&EPA
United States Industrial Environmental Research EPA-600/7-79-178i
Environmental Protection Laboratory November 1979
Agency Research Triangle Park NC 27711
Technology Assessment
Report for Industrial
Boiler Applications:
Flue Gas Desulfurization
Interagency
Energy/Environment
R&D Program Report
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EPA-600/7-79-178J
November 1979
Technology Assessment Report
for Industrial Boiler Applications:
Flue Gas Desulfurization
by
J.C. Dickerman and K.L. Johnson
Radian Corporation
P.O. Box 8650
Durham, North Carolina 27707
Contract No. 68-02-2608
Task No. 47
Program Element No. EHE624
EPA Project Officer: John E. Williams
Industrial Environmental Research Laboratory
Office of Environmental Engineering and Technology
Research Triangle Park, NC 27711
Prepared for
U.S. ENVIRONMENTAL PROTECTION AGENCY
Office of Research and Development
Washington, DC 20460
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PREFACE
The 1977 Amendments to the Clean Air Act required that emission standards
be developed for fossil-fuel-fired steam generators. Accordingly, the U.S.
Environmental Protection Agency (EPA) recently promulgated revisions to the 1971
new source performance standard (NSPS) for electric utility steam generating
units. Further, EPA has undertaken a study of industrial boilers with the
intent of proposing a NSPS for this category of sources. The study is being
directed by EPA's Office of Air Quality Planning and Standards, and technical
support is being provided by EPA's Office of Research and Development; As part
of this support, the Industrial Environmental Research Laboratory at Research
Triangle Park, N.C., prepared a series of technology assessment reports to aid
in determining the technological basis for the NSPS for industrial boilers.
This report is part of that series. The complete report series is listed below:
Title
The Population and Characteristics of Industrial/
Commercial Boilers
Technology Assessment Report for Industrial Boiler
Applications: Oil Cleaning
Technology Assessment Report for Industrial Boiler
Applications: Coal Cleaning and Low Sulfur Coal
Technology Assessment Report for Industrial Boiler
Applications: Synthetic Fuels
Technology Assessment Report for Industrial Boiler
Applications: Fluidized-Bed Combustion
Technology Assessment Report for Industrial Boiler
Applications: NO Combustion Modification
Technology Assessment Report for Industrial Boiler
Applications: NO Flue Gas Treatment
Technology Assessment Report for Industrial Boiler
Applications: Particulate Collection
Technology Assessment Report for Industrial Boiler
Applications: Flue Gas Desulfurization
Report No.
EPA-600/7-79-178a
EPA-600/7-79-178b
EPA-600/7-79-178c
EPA-600/7-79-178d
EPA-600/7-79-178e
EPA-600/7-79-178f
EPA-600/7-79-178g
EPA-600/7-79-178h
EPA-600/7-79-178i
These reports will be integrated along with other information in the document,
"Industrial Boilers - Background Information for Proposed Standards," which will
be issued by the Office of Air Quality Planning and Standards.
11
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CONTENTS
FIGURES
TABLES Xii
1. EXECUTIVE SUMMARY 1-1
1.1 Introduction 1-1
1.1.1 Background and Obj active 1-1
1.1.2 Approach 1-2
1.1.3 Conclusion . 1-2
1.1.4 Report Organization . 1-4
1.2 Emission Reduciton for Coal-Fired Industrial Boilers 1-5
1.2.1 Selection of Candidates for Best Control System 1-5
1.2.2 Cost Analysis of Candidate Systems 1-9
1.2.3 Energy Impacts of Candidate Control Systems 1-17
1.2.4 Environmental Impacts of Candidate Control Systems .. 1-20
2. EMISSION CONTROL TECHNIQUES 2-1
2.1 Principles of Control 2-1
2.2 Controls for Coal-Fired Boilers 2-6
2.2.1 Lime/Limestone Wet Scrubbing 2-7
2.2.1.1 System Description 2-7
2.2.2.2 System Performance 2-66
2.2.2 Double Alkali Process 2-79
2.2.2.1 System Description 2-80
2.2.2.2 System Performance 2-92
2.2.3 Wellman-Lord Sulfite Scrubbing Process 2-111
2.2.3.1 System Description 2-111
2.2.3.2 System Performance 2-128
111
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CONTENTS (Continued)
2.2.4 Magnesia Slurry Absorption Process „ 2-131
2.2.4.1 System Description 2-131
2.2.4.2 System Performance 2-142
2.2.5 Sodium Scrubbing 2-147
2.2.5.1 System Description 2-147
2.2.5.2 System Performance 2-159
2.2.6 Processes Under Development 2-160
2.2.6.1 Spray Drying 2-161
2.2.6.2 Citrate Buffered Absorption 2-171
2.2.6.3 Bergbau-Forschung/Foster Wheeler Process .. 2-177.
2.2.6.4 Atomics International Aqueous Carbonate
Process 2-186
2.2.6.5 Shell Flue Gas Desulfurization Process .... 2-195
2.2.6.6 Chiyoda Thoroughbred 121 Process 2-201
2.3 Controls for Oil-Fired Boilers 2-204
References 2-20.7
3. CANDIDATES FOR BEST SYSTEMS OF S02 EMISSION REDUCTION 3-1
3.1 Criteria for Selection of Best SOa Control Systems 3-3
3.2 Selection of Best Control Systems 3-7
3.2.1 Development Status 3-8
3.2.2 Performance 3-8
3.2.3 Applicability 3-10
3.2.4 Economic Considerations 3-10
3.2.5 Energy Considerations 3-15
3.2.6 Environmental Considerations 3-17
References . 3-20
4. COST ANALYSIS OF CANDIDATES FOR BEST EMISSION CONTROL SYSTEM .... 4-1
4.1 Contributors to Control Costs and Cost Bases 4-6
4.2 Control Costs for Coal-Fired Boilers 4-11
4.2.1 Interpretation of Results 4-11
4.2.2 Example Calculation 4-36
4.3 Costs to Control Oil-Fired Boilers 4-43
References 4-46
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CONTENTS (Continued)
5. ENERGY IMPACT OF CANDIDATES FOR BEST EMISSION CONTROL SYSTEMS ... 5-1
5.1 Introduction 5-1
5.2 Energy Impact of Controls for Coal-Fired Boilers 5-4
5.2.1 Sample Calculations 5-17
5.2.1.1 Calculate Raw Material Handling and
Preparation 5-20
5.2.1.2 Calculate Liquid Pumping Energy 5-20
5.2.1.3 Calculate Fan Energy 5-21
5.2.1.4 Calculate Waste Disposal Requirements 5-22
5.2.1.5 Calculate Utilities and Services 5-24
5.2.1.6 Calculation Summary 5-24
5.2.2 Methods to Reduce Energy Consumption 5-24
5.3 Impact of Controls for Oil-Fired Boilers 5-28
References 5-32
6. ENVIRONMENTAL IMPACT OF CANDIDATES FOR BEST EMISSION CONTROL
SYSTEMS 6-1
6.1 Introduction 6-1
6.2 Environmental Impact of Controls for Coal-Fired Boilers ... 6-1
6.2.1 Air Pollution 6-1
6.2.2 Water Pollution 6-13
6.2.3 Solid Waste 6-17
6.2.4 Environmental Impact on Modified Facilities 6-35
6.3 Impact of Controls for Oil-Fired Boilers 6-35
References 6-36
7. EMISSION SOURCE TEST DATA 7-1
7.1 Introduction 7-1
7.2 Emission Source Data for Coal-Fired Boilers 7-4
7.3 Data Presentation 7-6
7.3.1 Test Methods 7-10
7.4 Emission Source Test Data for Oil-Fired boilers 7-14
References 7-18
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CONTENTS (Continued)
ADDITIONAL FGD TOPICS 8-1
8.1 Partial Scrubbing , 8-1
8.1.1 Control Costs 8-1
8.1.2 Energy Requirements 8-8
8.1.3 Environmental Impacts 8-8
8.2 Limestone Scrubbing of Flue Gas from Five Percent
Sulfur Coal 8-17
8.2.1 Control Costs 8-18
8.2.2 Energy Requirements 8-18
8.2.3 Environmental Impacts 8-22
APPENDIX A - MATERIAL BALANCE CALCULATIONS A-l
APPENDIX B - COST SUMMARY TABLES B-l
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FIGURES
Number Page
1.2-1 FGD capital costs versus unit size (3.5% S coal, 90%'
removal) 1-11
1.2-2 FGD annualized costs versus unit size (3.5% S coal,
90% removal) 1-12
1.2-3 FGD capital costs versus unit size (0.6% S coal,
75% removal) 1-13
1.2-4 FGD annualized costs versus unit size (0.6% S coal,
75% removal) 1-14
1.2-5 Limestone process cost effectiveness 1-16
1.2-6 S02 emissions versus control level 1-20
1.2-7 Sludge production rates for the limestone FGD process ... 1-27
2.2.1-1 Process Flow Diagram Lime/Limestone Wet Scrubbing 2-9
2.2.1-2 Liquid-to-gas ratio and scrubber inlet pH versus pre-
dicted and measured S02 removal, spray tower with lime,
Shawnee plant 2-35
2.2.1-3 Liquid-to-gas ratio and scrubber gas velocity Versus pre-
dicted and measured S02 removal, TCA with lime, Shawnee
plant 2-36
2.2.1-4 L/G ratio and scrubber inlet pH versus predicted and
measured S02 removal - TCA with limestone - Shawnee
plant 2-37
2.2.1-5 L/G ratio versus percent S02 removal at various magnesium
ion concentrations TCA with limestone - Shawnee plant ... 2-38
2.2-.1-6 S02 removal versus L/G ratio, 170-MW horizontal module,
Mohave plant 2-39
2.2.1-7 Effect of liquid-to-gas ratio on S02 removal efficiency
with low sulfur coal at the Mohave power station 2-40
2.2.1-8 Gas velocity and slurry flow rate versus predicted and
measured S02 removal, spray tower with lime, Shawnee
plant 2-41
2.2.1-9 Scrubber inlet pH and liquid-to-gas ratio versus pre-
dicted and measured S02 removal, spray tower with lime,
Shawnee plant 2-44
2.2.1-10 Scrubber inlet pH and liquid-to-gas ratio versus pre-
dicted and measured S02 removal, TCA with lime, Shawnee
plant 2-45
Vll
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FIGURES (Continued)
2.2.1-11 Scrubber inlet pH versus percent S02 removal at various
magnesium ion concentrations TCA with limestone -
Shawnee plant 2-46
2.2.1-12 Scrubber inlet pH versus S02 removal for three L/G ratios -
TCA unit with limestone - Shawnee plant 2-47
2.2.1-13 Effect of inlet S02 concentration on S02 removal
efficiency for fixed design and operating conditions 2-49
2-2.1-14 Effect of liquid-to-gas ratio on S02 removal efficiency
with low sulfur coal at the Mohave power station 2-50
2.2.1-15 Dissolved alkalinity generated by addition of MgO 2-53
2.2.1-16 Effect of magnesium on S02 removal efficiency - TCA (no
spheres) with limestones 2-54
2.2.1-17 Effect of magnesium on S02 removal efficiency 2-52
2.2.1-18 170 MW horizontal S02 removal versus number of stages,
Mohave plant 2-56
2.2.1-19 S02 absorption efficiency for two scrubbers in series .... 2-57
2.2.1-20 R-C/Bahco Scrubber system flow diagram 2-68
2.2.1-21 S02 Removal efficiency as a function of lime stoichiometry 2-70
2.2.1-22 S02 removal efficiency as a function of limestone/S02 ....
