) and a heat of formation of 984 kcal/kg.
These simulations further assume that the anaerobic digester gas produced by a sewage treatment
plant is also used as Hynol feedstock. The ratio of digester gas to sludge produced by a sewage treatment
plant is taken to be 0.0149 kg-mol per kg of dry sludge. The digester gas Is assumed to contain 65% methane,
35% carbon dioxide, and 100 ppm hydrogen sulfide (H2S). Because H2S is the only contaminant of digester
gas, it is fed without preliminary cleanup, together with natural gas, to the ZnO desulfurization unit located
between the Hynol gasifier and reformer. Because natural gas and ADG both contain sulfur compounds in
addition to HgS, the first section of the desulfurization unit contains a catalyst where all sulfur compounds are
converted to H2S by reaction with H2 in the gasifier effluent.
The process configuration is initially assumed to be the same as that for the 90-atm optimized version
of the Hynol system discussed in Section 8. The digester gas is therefore compressed from 1 to 40.85 atm
prior to entering the process at the desulfurization unit. The simulations to be discussed assume that sludge
is fed in one of three amounts: the first simulation assumes that all of the woody biomass feedstock is replaced
with sludge; i.e., sludge is the only solid feedstock. The second and third simulations assume that both sludge
and woody biomass are fed, with intermediate and low displacement, using 60 and 20 kg of sludge,
respectively. In order to maintain a constant feed rate of total carbon in the solid mixture fed to the gasifier, the
amount of woody biomass is adjusted to 69.7 kg for the intermediate case and 97.23 kg for the low sludge
case, giving a total carbon content of 4.255 kg-atoms for each simulation. The amounts of H, O, N, and ash
fed are therefore different from the base-case Hynol simulations previously discussed, with the major effect
being the amount of inert nitrogen which is much higher in sludge than in wood. This affects the amount of
purge required, and the purge rate will therefore be examined as a primary operating variable . Another
important difference is the high ash content of sludge, which prohibits use of the unreacted carbon (which is
discharged as a mixture with the ash) as fuel to displace part of the natural gas used to preheat the reformer
combustion air. All of the unreacted carbon (assumed to be 13 mol% of the carbon fed) is therefore a waste
that must be disposed of with the ash. No credit for COz emission reduction due to disposal (or sequestration)
of that carbon is assumed.
12-1
-------
Case 1: Total Displacement of Woodv Biomass bv Sludge
This simulation is based on 161.28 kg of dry sludge fed to the gasifier with no woody biomass added.
The total carbon content is 4.255 kg-atoms, the hydrogen is 3.652 mols, the oxygen is 0.896 mol, the nitrogen
is 0.2315 mol, and the ash is 65.32 kg. Carbon conversion in the gasifier is taken to be 87 mol%. Where
necessary, steam is fed to the gasifer, via heat exchanger HE-1, in sufficient amount to eliminate the
equilibrium constraint on carbon conversion; i.e., equilibrium conversion = 100%. All operating pressures and
pressure drops are taken to be the same as the 90-atm optimization case of Section 8: 41.16 atm in the
gasifier, 32.37 atm in the reformer, 90 atm in the methanol converter, and 82 atm in the methanol condenser.
A total of 2.42 mols of digester gas is fed to the desulfurizer together with the same amount of natural gas used
in previous simulations (3.79 mols). The reformer is operated at 950°C with a steam/carbon ratio of 2.50,
requiring 3.918 mols of natural gas as fuel in addition to 1.2 mols of purge.
The concentration of H2S in the gasifier effluent will be 3530 ppm, and the temperature will be 800°C.
Special materials will be required for the gasifier and heat exchangers, between them and the desulfurization
unit to withstand the corrosive effects of H2S at high concentration.
Table 12-1 summarizes the data from this simulation. The results show a methanol yield of 9.799 mols
which compares to 7.202 mols methanol for the base case in which the same amount of natural gas is fed to
the process, but 1.432 mols more natural gas is required as reformer fuel (base case = 2.486 mols natural gas
as reformer fuel). The most problematical of those results is the high throughput of the methanol converter,
389 mols, compared to 84.6 mols for the base case; as a result of the high throughput in the methanol loop,
68.66 kWh more electric power is consumed than can be generated by the process and must be imported.
It is apparent that the optimal conditions for the base case are not optimal for the sludge case. This is partly
due to the addition of digester gas containing a relatively large amount of methane which increases the
methanol loop throughput as discussed in Section 7 (sensitivity analysis). Another factor is the high nitrogen
content of the sludge fed which accumulates in the system to 22% of the converter throughput (compared to
6.2% for the base case). The equilibrium conversion of the reactant gases is adversely affected by this dilution.
Re-Optimization of Operating Conditions for Sludoe Addition
Here, simulations will be discussed that are aimed at reducing the throughput of the methanol
synthesis loop and thus reducing the cost of methanol production and decreasing the electric power required
in order to obtain a closer match with the power that can be recovered within the system. Two variables are
manipulated for this re-optimization: the pressure of the methanol converter and the purge rate. These
variables are systematically adjusted while retaining the operating conditions of the gasifier loop at the same
values indicated above. As discussed in Section 8 (process optimization), the recycle rate of the methanol
synthesis loop-and thus, the electric power required by the compressors feeding the methanol loop-can be
reduced by lowering the pressure difference between the two loops; i.e., by reducing the pressure in the
methanol converter. (It could also be reduced, as discussed in Section 8, by reducing the amount of natural
gas or digester gas fed to the process, but this option will not be considered here.) The effect of the purge rate
has not been previously discussed because the low nitrogen content of woody biomass did not result in the
accumulation of nitrogen in the system to high levels. Increasing the purge rate will obviously reduce the
methanol yield since some of the methanol produced will be lost in amounts proportional to the purge rate.
The following operating conditions apply to all simulations of this case:
Sludge feed
Woody biomass feed
Digester gas feed to process
Natural gas feed to process
161.28 kg, dry
0
2.42 mols
3.79 mols
Gasifier ash (including unreacted carbon) 71.9 kg
12-2
-------
TABLE 12-1 STREAM DATA FOR HYNOL PROCESS UTILIZING SEWAGE SLUDGE FOR TOTAL DISPLACEMENT
OF WOODY BIOMASS
90-atm methanol converter, 1.2 mols purge rate
Stream composition, kg-mols
Gasifier feed
Gasifier
effluent
Feed to
MeOH loop
MeOH loop
recycle
MeOH con-
verter effluent
Purge
Crude
MeOH
mo
0.007
1.884
0.2442
0.100
2.61
0.0003
2.50
Ho
17.14
11.692
40.25
242.66
260.6
0.8295
0
CO
0.2218
0.3844
7.749
3.141
3.374
0.0107
0
CO,
0.1336
0.0671
2.563
1.891
2.05
0.00647
0
CH,
1.464
5.188
1.557
20.73
22.29
0.0709
0
N,
5.712
5.943
6.03
80.88
86.91
0.2765
0
CH,OH
0.115
0
0
1.628
11.55
0.0056
9.799
SUM
24.79
25.16
58.39
351
389
1.20
12.299
Deg. C
900
800
50
40
260
40
30
Atm
41.66
41.16
24.8
82.0
90.0
82.0
1.0
Steam fed to reformer, mols
27.64
Steam fed to gasifier, mols
0
Reformer duty, cal/sec
176880
Reformer throughput, mols
77.0
Heat exchanger HE-1 area, m2
0.932
Electric power required for gasifier loop compressor, kWh
74.05
Electric power required for methanol loop compressor, kWh
31.49
Recycle expander power recovered, kWh
8.13
-------
Simulation data of Table 12-2 show the primary effects of increasing the purge rate from 1.2 to 2.4
mols. This change in purge rate greatly reduced the throughput of the methanol converter (by 73%) and also
reduced the overall electric power consumption from 97.4 to 65.9 kWh. The methanol yield is also reduced,
from 9.799 to 9.38 mols. The simulation data summarized in Table 12-3 show the results of reducing the
pressure of the methanol converter from 90 to 70 atm while maintaining a purge rate of 2.4 mols. The effects
of this pressure change are to: (1) reduce the power consumption of the gasifier loop compressor from 67.6
to 54.2 kWh, (2) increase the converter throughput from 104 to 138 mols, and (3) reduce the overall electric
power consumption to 61.3 kWh.
The final simulation of this case, given in Table 12-4, was made with a purge rate of 3.6 mols while
maintaining the converter pressure at 70 atm. For this optimized case, the imported electric power is reduced
to 7.0 kWh or 24.3 kWh/kg of methanol produced. The methanol converter throughput is 11.2 mols per mol
of methanol produced which compares to 11.7 for the base case. Similarly, the gasifier and reformer ratios
of mols throughput per mol of methanol produced are 2.40 and 8.51, respectively, which compare to 2.06 and
7.08 mols/mol for the base case. A total of 1.28 mols of methanol is obtained per mol of total natural gas used
by the process as feedstock and as reformer fuel.
Case 2: Intermediate Displacement of Woodv Biomass with Sludae
This case considers the use of 60 kg of sludge and 69.7 kg of woody biomass (containing 10.5%
moisture) as gasifier solid co-feedstocks. The digester gas is fed in the same ratio to sludge as the preceding
case, or 0.9 mol. The methanol converter pressure is again 70 atm, at which the optimization procedure gives
a purge rate of 2.4 mols. Table 12-5 summarizes the data from this run. The ratio of methanol throughput per
mol of methanol product is 11.8; the ratios for the gasifier and the reformer are 2.29 and 7.88, respectively.
The yield of methanol per mol of total natural gas used by the process as feedstock and as reformer fuel is
1.157 mol MeOH/mo! NG. No electric power import is required for this case. The methanol yield and
equipment size are similar to the Hynol base case without sludge addition.
Case 3: Low Displacement of Woodv Biomass with Sludoe
If 20 kg of dry sludge and 97.23 kg of woody biomass are fed to the gasifier, the total carbon content
of the mixture is the same as the two cases considered above. With a methanol converter pressure of 70 atm
and a purge rate of 2.1 mols, there is again no requirement for imported power. Table 12-6 gives the data on
this run. The ratio of methanol produced to converter throughput is 11.7 and the corresponding ratios for the
gasifier and reformer are 2.27 and 7.68. The total of methanol produced per mol of total natural gas used is
1.095.
Case 4: Sludoe Without Dioester gas
If, contrary to all of the above examples, no digester gas is fed to the process, Table 12-7 shows the
resulting effects on stream compositions, reactor throughputs, and methanol yield. No woody biomass is fed
with the sludge for this case and a purge rate of 3 mols is used. Like the results of Case I, where digester gas
was used, the process cannot operate without the import of electric power; 7.7 kWh in this case.
This case was run also with digester gas (2.42 mols) used as fuel for the reformer furnace. Since the
process stream compositions are the same as given in Table 12-7, only the NG required for the furnace is
changed. Its value is thus reduced from 2.763 to 1.19 mols.
Performance Comparisons
Table 12-8 compares the performance of the cases discussed above and the base case where no
sludge is added. Use of sludge and digester gas, with or without woody biomass as co-feedstock, will give
methanol yields comparable to the base case Hynol system in which only woody biomass and natural gas are
12-4
-------
TABLE 12-2. STREAM DATA FOR HYNOL PROCESS UTILIZING SEWAGE SLUDGE FOR TOTAL DISPLACEMENT
OF WOODY BIOMASS
90-atm methanol converter, 2.4 mols purge rate
Stream composition, kg-mols
Gasifier feed
Gasifier
effluent
Feed to
MeOH loop
MeOH loop
recycle
MeOH con-
verter effluent
Purge
Crude
MeOH
H,0
0.0052
2.701
0.2232
0.0180
2.028
0.0006
2.00
H,
14.804
8.001
37.47
51.49
68.077
1.759
0
CO
0.6153
1.077
8.264
2.14
2.83
0.0731
0
CO,
1.029
0.3929
3.250
3.58
4.917
0.1222
0
CH,
1.154
5.126
1.312
4.014
5.327
0.1371
0
N,
2.50
2.730
2.817
8.69
11.506
0.2968
0
CH,OH
0.0941
0
0
0.327
9.81
0.0111
9.381
SUM
20.20
20.03
53.33
70.26
104
2.40
11.381
Deq. C
900
800
50
40
260
40
30
Atm
41.66
41.16
24.8
82.0
90.0
82.0
1.0
Steam fed to reformer, mols 29.4
Steam fed to gasifier, mols 0
Reformer duty, cal/sec 177760
Reformer throughput, mols 74.0
Heat exchanger HE-1 area, m2 0.741
Electric power required for gasifier loop compressor, kWh 67.57
Electric power required for methanol loop compressor, kWh 6.29
Recycle expander power recovered, kWh 7.96
-------
TABLE 12-3. STREAM DATA FOR HYNOL PROCESS UTILIZING SEWAGE SLUDGE FOR TOTAL DISPLACEMENT
OF WOODY BIOMASS
70-atm methanol converter, 2.4 mols purge rate
Stream composition, kg-mols
Gasifier feed
Gasifier
effluent
Feed to
MeOH loop
MeOH loop
recycle
MeOH con-
verter effluent
Purge
Crude
MeOH
H,0
0.4568
3.095
0.2257
0.034
2.15
0.0008
2.106
H,
14.852
8.175
37.67
74.75
91.35
1.730
0
CO
0.8027
1.253
8.413
4.04
4.94
0.0935
0
CO,
1.102
0.5123
3.396
5.55
6.93
0.1284
0
CH.
1.127
5.088
1.274
5.67
9.48
0.1313
0
N,
2.592
2.823
2.91
13.05
15.96
0.302
0
CH,OH
0.117
0
0
0.589
10.11
0.0136
9.40
SUM
21.05
20.95
53.89
103.7
138
2.40
11.506
Deq. C
900
800
50
40
260
40
30
Atm
41.66
41.16
24.8
82.0
90.0
82.0
1.0
Steam fed to reformer, mols 29.7
Steam fed to gasifier, mols 0.45
Reformer duty, cal/sec 178520
Reformer throughput, mols 75.24
Heat exchanger HE-1 area, m2 0.787
Electric power required for gasifier loop compressor, kWh 54.21
Electric power required for methanol loop compressor, kWh 11.98
Recycle expander power recovered, kWh 4.86
-------
TABLE 12-4. STREAM DATA FOR HYNOL PROCESS UTILIZING SEWAGE SLUDGE FOR TOTAL DISPLACEMENT
OF WOODY BIOMASS
70-atm methanol converter, 3.6 mols purge rate
Stream composition, kg-mols
Gasifier feed
Gasifier
effluent
Feed to
MeOH loop
MeOH loop
recycle
MeOH con-
verter effluent
Purge
Crude
MeOH
H,0
1.156
3.975
0.2250
0.0197
1.77
0.0011
1.74
H»
14.967
8.107
37.54
47.83
65.50
2.6156
0
CO
1.084
1.709
8.788
3.464
4.74
0.1893
0
CO,
1.670
0.904
3.833
5.336
7.55
0.2925
0
CH,
0.9914
4.956
1.178
3.168
4.35
0.1732
0
N,
1.761
1.993
2.080
5.628
7.71
0.3075
0
CH,OH
0.1196
0
0
0.382
9.53
0.0209
9.010
SUM
21.75
21.64
53.64
65.83
101
3.60
10.75
Deg. C
900
800
50
40
260
40
30
Atm
41.66
41.16
24.8
82.0
90.0
82.0
1.0
Steam fed to reformer, mols 30.5
Steam fed to gasifier, mols 1.15
Reformer duty, cal/sec 179070
Reformer throughput, mols 76.66
Heat exchanger HE-1 area, m2 0.826
Electric power required for gasifier loop compressor, kWh 53.64
Electric power required for methanol loop compressor, kWh 7.59
Recycle expander power recovered, kWh 4.85
-------
TABLE 12-5. STREAM DATA FOR HYNOL PROCESS UTILIZING AN INTERMEDIATE AMOUNT OF SEWAGE SLUDGE
FOR PARTIAL DISPLACEMENT OF WOODY BIOMASS
70-atm methanol converter, 2.4 mols purge rate
Stream composition, kg-mols
Gasifier feed
Gasifier
effluent
Feed to
MeOH loop
MeOH loop
recycle
MeOH con-
verter effluent
Purge
Crude
MeOH
mo
1.405
3.652
0.1794
0.0211
1.71
0.0008
1.68
H.
11.117
6.569
30.27
48.49
61.40
1.790
0
CO
1.7021
1.582
7.093
3.064
3.88
0.113
0
CO,
1.125
0.9476
3.088
4.910
6.40
0.181
0
CHt
0.8219
3.908
0.9672
3.59
4.55
0.132
0
N,
1.050
1.140
1.227
4.58
5.81
0.1691
0
CH,OH
0.0841
0
0
0.367
8.24
0.0135
7.777
SUM
16.30
17.80
42.82
65.0
92
2.40
9.457
Dep. C
900
800
50
40
260
40
30
Atm
41.66
41.16
24.8
82.0
70.0
62.0
1.0
Steam fed to reformer, mols 24.2
Steam fed to gasifier, mols 1.40
Reformer duty, cal/sec 142914
Reformer throughput, mols 61.34
Heat exchanger HE-1 area, m2 0.630
Electric power required for gasifier loop compressor, kWh 42.82
Electric power required for methanol loop compressor, kWh 7.50
Recycle expander power recovered, kWh 2.93
-------
used. The equipment throughputs and sizes are increased, but only by 6 %. Waste disposal is quite large,
mainly because of the high ash content of sludge and the fact that the un-gasified carbon discharged from the
gasifier cannot be utilized as fuel for the reformer and is added waste. The amount of waste is nevertheless
reduced from 161.3 to 71.9 kg, or 55 % compared to the amount of waste that would have to be landfilled were
the sludge not used for Hynol feedstock. Sulfur removed from the desulfurization unit (as S02 released during
regeneration of the ZnO sorbent) must also be recovered and disposed of. That step would most likely be
done by limestone scrubbing and disposal as calcium sulfate.
Electric power requirements for the four cases discussed are specified and compared with the base
case in Table 12-9. Heat exchanger duties are given in Table 12-10.