stoichiometry and slurry pumping rate 2-71
2.2.1-23 The effect of scrubber pressure drop on particulate
emission rates 2-73
2.2.1-24 A comparison of lime and limestone slurry settling rates . 2-74
2.2.1-25 The effect of operating time and slurry feed concentration
on centrifuge cake density 2-75
2.2.2-1 Simplified Flow Diagram for Sodium Double Alkali Process . 2-81
2.2.2-2 Data collection points and normal operating conditions ... 2-96
2.2.2-3 S02 removal versus scrubber effluent pH for the Envirotech/
Gadsby pilot plant with a two-stage absorber 2-97
2.2.2-4 S02 removal versus L/G ratio for the Envirotech/Gadsby
pilot plant with a single stage polysphere absorber 2-98
2.2.2-5 CEA/ADL Dual Alkali Process - S02 removal as a function of
pH - high sulfur coal 2-100
2.2.2-6 CEA/ADL Dual Alkali Process - S02 removal as a function of
pH - low/medium sulfur coal 2-101
2.2.2-7 Effect of feed stoichiometry on removal efficiency in the
Venturi/2-tray tower absorber for the EPA-ADL double-
alkali pilot program 2-102
viii
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FIGURES (Continued)
2.2.2-8 Performance Treads for the Firestone Double Alkali
system [[[ 2-104
2.2.2-9 Performance Trends for the GM Parma Double Alkali
system ...................... . ............................. 2-110
2.2.3-1 Process flow diagram Wellman-Lord Process ........... . ..... 2-112
2.2.3-2 Inlet and outlet S02 concentrations during run no. 1 ...... 2-128
2.2.3-3 Inlet and outlet S02 concentrations during run no. 2 ...... 2-129
2.2.3-4 Inlet and outlet S02 concentrations during run no. 3 ...... 2-129
2.2.3-5 S02 removal efficiencies during 12-day test ..... ........ .. 2-130
2.2.4-1 Process Flow Diagram for the Magnesia Slurry Absorption
Process [[[ 2-133
2.2.4-2 Effect of inlet SO concentration and venturi pressure drop
on SO 2 removal for the Mystic venturi absorber ............ 2-143
2.2.4-3 Effect of pressure drop on S02 removal for the Mystic
venturi absorber ............. ... .......................... 2-144
2.2.4-4 The effect of pH on S02 scrubbing efficiency .............. 2-146
2.2.5-1 Simplified Flow Diagram Sodium Scrubbing System ........... 2-149
2.2.5-2 Effect of pressure drop on SOz removal efficiency - venturi
with sodium carbonate (10 MW size) ........................ 2-156
2.2.5-3 Effect of liquid-to-gas ratio oil S02 removal efficiency ... 2-157
2.2.5-4 Effect of liquid-to-gas ratio on S02 removal efficiency -
TCA (no spheres) with sodium carbonate (10 MW size) ....... 2-158
2.2.6-1 Simplified Flow Diagram for Spray Drying Process .......... 2-164
2.2.6-2 Simplified Flow Diagram for Citrate/Phosphate Process ..... 2-173
2.2.6-3 Simplified Flow Diagram for Bergbau-Forschung Process ..... 2-179
2.2.6-4 FW/BF Dry Adsorption Process Adsorber Module Detail ....... 2-182
2.2.6-5 Simplified Flow Diagram for Atomics International Aqueous
Carbonate Process ......................................... 2-187
2.2.6-6 Block Flow Diagram - Shell Flue Gas Desulf urization Process 2-196
2.2.6-7 Simplified Flow Diagram for Chiyoda Thoroughbred 121
Process .............. . ............................ . ........ 2-202
4.0-1 Comparison of calculated and reported total capital
investment for the limestone process ...................... 4-3
4.0-2 Comparison of calculated and reported total capital
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FIGURES (Continued)
Page
4.0-3 Comparison of calculated and reported total capital
investment for the sodium scrubbing process (excludes
wastewater treating costs) 4-5
4.2-1 FGD capital costs versus unit size (3.5% S coal,
90% removal) 4-18
4.2-2 FGD annualized costs versus unit size (3.5% S coal,
90% removal) - 4-19
4.2-3 FGD capital costs versus unit size (0.6% S coal, 75%
removal) 4-20
4.2-4 FGD annualized costs versus unit size (0.6% S coal, 75%
removal) 4-21
4.2-5 FGD capital costs versus coal sulfur content (58.6 MW, 90%
removal) 4-23
4.2-6 FGD annualized costs versus coal sulfur content (58.6 MW,
90% removal) 4-24
4.2-7 FGD capital costs versus S02 removal (58.6 MW, 3.5% S coal) 4-25
4.2-8 FGD annualized costs versus S02 removal (58.6 MW,
3.5% S coal) „ . 4-26
4.2-9 FGD capital costs versus S02 removal (44 MW , 0.6% S coal). 4-28
4.2-10 FGD annualized costs versus S02 removal (44 MW , 0,6% S
coal) 4-29
4.2-11 Limestone process cost effectiveness 4-30
4.2-12 Sodium throwaway process cost effectiveness 4-31
4.2-13 Double-alkali process cost effectiveness 4-32
4.2-14 Spray drying process cost effectiveness 4-33
4.2-15 Wellman-Lord process cost effectiveness 4-34
4.3-1 Comparison of oil- and coal-fired FGD costs (Limestone
process, 44 MW ) 4-45
5.2-1 Energy consumption versus coal sulfur content 5-12
5.2-2 Energy consumption versus boiler size: high sulfur
eastern coal 5-13
5.2-3 Energy consumption versus boiler size: low sulfur
western coal 5-14
5.2-4 Percent energy consumption versus boiler size: high
sulfur eastern coal 5-15
X
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FIGURES (Continued)
Page
5.2-5 Percent energy consumption versus boiler size: low
sulfur western coal 5-16
5. 2-6 Energy consumption versus SC-2 removal 5-18
5.2-7 Energy consumption versus SOa removal 5-19
6.2-1 SC>2 emissions versus control level for standard boilers
firing 3.5% sulfur eastern coal 6-9
6.2-2 S02 emissions versus control level for standard boilers
firing 2.3% sulfur coal 6-10
6.2-3 S02 emissions versus control level for standard boilers
firing 0.6% sulfur western coal 6-11
6.2-4 S02 emissions versus control level 6-12
6.2-5 Sludge production rates for the limestone FGD process 6-21
7.4-1 Performance of Kureka Chemical's sodium throwaway process.. 7-15
7.4-2 Performance of Wellman-Lord process . 7-16
8.1-1 Sodium partial scrubbing capital investment costs 8-6
8.1-2 Sodium partial scrubbing annual costs 8-7
S.l-3 Sodium partial scrubbing energy consumption 8-12
8.1-4 Sodium partial scrubbing wastewater production 8-16
8.2-1 Capital and annualized costs for limestone scrubbing versus
coal sulfur content 8-22
XI
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TABLES
Number Page
1.2-1 FGD System Summary 1-6
1.2-2 Flue Gas Desulfurization Screening Criteria 1-5
1. 2-3 Range of FGD System Energy Requirements 1-17
1.2-4 Percentage Energy Consumptions for Nonregenerable Processes. 1-17
1.2-5 Comparison of spray drying process energy requirements (44
MWt, 0.6% S Coal, 75% removal) 1-18
1.2-6 Percentage Energy Consumption for Wellman-Lord Process 1-19
1.2-7 Aqueous Emissions From Sodium Throwaway Processes 1-22
1.2-8 Solid Waste Production for the Limestone FGD Process 1-25
1.2-9 Solid Waste Volumes For Limestone FGD Process (Ash-free
Basis) 1-26
1.2-10 Solid Waste Impact for the Double-Alkali Process (Ash-free
Basis ) 1-28
1.2-11 Solid Wastes From Spray Drying (Total Fly Ash + Alkali
Salts) 1-30
1. 2-12 Solid Waste by Origin l-3l
2.1-1 SO2 Emissions for Various Fuel Types 2-2
2.1-2 FGD System Summary 2-3
2.2.1-1 Estimated Quantity of Flue Gas Desulfurization Wastes and
Ash from Selected Coals - 1000 MWg Plant 2-13
2.2.1-2 Estimated Quantity of Sludge From Industrial Boiler
Limestone FGD Systems 2-14
2.2.1-3 Summary of New and Retrofit FGD Systems for U.S. Utility
Industry by Process 2-17
2.2.1-4 Summary of Operational Lime/Limestone FGD Systems for U.S.
Utilities as of March 1978 2-18
2.2.1-5 Summary of Lime/Limestone Systems under Construction for
U.S. Utilities as of March 1978 2-21
2.2.1-6 Summary of Planned Lime/Limestone FGD Systems as of 03/78... 2-24
2.2.1-7 Summary of Committed Lime/Limestone Systems for U.S. Indus-
trial Boilers as of March 1978 2-26
Xll
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TABLES (Continued)
Page
2.2.1-8 Summary of Japanese Lime/Limestone Installations 2-28
2.2.1-9 Typical Bahco Operating Conditions 2-67
2.2.1-10 Summary of Downtime During the Bahco System Testing 2-76
2.2.2-1 Full Scale Double-Alkali Systems in the U.S 2-84
2.2.2-2 Summary of Significant Operating Full Scale Sodium/Calcium
Alkali Systems in Japan 2-85
2.2.2-3 Performance of Dual Alkali FGD Facilities in the U.S 2-93
2.2.2-4 Key Operating Parameters and Results for Intensive Testing
at Parma 2-95
2.2.2-5 CEA/ADL Dual Alkali Process Viability Parameters 2-106
2.2.3-1 Summary of Operating Wellman-Lord Systems in the U.S 2-119
2.2.3-2 Summary of Operating Wellman-Lord Systems in Japan 2-120
2.2.3-3 Summary of Wellman-Lord Systems Planned in the U.S 2-121
2.2.3-4 Distribution of Wellman-Lord Sulfur Dioxide Removal Plants.. 2-122
2.2.4-1 Operating and Planned Magnesia Scrubbing Units on U.S.
Power Plants as of August 1978 2-137
2.2.4-2 Operating Magnesia Scrubbing Units on Japanese Power Plants
as of August 1978 2-137
2.2.5-1 Summary of Sodium FGD Process on U.S. Coal-Fired Industrial
and Utility Boilers 2-148
2.2.5-2 Summary of FGD Processes on U.S. Coal-Fired Utility Boilers. 2-152
2.2.5-3 Summary of FGD Processes Applied to U.S. Industrial Boilers. 2-152
2.2,5-4 Performance Data for Operating Sodium Scrubbing Systems 2-153
2.2.6-1 Summary of Commercial Dry Scrubbing Applications 2-165
2.3-1 Summary of United States Oil-Fired Industrial Boiler
Installations 2-206
3.1-1 Flue Gas Desulfurization Screening Criteria 3-4
3.2-1 Overall Status of Development of Candidates for Best Systems
of S02 Reduction , 3-9
3.2-2 Performance of Candidates for Best Systems for S02 Reduction 3-10
3.2-3 Applicability of Candidate Systems to Industrial Boilers.... 3-13
3.2-4 Preliminary FGD System Cost Estimates 3-14
3.2-5 FGD System Energy Requirements 3-16
3.2-6 FGD Systern Environmental Impacts 3-18
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TABLES (Continued)
Page
4.1-1 Indirect Capital Cost Factors 4-8
4.1-2 Values Used for Annual Cost Items 4-10
4.2-1 Limestone Process Cost Summary 4-12
4.2-2 Sodium Throwaway Process Cost Summary 4-13
4.2-3 Double-Alkali Process Cost Summary 4-14
4.2-4 Spray Drying Process Cost Summary 4-15
4.2-5 Wellman-Lord Process Cost Summary 4-16
4.2-6 Typical Increase in Capital Costs with Various Retrofit
Requirements - • 4-36
4. 2-7 Process Operating Conditions 4-37
4.2-8 Raw Material Handling Capital Costs 4-38
4. 2-9 SO 2 Scrubbing Capital Costs 4-38
4.2-10 Fan Capital Costs 4-39
4.2-11 Wastewater Pumps Capital Costs 4-39
4.2-12 Installed Costs 4-39
4.2-13 Capital Investment Costs for FGD Processes ' 4-41
4.2-14 Annualized Costs for FGD Processes 4-42
4.3-1 Cost Comparison of Oil- and Coal-Fired Limestone FGD Costs.. 4-43
5.1-1 Range of FGD System Energy Requirements 5rl
5.2-1 Material and Energy Balance Assumptions and Design Bases.... 5-5
5.2-2 Energy Requirements for the Limestone FGD Process 5-6
5.2-3 Energy Requirements for the Sodium Throwaway FGD Process.... 5-7
5.2-4 Energy Requirements for the Double-Alkali Process 5-8
5.2-5 Spray Drying Energy Requirements 5-9
5.2-6 Energy Requirements for the Wellman-Lord Process 5-10
5.2-7 Percentage Energy Consumptions for Nonregenerable Processes. 5-25
5.2-8 Comparison of Spray Drying Process Energy Requirements 5-26
5.2-9 Percentage Energy Consumption for Wellman-Lord Process 5-27
5.3-1 Limestone Process Energy Requirements for Residual Oil
Application. 5-29
5.3-2 Wellman-Lord Process Energy Requirements for Residual Oil
Application.....,,.,.,,...,,.... „....,,....„.........„. 5-31
XIV
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TABLES (Continued)
Page
6.2-1 Air Pollution Impacts of S02 Control Techniques for Coal-
Fired Boilers 6-4
6.2-2 Air Pollution Impacts of S02 Control Techniques for Coal-
Fired Boilers 6-5
6.2-3 Air Pollution Impacts from "Best" S02 Control Techniques
for Coal-Fired Boilers 6-6
6.2-4 Air Pollution Impacts from "Best" S02 Control Techniques
for Coal-Fired Boilers 6-7
6.2-5 Air Pollution Impacts from "Best" S02 Control Techniques
for Coal-Fired Boilers 6-8
6. 2-6 Physical Properties of Waste Products 6-13
6. 2-7 Wellman-Lord Prescrubber Discharges . 6-14
6.2-8 Water Pollution Impacts for the Sodium Throwaway System.... 6-16
6.2-9 Solid Waste Impact for the Limestone FGD Process 6-19
6.2-10 Solid Waste Volumes for the Limestone FGD Process 6-20
6.2-11 Solid Waste Impact for the Double-Alkali Process (Ash-free
Basis) 6-23
6.2-12 Effect of Chemical Fixation on Bulk Density 6-22
6.2-13 Solid Wastes from Spray Drying (Total Fly Ash + Alkali
Salts) 6-25
6 . 2-14 Solid Wastes by Origin. . . . , 6-26
6.2-15 Coefficients of Permeability for FGD Sludges.. 6-29
6. 2-16 Permeabilities of FGD Sludges . 6-29
6.2-17 Coefficients of Permeability for FGD Sludges Treated with
Fly Ash and/or Cement . 6-31
6.2-18 Effect of Sludge Treatment on Permeability 6-31
6.2-19 Equilibrium Concentrations of Trace Elements in FGD Sludge
Leachate (in ppm) 6-32
6.2-20 Levels of Chemical Species in FGD Sludge Liquors and
Elutriates 6-33
6.2-21 Solid Waste Production for the Wellman-Lord Process 6-34
7.1-1 Source Monitoring Data from Industrial Boiler FGD Systems.. 7-2
7.1-2 Continuous Monitoring Data from Utility Boiler FGD Systems. 7-3
7 . 3-1 SOz Removal Efficiency Data 7-7
7.3-2 Lime Test Data Summary - Rickenbacker Air Force Base -
R-C/Bahco Lime Scrubbing System 7-8
XV
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TABLES (Continued)
7.3-3 Limestone Test Data Summary - Rickenbacker Air Force Base - 7-9
R-C/Bahco Scrubbing System
7.3-4 Coal-Fired Emission Source Data Firestone Tire and Rubber - 7-11
FMC Double-Alkali FGD System
7.4-1 Oil-Fired Emission Source Data Firestone Tire and Rubber -
FMC Double-Alkali FGD System 7-14
8.1-1 Sodium Throwaway Partial Scrubbing Cost Summary 8-3
8.1-2 Double-Alkali Partial Scrubbing Cost Summary 8-4
8.1-3 Limestone Partial Scrubbing Cost Summary 8-5
8.1-4 Sodium Throwaway Partial Scrubbing Energy Requirements 8-9
8.1-5 Double-Alkali Partial Scrubbing Energy Requirements 8-10
8.1-6 Limestone Partial Scrubbing Energy Requirements 8-11
8.1-7 Sodium Throwaway Partial Scrubbing Wastewater Production
Rates 8-13
8.1-8 Double-Alkali Partial Scrubbing Sludge Production Rates 8-14
8.1-9 Limestone Partial Scrubbing Sludge Production Rates 8-15
8.2-1 Assumptions and Bases for Material and Energy Balances 8-17
8.2-2 Cost Summary for Limestone Scrubbing 8-19
8.2-3 Energy Summary for Limestone Systems 8-21
8.2-4 Solid Waste Summary for Limestone Systems 8-22
XVI
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ACKNOWLEDGMENT
The authors wish to acknowledge the key contributions made by the following
people in the conduct of this evaluation. Overall guidance and coordination
of this evaluation with other eveluations of industrial boiler control systems
being conducted by EPA's Industrial Environmental and Research Laboratory
(IERL) was performed by Mr. J. David Mobley. Mr. Charles Sedman provided
overall project guidance for EPA's Office of Air Quality Planning & Standards
(OAQPS). In addition, valuable input was received from the reviews of
draft materials by several owners and operators of industrial boiler flue
gas desulfurization systems, and their contribution is greatly appreciated.