12-9
-------
TABLE 12-6. STREAM DATA FOR HYNOL PROCESS UTILIZING A SMALL AMOUNT OF SEWAGE SLUDGE FOR
PARTIAL DISPLACEMENT OF WOODY BIOMASS
70-atm methanol converter, 2.1 mols purge rate
Stream composition, kg-mols
Gasifier feed
Gasifier
effluent
Feed to
MeOH loop
MeOH loop
recycle
MeOH con-
verter effluent
Purge
Crude
MeOH
H,0
1.604
3.627
0.1616
0.0201
1.64
0.0007
1.618
H,
9.627
5.993
27.413
46.02
57.25
1.592
0
CO
0.5893
1.5718
6.466
2.817
3.50
0.0975
0
CO,
0.9730
1.024
2.848
4.65
5.96
0.1609
0
CH„
0.7373
3.480
0.8719
3.53
4.40
0.1219
0
N,
0.6984
0.7318
0.8191
3.34
4.16
0.115
0
CH,OH
0.0708
0
0
0.339
7.66
0.0117
7.237
SUM
14.30
16.43
38.58
60.72
84.58
2.10
8.855
Deg. C
900
800
50
40
260
40
30
Atm
41.66
41.16
24.8
62.0
70.0
62.0
1.0
Steam fed to reformer, mols 21.83
Steam fed to gasifier, mols 1.60
Reformer duty, cal/sec 128805
Reformer throughput, mols 55.56
Heat exchanger HE-1 area, m2 0.5506
Electric power required for gasifier loop compressor, kWh 38.58
Electric power required for methanol loop compressor, kWh 6.99
Recycle expander power recovered, kWh 2.50
-------
TABLE 12-7. STREAM DATA FOR HYNOL PROCESS UTILIZING SEWAGE SLUDGE FOR FULL DISPLACEMENT
OF WOODY BIOMASS WITHOUT DIGESTER GAS ADDITION
70-atm methanol converter, 3.0 mols purge rate
Stream composition, kq-mols
Gasifier feed
Gasifier
effluent
Feed to
MeOH loop
MeOH loop
recycle
MeOH con-
verter effluent
Purge
Crude
MeOH
HoO
0.0058
2.416
0.1883
0.0260
1.496
0.0009
1.463
H,
14.695
8.235
32.72
65.76
82.768
2.302
0
CO
0.5811
0.8014
6.460
2.60
3.273
0.0910
0
o
o
V)
0.6198
0.2540
2.165
2.77
3.560
0.0971
0
CH,
1.1743
5.13
1.3719
5.25
6.623
0.1839
0
N,
1.9674
2.199
2.2861
8.803
11.087
0.3081
0
CH,OH
0.107
0
0
0.479
7.643
0.01676
7.041
SUM
19.15
19.04
45.19
85.7
116
3.00
8.504
Deg. C
900
800
50
40
260
40
30
Atm
41.66
41.16
24.8
62.0
70.0
62.0
1.0
Steam fed to reformer, mols 22.60
Steam fed to gasifier, mols 0
Reformer duty, cal/sec 145016
Reformer throughput, mols 60.55
Heat exchanger HE-1 area, mz 0.7226
Electric power required for gasifier loop compressor, kWh 45.26
Electric power required for methanol loop compressor, kWh 9.91
Recycle expander power recovered, kWh 7.54
-------
TABLE 12-8. PERFORMANCE SUMMARY OF SEWAGE SLUDGE SIMULATIONS
Base
case
Sludge
only
Intermed.
sludge
Low
sludge
Sludge
w/o ADG
Biomass feed (incl. moisture and ash), kg
111
0
69.7
97.23
0
Sludge feed (dry, with ash), kg
0
161.28
60
20
161.28
ADG feed, mols
0
2.42
0.9
0.3
0
NG feed to process, mols
3.79
3.79
3.79
3.79
3.79
CH4 feed as ADG, mols
0
1.573
0.585
0.195
0
NG used as reformer fuel, mols
2.486
3.250
2.930
2.815
2.763
Methanol produced, mols
7.202
9.010
7.777
7.237
7.041
Gasifier throughput per mol MeOH
2.06
2.40
2.29
2.27
2.64
Reformer throughput per mol MeOH
7.08
8.51
7.89
7.68
8.60
Converter throughput per mol MeOH
11.7
11.2
11.8
11.7
16.5
Total throughput per mol MeOH
20.9
22.1
22.0
21.6
27.8
Mols MeOH/mol total NG
1.147
1.280
1.157
1.096
1.074
Mols MeOH/mol total CH4
1.147
1.046
1.064
1.064
1.074
Waste produced (incl. carbon), kg
0.764
71.9
31.3
15.3
71.9
Thermal efficiency, %
68.4
66.1
66.3
65.7
63.5
Fuel cycle C02 emission reduction, kg
724
968
792
713
700
12-12
-------
TABLE 12-9. SUMMARY OF ENERGY REQUIREMENTS FOR SEWAGE SLUDGE SIMULATIONS
hp
w
Base case
Case I
Sludge only
Case 2
Intermed. sludge
Case 3
Low sludge
Case 4
Sludge w/o ADG
NG required as reformer fuel, mols
2.486
3.250
2.93
2.817
2.763
Electric power required (kWh) as:
Gasifier loop compressor
45.89
53.64
42.82
38.58
45.26
Methanol loop compressor
5.63
7.59
7.50
6.99
9.91
SPR furnace air blower
11.0
16.5
13.73
12.67
13.7
ADG compressor
0
12.45
4.63
1.54
0
NG compressor
1.15
1.15
1.15
1.15
1.15
HP BFW pumps
0
2.44
0
0
0
IP BFW pumps
2.0
1.92
2.25
1.92
2.04
Drier recycle fan
1.2
1.2
1.2
1.2
1.2
C ooling water pumps
2.1
2.1
1.99
1.67
2.3
Other
1.7
1.7
1.7
1.7
1.7
Total electric power required, kWh
70.7
100.7
77.0
67.3
77.2
Power available (kWh) from:
HRSG-2
0
20.97
0
0
0
HRSG-1
57.0
57.0
66.32
60.0
52.44
HRSG-3
4.6
0
0
0
0
Recycle expander
3.7
4.85
2.93
2.50
7.54
NG fuel expander
6.39
10.91
8.96
8.21
9.49
Total power available, kWh
71.7
93.7
78.2
70.7
69.5
-------
TABLE 12-10. HEAT EXCHANGER DUTIES AND AREAS FOR HVNOL PROCESS UTILIZING SEWAGE TREATMENT SLUDGE,
DIGESTER GAS , AND WOODY BIOMASS AS CO-FEEDSTOCKS
Heal
exchnger ID
Sludge only
Duty, cat/sec Area, m'
Intermediate sludge
Duty, cat/sec Area, m2
Low sludge
Duty, cal/sec Area, m*
Sludge without digester gas
Duty, cal/sec Area, m*
SH-G1
5761
0.0664
7285
0.0908
6266
0.0761
5025
0.0571
SH-G2
7173
0.1395
2870
0.0632
2808
0.0592
2907
0.0510
SH-G3
2687
0.1004
2337
0.0807
2237
0.0736
3094
0.0764
HRSG-1
12800
0.3446
11130
0.2984
10660
0.2848
14736
0.368
PGPH-1
20670
1.621
16710
2.429
15250
1.819
16631
2.239
PGPH-2
29065
3.779
23200
1.291
21070
1.117
22480
1.308
HE-1
38400
0.8263
30790
0.630
28780
0.5506
32789
0.7226
SH-S2
15743
0.3168
n/a
n/a
n/a
n/a
n/a
n/a
SH-S1
5930
0.1124
11160
0.1946
9960
0.1680
12838
0.2473
HRSG-2
47360
1.170
52600
1.391
46950
1.138
35885
0.8239
HE-6
16110
0.640
13300
1.087
12320
0.790
11534
0.396
HRSG-3
15337
1.086
15337
1.949
15337
1.822
15337
1.00
Reboiler
28311
1.67
24490
1.47
22800
1.40
22145
1.35
HE-2
2082
0.1167
1840
0.1077
1730
0.100
1654
0.0953
HE-3
31300
6.755
21600
5.37
19970
4.658
24160
5.50
HE-5
1362
0.0811
998
0.0647
871
0.0557
645
0.0428
APH-1
20480
2.823
19360
3.53
17860
2.435
17780
2.578
FGPH-1
943
0.0231
943
0.0228
943
0.0226
943
0.0231
FGPH-2
6592
0.210
5260
0.238
4618
0.185
5355
0.226
HE-4
43260
6.172
39030
5.563
35970
5.114
46942
6.810
Cond-3
36530
6.59
32560
5.72
30356
5.31
32740
7.03
-------
SECTION 13
NATURAL GAS AS CO-FEEDSTOCK FOR
THE BCL METHANOL PROCESS
Introduction
The BCL process, as discussed in Section 4, utilizes biomass exclusively as feedstock for the
production of methanol or hydrogen. It is a conceptual design prepared by Princeton University that would
utilize the biomass gasifier developed by the Battelle Columbus Laboratory as the first step of the process.
A conventional steam reformer, a shift reactor, and a Selexol unit were added by Princeton to the BCL gasifier
to produce methanol as a final product for use as transportation fuel. Of the five gasifier types that they
evaluated, this conceptual arrangement was found to yield the lowest production cost.
As discussed in Section 1, there are two major barriers to the use of biomass alone as feedstock for
production of transportation fuel. The first barrier is the high production cost of fuel methanol ($12.4/GJ)
relative to gasoline ($4.5/GJ). The second barrier is the limited supply of biomass that might be produced as
dedicated energy crops for conversion to transportation fuel; the amount of fuel that could be produced from
that biomass would not supply the major portion of the U.S. vehicle population. In Section 5 it has been shown
that the use of natural gas as co-feedstock for the Hynol process will leverage the amount of transportation fuel
that can be produced from a given biomass supply and that the additional displacement of petroleum fuels thus
achieved also results in greater reduction of greenhouse gas emissions from the vehicle population in spite
of the C02 emissions resulting from the natural gas used. Natural gas also increases the size of the plant
producing that fuel, providing further cost reduction due to economy of scale.
This section will examine the feasibility of utilizing natural gas as co-feedstock with the BCL process
and compare that approach with the performance projections of the Hynol process discussed earlier in this
report.
BCL PrQcess Modification for Ngturgl Ggs Addition
Figure 13-1 is a block diagram of the BCL process modified to accept natural gas as co-feedstock.
The biomass gasifier and char combustor remain the same as that discussed in Section 4, Figure 4-1. The
changes consist of enlargement of the steam reformer, synthesis-gas compressor, and methanol converter
to accommodate the increased throughput due to the addition of natural gas to the reformer. As was the case
in all prior simulations of the Hynol process, a biomass feed of 100 kg (dry basis) containing 10% moisture is
the basis for this simulation. The amount of natural gas added is the same as the total amount used for the
optimized 90-atm Hynol simulations (3.79 mols as process feed and 2.486 mols as reformer fuel, 6.276 mols
total). Steam is also fed to the reformer in sufficient amount to provide a steam/carbon ratio of 2.5 entering the
reformer. Two methods of adding natural gas were examined. In the first method, all of the natural gas is
assumed to be added to the reformer as part of the process gas entering that unit; the fuel needed for the
reformer furnace in that case is obtained entirely as purge from the methanol converter. The second method
13-1
-------
Natural gas
Desulfurizer
Biomass
gasifier/combustor
Quench
HRSG-1
Steam
CH,OH
Steam to
biomass
drier
Distillation
Reformer
HRSG-2
BFW
Methanol
converter
Crude
methanol
Figure 13-1. Block diagram for BCL process modified for natural gas addition.
13-2
-------
of adding the natural gas assumes that part of it is added to the process stream entering the reformer and the
remainder is used as fuel for the reformer furnace (together with a smaller amount of purge gas than that
required for first case).
Neither case was found to require a shift converter following the reformer or a Selexol unit for C02
removal; those parts of the BCL process are therefore eliminated. The Lurgi type methanol converter is used
here instead of the ICI type assumed by Princeton for the BCL process; a lower gas recycle in the methanol
loop is not required for temperature control in the methanol converter (it is controlled by boiling water; the
amount of gas recycled is therefore independent of the converter cooling requirement).
As discussed in Section 4, 3.15 mols of steam at 310 °C is extracted in HRSG-1 from the hot gas
exiting the biomass gasifier, and the gases are then quenched from 355 to 82 °C with water. The particulate
and tars are assumed removed in that step. The gases are then compressed in three stages (to 2,5.7, and
16.9 atm) with interstage cooling to 40 °C. Steam is then added (generated by HRSG-1, HRSG-2 after the
reformer, and the steam boiler on the methanol converter). The natural gas is desulfurized with ZnO, then
expanded in a turbine to 16.9 atm and fed into the process stream entering the reformer, at which point the
steam/carbon ratio is 2.50. The process stream enters the preheater (located within the reformer combustion
flue gas stream) and then enters the catalytic reformer. This reformer is assumed to operate at the same
temperature (B67°C) and pressure (14.1 atm) as the BCL process. The pressure drop, however, is taken to
be 2.5 atm, which is the low end of the range specified by ICI for steam reformers. After reforming, high
temperature steam is raised in HRSG-2 for the reformer and, for electric power production, that step is
followed by a series of heat exchangers that provide the remaining energy required to operate the process:
(i) heating the boiler feed water for the methanol converter's steam boiler, (ii) generating steam for the biomass
drier, and (iii) generating low pressure steam for the biomass gasifier. The gases leaving that section at 145°C
then pass through the distillation reboiler for separation of water from the crude methanol, a preheater for the
methanol column feed, a natural gas preheater, and finally, a condenser.
The pressure drops in the reformer-feed preheater and all heat exchangers following the reformer are
allocated on the same basis as that used for the Hynol optimization: the pressure drop across each heat
exchanger was taken to be proportional to the heat exchange area calculated by Aspen using a proportionality
constant of 0.716. The resulting pressure of the synthesis gas entering the main compressor is 8.09 atm for
Case I and 9.13 atm tor Case 2. The methanol converter feed stream is compressed to104.5 atm as for the
BCL process; the methanol converter pressure, however, is taken to be 104 atm because the pressure drop
of the Lurgi converter is less than the 8 atm required for the ICI unit. (This 8-atm pressure drop is allocated
to the large syngas preheater/heat-exchanger and crude methanol condenser that follow methanol synthesis.)
In addition to the electric power produced by steam generated from the high temperature process
gases leaving the reformer, a large amount of power is recovered from the hot flue gas leaving the reformer
furnace (in addition to that required for preheating the furnace air and the reformer feed stream). The amount
of steam thus generated is dependent on the total enthalpy of the furnace fuel. In both cases, the amount of
purge fed to the reformer as fuel was adjusted to allow the total electric power generated to balance the total
power required by the system. Unlike the original BCL process, no electric power import is required by the
modified process.
Tables 13-1 and 13-2 summarize the stream compositions for the two methods of natural gas addition
described above. Both simulations yield less methanol than the Hynol system (6.511 for Case 1, and 6.781
mols for Case 2, compared to 7.202 mols for Hynol). The results thus indicate that part of the natural gas
should be used as fuel for the reformer for maximum methanol yield; the remaining discussion of the modified
BCL system therefore focuses on Case 2. The split of natural gas used for Case 2 is 4.75 mols for process
feed and 1.526 mols for reformer fuel.
13-3
-------
TABLE 13-1. PRINCIPAL STREAM COMPOSITIONS FOR BCL PROCESS MODIFIED FOR
NATURAL GAS ADDITION. TOTAL NATURAL GAS USED = 6.276 mols
Option 1: All natural gas used as process feed
Stream composition, mols
Stream
No. 1
Stream
No. 2
Stream
No. 3
Stream
No. 4
Stream
No. 5
Stream
No. 6
Stream
No. 7
Stream
No. 8
Stream
No. 9
H20
1.645
2.922
23.70
14.35
0.302
0.0674
0.0733
1.524
0.0028
h2
0.776
0.776
0.776
22.75
22.75
22.75
40.81
26.18
8.101
CO
1.726
1.726
1.726
5.595
5.595
5.595
6.609
1.472
0.4549
C02
0.416
0.416
0.429
3.16 7
3.167
3.167
6.158
4.707
1.341
o
X
0.588
0.588
6.531
0.7166
0.7166
0.7166
2.270
2.270
0.6967
C2H4
0.194
0.194
0.194
0
0
0
0
0
0
c2h6
0.027
0.027
0.2027
0
0
0
0
0
0
n2
0.016
0.016
0.1603
0.1603
0.1603
0.1603
0.513
0.513
0.1583
ch3oh
0
0
0
0
0
0
0.100
6.688
0.0449
SUM
5.388
6.665
33.72
46.74
32.69
32.46
56.54
43.35
10.80
K
1200
355
651
1140
313
424
380
533
313
Atm
1.0
1.0
16.92
14.11
8.09
104.5
104.5
104
96
Natural gas used as process co-feedstock, mols 6.276
Natural gas used as reformer fuel, mols 0
Steam fed to reformer, mols 23.65
Steam/carbon ratio in reformer 2.50
Reformer duty, cal/sec 119667
Methanol product, mols 6.511
13-4
-------
TABLE 13-2. PRINCIPAL STREAM COMPOSITIONS FOR BCL PROCESS MODIFIED FOR
NATURAL GAS ADDITION. TOTAL NATURAL GAS USED = 6.276 mols
Option 2: 24.3% of natural gas used as reformer fuel
Stream composition, mols
Stream
No. 1
Stream
No. 2
Stream
No. 3
Stream
No. 4
Stream
No. 5
Stream
No. 6
Stream
No. 7
Stream
No. 8
Stream
No. 9
HzO
1.645
2.922
19.85
12.24
0.2193
0.0552
0.0736
2.330
0.0012
h2
0.776
0.776
0.776
18.36
18.36
18.356
56.94
41.09
2.493
CO
1.726
1.726
1.726
4.637
4.637
4.637
6.131
1.591
0.0965
C02
0.416
0.416
0.4255
2.774
2.774
2.774
7.337
5.081
0.2949
ch4
0.588
0.588
5.086
0.535
0.535
0.535
8.247
8.249
0.4984
c2h<
0.194
0.194
0.194
0
0
0
0
0
0
c2h6
0.027
0.027
0.160
0
0
0
0
0
0
n2
0.016
0.016
0.1253
0.1253
0.1253
0.1253
2.005
2.005
0.1215
CHaOH
0
0
0
0
0
0
0.2192
7.015
0.0141
SUM
5.388
6.665
28.34
38.67
26.65
26.48
80.95
67.36
3.52
K
1200
355
858
1140
313
313
319
533
313
Atm
1.0
1.0
16.89
14.11
9.13
104.5
104.5
104
96
Natural gas used as process co-feedstock, mols 4.75
Natural gas used as reformer fuel, mols 1.526
Steam fed to reformer, mols 19.8
Steam/carbon ratio In reformer 2.50
Reformer duty, cal/sec 95379
Methanol product, mols 6.781
13-5
-------
Figure 13-2 and Table 13-3 give Ihe details of the modifications to the original BCL system that prepare
the process gases produced by the biomass gasifier prior to natural gas addition. Because a higher pressure
drop is taken in the preheater and reformer, the pressure of the steam and natural gas fed is higher than the
original BCL process (16.9 atm), but exit the reformer at the same assumed pressure of 14.1 atm.