XV11
-------
SECTION 1
EXECUTIVE SUMMARY
1.1 INTRODUCTION
1.1.1 Background and Objective
The Clean Air Act Amendments of 1977 require the Environmental Protection
Agency to coordinate and lead the development and implementation of regulations
on air pollution. These include standards of performance for new and modified
sources of pollution. Specifically mentioned in the Act are fossil fired
steam generators. Accordingly, EPA has undertaken a study of industrial
boilers with intent to establish emission control levels based upon the results
of this and other studies.
This report presents the results of a study conducted to evaluate the
applicability of various flue gas desulfurization (FGD) technologies for
treating S02 emissions produced from small sized industrial boilers. Results
of this evaluation will be used by the EPA to establish regulatory alternatives
for small industrial boilers. Factors that were considered in evaluating the
applicability of FGD technologies to industrial boilers included development
status, capital and operating costs, energy requirements, environmental impacts,
and performance and operating data.
1.1.2 Approach
In order to complete the objective of this study, a multiphased project
approach was used.- First, a comprehensive list of FGD processes was reviewed
which included processes in commercial use, processes under development, and
processes for which development efforts have ceased. Process status reports
were prepared for eleven of these processes; those which are currently commer-
cially used or are undergoing major demonstration efforts. Status reports for
1-1
-------
each process included detailed process descriptions, and discussions of design
considerations and process performances.
The second phase involved selecting the candidate processes that appeared
to be best suited to industrial boiler applications from the list of eleven
processes for which the process status reports were prepared. These
candidate technologies were compared using three emission control levels
labelled "moderate, intermediate, and stringent" which correspond to SC>2
removal levels of 75, 85 and 90 percent. The control levels were chosen to
form a basis of comparison of the control technologies considering performance,
costs, energy, and non-air environmental impacts. A series of material and
energy balance calculations were then prepared for the candidate processes
to assess the environmental and energy impacts associated with process size,
fuel sulfur content, and percent S02 removal. Finally, a series of capital
and operating cost estimates were performed to assess the economic impact of
the candidate FGD processes as functions of the above three variables.
From these comparisons, candidate "best technologies"for control of S02
were recommended for consideration in subsequent industrial boiler studies.
These "best technology" recommendations do not consider combinations of
technologies to remove more than one pollutant and have not undergone the
detailed environmental, cost, and energy impact assessments necessary for
regulatory action. Therefore, the levels of "moderate, intermediate, and
stringent" and the recommendation of "best technology" for individual
pollutants are not to be construed as indicative of the regulations that
will be developed for industrial boilers. EPA will perform rigorous
examination of several comprehensive regulatory options before any decisions
are made regarding the standards for emissions from industrial boilers.
1.1.3 Conclusions
The major conclusion of this study is that there is no "best" FGD process
for application to industrial boilers. Each of the five candidate processes
1-2
-------
has its own advantages and disadvantages which could serve to make it the
"best" process for a specific application.
For the small sized applications, 8.8 MW (SOxlO6 Btu/hr), the sodium
throwaway process has the lowest annualized costs, is the lowest energy user,
and produces a small aqueous waste stream which is generally treated in
existing wastewater treating facilities. Currently, this process is the process
of choice since 102 industrial boiler FGD systems (over 75 percent) are of the
sodium throwaway type. However, there may be some areas where this process
cannot be used because of regulations on the discharge of dissolved solids.
In those cases either the limestone or double alkali process would probably
become the process of choice.
The limestone and double alkali processes are considered together as
their costs and environmental impacts are very similar. The double alkali
process will have lower energy requirements than the lime/limestone process
due to its more alkaline liquor and lower liquor circulation rate. The major
disadvantage of these processes is the necessity of sludge disposal.
The spray drying/baghouse FGD system was considered only for low sulfur
coal application due to data availability. Capital costs of this process were
greater than the other throwaway processes due primarily to the fact that
baghouse costs were included as part of the FGD system. Spray drying annualized
costs, however, compared quite well with the other processes due primarily to
the decreased disposal costs associated with handling a dry waste product. If
baghouse costs were not included as part of this process, it would become the
least expensive alternative. Future applications of this technology will
probably be dependent upon the results of the five units that are currently
under construction on both industrial and utility boilers.
The Wellman-Lord process will probably have limited applications for the
process sizes evaluated in this study. This is due to its process costs,
complexity in comparison to the other systems, and energy requirements.
1-3
-------
However, for some applications with a strong by-product market, or with severe
environmental regulations limiting the disposal of waste products from the
other processes, the Wellman-Lord process may be attractive.
1.1.4 Report Organization
Section 2 of this report presents the process status reports for the
eleven commercial and developing FGD processes that were evaluated for their
potential to control SOz emissions from industrial boilers. Each of the
processes are described and discussed with regard to their development status,
design considerations, operating procedures, and system performance. Section
3 presents the methodology and results of a process screening that was used to
select the five candidate processes for best control system for industrial
boiler applications.
Section 4 presents the results of the cost analysis and discusses the
effects of process size, fuel sulfur content, and S02 removal level on process
costs. Sections 5 and 6 present process energy requirements and environmental
impacts, respectively. Information presented in these sections illustrates
the effects of process size, fuel sulfur content, and SOa removal levels on
process energy requirements and environmental impacts. Section 7 presents the
performance data that were obtained from industrial boiler FGD systems and
Section 8 presents the impacts of partial scrubbing for the candidate FGD
systems.
There are also two Appendices included with this report. Appendix A
presents the results of the material and energy balance calculations for each
of the processes and illustrates the variation in process stream flow rates
with unit size, S02 removal level, and fuel sulfur content. Appendix B pre-
sents the results of the capital and operating cost estimates.
The remainder of this section presents a summary of the results and
highlights of the information presented in the rest of the report.
1-4
-------
1.2 EMISSION REDUCTION FOR COAL-FIRED INDUSTRIAL BOILERS
There are currently some 100 FGD processes that are in various stages of
development including processes in early developmental stages and those for
which development efforts have ceased. Of these processes, there are five
that are in commercial use today in the United Statues. In addition, there
are six that are currently at the demonstration stage. It is felt that these
eleven systems will be used for the majority of near-term FGD applications to
both utility and industrial boilers. Consequently, they are discussed in
detail in Section 2 of this report. Table 1.2-1 presents a summary of the
development status and industrial boiler applicability of these eleven processes.
1.2.1 Selection of Candidates for Best Control System
In order to select the candidate control systems, a set of evaluation or
screening criteria were established to provide an objective and consistent
means of comparing the processes and to insure that the same factors were
considered for each process. The screening criteria were then applied to each
process and the results were compared and used to select the processes that
appeared to be best suited for near-term industrial boiler applications. The
criteria used for this screening are listed in Table 1.2-2.
TABLE 1.2-2. FLUE GAS DESULFURIZATION SCREENING CRITERIA
1. Status of Development
• Overall Process Development
• Availability of Data
2. Performance
• SO2 Removal
• Reliability
3. Applicability
• Simplicity
• Flexibility
• Controllability
• By-Product Marketability
4. Economic Considerations
• Capital Investment Costs
• Operating Costs
5. Energy Considerations
• Liquid Pumping Requirements
• System Pressure Drop
• Regeneration Energy
• Requirement for Reducing Gas
6. Environmental Considerations
• Multipollutant Control
• Secondary Pollutant Emissions
1-5
-------
Process
Development status
No. of operational plants
Industrial Utility
Applicability to industrial boilers
I
CTN
Lime/Limes tone
Double Alkali
Wellman-Lord
Magnesium Oxide
Sodium Scrubbing
Spray Drying
Citrate/Phosphate
Commercial industrial and
utility applications.
Commercial industrial appli-
cations - a 280 MWe utility
application is planned.
Commercial applications for
tail gas treating. A 115 MW£
utility demonstration test
has been completed.
Commercial utility applica-
tions. No planned
industrial applications.
Commercial industrial and
utility applications.
Pilot-scale. Commercial
industrial utility systems
are under construction.
1 MW pilot-scale. A 64 MWg
industrial boiler applica-
tion is planned.
28
21
119
Generally applicable. Possible
limitations due to solids disposal
land requirements.
Generally applicable. Has demon-
strated relatively reliable
operations. Possible limitations
due to solids disposal land
requirements.
Generally applicable. Process costs
and complexity will limit applica-
tions to small boilers. Has demon-
strated good reliability.
Process complexity will limit applica-
tions for industrial boilers. Long
term reliability not demonstrated.
Generally applicable. Possible limita-
tions due to sorbent availability and
cost, and water treatment.
S02 removal may be limited for lime
based high sulfur coal applications.
System is generally applicable except
for land requirements for solids dis-
posal. High reliability is claimed
but undemonstrated.
Applicability to small boilers will
be limited by overall complexity and
the need for a reducing gas to
produce HjS.
-------
TABLE 1.2-1. (Continued)
Process
Development status
No. of operational plants
Industrial Utility
Applicability to Industrial boilers
Bergbau-Forschung/
Foster Wheeler
Atomics International/
Aqueous Carbonate
Process
Shell/UOP
Chiyoda 121
20 MW demonstration in U.S.
and a 45 MW demonstration
in Germany.
1.25 MW nonintegrated pilot
plant. A 100 MW utility
demonstration is planned.
0.6 MW pilot plant in U.S.
on coal-fired boiler.
40 MW in Japan on oil-
fired boiler.
Small-scale pilot plant.
A 20 MW utility demon-
stration is planned.
Applicability will be limited by
overall complexity and the require-
ment for extensive solids handling
equipment.
Applicability will be limited by
overall complexity for small boiler
applications. Use of unfamiliar
technology in the reducing reactor may
hinder process acceptability.
Applicability will be limited by
overall complexity and the require-
ment for hydrogen for regeneration.
Generally applicable. Possible
limitations due to solids disposal
land requirements in cases where
by-product gypsum marketability is
not feasible.
-------
Results of the process screening were that five FGD processes were
selected as candidate systems for application to industrial boilers. The
processes selected were:
Sodium Scrubbing
Double Alkali
Lime/Limestone
Spray Drying
Wellman-Lord
The first three processes are all currently used to control S02 emissions
from industrial boilers throughout the United States. These processes are
the ones of choice as evidenced by the fact that 118 of the 132 operating
industrial boiler FGD systems are of these process types. The remaining
operating industrial boiler FGD systems use ammonia process waste waters as a
sorbent and are predominately found in sugar processing plants. Lime/limestone,
double alkali, and sodium scrubbing processes also appear to be the processes
of choice for future installations as evidenced by the fact that 36 out of 39
systems in the planning or construction stages are of these process types.
The spray drying process was selected as a candidate technology for
industrial boiler applications due to its potential for widespread use as
evidenced by the large amount of interest expressed in this rapidly developing
process. Presently, there are no full-scale spray drying FGD systems in
operation; however, orders have been placed for five commerical spray drying
systems, two of which are industrial boiler applications. Data on spray
drying is available only from pilot scale units. This data will be used in
later sections to compare the energy, environmental, and economic impacts of
this process versus the other candidate processes.
The Wellman-Lord process was selected to compare the impacts of a regener-
able FGD process against the other candidate processes which are all throwaway
systems. There are currently no regenerable systems in operation on small
1-J
-------
industrial boilers, however, a demonstration of the Citrate Process is sche-
duled for operation in the near future. The Wellman-Lord process was selected
as a candidate process over the Citrate Process primarily because of its
increased development status and availability of data.
1.2.2 Cost Analysis of Candidate Systems
Process costs are based on mid-1978 dollars and were evaluated as a
function of process size, fuel sulfur content, and S02 removal for the five
candidate FGD processes. The general approach used in developing the process
costs consisted of four main steps. First, a series of material and energy
balance calculations were performed for each process to define process stream
flow rates and energy requirements as functions of unit size, SO2 removal,
and the amount of sulfur in the coal. The energy requirements for each
process are presented and discussed in Section 5 and the material balance
results are presented in Appendix A. Second, each of the FGD processes were
divided into a number of process modules which represented separate processing
areas. Third, equipment sizes were then developed for each process module
based on the results of the material and energy balances. Finally, capital
cost estimates were prepared by contacting process equipment vendors for
price quotations in the size range of the standard industrial boilers for
this study. Except for the spray drying process, particulate control equipment
costs were not included in this study.
The wastewater treating processing area for the sodium throwaway process,
consisting of oxidation and pH neutralization, was assumed to be associated
with the boiler or plant in question. Therefore, wastewater treatment appears
as an operating cost only. Similarly, solids disposal for the other non-
regenerable processes was assumed to be contracted out for offsite disposal
and it, too, appears only as an operating cost.
1-9
-------
Direct capital investment costs for the FGD processes ranged from a low
of $187,000 for a sodium throwaway process (8.8 MWt (30xl06 Btu/hr), 75
percent removal, 0.6 percent sulfur) to a high of $2,573,000 for a Wellman-
Lord process (58.6 MWt (200xl06 Btu/hr), 90 percent removal, 3.5 percent
sulfur coal). When indirect capital expenses are added, the total capital
investment costs for these two cases become $394,000 and $4,233,000 respec-
tively. From a capital cost standpoint, the sodium throwaway process appears
to be the least costly process, and the Wellman-Lord the most costly process
for all of the cases considered.
With regard to annualized costs, the relative rankings of the FGD
processes remain the same as with the capital costs for all cases considered.
The sodium throwaway process again emerged as the least costly alternative.
It should be noted however, that part of the low costs for this process
are attributable to the relatively simple waste treating process consisting
of oxidation plus pH neutralization. If a more elaborate water treating
scheme were required, possibly to comply with zero discharge regulations,
process costs would increase accordingly and could even make this process
the most costly alternative. Figures 1.2-1 and 1.2-2 illustrate graphically
the relative capital and annualized costs of the candidate processes applied
to boilers burning a 3.5 percent sulfur coal and Figures 1.2-3 and 1.2-4
illustrate the same costs for 0.6 percent sulfur coal applications.
The cost effectiveness of the various FGD processes was also determined
as part of this study. Cost effectiveness was defined as dollars per
kilogram of removed S02 ($/kg S02) and was calculated by dividing the annualized
process costs by the kilograms of S02 removed in a year assuming a 60 percent
load factor. Results of these calculations show that both coal sulfur content
and process size significantly affect the cost effectiveness of an FGD process.
--Throughout this report, boiler sizes will be expressed as watts of heat
input. To avoid confusion with electrical output watts, heat input energy
will be expressed as MW and electrical output energy will be expressed as MW
1-10
-------
4000
w
V-i
nl
CO
o
u
ex
rt
o
3000
2000
10GC
Double Alkali
Sodium Throwaway
29.2
(100)
58,6
(200)
87.9
(300)
118
(400)
Size in MWfc(106 Btu/hr)
Figure 1.2-1.
FGD capital costs versus unit size,
(3.5% S coal, 90% removal)
1-11
-------
2000
u>
o
•X3
CO
4-1
CO
-------
ca
rH
i-l
o
en
O
w
o
o
3000
•Spray Drying (lime)
2000
Spray Drying (Sodium)
-H
u
1000
Limestone
Sodium Throwaway
29.3
(100)
58.6
(200)
Size in
87.9
(300)
(106Btu/hr)
118
(400)
Figure 1.2-3.
FGD capital costs versus unit size.
(0.6% S coal, 75% removal)
1-13
-------
1500
CO
)-l
CO
CO
CO
O
T)
0)
' N,
'3
C
1000
500
— Spray Drying
Limestone
Sodium Throwaway
29.3
(100)
58.6
(200)
87.9
(300)
118
(400)
Size in MWfc (10eBtu/hr)
Figure 1.2-4. FGD annualized costs versus unit size.
(0.6% S coal, 75% removal)
1-14
-------
For a given size system, cost effectiveness increases with an increasing coal
sulfur content. For a fixed coal sulfur content, cost effectiveness increases
with increasing process size. Consequently, the most cost effective systems
are those designed for the large sized boilers burning high sulfur coal, and
the least cost effective systems are those designed for the small boilers
burning low sulfur coal. Figure 1.2-5 illustrates these effects for the
limestone processes. Curves developed for the other process showed similar
effects.
1.2.3 Energy Impacts of Candidate Control Systems
Process energy requirements were evaluated as a function of process
size, fuel sulfur content, and level of SOo control. Results of these calcu-
lations, shown in Table 1.2-3, indicate that the process energy penalties
range from about 1/2 to 6 percent of the gross heat input to the boiler when
no reheat is required, and from about 2 to 8 percent including stack gas
reheat. Since the spray drying process exhausts its stock gas at 175 F, no
reheat energy penalties were charged for this process. The larger energy
consumption for the Wellman-Lord process is due to the steam and methane
requirements for the regeneration and SO^ reduction portions of the process.
A summary of the relative percentage of the energy requirements of each
process as compared to the overall energy requirement for the throwaway
FGD processes applied to a boiler burning 3.5 percent sulfur coal is presented
in Table 1.2-4. This table shows that fans are the largest energy consumer
for each of the wet throwaway processes when stack gas reheat energy require-
ments are not included. However, when reheat energy requirements are included,
they become the dominant energy consuming portion of these processes. The
variations in energy requirements for these processes are due to different
levels of sulfur in the coal, different levels of SO^ control, and to a
smaller extent, unit size.
1-15
-------
01
I
CD
Pi
CM
o
C/3
00
@ 90% S02 removal
CO
10
(L>
a
(U
a
a)
w
o
o
0.6% Sulfur Coal
.03.5% Sulfur Coal
14.6 273 ?
(50) (100) (150)
Size in MW (106 Btu/hr)
(200)
Figure 1.2-5. Limestone process cost effectiveness
1-16
-------
TABLE 1.2-3. RANGE OF FGD SYSTEM ENERGY REQUIREMENTS*
SOo control method
Energy Requirement Energy Requirement
Not Including Reheat Including Reheat
Limestone
Sodium Scrubbing
Double Alkali
Spray Drying
Wellman-Lord
0.9-1.8
0.5-0.8
0.5-0.6
0.5-0.8
1.6-6.2
2.4-3.5
2.0-2.6
2.0-2.3
0.5-0.8
3.2-7.9
* Energy Requirements expressed as percent of net heat input to boiler.
TABLE 1.2-4. PERCENTAGE ENERGY CONSUMPTIONS FOR NONREGENERABLE PROCESSES
(58.6 MWt Boiler, 90% Removal, Eastern Coal)
Source of
energy
consumption
Limestone
Percent
kW of total
Double Alkali
Percent
kw of total
Sodium TA
Percent
kw of total
Raw materials
handling and
preparation
Liquid pumps
Fans
Disposal pumps
Utilities and
services
TOTAL
Reheat Steam
114.2 13
317.4 35
471.7 51
8.6 1
911.9
884.0
1795.9
12.7 1
63.2 21
222.0 72
9.3 3
307.2
884.0
1191.2
9.3
40.7
222.0
160.4
9.3
441.7
884.0
2
9
51
36
2
1325.7
1-17
-------
Energy requirements for the spray drying process are quite low, and
compare well with the other nonregenerable processes when no reheat is included.
Table 1.2-5 illustrates this for a low sulfur western coal case. If reheat is
required, the spray drying process has a significant advantage since it exhausts
its flue gas at 175°F and requires no reheat.
TABLE 1.2-5. COMPARISON OF SPRAY DRYING PROCESS ENERGY REQUIREMENTS
(44 MW , 0.6% S coal, 75% Removal)
Source of
energy
consumption
Raw Material Handling
Liquid Pumping
Fans
Disposal Pumps
Atomization
Utilities & Services
Total
Reheat steam
Limestone
kW
15
133
292
-
-
8
448
775
1223
Percent
of total
3
30
65
-
-
2
Sodium
TA
Percent
kW of total
1
29
195
21
-
9
254
775
1029
-
11
78
8
-
3
Lime
kW
1.
0.
140
-
120
8
270
—
270
Spray Drying
Percent
of total
5 1
5
52
-
44
3
For the Wellman-Lord process, the relative amounts of energy consumed by
the various process areas varies depending upon the sulfur content of the coal
being burned as shown in Table 1.2-6. However, for both the eastern 3.5
percent sulfur coal case and the western 0.6 percent sulfur coal cases, the
regeneration and sulfur production areas are the major energy users. It is
doubtful that the energy requirements of the regeneration processing area can
be reduced since double effect evaporators were assumed for these calculations
which are some 45 percent more energy efficient than single effect evaporators.
If reheat energy requirements are included, they become the major energy
consumer for the low-sulfur western coal case and serve to increase the energy
requirements of both cases by about 900 kW .
1-18
-------
TABLE 1.2-6. PERCENTAGE ENERGY CONSUMPTION FOR WELLMAN-LORD PROCESS
(58.6 MW 90% S02 Removal)
Source of
energy
consumption
Raw materials
handling and preparation
Pumps
Fans
Process steam
Methane
Utilities and services
Total
Reheat steam
k
1
42
208
469
219
10
950
900
1850
.9
.2
.6
.0
.0
.2
.9
.0
.9
Western coal
(0.6%S)
Percent of total
<1
4
22
49
23
1
48.6
Eastern coal
k
9
82
205
2220
1048
9
3575
884
4459
.3
.9
.5
.0
.0
.6
.3
.0
.3
Percent of
-------
PP
tr>
o
00
c
o
•H
cn
en
•H
I
O
to
2400(5.6)->
2100(4.9)—
1800(4.2)_
1500(3.5)_
1200(2.8)-
900(2.1)-
600(1.4) -
300(0.7)
All Boiler Sizes Firing
High Sulfur Eastern Coal
All Boiler Sizes Fifing
Medium Sulfur Coal
All Boiler Sizes Firing
Low Sulfur Western Coal
I
25
50 5(
I
75
I I I
85 90 100
S02 Control Level (%)
Figure 1.2-6. S02 emissions versus control level.
1-20
-------
With regard to water pollution, only the sodium throwaway process should
produce a significant environmental impact. Good design and operating prac-
tices for the limestone and double-alkali processes include dewatering the
sludge and recycling the supernatant liquid. Dewatered sludges, however,
contain up to 50 percent water which under proper design will be contained in
the ultimate disposal site; be it a pond or landfill. Consequently, there
should be essentially no water emissions from these systems except for times
of process upsets.
The aqueous waste stream from the prescrubber of the Wellman-Lord process
will be characterized by a low pH which results from the chlorides that are
removed from the gas stream. However, except for the high chloride concentra-
tions and low pH, the quality of the prescrubber discharge will be very
similar to that of the boiler ash sluice water. Since this stream has been
estimated to be approximately one percent of the ash sluicing requirements
for a power plant, it can be used for ash sluicing where it will become
diluted and neutralized with the other ash sluice water. Consequently, water
emissions from the Wellman-Lord prescrubber stream should thus be limited to
intermittent discharges from the ash pond.
The aqueous stream from a sodium throwaway system will contain about
five percent dissolved solids. In these systems, the absorbed S02 reacts to
form NA SO and Na2S04 which are removed from the system as dissolved solids
2 0
in an aqueous waste. Consequently, the amount of aqueous emissions is directly
related to both the S02 control level and the coal's sulfur concentration.
Discharge rates and average stream compositions for the cases considered in
this study are given in Table 1.2-7.
Common water treating practice for sodium throwaway systems in use today
is to discharge their wastes to an evaporation pond or to an existing centra-
lized water treating plant. Of the 102 sodium scrubbing systems in use
today, about 80 use evaporation ponds and 10 use centralized water treating
for disposal of their FGD wastes. The remaining systems use varied approaches
ranging from discharge to city sewers to deep mine injection.
1-21
-------
I
t-o
ho
TABLE 1.2-7. AQUEOUS EMISSIONS FROM SODIUM THROWAWAY PROCESSES
Control
Boiler size and type level
8.8 MWt(30 X 106Btu/hr)
Underfeed stoker
22 MWt(75 X 10sBtu/hr)
Chaingrate Stoker
44 MWt(150 X 105Btu/hr)
Spreader Stoker
58.6 Mwt(200 X 106Btu/hr)
Pulverized coal
90
85
75
56
90
75
56
90
75
56
90
85
75
56
3.5% S eastern coal
Jl/sec
0.85
0.79
0.71
0.52
2.18
1.77
1.26
4.38
3.53
2.45
5.80
5.26
4.71
3.34
(gpm)
13.4
12.5
11.2
8.4
34.6
28.1
19.9
69.4
55.9
38.8
92.0
83.4
74.7
52.9
2.3% S coal 0.6% S western coal
£/sec (gpm) &/sec
0.17
0.16
0.15
_ _ -
1.28 20.3 0.46
0.37
_ — —
0.92
0.75
- - -
1.23
1.11
0.99
_ _ -
(gpm)
2.7
2.6
2.4
—
7.3
5.9
—
14.6
11.8
—
19.5
17.6
15.7
—
118 MWt(400 X 105Btu/hr)
Pulverized coal
90
11.7
185
6.82 108
2.46
38.9
Avg.
Avg.
Dissolved Solid Compositions
IDS Concentration (wt. %)
Na
Na
NA
2
2
2
S03
so,,
C03
=
=
=
5
77
9
14
percent
percent
percent
73
13
14
percent
percent
percent
5
69
17
14
percent
percent
percent
5
*Based on material balance calculations provided in Appendix A.
-------
For purposes of this evaluation, onsite treatment of sodium system
aqueous wastes using a basic water treating scheme of sulfite oxidation and
pH neutralization was selected as the treatment method for evaluation.
Although evaporation ponds are currently used in the majority of sodium
system applications, their use is limited to certain geographic areas of the
country where the annual evaporation rate exceeds the annual rainfall. The
water treatment system selected for this evaluation will result in a sodium
sulfate stream which must be disposed.
The major solid waste impacts from the five candidate processes result
from the sludges produced in the limestone and double-alkali processes and
the dry solid produced in the spray drying process. This assumes, of course,
that a market exists for the sulfur and H^SO^ produced in the Wellman-Lord
process. A solid purge stream of ^280^ is also produced in the Wellman-Lord
process, but the stream is relatively small and should not constitute a major
solid waste impact, especially for the size applications under consideration
in this evaluation.
Both the limestone and double alkali sludges are composed primarily of
calcium sulfite and sulfate salts. Significant amounts of fly ash may also
be present, depending on the method of upstream particulate control in use.
For this study, upstream particulate removal was assumed and sludge production
rates are given on an ash free basis. The sludges are relatively inert and
can be disposed of in an environmentally acceptable manner. The disposal
methods currently in use are lined and unlined ponding and landfilling of
treated and untreated materials.
The dry solid waste product from the spray drying process will consist
primarily of calcium or sodium salts, depending upon the type of alkali used
as the S02 sorbent. Significant amounts of fly ash will also be present
since the solids collection device associated with the spray drier, probably
a baghouse, will remove the particulates generated from the coal combustion
process along with the spray drying solid wastes. Upstream particulate
I--:
-------
removal is not practical for this process since the spray dryer1s performance
is not adversely affected by the presence of fly ash and dual particulate
removal units would be unattractive from both an energy and economic view.
As with the sodium throwaway system, all of the S02 absorbed from the
flue gas by a limestone system must leave the process in a waste stream, in
this case as a waste sludge. Consequently, the amount of sludge produced by a
limestone system is proportional to the sulfur content of the coal and the SC>2
removal level. Table 1.2-8 presents the results of the limestone process
material balance calculations and shows the variation in sludge production
with coal sulfur content and SO2 removal. A 50 percent solids sludge was
assumed for these calculations.
The volume of sludge produced is also important as the sludge volume
will determine the size of the holding pond or landfill area. Table 1.2-9
presents the results of calculations to estimate the sludge volumes produced
by a limestone process for the standard sized boilers. The difference in
sludge densities for the two coals is due to the higher oxidation for the
western coal cases. Results are presented in units of cc/sec, Ib/hr, and
acre-feet/30 years. The last category, acre-feet/30 years gives an indication
of the total volume of sludge to be handled over the life of the plant assuming
a 30 year life and an onstream factor of 60 percent. Figure 1.2-7 illustrates
the results of these calculations graphically and shows the variation in
sludge production with coal sulfur content, boiler size, and level of removal.
As one would expect, sludge production increases with all of these factors.
The quantity of sludge produced from the double alkali process will also
vary with sulfur removal and fuel sulfur content. Table 1.2-10 presents the
solid waste impacts for the double-alkali process. The major difference
between the amount of sludge produced from the double-alkali and limestone
systems is that the limestone system stoichiometry is based upon 1.2 moles
sorbent per mole SO 2 removed whereas the double-alkali system stoichiometry
is based upon 1.0 moles sorbent per mole S02 removed.