Figure 13-3 and Table 13-4 show the details of the reforming and power block which is the second
major step of the process. After steam and natural gas addition, the process stream is preheated in a heat
exchanger in the furnace flue gas and enters the reforming reactor at 585 °C. Gases leaving the reformer at
867 °C pass through a series of heat exchangers that generate steam for the reformer and for generation of
part of the electric power required to operate the system. Additional power is produced from steam generated
by the heat recovered from the hot flue gas downstream from the reformer/preheater. Three final heat
exchangers in this block recover energy to heat the boiler feed water for the methanol-converter steam boiler,
to generate low pressure steam for the biomass drier, and to generate the steam needed by the BCL biomass
gasifier.
The process gas stream leaves the reformer block at 145 °C and enters the reboiler of the distillation
block as shown in Figure 13-4 and Table 13-5. This arrangement for recovery of heat for distillation following
reforming and preceding methanol synthesis follows conventional practice for methanol production. The crude
methanol condenser, HE-3, is also the main initial preheater for boiler feed water for the heat recovery steam
generators, providing a water temperature of 90°C while cooling the process stream to 50°C before
compression and methanol synthesis. Compression occurs in three stages (to 20, 40, and 104.5 atm) with
interstage cooling to 40 °C to enter the methanol synthesis block.
The methanol synthesis block consists of a Lurgi-type converter, a heat exchanger to preheat the
syngas with hot converter effluent, a condenser for crude methanol, and a recycle compressor. This
arrangement, Figure 13-5 and Table 13-6, is again conventional practice for the Lurgi system. The
temperature of the exothermic methanol synthesis reaction in the catalytic converter is controlled by boiling
water surrounding the vertical tubes containing the catalyst. This boiler maintains the temperature within the
catalyst to the 260°C optimum for conversion to methanol and produces a large amount of medium pressure
steam from which electric power is extracted in the reforming block before the steam is fed to the reformer.
The recycle ratio for this system is only 2.06.
Heat exchanger duties and areas are given in Table 13-7. The number and types correspond closely
with those required for the Hynol process with the notable exception of the high temperature ceramic heat
exchanger needed for the latter process.
Performance Evaluation
Table 13-8 shows the electric power requirements and the power that can be recovered within the
process for both cases of natural gas addition. As indicated earlier, a condition assumed for these simulations
was that the system would be configured so that the electric power was balanced and no import would be
required. Although both cases are balanced, the second case requires 19% less power which is a major factor
contributing to its higher methanol yield (natural gas used directly for reformer fuel does not have to be
reformed to syngas and compressed before coming back as furnace fuel). In all cases where pressure
reduction is necessary, such as the use of converter purge for reformer fuel or feeding natural gas as either
feedstock or furnace fuel, it is assumed that a turbine-generator Is used to recover that energy as well as to
facilitate pressure reduction. This is a substantial amount of energy where the purge gas is used for reformer
fuel.
In addition to the elimination of the shift reactor and Selexol unit that are part of the original BCL
process, the use of natural gas also provides sufficient low pressure steam for biomass drying (via HRSG-3
in the reformer block) and therefore allows the recovery of high pressure steam from the flue gas leaving the
13-6
-------
C-1
H*0
2.778 mo 15
Quencher
c-z
H.0
0.0929 mol
HRSG-1
TG-I
Gasifier
Natural gas
Figure 13-2. Details of the gas preparation block for BCL process modified for natural gas addition.
13-7
-------
TABLE 13-3. STREAM DATA FOR BCL PROCESS MODIFIED FOR NATURAL GAS ADDITION.
GAS PREPARATION BLOCK (FIGURE 13-2)
Stream
No.
I.D.
Flow,
kg-mol
Temp.,
°C
Pressure,
atm
1
Gasifier effluent
5.388
927
1
2
Heat recovery steam generator HRSG-1 effluent
5.388
344
1
3
Boiler feed water
3.15
90
97.3
4
HP steam to reformer block
3.15
310
97.3
5
Quench water
2.018
90
1
6
Quencher effluent
6.681
82
1
7
Quench condensate water
0.741
82
1
8
Second stage compressor feed
3.886
40
2.0
9
Third stage compressor feed
3.793
40
5.7
10
Third stage compressor effluent
3.793
158
16.9
11
Natural gas
4.75
90
28.5
12
Desulfurized natural gas
4.75
268
16.9
13
Process gas feed to reformer
28.34
382
16.9
30
I.P. steam from HRSG-2 and reheater SH-RH
19.8
527
16.9
13-8
-------
283* C
Purge
HRSG-4
From gas
preparation
block
TG-5
SH-S1
©~
TG-2
©~
TG-3
ME-S
To distillation
block
43
Figure 13-3. Details of the reformer and power block for modified BCL process.
13-9
-------
TABLE 13-4. STREAM DATA FOR BCL PROCESS MODIFIED FOR NATURAL GAS ADDITION.
REFORMER AND POWER BLOCK (FIGURE 13-3)
Stream
No.
I.D.
Flow,
kg-mol
Temp.,
°C
Pressure,
atm
4
Steam from HRSG-1
3.15
310
97.3
10
Compressor effluent from gas preparation block
3.837
158
16.9
12
Desulfurized natural gas
4.75
268
16.9
13
Process gas feed to reformer
28.34
382
16.9
16
Preheated feed to reformer
28.34
585
16.6
17
Reformer effluent
38.67
867
14.1
18
Steam reheater SH-RH hot-side effluent
38.67
686
13.96
19
Steam superheater SH-S1 hot-6ide effluent
38.67
627
13.88
20
HRSG-2 hot-side effluent
38.67
517
13.76
21
HE-6, BFW heater for MSR, hot-side effluent
38.67
365
13.59
22
Biomass drier steam generator, hot-side effluent
38.67
185
13.15
23
Gasifier steam generator, HE-G, hot-side effluent
38.67
145
12.84
24
H.P. BFW for HRSG-2
3.70
90
97.3
25
H.P. steam
3.70
310
97.3
26
Superheated steam
6.85
527
97.3
27
Turbine-generator TG-3 condensate water
3.48
30
0.05
2B
Steam from methanol converter
16.44
250
39.2
29
Turbine-generator TG-2 and TG-4 effluent
19.8
205
17.2
30
Steam feed to reformer
19.8
527
16.9
31
Air for reformer furnace
30.0
25
1.0
32
Air to preheater
30.0
69
1.5
33
Preheated air to furnace
30.0
260
1.0
34
Natural gas feed for furnace fuel
1.526
25
1.5
35
Preheated fuel gas for furnace
5.046
310
1.0
36
Boiler feed water from HE-3
9.17
90
97.3
37
H.P. steam to superheater
9.17
310
97.3
38
Steam to turbine-generator TG-5
9.17
527
97.3
39
Condensate water
9.17
30
0.05
76
Purge from methanol synthesis block TG-6
3.52
40
96.0
13-10
-------
c-s
From reformer
block
c-1
Reboiler
63
C-3
S2|
—0—~
56
Figure 13-4. Details of the distillation block for modified BCL process.
13-11
-------
No.
23
50
51
52
53
54
56
57
58
59
60
61
62
63
64
65
66
67
TABLE 13-5. STREAM DATA FOR THE MODIFIED BCL PROCESS.
DISTILLATION BLOCK (FIGURE 13-4)
I.D.
Flow,
kg-mol
Temp.,
°C
Pressure,
atm
Feed from reformer block
38.67
145
12.84
Reboiler hot-side effluent
38.67
115
11.36
Reboiler column feed
9.091
76
1.0
Preheated feed to distillation column
9.091
65
1.0
Hot-side effluent from column feed heater HE-2
38.67
114
11.27
Condensate water
7.75
115
11.27
Natural gas feed
6.276
25
28.5
Preheated natural gas
6.276
90
28.5
Condenser hot-side feed
30.93
110
11.18
Boiler feed water
40.5
25
3.0
Boiler feed water
40.5
90
2.0
Condensate water
4.15
40
9.13
First stage compressor feed
26.65
40
9.13
Second stage compressor feed
26.53
40
20
Third stage compressor feed
26.48
40
40
Feed to methanol converter
26.48
40
104.5
Column bottoms (water)
2.31
103
1.0
Methanol product
6.781
30
1.0
13-12
-------
Methanol
converter
72
Synthesis gas
from distillation
block
HE-7
HE-8
78
|7*i
Purge gas
to reformer
furnace
[75]-—~ Crude
Figure 13-5. Methanol synthesis block for modified BCL process.
13-13
-------
27
40
65
70
71
72
73
74
75
76
77
78
TABLE 13-6. STREAM DATA FOR MODIFIED BCL PROCESS.
METHANOL SYNTHESIS BLOCK (FIGURE 13-5)
I.D.