1-24
-------
TABLE 1.2-8.
i
ro
SOLID WASTE PRODUCTION FOR THE LIMESTONE FGD PROCESS
(Ash-Free Basis, 50 percent solids)
Percent
Boiler size and type removal g/s
8.8 MWt(30xl06 Btu/hr)
Underfeed Stoker
22 MWt(75xl05 Btu/hr)
Chaingrate Stoker
44 MWt(150xl06 Btu/hr)
Spreader Stoker
58.6 MWt(200xl06Bbu/hr)
Pulverized Coal
90
85
75
90
75
90
75
90
85
75
89.8
84.7
74.9
226.5
190.2
457.0
380.6
606.8
576,1
507.0
3.5% S eastern coal
(Ib/hr)
(712)
(672)
(594)
(1796)
(1508)
(3624)
(3018)
(4812)
(4568)
(4020)
&/min
2.7
2.7
2.3
6.8
5.7
13.6
11.4
18.2
17.4
15.2
(gal/min)
(0.7)
(0.7)
(0.6)
(1.8)
(1,5)
(3.6)
(3.0)
(4.8)
(4.6)
(4.0)
g/s
20
19
17
52
43
104
87
139
131
116
.9
.4
.4
.2
.6
.7
.3
.5
.9
.0
0.6% S western coal
(Ib/hr)
(166)
(154)
(138)
(414)
(346)
(830)
(692)
(1106)
(1046)
(920)
Jl/min
0.61
0.57
0.53
1.5
1.1
3.0
2.7
4.2
3.8
3.4
(gal/min)
(0
(0
(0
(0
'(0
(0
(0
(1
(1
(0
,16)
-15)
-14)
.4)
.3)
.8)
.7)
-1)
.0)
.9)
-------
i
"O
TABLE 1.2-9. SOLID WASTE VOLUMES FOR THE LIMESTONE FGD PROCESS
(Ash-Free Basis, 50 percent solids)
Percent
Boiler size and type removal
8.8 MWt(30xl06 Btu/hr)
Underfeed Stoker
22 MWt(75xl06 Btu/hr)
Chaingrate Stoker
44 MWt(150xl06 Btu/hr)
Spreader Stoker
58.6 MWt(200xl06 Btu/hr)
Pulverized coal
90
85
75
90
75
90
75
90
85
75
==^=^==
3.5% S eastern coal*
Sludge volume
cc/sec
69.1
65.2
57.6
174.2
146.3
351.5
292.8
466.8
443.2
390.0
(ftVhr)
8.8
8.3
7.3
22.1
18.6
44.6
37.2
59.3
56.3
49.5
(acre-ft/
30 yrs)
31.8
30.0
26.4
80.0
67.4
161.4
134.6
214 ..6
203.8
179.2
0.6% S western coalt
Sludge volume
cc/sec
13.8
12.8
11.5
34.6
28.9
69.3
57.8
92.4
87.4
76.8
(ftd/hr)
1.8
1.6
1.5
4.4
3.7
8.8
7.4
11.8
11.1
9.8
(acre-ft/
30 yrs)
6.6
5.8
5.4
16.0
13.4
32 ;0
26.8
42.8
40.2
35.4
*Eastern sludge bulk density =1.30 g/cc (81.2 lb/ft3) j
^Western sludge bulk density =1.51 g/cc (94.0 lb/ft3) I
are average values from Table 6.2-12.
-------
0)
M
U
QJ
CJ
U
C
O
•H
4J
u
3
TD
3
500(230)
400(184)—
300(138)-
200(92) _
100(46)-—
Eastern Coal
90% Removal
Western Coal
90% Removal
Eastern Coal
75% Removal
Western Coal
75% Removal
22
44
58.6
Boiler Size, MW
Figure 1.2-7-
Sludge production rates for the limestone
FGD process (50 percent solids in sludge)
1-27
-------
1
N3
00
TABLE 1.2-10. SOLID WASTE IMPACT FOR THE DOUBLE ALKALI PROCESS
(Ash-free Basis, 50 percent solids)
Percenr
3.5% S eastern coal
0.6% S western coal
2.3% S coal
Boiler Size and type remova . g/s (Ib/hr) Jl/min (gal/min)
8.8 MWt(30 X 106 Btu/hr) 90 78.7
Underfeed Stoker
(624) 2.3
22 MWt(75 X 106 Btu/hr) 90 200.0 (1584) 5.7
Chaingrate Stoker
58.6 MWt(200 X 10G Btu/hr) 90 506.0 (401?.) 14.4
Pulverized Coal
118 MWt(400 X 10S Btu/hr) 90 1065.6 (8450) 30.7
Pulverized Coal
(0.6)
g/s "(Ib/hr) t/min (gal/min)
15.4 (122) 0.4 (0.1)
g/s (Ib/hr) t/roin (gal/min)
(1.5) 38.1 (302) 1.1 (0.3) 112.0 (8S8) 3.0 (0.8)
(3.8) 101.6 (806) 2.7 (0.7)
(8.1) 203.5 (1614) 5.3 (1.4) 596.8 (4732; 17-1
-------
All of the S02 absorbed in a spray dryer must also exit the process as a
waste stream, in this case as a solid salt. Table 1.2-11 presents the quantity
of solid wastes produced by the spray drying system as a function of S02
removal, coal type, and process size. Solid waste quantities from this
process are, however, a combination of spray dryer solids and fly ash generated
from coal combustion. It is interesting to note that for the cases considered
in this evaluation, the majority of solid wastes from this process resulted
from fly ash and not from the removal of SO2. Table 1.2-12 presents a break-
down of the origin of solid waste material for each of the spray drier cases.
Results of pilot plant testing reported by Basin Electric were that the
spray drying product produced from the coals they tested handled as well as
fly ash and would not require special handling equipment other than the
conventional dry handling equipment used for fly ash. Disposal methods
planned for the systems they have under construction are disposal in depleted
mines and landfill after mixing with conventional scrubber sludge. The two
spray drying systems under construction for industrial boiler applications
both plan to truck the waste solids to an offsite landfill area.
1-29
-------
TABLE 1.2-11. SOLID WASTES FROM SPRAY DRYING
(Total Fly Ash + Alkali Salts)
Boiler size
MWt (106 Btu/hr)
44
44
44
58.6
17.6
44
44
44
w 118
o
118
22
(150)
(150)
(150)
(200)
(60)
(150)
(150)
(150)
(400)
(400)
(75)
% S
Coal
0.6
0.6
0.6
0.6
0.6
0.6
0.6
0.6
0.6
2.3
2.3
SOz
Removal
90
75
50
75
75
90
75
50
70
70
70
Solids
Sorbent
Sodium
Sodium
Sodium
Sodium
Sodium
Lime
Lime
Lime
Lime
Lime
Lime
g/sec
132
114
99
174
29
134
114
96
344
729
85
(lb/hr)
(1047)
(904)
(785)
(1378)
(227)
(1066)
(905?
(764)
(2725)
(5782)
(675)
cc/sec
111
96
83
146
24
113
96
81
289
613
71
Volume
(ft°/hr)
(14.1)
(12.2)
(10.6)
(18.5)
(3.1)
(14.3)
(12.2)
(10.3)
(36.7)
(77.9)
(9.1)
a
acre-ft/15 years
25.5
22,1
19.2
33.5
5.6
25.9
22.0
18.6
66.5
141.0
16.5
a - Based on density = 1.19 g/cc (74.2 lb/ft3) Reference 14
-------
TABLE 1.2-12. SOLID WASTE BY ORIGIN
Amount of solid waste
MW (106
44
44
44
58.6
17.6
44
44
44
118
| — i
1 118
UJ
M 22
Btu/hr)
(150)
(150)
(150)
(200)
(60)
(150)
(150)
(150)
(400)
(400)
(75)
% S
Coal
0.6
0.6
0.6
0.6
0.6
0.6
0.6
0.6
0.6
2.3
2.3
SO 2
Removal
90
75
50
75
75
90
75
50
70
70
70
Type
sorbent
Sodium
Sodium
Sod ium
Sodium
Sodium
Lime
Lime
Lime
Lime
Lime
Lime
g/s
69
69
69
113
11
69
69
69
226
401
24
Fly ash
(Ib/hr)
(546)
(546)
(546)
(896)
(84)
(546)
(546)
(546)
(1791)
(3182)
(187)
Percent
of total
52
60
70
65
37
51
60
71
66
55
27
Desulf urization
g/s
63
43
30
61
18
66
45
27
118
328
61
products
(Ib/hr)
(501)
(341)
(239)
(482)
(143)
(520)
(359)
(218)
(934)
(2600)
(488)
Percent
of total
48
40
30
35
49
40
29
34
45
-------
SECTION 2
EMISSION CONTROL TECHNIQUES
2.1 PRINCIPLES OF CONTROL
There are several methods that may be used to reduce S02 emissions
from industrial boilers to comply with current emission standards. SOa
reduction methods that are applicable for a given site will be dependent
upon the required SOa removal or emission levels. Potential SOa removal
methods are: 1) use of low sulfur coal, 2) physical or chemical coal
cleaning, 3) use of synthetic fuels from gasification or liquefaction
plants, 4) fluid bed combustion, and 5) flue gas desulfurization. This
report discusses flue gas desulfurization (FGD) as a method of controlling
emissions.
The sulfur content of the fuel fired in a boiler dictates the amount of
emissions produced at a given installation. Fuels of concern for this
study are coal and residual oil. In general, light fuel oils and natural
gas are treated to remove sulfur prior to combustion. Table 2.1-1 lists
the proximate analyses for the standard fuels to be considered in this
evaluation. This table also presents the relative amount of uncontrolled
S02 emissions that would result from each assuming that 95 percent of
the fuel's sulfur content is converted to S02.
There are currently some 100 FGD processes that are in various stages of
development including processes in early developmental stages and those for
which development efforts have ceased.1 Of these processes, there are five
that are in commercial use today in the United States. In addition, there
are six that are currently at an advanced demonstration stage. Table 2.1-2
presents a summary of these eleven processes with regard to their develop-
ment status and applicability to industrial boilers. It is felt that these
2-1
-------
TABLE 2.1-1. S02 EMISSIONS FOR VARIOUS FUEL TYPES
I
1-0
a.
Composition
HHV kJ/kg
(Btu/lb)
Percent S
Percent ash
Uncontrolled SOz Emissions
(lb/106 Btu)
SOa emissions relative to
high sulfur eastern coal
High sulfur
eastern
27,447
(11,800)
3.5
10.6
5.61
1
Fuel type
Low sulfur Subbituminous Residual
eastern western oil
32,099
(13,800)
0.9
6.9
1.24
0.22
22,323
(9,600)
0.6
5.4
1.18
0.21
41,714b
(149,800)
3.0
0.1
3.04
0.54
Distillate
oil
38,706b
(139,000)
0.5
-
•0.48
0.08
Composition based on data provided by PEDCo as a basis for this study.
HHV expressed as kJ/kg(Btu/gal)
-------
TABLE.. 2.1-2 . FGD SYSTEM SUMMARY
Process
Development status
No. of operational plants
Industrial Utility
Applicability to industrial boilers
fo
I
UJ
Lime/Limestone
Double Alkali
Wellman-Lord
Magnesium Oxide
Sodium Scrubbing
Spray Drying
Citrate/Phosphate
Commercial industrial and
utility applications.
Commercial industrial appli-
cations - a 280 MWe utility
application is planned.
Commercial applications for
tail gas treating. A 115 MWfi
utility demonstration test
has been completed.
Commercial utility applica-
tions. No planned
industrial applications.
Commercial industrial and
utility applications.
Pilot-scale. Commercial
industrial utility systems
are.under construction.
1 MW pilot-scale. A 64 MWe
industrial boiler applica-
tion is planned.
28
21
119
Generally applicable. Possible
limitations due to solids disposal
land requirements.
Generally applicable. Has demon-
strated relatively reliable
operations. Possible limitations
due to solids disposal land
requirements.
Generally applicable. Process costs
and complexity will limit applica-
tions to small boilers. Has demon-
strated good reliability.
Process complexity will limit applica-
tions for industrial boilers. Long
term reliability not demonstrated.
Generally applicable. Possible limita-
tions due to sorbent availability and
cost, and water treatment.
S02 removal may be limited for lime
based high sulfur coal applications.
System is generally applicable except
for land requirements for solids 'dis-
posal. High reliability is claimed
but undemonstrated.
Applicability to small boilers will
be limited by overall complexity and
the need for a reducing gas to
produce H2S.
-------
TABLE 2.1-2. (Continued)
Process
Development status
No. of operational plants
Industrial Utility
Applicability to industrial boilers
Bergbau-Forschung/
Foster Wheeler
Atomics International/
Aqueous Carbonate
Process
Shell/UOP
Chiyoda 121
20 MW demonstration in U.S.
and a 45 MW demonstration
in Germany.
1.25 MW nonintegrated pilot
plant. A 100 MW utility
demonstration is planned.
0..6 MW pilot plant in U.S.
on coal-fired boiler.
40 MW in Japan on oil-
fired boiler.
Small-scale pilot plant.
A 20 MW utility demon-
stration ia planned.
Applicability will be limited by
overall complexity and the require-
ment for extensive solids handling
equipment.
Applicability will be limited by
overall complexity for small boiler
applications. Use of unfamiliar
technology in the reducing reactor may
hinder process acceptability.
Applicability will be limited by-
overall complexity and the require-
ment for hydrogen for regeneration.
Generally applicable. Possible
limitations due to solids disposal
land requirements in cases where
by-product gypsum marketability is
not feasible.
-------
eleven systems will be used for the majority of near-term FGD applications
to both utility and industrial boilers. Consequently, they are evaluated
in the following sections of this chapter.
Much of the data base for FGD systems, as presented in this chapter,
has been collected for FGD systems applied to utility boilers as opposed
to industry boilers. In general, utility operating data for FGD systems
can be used to represent the operations of industrial boiler systems
since the actual design of the boilers are very similar. However, there
are some operating differences between utility and industrial boilers
which may affect the operation of an FGD system when applied to industrial
boilers. These will be discussed where appropriate in the following
sections. Most notable of these are: 1) stoker-fired industrial boilers
generally operate with higher quantities of excess air than pulverized
coal boilers, 2) industrial boilers may have larger boiler load swings as
they are operated to follow process operation, and 3) utility boiler
operators are generally higher paid and better trained than industrial
boiler operators. Another important difference is that when industrial
boilers are down, replacement power or steam cannot generally be drawn
from a grid as can be done for utility boilers. Consequently, a less
complex, more reliable, and highly automated FGD system may be more
suitable for industrial boiler applications.^>5
The higher amount of excess air used in stoker-fired boilers (may
vary from 30-50 percent) will result in a higher oxygen content in the
flue gas. Pulverized coal boilers for both utility and industrial
applications are generally designed for 20-25 percent excess air. In
some FGD systems, the higher oxygen concentration can adversely affect
the rate of oxidation of the S0? scrubbing liquor. The dual alkali and
Wellman-Lord systems are perhaps the most affected by this phenomenon.