Flow,
kg-mol
Temp.,
°C
Steam to turbine-generator TG-3, reformer block
16.44
250
Boiler feed water from HE-6
16.44
246
Synthesis gas feed from distillation block compressor
26.48
40
Converter feed preheater cold-side inlet
81.13
70
Preheated converter feed
81.13
225
Converter effluent
67.54
260
Preheater hot-side effluent
67.54
129
Condenser effluent
67.54
40
Crude methanol
9.091
40
Purge gas to reformer furnace
3.52
40
Recycle compressor feed
54.65
40
Condenser cooling water
800
25
13-14
-------
TABLE 13-7. HEAT EXCHANGER DUTIES AND AREAS FOR MODIFIED BCL PROCESS
(OPTION 2: NATURAL GAS USED AS BOTH PROCESS FEED AND AS REFORMER FUEL)
Unit ID
Description
Duty, cal/sec
Area, m2
HRSG-1
Heat recovery steam generator
8830
0.1240
C-1
Compressor effluent cooler
9785
2.047
C-2
Compressor effluent cooler
1246
0.0813
PGPH
Process gas preheater
16474
0.617
SPR
Steam pyrolysis reactor (reformer)
95329
na
HE-RH
Steam reheater, for reformer feed
17140
0.208
SH-S1
Steam superheater
5413
0.116
HRSG-2
Steam generator for reformer
9871
0.165
HE-6
Boiler feed water heater for MSR
13450
0.243
HRSG-3
Steam generator for biomass drier
15337
0.615
HE-G
Steam generator for gasifier
4253
0.425
Reboiler
Distillation column reboiler
2184
2.06
HE-2
Crude methanol preheater
1845
0.141
HE-5
Natural gas preheater
1404
0.133
HE-3
Condenser
13134
2.85
C-3
Compressor effluent cooler
5208
0.617
C-4
Compressor effluent cooler
4397
0.534
C-5
Compressor effluent cooler
6241
0.618
HE-7
Syngas preheater for methanol converter
30398
3.896
MSR
Methanol synthesis reactor
33464
na
HE-8
Crude methanol condenser
28730
3.91
HRSG-4
Steam generator for turbine-generator TG-5
29730
0.575
SH-SPR
Steam superheater for TG-5
4939
0.0153
APH
Air preheater for reformer furnace
11256
1.651
PGPH
Fuel gas preheater for reformer furnace
3052
0.208
13-15
-------
TABLE 13-8. ELECTRIC POWER UTILIZATION FOR BCL PROCESS
MODIFIED FOR NATURAL GAS ADDITION
Mode of natural gas addition
Reformer
feed only
Reformer feed
and furnace fuel
Power requirements, kWh
Methanol converter feed compressor
82.64
63.74
Feed compressor, gas preparation block
13.16
13.16
Reformer furnace air blower
11.51
10.56
Pumps
6.10
5.47
Air blower, biomass char combustor
2.17
2.17
Recycle compressor, methanol loop
1.79
4.15
Steam recycle blower, biomass drier
1.2
1.2
Total power required
118.6
100.45
Power available, kWh
TG-1, natural gas expander (to process)
1.67
1.37
TG-2, reformer steam turbine
10.29
5.93
TG-3, condensing steam turbine
20.21
17.50
TG-4, MSR-reformer steam turbine
10.41
9.61
TG-5, Reformer furnace steam turbine, HRSG-4
47.3
48.04
Biomass char combustor HRSG
9.88
9.88
Purge gas expander
19.31
6.30
Natural gas expander (to furnace fuel)
0
1.90
Total power available, kWh
119.1
100.5
13-16
-------
char combustor {that high temperature flue gas was otherwise required for biomass drying). The reconfigured
char combustor, shown in Figure 13-6, provides 9.8B kW of net electric power from that steam.
The ratio of natural gas used as reformer fuel and feedstock discussed above was an arbitrary
selection. This is not necessarily the optimum ratio; a search for an optimum may show a closer approach
to the Hynol yield (Option 2 is only 6 % less than Hynol). A critical variable in this comparison is the pressure
drops assumed for the various components that determine the pressure entering the synthesis-gas compressor
feeding the methanol converter system. A lower cumulative pressure drop reduces the electric power
requirement and therefore requires less purge for reformer fuel which, in turn, allows a higher methanol yield
to be obtained. Pressure drops could be reduced by heat exchanger designs that trade size and cost for
energy consumption. The optimum system is a function of the relative current costs of fuel energy and capital
as well as the methanol yield. A lower pressure assumed for the methanol converter may also enhance
methanol yield by requiring less electric power for syngas compression.
A comparison between tie optimized Hynol process and the modified BCL system discussed above
is summarized in Table 13-9. The comparisons show a higher performance by Hynol in terms of thermal
efficiency as well as methanol yield, although the advantage relative to the use of biomass and natural gas in
separate BCL and steam reforming processes is seen to be less when compared to the use of those
feedstocks in a single, modified BCL system. The overall C02 emission reduction by Hynol is about 10%
greater than the modified BCL system, which in turn is 10 % greater than the BCL plus steam reforming option.
The difference between energy ratio and thermal efficiency is, in both process, due to the subtraction of power
obtained from natural gas expansion for the thermal efficiency calculation; otherwise, the two parameters would
be equal because no power is imported.
The last two rows of Table 13-9 relate to cost. A cost analysis was not made on the modified BCL
system because it is not yet optimized, but an index of relative capital cost is shown by the ratios of the sum
of throughputs for the main equipment items. This comparison indicates a slightly lower capital cost may be
possible for the modified BCL system relative to Hynol, and the cost per tonne of C02 emission reduction could
also be about 5% lower.
A major advantage of the modified BCL system is that it does not require development or
demonstration-each of the unit operations is conventional technology or, in the case of the biomass gasifier,
has been demonstrated. This Is not true of Hynol, and a major barrier to the Hynol process is the fact that it
has not been demonstrated as technically feasible with regard to the operation of the gasifier, the high
temperature heat exchanger, the avoidance of tar formation, or the capability to control its operation as a fully
integrated system using gasifier feed recycled from the methanol converter. The BCL system could therefore
be operational in a very short time, whereas Hynol would require a long period of expensive development and
demonstration. It is also clear that the use of natural gas as co-feedstock in a single process is not confined
to the BCL-type gasifier or the Hynol gasifier: any of the other demonstrated biomass gasifiers could be used
in combination with a conventional steam reformer and a methanol converter to reduce the cost of methanol
production and improve the overall COz emission-reduction potential.
Finally, a system using natural gas as co-feedstock with a BCL gasifier or any other gasifier type can
be designed for a broad range of natural gas to biomass ratios, whereas the Hynot process has-because of
its recycle stream and gasifier energy balance requirement~a relatively fixed ratio of natural gas to biomass.
A BCL gasifier could therefore be constructed to operate with a lower ratio of natural gas to biomass; the
minimum ratio, yet to be determined, would be determined by the amount of natural gas necessary to allow the
elimination of the shift reactor and avoid electric power import.
13-17
-------
C02 1.092 mols
0? 0.1953
N2 4.898
H,0 0.5808
6.7661
131°C
Steam
310° C
Air 6.2 mols
332°C
2.2 kW
0.345 kW
694°C
10.22 kW
527°C
798°C
0.05 atm
Sand
Char
combustor
Sand
Char
Figure 13-6. Modification of BCL char combustor for steam recovery.
13-18
-------
TABLE 13-9. COMPARISON OF PERFORMANCE PARAMETERS FOR MODIFIED
BCL PROCESS AND OPTIMIZED HYNOL PROCESS USING THE SAME AMOUNT
OF BIOMASS AND NATURAL GAS FEEDSTOCKS
Basis: 100 kg biomass feed, 87% assumed carbon conversion
Process parameter
Methanol process
Hynol
Modified BCL
Natural gas used as process feed, mols
3.79
4.75
Natural gas used as reformer fuel, mols
2.486
1.526
C02 emission from natural gas, kg
277.5
277.5
Methanol produced, mols
7.202
6.781
Gasoline displaced, gallons
94.4
88.9
Tailpipe C02 reduction per vehicle, %
67.5
65.3
Energy ratio
0.69
0.65
Thermal efficiency, %
68.4
64.7
Total fuel-cycle C02 emission reduction, kg
724
660
Ratio, Hynol fuel-cycle emissions reduction/
BCL + steam reforming reduction
1.202
1.099
Ratio, sum of gasifier, reformer, and
methanol converter throughputs/methano!
product
20.9
17.4
Sum above per tonne of fuel cycle C02
emission reduction
208
197
13-19
-------
SECTION 14
REFERENCES
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Blok, K.; Turkenburg, W.C.; Hendricks, C.A.; and Steinberg, M. "Proceedings of the First International
Conference on Carbon Dioxide Removal," Pergamon Press, ISSN 0196-8904 (1992).
Buckingham, P.A.; Cobb, D.D.; Leavitt, A.A.; and Snyder, W.G. "Coal-to-Methanol: An Engineering
Evaluation of Texaco Gasification and ICI Methanol-Synthesis Route," Fluor Constructors and
Engineers, Inc. prepared for the Electric Power Research Institute, EPRI report No. AP-1962, Palo Alto,
CA (1981).
DeLuchi, M.A.; Sperling, D.; and Johnson, R.A. "A Comparative Analysis of Future Transportation Fuels,"
Research Report VCB-ITS-RR-87-13, Institute of Transportation Studies, University of California,
Berkeley, CA (October 1987).
DeLuchi, M.A. "Emissions of Greenhouse Gases from the Use of Transportation Fuels and Electricity,"
Center for Transportation Research, Argonne National Laboratory, Argonne, IL, Report ANUESD/TM-22
Vol. 1, p.21 (November 1993).
Directed Technologies, Hydrogen & Fuel Cell Letter, XII, 1, p.5 (January 1997).
Doctor, R.D.; Molburg, J.C.; and Thimmapuram, P.R. "KRW Oxygen-Blown Gasification Combined Cycle:
Carbon Dioxide Recovery, Transport, and Disposal," Argonne National Laboratory, Argonne, IL, Report
ANL/ESD-34 (August 1996).
Dong, V.; and Cole, E. "Evaluation of Biomass Reactivity in Hydrogasification for the Hynol Process," EPA-
600/R-96-071 (NTIS PB96-187638) (June 1996).
Dong, Y.; and Karwowski, J. "Methane-Steam Reaction over Nickel Catalysts in the Hynol Process," EPA-
600/R-97/093 (NTIS PB98-100480) (September 1997).
Espino, R.L.; and Robbins, J.L. "Fuel and Fuel Reforming Options for Fuel Cell Vehicles," presented at 30th
International Symposium on Automotive Technology & Automation, Florence, Italy (June 1997).