When the oxidation rate is high, significant amounts of sodium sulfate
are formed in these systems which are difficult to regenerate and in
many cases must be removed as a waste stream;
Wide load swings in industrial boilers are due to two main factors:
1) the need to follow process operations, and 2) load swings that occur due
2-5
-------
to the small boiler sizes. Boilers that provide steam for batch process
operations may change load from near capacity to reserve standby in a matter
of minutes. Apparently, this practice is dependent upon the industry and
the site specific operations and no generalizations can be drawn. The small
size of the industrial boilers also impacts their load swings since small
boilers have less repetition of individual equipment. For example, in a
small boiler with only two burners, if one of the burners goes out, a rapid
loss of up to half the load can result. However, in a large boiler, loss of
a single burner will not significantly affect the total load. The important
point is that FGD systems applied to industrial boilers should be able to
operate under rapidly changing load conditions.
2.2 CONTROLS FOR COAL-FIRED BOILERS
The eleven FGD systems summarized in Table 2.1-2 have been evaluated for
possible application to control SOz emissions from small industrial boilers.
These systems were evaluated as it was felt that they will be used for the
majority of near-term FGD applications due to their relatively advanced state
of development. The systems described in this chapter are listed below
according to their current status of development.
1) Commercially Applied Processes
• Lime/Limestone
Double Alkali
• Wellman-Lord
Magnesium Oxide
Sodium Scrubbing
2) Developing Processes
• Spray Drying
Citrate/Phosphate
Bergbau Forschung/Foster Wheeler
2-6
-------
• Atomics International - Aqueous Carbonate
• Shell/UOP
• Chiyoda 121
The following discussions present a description of each of the above
systems and evaluate them with regard to their development status, industrial
boiler applicability, design and operating considerations, and system per-
formance.
2.2.1 Lime/Limestone Wet Scrubbing
2.2.1.1 System Description—
A. System—The lime/limestone FGD process uses a slurry of calcium
oxide or calcium carbonate to absorb S02 in a wet scrubber. The chemistry is
quite complex, involving many side reactions. The overall reactions are those
of S02 with lime (CaO) or limestone (CaC03) to form calcium sulfite (CaS03)
with some oxidation of the sulfite to form calcium sulfate (CaSOit). Alkaline
fly ashes, if present, also contribute alkalinity in the form of soluble cal-
cium, magnesium, and/or sodium oxides which participate in the SC>2 removal
reactions.
The calcium sulfite and sulfate crystals precipitate in a reaction
vessel or hold tank which is designed to provide adequate residence time for
solids precipitation as well as for dissolution of the alkaline additive.
The hold tank effluent is recycled to the scrubber to absorb additional SOz.
A slipstream from the hold tank is sent to a solid-liquid separator to remove
the precipitated solids from the system. The waste solids are generally
disposed of by ponding or landfill.
The basic design of a lime/limestone scrubbing system can be divided
into the following process areas:
1) S02 absorption
2) Solids separation
3) Solids disposal
2-7
-------
A simplified flow diagram is presented in Figure 2.2.1-1.
1) S02 absorption - Absorption of S02' takes place in a wet scrubber
using lime or limestone in a circulating slurry . Particulates can be removed
in the S02 absorber or ahead of the absorber by an electrostatic precipita-
tor, wet scrubber, baghouse, or mechanical collector. The selection of a
method for removal of particulates is based on economics and operational
reliability. Removing particulates in the S02 absorber increases the solids
load in the S02 scrubbing system.
After leaving the particulate removal device, the flue gas enters the
wet S02 scrubber where absorption occurs. The overall reactions of gaseous
S02 with the alkaline slurry yielding CaS03«%H20 are shown in Equations
2.2.1-1 and 2.2.1-2.
For lime systems:
S02(. x + CaO,g) + %H20 -> CaS03'%H20(s) (2.2.1-1)
For limestone systems:
S°2(g) + CaC°3(s) + ^H2° "*" CaS03'%H20(s) + C02( , (2.2.1-2)
The solid sulfite is only very slightly soluble in the scrubbing liquor and
thus will precipitate to form an inert solid for disposal. In the lime sys-
tem some C02 may also be absorbed from the flue gas and will react in a
similar fashion to form solid calcium carbonate.
In most cases some oxygen will also be absorbed from the flue gas or
surrounding atmosphere. This leads to oxidation of absorbed SOz and precipi-
tation of solid CaSOit^HaO. The overall reaction for this step is as follows:
2-?
-------
S02 ABSORBER
REHEATER
FAN
TO STACK
FLUE GAS >•
I STEAM
MAKE-UP WATER
LIME
OR
LIME
STONE"
LIME
SLAKER
V V V V
CRUSHING SLURRY
AND
GRINDING
EFFLUENT HOLD TANK
SECOND STAGE
SOLID-LIQUID
SEPARATOR
OR
SETTLING POND
SOLID-LIQUID
SEPARATOR
SOLID WASTE
Figure 2.2.1-1. Process Flow Diagram Lime/Limestone Wet Scrubbing.
-------
For lime systems:
S02 + %02, , + CaO(s) + 2H20 -*• CaSOi,«2H20(s) (2.2.1-3)
For limestone systems:
S02( , + %02( , + CaC03(s) + 2H20 -»• CaS04-2H20(s) + C02 (g) (2.2.1-4)
The extent of oxidation can vary considerably, normally ranging anywhere from
almost zero to 40 percent. In some systems treating dilute S02 flue gas
streams, sulfite oxidations as high as 90 percent have been observed. The
actual mechanism for sulfite oxidation is not completely understood, although
the rate is known to be a strong function of oxygen concentration in the flue
gas and liquor pH. It may also be increased by trace quantities of catalyst
in fly ash entering the system.
Various types of gas-liquid contactors can be used as the SC>2 absorber.
These differ in S02 removal efficiency as well as operating reliability.
General types of contactors that have been used for S02 removal include:
venturi scrubbers,
• packed towers,
spray towers (horizontal and vertical),
• tray towers,
• grid towers, and
mobile bed absorbers (such as marble bed and turbulent contact
absorber (TCA)).
The liquid to gas ratio (L/G) generally varies widely depending mainly upon
the type of contactor, the flue gas S02 concentration, and the required
removal efficiency. A discussion of contactor types is found in Section
2.2.1.ID, Factors Affecting Performance. Simple impingement devices are
placed downstream from the absorber to remove mist entrained in flue gas.
2-10
-------
The effluent hold tank or reaction tank receives the lime or limestone
feed slurry and absorber effluent. In addition, settling pond water and
clarifier overflow can be sent to the hold tank. The tank is equipped with
an agitator for uniform composition. The volume of the hold tank is sized
to allow adequate residence time for calcium sulfite and sulfate precipita-
tion and for lime or limestone dissolution. Too little residence time in
the hold tank can cause nucleation of calcium sulfite and sulfate solids in
the scrubber, resulting in scaling.
Design of the reaction tank is specific for each installation and is
dependent upon the process precipitation rates and the lime or limestone
dissolution rates. Bechtel reports that a liquid residence time of between
5-15 minutes has been tested at Shawnee and used in some full-scale systems.6
The feed material for a lime scrubbing process is usually produced by
calcining limestone. Feed for a limestone process generally comes directly
from the quarry and is then reduced in size by crushing and grinding. The
lime or limestone is mixed with water to make a 25-60 percent solids slurry.
2) Solids separation - A continuous slurry stream of 10-15 percent
solids is removed from the hold tank and recycled to the absorber. The flow
rate of this stream is generally dependent on the quantity of S02 to be
removed from the flue gas. A bleed stream is removed from this recycle
stream and is dewatered to minimize the area needed for sludge disposal.
Dewatering can be accomplished in a variety of ways depending on the loca-
tion of the disposal site and the method of disposal used.
For systems with on-site ponding, the bleed stream may be pumped direct-
ly from the effluent hold tank to the pond. The supernatant liquour may then
be recycled back to the hold tank. Further thickening of the sludge can be
achieved depending on the physical properties of the sludge. A thickening
2-11
-------
device such as clarifier can be used to increase the solids content to about
30 percent. A vacuum filter or centrifuge could be used to raise the solids
content to about 50 percent or higher. A combination of vacuum filtering
and forced oxidation of the precipitated solids can increase the solids con-
tent to 85 percent or greater. Further details are given in Section 2.2.1.ID,
Factors Affecting Performance.
3) Solids disposal - Sludge disposal is one of the main disadvantages
of lime/limestone FGD systems in comparison to recovery processes. While moot
sludge disposal studies have been aimed at utility applications where the mag-
nitude of sludge production imposes large area requirements as illustrated
in Table 2.2.1-1, it might also be useful to identify the disposal procedure
of an industrial user of lime/limestone stystem. Rickenbacker Air Force Base
discharges its thickened sludge to a 5-acre lined pond located about 700
feet from the absorber.7 The pond is approximately 450 feet long, 250 feet
wide and 12 feet deep with a pond life expectancy of well over five years at
a coal combustion rate of 200,000 tons of 5 percent sulfur coal. Estimates of
sludge production rates for typical industrial boiler sized limestone FGD
systems are given in Table 2.2.1-2.
d
Dewatered sludge is generally sent to a pond or landfill for disposal.
If land is available on-site, a large pond is usually constructed to settle
out the solids and reuse the supernatant liquor. Both of the industrial
boiler facilities that use a lime/limestone FGD process dispose of their de-
watered sludge in ponds. Commercial "stabilization" methods are currently
in use at some utilities to convert the sludge to structurally stable landfill
material with low permeability.8'9 These methods could be used when on-site
disposal is not possible. The stabilized material can then be trucked to an
off-site area for landfill. The stabilized material can then be trucked to
an off-site area for landfill. Similarly, combined disposal of ash and sludge
in certain cases results in a structurally stable disposal material.10 Other
disposal alternatives currently under study by EPA include deep mine filling,
strip mine filling and ocean dumping 11 They are being evaluated both from
economic and environmental standpoints.
2-12
-------
TABLE 2.2.1-1. ESTIMATED QUANTITY OF FLUE GAS DESULFURIZATION
WASTES AND ASH FROM SELECTED COALS - 1000 MWe PLANT 12
ABC
Ash Content of Coal (%) 15 8 15
Sulfur Content of Coal (%) 3.5 0.8 2.0
Heating Value of Coal
(Btu/lb) 12,500 8,500 12,500
Annual Coal Use
(million tph) 2.63 3.85 2.63
Sulfur Emission Standards 1.2 lb/106 Btu 50%. removal 1.2 lb/106 Btu
Annual Sludge Production
(tpy dry) 315,000 35,000 160,000
Annual Ash Production
(tpy dry) 395,000 310,000 395,000
Annual Total Solid Wastes
(tpy dry) 710,000 345,000 555,000
2-13
-------
ho
I
TABLE 2.2.1-2. ESTIMATED QUANTITY OF SLUDGE FROM INDUSTRIAL BOILER LIMESTONE FGD SYSTEMS
(Ash Free Basis, 90 Percent SO Removal)
Boiler Size
(106BTU/hr)
30
30
75
75
150
150
200
200
Coal Sulfur
Content (%)
3.5
0.6
3.5
0.6
3.5
0.6
3.5
0.6
Mass Production Rate
g/sec
90
21
227
52
457
105
607
140
(Ib/hr)
(712)
(166)
(1796)
(414)
(3624)
(830)
(4812)
(1106)
Volume Production
cc/sec
69
14
174
35
352
69
467
92
(ft3/hr)
(8.8)
(1.8)
(22.1)
(4.4)
(44.6)
(8.8)
(59.3)
(12)
-------
B. Development status—Lime/limestone FGD systems have been commer-
cially demonstrated at several utility and industrial boiler installations
in the U.S., Japan, and several other countries. The wide use of lime/
limestone FGD systems is due primarily to the fact that these systems are
the most technically advanced and generally the most economically attractive,
at least for utility situations.
Lime/limestone scrubbing was first used to control S02 emissions on
commercial boiler pilot plants in England about 40 years ago. This led to
full-scale gas washing plants which proved successful in removing S02 and
dust from stack gas. The lime/limestone process was also the first S02
stack gas scrubbing technology used in this country, mainly due to the fact
that there was more experience behind the process and that it produced a
throwaway sludge rather than a marketable by-product which could require a
considerable marketing effort in many instances to dispose of. The trend
to using lime/limestone scrubbing for SOz removal is strong today due to
rapid progress in coping with the many process problems and a clearer under-
standing of process economics.
1) Commercial applications - The commercial use of lime/limestone
systems for coal-fired boilers is addressed in three categories:
' Utility boilers in the U.S.
Industrial boilers in the U.S.
' Foreign applications
Utility installations - Lime/limestone systmes are the most widely used form
of FGD selected by the U.S. electric utility industry. As of March, 1978,
some 34,000 MW of coal-fired electrical generating capacity in the United
States had been commited to lime/limestone scrubbing. This figure includes
28 facilities in operation, 35 under construction, and another 16 in the
planning stages (i.e., contract awarded, letter of intent signed, or
requesting/evaluating bids).
2-15
-------
Wide use of lime/limestone scrubbing is true for both new and retrofit cases.
As shown in Table 2.2.1-3, 94 percent of the total MW of controlled capacity
of new systems currently operational is by lime/limestone processes. For
operational retrofit situations, 84 percent of the total capacity is
controlled by lime/limestone. Data are also summarized for units under
construction and in the planning stages. Operating lime/limestone scrubbing
units on United States power plants are shown in Table 2.2.1-4, units under
construction in Table 2.2.1-5, and planned units in Table 2.2.1-6. It might
also be noted that an additional 11 lime/limestone systems which were opera-
tional within the last five years are currently in a "terminated" state for
one reason or another.
Industrial installations - Industrial usage of lime/limestone FGD systems
in the U.S. has been quite limited. As of January 1979 there were two opera-
tional and two planned lime/limestone system with a combined capacity of
489,000 scfm. This represents less than 5 percent of the total controlled
(operational, under construction, and planned) U.S. industrial boiler capa-
city. The lime/limestone systems are summarized in Table 2.2.1-7.