Feldman, H.F.; Paisley, M.A.; Appelbaum, H.R.; and Taylor, D.R. "Conversion of Forest Residues to
Methane-Rich Gas in a High-Throughput Gasifier," Report No. PNL-6570, UC-245, DOE Pacific
Northwest Laboratory (May 1988).
Garg, M.; Piskorz, J.; Scott, D.; and Radlein, D. 'The Hydrogasification of Wood," Ind. Eng. Chem. Res.,
Vol. 27, No. 2 pp. 256-264 (1988).
14-1
-------
Gordon, D. "Sustainable Transportation: What Do We Mean and How Do We Get There?," in
Transportation and Energy: Strategies for a Sustainable Transportation System, D. Sperling and S.A.
Shaheen, Eds. American Council for an Energy-Efficient Economy, Berkeley, CA (1995).
Graham, R.L "An Analysis of the Potential Land Base for Energy Crops in the Conterminous United
States," Biomass and Bioenergy, 6 3, pp. 175-189(1994).
Graham, R.L; Lichtenberg, E.; Roningen, V.O.; Shapouri, H.; and Walsh, M.E. "The Economics of Biomass
Production in the United States," Proceedings: Second Biomass Conference of the Americas, pp. 1314-
1323(1995).
Herzog, H.J. "Carbon Dioxide Removal," Proceedings of the Third International Conference on Carbon
Dioxide Removal, Cambridge, MA, Elsevier Science, ISBN 0080428401 (1997).
Intergovernmental Panel on Climate Change (IPCC) "Climate Change 1995, Impacts, Adaptation and
Mitigation of Climate Change: Scientific-Technical Analyses," Cambridge University Press, ISBN 0-521-
56437-9 (1996).
Katofsky, R.E. "The Production of Fluid Fuels from Biomass," Report No. PU/CEES 279, Center for Energy
and Environmental Studies, Princeton University (June 1993).
Kelly, H.; and Williams, R.H. "Fuel Cells and the Future of the US Automobile," Office of Technology
Assessment, Washington, D.C., 1992.
Liinanki, L.; Horvath, A.; Lehtovaara, A.; and Lindgren, G. 'The Development of a Biomass Based
Simplified IGCC Process," presented at the 13th EPRI Conference on Gasification Power Plants, San
Francisco, CA (October 1994).
Marrison, C.I.; and Larson, E.D. "Cost Versus Scale for Advanced Plantation-Based Biomass Energy
Systems in the U.S.A. and Brazil," Proceedings: Second Biomass Conference of the Americas, pp.
1272-1290(1995).
Michaelis, L. "Alternative Fuels and Greenhouse Gas Emission Policy," in Transportation and Energy:
Strategies for a Sustainable Transportation System, D. Sperling and S.A. Shaheen Eds., American
Council for an Energy-Efficient Economy, Berkeley, CA (1995).
Sandelli, G.J. "Demonstration of Fuel Cells to Recover Energy from Landfill Gas," EPA-600/R-92-007
(NTIS PB92-137520) (January 1992).
Steinberg, M.; and Dong, Y. "Process and Apparatus for the Production of Methanol from Condensed
Carbonaceous Material," U.S. patent No. 5,344,848 (September 6, 1994a).
Steinberg, M.; and Dong, Y. "Hynol--An Economical Process for Methanol Production from Biomass and
Natural Gas with Reduced C02 Emissions," Proceedings: 10th World Hydrogen Energy Conference,
Vol. I, pp. 495-504 (1994b).
Thomas, C.E.; James, B.D.; Baum, G.N.; Lomax, F.D.; and Kuhn, I.F. "Comparison of Onboard Hydrogen
Storage for Fuel Cell Vehicles," Directed Technologies, Inc. Report for subcontract 47-2-R31148, Task
4.2 to Ford Motor Co. (May 1996).
Thorneloe, S. U.S. EPA, Personal communication (November 1992).
14-2
-------
Unnasch, S. "Hynol Process Engineering: Process Configuration, Site Plan, and Equipment Design," EPA-
600/R-96-006 {NTIS PB96-167549) (February 1996).
U.S. Environmental Protection Agency "Reducing Risk: Setting Priorities and Strategies for Environmental
Protection," Science Advisory Board, SAB-EC-90-021 (1990).
Williams, R.H.; Larson, E.D.; Katofsky, R.E.; and Chen, J. "Methanol and Hydrogen from Biomass for
Transportation," presented at Bioresources '94, Biomass Resources: A Means to Sustainable
Development, Bangalore, India (1994).
14-3
-------
APPENDIX A
MODIFICATION OF THE BCL PROCESS PROPOSED BY KATOFSKY
Ryan Katofsky (1993) published a report that evaluated five processes for production of methanol or
hydrogen from biomass. Unlike Hynol, none of those processes utilize natural gas as co-feedstock-either
methanol or hydrogen is produced from biomass only. The primary difference between the process options
considered by Katofsky was the gasifier type. Of the five processes, the best overall performance and lowest
cost was determined to be the option utilizing the BCL gasifier. For that reason, the BCL process as outlined
by Katofsky is adapted in this report as a basis for comparison with the Hynol process. Section 4 of this report
discusses a modified version of Katofsky's BCL process that contains modifications necessary to properly
compare BCL with systems that use natural gas. The purpose of this Appendix is to discuss the reasons for
those modifications which concern apparent deficiencies in the energy balances and biomass requirements.
The basic data for the BCL process for production of methanol from biomass, as given on page 279
of Katofsky's 1993 report, do not identify the energy requirements of the individual process steps and do not
indicate the amount of biomass required for the biomass combustor that is part of the gasification system. The
amount of dry biomass (containing 10 wt% moisture) ted to the process is given as 1673 tonnes/day and the
methanol yield is given as 1229 kg mols/hr. It is not specified whether the 1673 tonnes of biomass includes
the biomass fed to the combustor or any additional biomass needed to generate electric power not produced
within the process. The process gases entering the system from the gasifier total 4005 kg mols containing
30.54 mol% H20,14.4% H2,32.03% CO, 7.72% CO^ 10.91% CH„, 3.6% C2H4, 0.5% C2H6, and 0.3% N2. Figure
A-1a, b is a flow sheet calculated for this report using Aspen Plus that includes the energy requirements for
each process step, excluding the gasifier, based on the process conditions specified by Katofsky.
Figure A-2 is an energy balance on the reformer unit. It assumes the same stream flows shown in the
original data of Katofsky (1993). Given the amounts of fuel gas and purge gas extracted from the process,
specified by those data as fuel for the reformer combustor, the indicated heat duty is only 80.9 % of that
required by the reformer in Figure A-1a (1384 kcal/sec, including 10124 kcal for the reformer and 3717 kcal
for the process gas preheater) for that part of the system. Katofsky's data do not include the duty of the
process gas preheater between streams 7 and 8. This is one of two problems that require correction in the
Katofsky analysis. The other problem concerns the energy balance on the gasifier/combustor unit.
Figure A-3 is a material balance on the gasifier, based on Katofsky's data for 1 hour of operation,
given the composition of the biomass (51.06 wt% C, 41.66%, 0.0625% H) and the carbon conversion (88%)
that he assumed. The carbon and oxygen balances shown in Figure A-3 indicate that Katofsky's (1993) flow
sheet does not include the biomass used as extra fuel for the gasifier combustor and that the 4005 mols of
gasification products fed to the process, as given by Katofsky for a 1673 tonnes/day plant, is based only on
the biomass fed to the gasifier, not the process as a whole.
Figure A-4 is a complete material and energy balance on the BCL gasifier/combustor unit as calculated
for this report. It is based on 100 kg of dry biomass fed to that system (smaller units are used to reduce the
size of the numbers displayed). These results show that, for every 100 kg of dry biomass fed to the process,
15.52 kg is required by the combustor, in addition to the unreacted char from the gasifier, to satisfy the energy
requirements of the gasifier/combustor system.
A-1
-------
1804 kW
31-
-1124
kcal/sec
2855 kW
HzO
kcal/sec
Quench
Reformer
+10124
kcal/sec
HRSG-2
-7690
kcal/sec
HRSG-1
-5995
kcal/sec
40
Shift
0 kcal/sec
Gasifier
Combustor
Biomass
Biomass
Figure A-1a. BCL process flow sheet as established for this report (continued).
A-2
-------
Methanol
synthesis
Okcal/sec
n—
1191 kW
-206
kcat/sec
1776
kcal/sec
Condenser
-1170
kcal/sec
CHaOH
3011 kW
2720 kW
-1417
kcal/sec
Distillation
-651
kcal/sec
H20
Selexol
3193 kW
-8416
kcal/sec
HjO + CO.
Figure A-1b. BCL process flow sheet as established for this report (continuation of Figure A-la).