One of the most significant lime/limestone applications is the Bahco sys-
tem installed at the Rickenbacker Air Force Base (RAFB) in Columbus, Onio.
he RAFB installation is the first Bahco system installed in the United
States. It is designed to treat flue gases from an equivalent coal-firing
rate of 58.6 MW (200 x 105 Btu/hr) under full load conditions as well as
flue gases from summer load conditions, about 5.8 MW (20 x 106 Btu/hr).
EPA has recently completed a two year test program on this system. Although
no continuous monitoring data are available from these tests, the discrete
data sets from the various factorial runs showed the system to consistently
exceed design specifications. Data on process operations for this system
will be presented in the appropriate sections of this chapter.19
2-16
-------
TABLE 2.2.1-3. SUMMARY OF NEW AND RETROFIT FGD SYSTEMS FOR U.S. UTILITY INDUSTRY BY PROCESS
1t
I
I—1
-~J
Process
Lime
Lime/Alkaline Fly Ash
Lime/Limes tone
Limestone
Subtotal - Lime/Limestone
Aqueous Carbonate
Aqueous Carbonate/Fab. Filter
Double Alkali
Magnesium Oxide
Not Selected
Regenerable Not Selected
Sodium Carbonate
Wellman Lord
Wellman Lord/Allied Chemical
Totals
Lime/Limestone % of Total MW
New or
retrofit
N
R
N
R
N
R
N
R
N
R
N
R
N
R
N
R
N
R
N
R
N
R
N
R
N
R
N
R
N
R
N
R
Operational
No.
It
8
3
0
0
2
8
3
15.
13.
0
0
0
0
0
0
0
1
0
0
0
0
1
2
0
0
1
1
17.
17.
MW
2450
1650
1170
0
0
20
4443
790
8063.
2460.
()
0
0
0
0
0
0
120
0
0
0
0
125
250
0
0
375
115
8563.
2945.
94
84
Construction
No.
10
0
1
0
0
0
23
1
34.
1.
0
0
0
0
2
1
0
0
0
0
0
0
1
0
1
1
0
1
38.
4.
MW
4565
0
500
0
0
0
9620
425
14685.
425.
0
0
0
0
825
277
0
0
0
0
0
0
509
0
500
180
0
340
16519.
1222.
89
35
Contract
awarded
No.
2
0
2
0
0
0
9
1
13.
1.
0
1
1
0
0
0
0
0
1
0
0
0
0
0
0
0
0
0
15.
2.
MW
1425
0
1400
0
0
0
4783
575
7608.
575.
0
100
400
0
0
0
0
0
300
0
0
0
0
0
0
0
0
0
8308.
675.
92
85
Planned
No.
0
2
1
3
0
0
5
0
6.
5.
0
0
0
0
0
0
0
3
18
4
0
1
1
0
1
0
0
0
26.
13.
MW
0
660
527
579
0
0
2880
0
3407.
1239.
0
0
0
0
0
0
0
726
9500
2100
0
650
125
0
500
0
0
0
13532.
4715.
25
26
Total No.
of plants
No.
16
10
7
3
0
2
45
5
68.
20.
0
1
1
0
2
1
0
4
19
4
0
1
3
2
2
1
1
2
96.
36.
MW
8440
2310
3597
579
0
20
21726
1790
33763.
4699.
0
100
400
0
825
277
0
846
9800
2100
0
650
759
250
1000
180
375
455
46922.
9557.
72
49
-------
TABLE 2.2.1-4. SUMMARY OF OPERATIONAL LIME/LIMESTONE FGD SYSTEMS FOR U.S. UTILITIES
AS OF MARCH 1978 15
ro
I
oo
New or
Process Utility company/power station retrofit
Lime Columbus & Southern Ohio Elec. N
Conesvllle No. 5
Columbus & Southern Ohio Elec. N
Conesville No. 6
Duquesne Light R
Phillips Power Station
Duquesne Light R
Elrama Power Station
Kansas City Power & Light R
Hawthorn No. 3
Kansas City Power & Light R
Hawthorn No . 4
Kentucky Utilities R
Green River, Units 1, 2, and 3
Louisville Gas & Electric R
Paddy's Run No. 6
Louisville Gas & Electric R
Cane Run No. 4
Louisville Gas & Electric R
Cane Run No. 5
Minnkota Power Coop. N
Milton R. Young No. 2
Montana Power Company N
Size of FGD
unit (MW)
400
400
410
510
140
100
64
65
178
183
450
360
Start up Years
Vendor mo/yr Operational
UOP
UOP
Chemico
Chemico
Combustion Engineering
Combustion Engineering
American Air Filter
Combustion Engineering
American Air Filter
Combustion Engineering
AOL/Combustion Equipment Assoc .
ADL/Combustlon Equipment Assoc.
1/77
4/78
7/73
10/75
11/72
8/72
9/75
4/73
8/76
12/77
9/77
7/76
1
%
5
3
6
6
3
5%
2
1
1
2
Colstrip No. 2
-------
TABLE 2.2.1-4. (Continued)
t-o
I
Process Utility company/power station
Lime Montana Power Company
(Cont'd) Colstrip No. 1
Pennsylvania Power Co.
Bruce Mansfield No. 1
Pennsylvania Power Co.
Bruce Mansfield No. 2
Limestone Arizona Public Service
Cholla No. 1
Indianapolis Power & Light Co.
Petersburg No. 3
Kansas City Power & Light
LaCygne No. 1
Kansas Power & Light
Lawrence No. 4
Kansas Power & Light
Lawrence No . 5
Northern States Power Co.
Sherburne County Station No. 1
Northern States Power Co.
Sherburne County Station No. 2
South Carolina Public Service
Winyah No. 2
Springfield City Utilities
New or
retrofit
N
N
N
R
N
N
R
N
N
N
N
N
Size of FGD
unit (MW)
360
825
825
115
530
820
125
400
710
710
280
200
Vendor
ADL/Combustion Equipment Assoc.
Chemico
Chemico
Research-Co ttrell
UOP
Babcock. & Wilcox
Combustion Engineering
Combustion Engineering
Combustion Engineering
Combustion Engineering
Babcock & Wilcox
UOP
Start up Years
mo/yr operational.
11/75
4/76
7/77
10/73
10/77
2/73
12/68
11/71
3/76
4/77
7/77
4/77
3
2%
1
5
1
5%
10
7
2%
1%
1
1%
Southwest No. 1
-------
TABLE 2.2.1-4. (Continued)
New or Size of FGD
Process Utility company/power station retrofit unit (MW) Vendor
Limestone Tennessee Valley Authority R 550 TVA
(Cont'd) Widows Creek No. 8
Texas Utilities Co. N 793 Research Cottrell
Martin Lake No. 1
Lime/ Tennessee Valley Authority R 10 Chemico
Limestone Shawnee No. 10B
Tennessee Valley Authority R 10 UOP
Shawnee No. 10A
Start up
mo/yr
5/77
8/77
tt/12
It/12
Years
operational
1
1
6%
6%
ho
1
-------
I
to
TABLE 2.2.1-5. SUMMARY
FOR U,
OF LIME/LIMESTONE SYSTEMS UNDER CONSTRUCTION
S. UTILITIES AS OF MARCH 197816
Process Utility company/power station
Lime Allegheny Power System
Pleasants No. 1
Allegheny Power System
Pleasants No. 2
Big Rivers Electric Coop Corp.
Reid Steam Station No. 2
Big Rivers Electric Coop Corp.
Reid Steam Station No. 3
Cooperative Power Association
Coal Creek No. 1
Cooperative Power Association
Coal Creek No. 2
Louisville Gas & Electric
Mill Creek No. 3
Louisville Gas & Electric
Mill Creek No. It
Minnesota Power & Light Co.
Clay Boswell No. 4
Utah Power & Light Co.
Emery No . 1
Utah Power & Light Co.
Huntington No. 1
Limestone Alabama Electric Coop
New or
retrofit
N
N
N
N
N
N
N
N
N
N
N
N
Size of FGD
unit (MW)
625
625
250
240
545
545
425
495
500
400
415
225
Vendor
Babcock & Wilcox
Babcock & Wilcox
American Air Filter
American Air Filter
Combustion Engineering
Combustion Engineering
American Air Filter
American Air Filter
Peabody Engineering
Chemico
Chemico
Peabody Engineering
Start up
mo/yr
3/79
3/80
12/79
12/80
2/79
11/79
7/78
6/80
5/80
12/78
5/78
6/78
Tombigbee No. 2
-------
TABLE 2.2.1-5. (Continued)
I
ro
Process Utility company/ power station
Limestone Alabama Electric Coop
(Cont'd) Tombigbee No. 3
Arizona Electric Power Coop
Apache No. 3
Arizona Electric Power Coop
Apache No. 2
Arizona Public Service
Cholla No. 2
Basin Electric Power Coop
Laramle River No. 1
Basin Electric Power Coop
Laramie River No. 2
Board of Municipal Utilities
Sikeston Power Station
Brazos Electric Power Coop
San Miguel No. 1
Central Illinois Light Co.
Duck Creek No. 1
Colorado Ute Electric Assn.
Craig No. 1
Colorado Ute Electric Assn.
Craig No. 2
Commonwealth Edison
New or
retrofit
N
N
N
N
N
N
N
N
N
N
N
R
Size of FGD
unit (MW)
225
200
200
250
550
550
235
400
400
450
450
425
Vendor
Peabody Engineering
Research Cottrell
Research Cottrell
Research Cottrell
Research Cottrell
Research Cottrell
Babcock & Wilcox
Babcock & Wilcox
Riley Stocker/Environeering
Peabody Engineering
Peabody Engineering
UOP
Start up
mo/yr
6/79
4/79
6/78
6/78
4/80
10/80
6/81
6/80
8/78
3/79
3/79
12/79
Powerton No. 51
-------
TABLE 2.2.1-5. (Continued)
NO
I
ro
Process Utility company /power station
Limestone Kansas Power & Light
(Cont'd) Jeffery No. 1
Kansas Power & Light
Jeffery No. 2
Salt River Project
Coronado No. 1
Salt River Project
Coronado No. 2
Southern Illinois Power Coop
Marion No. 4
Southern Mississippi Electric
R. D. Morrow No, 1
Southern Mississippi Electric
R. D. Morrow No. 2
Texas Power and Light Co.
Sandow No. 4
Texas Utilities Co.
Martin Lake No. 2
Texas Utilities Co.
Martin Lake No. 3
Texas Utilities Co.
New or
retrofit
N
N
N
N
N
N
N
N
N
N
N
Size of FGD
unit (MW)
680
680
350
350
184
180
180
545
793
793
750
Vendor
Combustion Engineering
Combustion Engineering
Pullman Kellogg
Pullman Kellogg
Babcock & Wilcox
Riley Stoker/Environeering
Riley Stoker/Environeering
Combustion Engineering
Research Cottrell
Research Cottrell
Chemico
Start up
mo/yr
6/78
6/80
4/79
4/80
6/78
5/78
8/78
7/80
2/78
12/78
2/78
Monticello No. 3
-------
TABLE 2.2.1-6. SUMMARY OF PLANNED LIME/LIMESTONE FGD SYSTEMS AS OF 03/78
1 7
Process
Limes tone
Lime
Limestone
Limestone
Limestone
l-o
M Lime/
£- Alkaline
Fly Ash
Lime/
Alkaline
Fly Ash
Limestone
Limestone
Lime
Limestone
Utility company/power station
Contract Awarded
Arizona Public Service
Cholla No. 4
Cincinnati Gas i Electric Co.
East Bend No. 2
Hoosier Cooperative
Merom No. 1
Hoosier Cooperative
Merom No. 2
Lakeland Utilities
Mclntosh Power Plant Unit No. 3
Montana Power Co.
Colstrip No. 3
Montana Power Co,
Colstrip No. 4
Northern States Power Co.
Sherburne County Station No. 3
Northern States Power Co.
Sherburne County Station No. 4
Pennsylvania Power Co.
Bruce Mansfield No. 3
Springfield Water Light & Power
New of
retrofit
N
N
N
N
.N
N
N
N
N
N
Size of FGD
unit (MW)
350
600
490
490
350
700
700
860
860
825
190
Vendor
Research Cottrell
Babcock & Wllcox
Not selected
Not selected
Not selected
ADL/Combustion Equip. Associate
ADL/Combustion Equip. Associate
Combustion Engineering
Combustion Engineering
Pullman Kellogg
Research Cottrell
Start up
mo/yr
6/80
1/81
12/80
10/81
10/81
7/80
7/81
5/81
5/83
4/80
7/80
Dallman No. 3
-------
TABLE 2.2.1-6. (Continued)
Process
Limestone
Limestone
Limestone
Lime/
Alkaline
Fly Ash
Utility company/power station
Tennessee Valley Authority
Widows Creek No. 7
Texas Municipal Power Agency
Gibbons Creek Unit No. 1
Texas Utilities Co.
Martin Lake No. 4
Letter of Intent Signed
Wisconsin Power & Light Co.
Columbia No. 2
Requesting/Evaluating Bids
New or Size of FGD
retrofit unit (MW) Vendor
R 575 Combustion Engineering
N 400 Combustion Engineering
N 793 Research Cottrell
N 527 Chemico
Start up
mo/yr
1/82
11/82
1/80
Limestone Indianapolis Power & Light Co.
Petersburg No. 4
530
Not Selected
4/82
-------
TABLE 2.2.1-7
SUMMARY OF COMMITTED LIME/LIMESTONE SYSTEMS FOR U.S. INDUSTRIAL BOILERS
AS OF MARCH 197818
Process
Lime
Lime
Lime/
Limestone
Limestone
Vendor
Koch
Engineering
Carborundum
Environ.
Systems, Ltd.
Research
Cottrell-Bahco
Company/ New or Size of FGD unit
Location retrofit s.cfm MW
Armco Steel R 84,000 42
Middle town, OH
Carborundum R 30,000 15
Abrasives
Buffalo, NY
Rickenbacker R 55,000 27
Air Force Base
Columbus , OH
Dupont N 320,000 160
Texas
Startup
characteristics mo/yr
Coal 0.8% sulfur 1975
Coal 2.2% sulfur 1980
Coal 3.6% sulfur 1976
Coal 0.5% sulfur 1982
-------
Foreign installations - Foreign applications of lime/limestone scrubbing
also serve to illustrate its commercial availability. The most widespread
use of this technology is in Japan, with a few facilities also in Germany,
Sweden, Russia, and England.
Japanese wet lime/limestone FGD installations by the end of 1977 totalled
94 plants with a combined flue gas capacity of 43.4. X 106 Nm3/hr, equivalent
to about 13,000 MW or about 50% of all Japanese FGD capacity. This total
e
includes industrial and utility boilers, sintering plants and sulfuric acid
plants.20 According to a report for EPA prepared by Dr. Jumpei Ando, lime/
limestone FGD applications are distributed as shown in Table 2.2.1-8. The emph-
asis in Japan is on oil-fired boilers. Practically all the processes make
? 9
salable gypsum, and "operabilities" are reported to exceed 95 percent.