A-3
-------
Process gas from gasifier
577 K, 15.5 Bar
HjO
H,
CO
CO,
CH„
C2H4
CjHj
N,
2057 kgmols
542
1204
290
410
135
19
11
co}
263 kgmols
H-0
556
Ns
1777
02
60.4
2656
0
4668
Fuel gas from gasifier
355 K, 1.01 Bar
765
kcat/sec
HjO
Hj
CO
C02
CH,
c?h4
c?h6
n,
135 kgmols
36
79
19
27
9
1
1
307
Steam
458 kgmols
523 K, 15.5 Bar
479
kcal/sec
+3014
kcal/sec
850 K
14.8 atm
/
391 K
\
526 K
608 K
1089 K
+8194
kcal/sec
2239 K
1.5 Bar
342 K 786 kW
553 K
473 K
Air
2235 kgmols
298 K, 1 Bar
. Purge gas from
methanol synthesis
300 K, 96.78 Bar
H20 0 kgmols
H, 116
CO 11
CO, 5
CH« 101
N2 10
CH.OH 1
244
J
Reformer exit
1140 K, 14.5 Bar
HjO
h2
CO
C02
CH4
n2
1483 kgmols
2468
1355
731
126
11
6174
Figure A-2. Energy balance on BCL reformer based on Katofsky flowsheet data •
A-4
-------
Process gas to reformer
1200 K, 1.01 Bar
Dry biomass 62737 kg
Biomass moisture (10%)
6970 kg
Gasifier
kamols
Mgl %
h2o
1223
30.54
h2
577
14.40
CO
1283
32.03
C02
309
7.72
CH4
437
10.91
c2h4
144
3.60
C2H6
20
0.50
n2
12
0.30
4005
100.00
Steam 1098 mols
Material balance on gasifier:
Given: 1673 TPD biomass fed with 10% moisture, 88% carbon conversion
Basis: one hour
Dry biomass = 1673(0.9)1000/24 = 62737 kg
Biomass moisture = 1673(0.1)1000/24 = 6970 kg
Carbon out = 1283 + 309 + 437 + 2(144) + 2(20) = 2357 kg atoms
Carbon in = 62737 (0.5106)0.88/12 = 2357 kg atoms
Oxygen out = 1223 + 1283 + 2(309) = 3124 kg mols
Oxygen in - 62737(0.4136)/16 + (6970/18) + 1098 = 3107 kg mols
Figure A-3. BCL gasifier material balance.
A-5
-------
Gasifier Energy Balance
Sensible heat in gasifier effluent at 927°C:
[1.647(9.15) + 0.777(7.1) + 1.728(7.55) + 0.416(11.75) + 0.589(14.2) + 0.194(19) +
0.027(21) + 0.016(7.4)](927-25) = 46233 kcal
Specific heat of carbon (Perry) = 0.314 cal/g °C
Sensible heat out with char: 0.4314(12)(927-25)0.314 = 1466 kcal
Sensible heat in with steam at 157°C: 1.478(8.1)(157-25) = 1580 kcal
Heat of Reaction:
Product heat of formation:
1.647(-57800) + 1,728(-26416) + 0.416(-94080) + 0.589(-17890) + 0.194(-20240) +
0.027(+12500) = -194107 kcal
Feed heat of formation: (biomass heat of formation = -1336.3 kcal/kg, MAF)
1.478(-57800) + (9.387/18)(-68300) + [84.48 - (84.48 x 0.00792)](-1336.3) = -233043 kcal
Heat of reaction: -194107 - (-233043) = +38936 kcal
Energy balance:
Sensible heat in with dry biomass + sensible heat in with biomass moisture + sensible
heat in with steam - heat of reaction + heat added with sand = sensible heat out with
process gas + sensible heat out with char.
4275 + 704 + 1580 - 38936 + heat added with sand = 46233 + 1466
Heat added with sand = 80076 kcal
Combustor Energy Balance
Char fed from gasifier at 927 °C = 0.4314 kgmol, sensible heat = 1466 kcal
Dry biomass = 15.52 kg
Biomass carbon - 15.52(0.5106)/12 = 0.660 kgatom
Biomass oxygen = 15.52(0.4166)/32 = 0.202 kgmol
Biomass hydrogen = 15.52(0.0625)/2 = 0.485 kgmol
Theoretical oxygen requirement:
0.4314 + 0.660 + (0.485/2) - 0.202 = 1.132 kgmols
A-6
-------
for 15% excess air: 1.132(1.15) = 1.30 kgmols
N2 = 1.30(79/21) = 4.890 kgmols
Air = 4.89 + 1.30 = 6.19 kgmols
Sensible heat in with air; 6.19(7.38)(727-25) = 32069 kcal
Flue gas sensible heat:
[1.092(12.1) + 0.1953(8.0) + 4.898(7.58) + 0.5808(9.4)](1114-25) = 62467 kca!
Heat of reaction:
1.092(-94050) + 0.5808(-57800) - 15.52(-1336.3) - (1.724/18)(-68300) = -108992 kcal
Energy balance:
Sensible heat in with air + sensible heat in with char + sensible heat in with biomass -
heat of reaction = heat added to sand + sensible heat out with flue gas
1466 + 32069 + 0 - (-108992) = 80076 + 62467
142527 = 142543
A-7
-------
APPENDIX B
ENERGY REQUIREMENTS FOR DRYING BIOMASS
Case 1: BCL Process Using Flue Gas and Heated Air as Heat Source (Figure 4-7)
Assume initial biomass moisture content = 48.4 wt%, dry to 10 wt%;
air discharged at 80 °C, flue gas discharged at 120 °C
Basis: 100 kg MAF biomass, 83.4 kg biomass moisture; dry to 11.1 kg biomass moisture;
H20 evaporated = 82.7 kg
Specific heat of wood = 0.57 cal/g °C
Heat 100 kg wood from 25 to 100 °C: 100 x 0.57(100-25) = 4275 kcal
Heat 93.8 kg H20 from 25 to 100 °C: 93.8 x 1.0(100-25) = 7035 kcal
Evaporate 82.7 kg H20 at 100 °C: 82.7(2.2)970.3(0.252) = 44487 kcal
Heat 82.7 kg H20 from 100 to 120 °C: (82.7/18)8.05(120-100) = 740 kcal
Total drier heat input required = 4275 + 7035 + 44487 + 740 = 56537 kcal
= 15,705 cal/sec
Case 2: Hvnol Process Using Low Pressure Steam as Heat Source (Figure 5-7^
Assume initial biomass moisture content = 48.4 wt%, dry to 10 wt%; moisture evaporated = 82.7 kg
Heat 100 kg wood from 25 to 105 °C: 100 x 0.57(105-25) = 4560 kcal
Heat 93.8 kg H20 from 25 to 100 °C: 93.8 x 1.0(100-25) = 7035 kcal
Evaporate 82.7 kg H20 at 100 °C: 82.7(2.2)970.3(0.252) = 44487 kcal
Heat 82.7 kg H,0 from 100 to 105 °C: (82.7/18)8.05(105-100) = 185 kcal
Heat added from steam recyle blower: 1.22 kWh x 860.5 kcal/kWh = 1050 kcal
Total drier heat input required = 4560 + 7035 + 44487 + 185 -1050 = 55217 kcal
= 15,337 cal/sec
B-1
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TECHNICAL REPORT DATA
(Please read Instructions on the reverse before comp!
1. REPORT NO. 2.
EPA-600/R-97-153
4. TITLE AND SUBTITLE
Hynol Process Evaluation
5. RCCOBT DATE
December 1997
6. PERFORMING ORGANIZATION CODE
7. AUTHOR(S)
Robert H. Borgwardt
8. PERFORMING ORGANIZATION REPORT NO.
9. PERFORMING ORGANIZATION NAME ANO ADORESS
See Block 12
10. PROGRAM ELEMENT NO.
11. CONTRACT/GRANT NO.
NA (Inhouse)
12. SPONSORING AGENCY NAME ANO ADDRESS
EPA, Office of Research and Development
Air Pollution Prevention and Control Division
Research Triangle Park, NC 27711
13. TYPE OF REPORT AND PERIOD COVERED
Final; 6/95 - 7/97
14. SPONSORING AGENCY CODE
EPA/600/13
15.supplementary notes APPCD project officer is Robert H. Borgwardt, Mail Drop S3,
919/541-2336.
i6. ABSTRACTj^e report examines process alternatives for the optimal use of natural gas
and biomass for production of fuel-cell vehicle (FCV) fuel, emphasizing maximum
displacement of petroleum and maximum reduction of overall fuel-cycle carbon di-
oxide (C02) emissions at least cost. Three routes are evaluated: (1) production of
methanol from biomass and from natural gas by independent processes, (2) produc-
tion of methanol or hydrogen by hydrogasification of biomass using natural gas as co-
feedstock supplemented with with, and without, the use of carbonaceous municipal
wastes as co-feedstocks, and (3) production of methanol or hydrogen by addition of
natural gas to a biomass-to-methanol process originally designed for biomass only.
The results show that the combined use of natural gas and biomass in a single pro-
cess can reduce net fuel-cycle C02 emissions by 20% relative to separate systems
and reduce the cost of fuel production to a range competitive with the current cost of
gasoline. A plant optimized for efficiency and size, with 25% of the feedstock energy
consisting of biomass, should be able to produce methanol at a cost of $0.42/gal
($6.09/GJ), or hydrogen at $5.98/GJ. This technology represents a cost-
effective use of biomass as a source of fuel energy.
17. KEY WOROS AND DOCUMENT ANALYSIS
a. DESCRIPTORS
b.lDENTIFIERS/OPEN ENDED TERMS
c. COSATI Field/Gioup
Pollution Natural Gas
Fuel Cells Biomass
Motor Vehicles Carbinols
Greenhouse Effect Wastes
Carbon Dioxide
Hydrogen
Pollution Control
Stationary Sources
Hynol Process
Methanol
Hydrogasification
13 B 21D
10B 08A.08C
13F 07C
14A 14 G
07B
18. DISTRIBUTION STATEMENT
Release to Public
19. SECURITY CLASS (This Report)
Unclassified
21. NO. OF PAGES
206
20. SECURITY CLASS (This page)
Unclassified
22. PRICE
EPA Form 2220-1 (9-731
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