2) Recent Improvements - Lime/limestone FGD technology, as sum-
marized in this report, can be regarded as established and commercially
available. However, attempts are constantly being made to improve the
overall performance of the technology and to increase its acceptability.
Current research in the areas of mass transfer additives and cocurrent
scrubbing offer the potential for improved process performance and cost
savings, and ongoing work in the area of forced oxidation has the potential
for improving the process solid waste characteristics thus increasing the
overall acceptability of the technology. Each of these areas of recent
developments is briefly described in this section.
Mass transfer additives - Additives such as MgO, NaaCOs, and organic acids
can be used to improve mass transfer in lime/limestone systems, thus improv-
ing SOa removal, increasing limestone utilization, and permitting the use of
2-27
-------
TABLE 2.2.1-8. SUMMARY OF JAPANESE LIME/LIMESTONE INSTALLATIONS
21
Type
Installation
Utility Boiler
Utility Boiler
Industrial Boiler
Industrial Boiler
Chemical Plant
Fuel
Oil
Coal
Oil
Coal
Total
Capacity
106Nm3/hr
24.
3.
3.
0.
11.
43.
1
6
7
8
2
4
MWe
7530
1125
1150
250
3500
13555
Percent of
Total Lime/
Limestone
55.
8.
8.
1.
25.
100
5
3
6
8
8
Capacity
2-28
-------
more simple scrubber types (such as cocurrent scrubbers). S02 removal can be
increased from 85 to 95 percent through use of these additives.2 The im-
provement in limestone utilization also improves process operability, since
excess limestone (which can cause plugging of the scrubber mist eliminator)
is minimized. The use of additive also enhances the effects of dissolved
alkalinity by increasing the concentrations of HC03 and S03 which react
with S02 as shown below:
HC03~ + S02 + C02 + HS03~ 2.2.1-5
S03= + S02 + H20 -> 2 HS03~ 2.2.1-6
Alkali additives, MgO, NaaCOs , and NH3 , have received the most attention
to date. MgO has been tested extensively with lime and limestone scrubbing
in both pilot and full scale systems. Na2C03 and NH3 have been used in
double alkali processes and have been patented as additives for limestone-
+ 2+
slurry scrubbers. The effects of Na and Mg addition on total dissolved
alkalinity and improved scrubber performance have been calculated and
reported by Rochelle.24 It is believed that NH3 probably behaves similarly
to Na . When the soluble alkali carbonates or oxides are added to a slurry
scrubbing system they accumulate in solution primarily as the soluble sulfate
Chloride concentrations can reduce the apparent effectiveness of
alkali additives by reacting with the additive to form inactive chloride
salts. For a given alkali concentration, dissolved chloride reduces ac-
cumulated sulfate ion by 1 mole of sulfate per 2 moles of chloride. Con-
sequently, there is little effect on dissolved alkalinity until an amount
of alkali has been added equivalent to the chloride concentrations. In
some scrubbing installations the alkali requirement will be dictated by
the chloride level and most of the alkali will be present as the inactive
chloride.
Buffer additives are also being evaluated for their effectiveness
in enhancing the liquid phase mass transfer of limestone scrubbing systems,
2 5
tion. , and recent testing by EPA has demonstrated the effectiveness of
Several carboxylic acids have been proposed by Rochelle for this applica-
h;
adipic acid as a buffer additive.26
2-29
-------
Recent results of EPA's Shawnee Test Facility, reported at the
fifth FGD Symposium indicated that adipic acid concentrations of 700 to
1500 ppm increased S0~ removals of the system to greater than 90 percent
while decreasing the annual revenue requirements by 0.3 to 0.4 mill/kwh.
Although this system has experienced decomposition of the adipic acid ad-
ditive, the overall economics are still favorable. Long-term effects of
organic buffers on sludge formation and stability are still under investi-
2 7
gation.
Concurrent Scrubbing - Cocurrent scrubbers differ from standard counter
current scrubbers in that both the gas to be scrubbed and the scrubbing
liquor enter the top of the scrubber and flow in the same direction through
the absorber. Advantages of this type scrubber are that higher gas vel-
ocities and lower pressure drops occur than the conventional countercurrent
scrubbers which can result in lower capital and operating costs. Flue gas
from the bottom of the scrubber experiences a change in direction to "knock-
out" the majority of entrained scrubbing liquor before going to a mist
eliminator. The actual scrubber may use sprays or packing for gas/liquid
contact.
Higher gas velocities through the scrubber are possible; therefore,
smaller diameter absorbers may be used. Since spray towers are generally
designed for length to diameter (L/D) ratios of about 2 for good gas dis-
tribution, cocurrent operations are attractive from a liquid pumping height
standpoint. Reduced pressure drop will result in fan cost savings and in
lower energy requirements and operating costs.
In addition, the cocurrent scrubber concept should reduce the effect
of plugged nozzles on poor gas distribution. In countercurrent designs
2-30
-------
the gas tends to flow where the pressure drop is lowest (i.e., in the area
of a plugged nozzle) which results in reduced mass transfer. However, in
the cocurrent scrubber, the liquid flow imparts additional energy to the gas
so gas will not tend to flow toward plugged nozzles. Instead, the gas flow
should be higher in regions where the liquid flow is higher.
Forced oxidation - Forced oxidation systems are presently being marketed
although further development work is being performed by EPA and others. As
the name implies, forced oxidation systems increase the amount of sulfite
which is oxidized to sulfate. This is generally done by sparging air into
the hold tank. Potential advantages include improved waste solid character-
istics and better limestone utilizations. If oxidations approaching 100
percent are achieved, the gypsum waste solid may be suitable for wallboard
2. 8
production.
Gypsum solids are more easily dewatered than sludges with a high
concentration of calcium sulfite since the gypsum solids are larger, denser,
and have a faster settling rate. Gypsum solids have been shown to be suitable
for landfill disposal. A liquor purge may be necessary to control soluble
species in systems that have waste products with a high solids content, but
the purge would be no larger than the amount of liquor discharged with the
r\ Q
waste solids from a system that does not use forced oxidation.
One disadvantage of forced oxidation is the elimination of CaSOs solids
so that they are not available for mass transfer. The use of additives to
enhance mass transfer may thus be attractive when used in conjunction with
forced oxidation. Since forced oxidation also minimizes additive losses in
the waste solids, forced oxidation and mass transfer additives appear to be
particularly attractive when used together.30
C. Applicability—A major factor to be considered when assessing a
lime/limestone system's applicability is available land. The area required
for sludge disposal can be especially important when the industrial plant
is located in a densely populated area such as the East Coast. Estimates of
sludge disposal requirements were presented previously in Section 2.2.I.I.C.
2-31
-------
In some cases the plant might be required to pay for off-site transportation
and disposal. Disposal problems and costs can be limiting factors for both
new and retrofit applications.
Sorbent availability should not limit process applicability since
limestone is readily available material as is reflected by its low price of
approximately $7/ton.31 Sorbent price, however, will be dependent on trans-
portation costs. Production of limestone in the United States was approxi-
mately 630 million tons in 1970. Since limestone is mined in many areas of
the U.S., it will be reasonably close to many industrial boiler installations.
Lime is produced by calcining high-calcium limestone in kilns at about
1800-2300°F to drive off C02 and form CaO. The heat of reaction required
to convert CaC03 to CaO is about 2.8 x 106 Btu/ton of CaO. In practice,
heat input may vary from 4 to 10 x 106 Btu/ton of lime. Total lime consump-
tion by the United States in 1975 (including limes processed within captive
facilities) was 19.2 million tons.32 Over 80 percent of the lime used in
the United States is by chemical and industrial manufacturing, mostly as
quicklime. Since limestone is used to make lime and is readily available,
lime also can be obtained fairly easily. The cost of lime is presently
about $40/ton 33 with a large portion of the cost being due to the fuel
costs associated with lime production.
As presented in the preceding section, there are considerably more
applications of lime/limestone scrubbing in the utility sector than in the
industrial sector both in this country and abroad. However, the process
is applicable to small boiler operations as well. Considerations affecting
retrofittability and by-product utilization for industrial applications are
similar to those for utility applications.
D. Factors affecting performance—
1) Design and operating considerations - Although lime/limestone
scrubbing systems tend to be somewhat simple mechanically, they are more
complex from a chemical process point of view. The removal of SOa from flue
gas in a lime/limestone scrubbing process is a gas-liquid-solid mass trans-
2-32
-------
fer phenomenon. First, sulfur dioxide must be transferred from the gas phase
to the liquid phase in a scrubber. Then the sulfur species must be precipi-
tated from the scrubbing liquor as insoluble calcium salts and disposed of.
Most lime/limestone systems tend to be at least partially liquid film
mass transfer limited. Only at low 862 gas loadings do these systems
approach gas film limiting situations. Liquid film resistance can be reduced
by increasing the amount of liquid phase alkalinity available in the scrubber.
As a rule, a large portion of the alkalinity required for SOz removal is
derived from solids dissolution in the scrubber. Since solid-liquid reac-
tions tend to be significantly slower than do liquid-liquid reactions, it is
advantageous to minimize the amount of solids dissolution required by maxi-
mizing the soluble liquid phase alkalinity in the scrubber feed liquor.
This is why systems which operate with high magnesium and sodium concentra-
tions but low chloride levels exhibit higher SOa removals than systems which
are lower in soluble alkalinity.3 "*
Gas distribution can also be a major problem, particularly in large
commercial units. Poor gas distribution will minimize the effective inter-
facial mass transfer area. Portions of the scrubber can become liquid phase
alkalinity limited even though the total alkalinity entering the scrubber is
sufficient for good SOa removal. Thus, the potential impact of gas distri-
bution problems should be seriously considered in both process design and
analysis.
In addition to mass transfer and gas distribution, several other
design and operating variables should be considered in the design of a
lime/limestone FGD system. The effect of the following variables, both on
SOz absorption efficiency and on overall process operations are summarized
in this section:
• L/G Ratio
• Slurry pH
" Inlet S02 Concentration
2-33
-------
• Reaction Tank Design
• Effects of Soluble Species
• Number of Contact Stages
• Ash Removal
• Oxidation and Scale Prevention
• Water Balance
• Reheat
• Sludge Dewatering and Disposal
• Erosion and Corrosion
L/G ratio - For both lime and limestone systems higher SO/ removal efficiencies
are achieved at greater L/G ratios up to the point where flooding and poor gas
distribution occurs. This trend is documented by performance data from
several different scrubber operations. Data from the Shawnee test site which
represents an eastern high-sulfur coal (i.e. , high inlet SOz concentration)
is shown in Figures 2.2.1-2 through 2.2.1-5. Results from a western plant
are presented in Figures 2.2.1-6 and 2.2.1-7- While these data represent
numerous scrubber types and operating conditions as well as lime and lime-
stone systems, the increasing trend of SOa removal with increasing L/G is
evident in all.
Higher L/G's can be obtained by either increasing the liquor flow or reducing
the gas flow through the scrubber. The first approach entails more pumping
capacity, pipes, and slurry. The latter approach requires additional scrub-
bing capacity. The first method, illustrated in Figures 2.2.1-2 through
2.2.1-7, is more commonly used by designers for cost reasons. Shawnee spray
tower data illustrate the effects of decreasing the gas velocity (Figure
2.2.1-8). Decreasing the gas velocity is much less effective for a packed
or turbulent type contactor than a spray tower because increased agitation
and mixing offset the decreased L/G ratio for these contactors.
Slurry pH - At higher pH levels (i.e. , greater lime or limestone feed stoich-
iometry), increased S0? removal efficiency occurs because more alkali is available
and higher dissolution rates are achieved. However, scaling will occur if
the pH becomes too high. Therefore, careful pH monitoring and control is
2-34
-------
LIQUID TO GAS RATIO, 1/mJ
TOO
10
SCRUBBER INLET pH FOR
FACTORIAL TESTS
90
t 80
ce
o
oj
o
o
o:
70
60
50
40
20
I
"SPRAY TOWER GAS VELOCITY = 2.3m/s(7.5 ft/sec)
SO, INLET CONCENTRATION = 2,500 - 3,500 ppm
EFFECTIVE LIQUOR Mg++CONCENTRATION = 0 ppm
LIQUOR Ci'CONCENTRATION = 8,000 - 13,000 ppm
30 40 50 60
LIQUID TO GAS RATIO, gal/mcf
70
80
Figure 2.2.1-2.
Liquid-to-gas ratio and scrubber inlet pH
versus predicted and measured SOa removal,
spray tower with lime, Shawnee plant.
2-35
-------
LIQUID TO GAS RATIO,l/mj
TOO
10
UJ
O
90
80
7°
60
50
40
TCA GAS VELOCITY FOR
FACTORIAL TESTS
O 3.8 m/S (12.5 ft/sec)
A 3.2 m/S (10.4)
D 2.5 m/S (8.3)
SCRUBBER INLET pH - 7.9 - 8.1
SO, INLET CONCENTRATION = 2,200 - 2,800 ppm
TOTAL HEIGHT .OF SPHERES = 0.38 m (15.0 in)
EFFECTIVE LIQUOR Mg++CONCENTRATION = 0 ppm
20 30 40 50 60
LIQUID TO GAS RATIO, gal/mcf
70
80
Figure 2.2.1-3.
Liquid-to-gas ratio and scrubber gas velocity
versus predicted and measured S02 removal,
TCA with lime, Shawnee plant."
3C
2-36
-------
100
90
80
L/6 1/iT
6 7
10
oc
CM
ee
ui
Q.
70
50
40
30
SCRUBBER INLET pH
• pH-5.8
O pH-5.7-5.9
D pH-5.4-5.6
a pH«5.1-5.3
LONG-TERM TEST
FACTORIAL TESTS
FACTORIAL TESTS
FACTORIAL TESTS
SCRUBBER GAS VELOCITY 3.17 m/s (10.4 ft/sec'
TOTAL HEIGHT OF SPHERES=38.1 cm (15.0 in.)
EFFECTIVE LIQUOR Mg"1"1- CONCENTRATIONS ppm
INLET SO? CONCENTRATION=2,400-2,900 ppm
LIQUOR CiCONCENTRATION=3,000-7,000 ppm
20 30 40 50 60
L/6, gal/mcf
70
80
Figure 2.2.1-4.
L/G ratio and scrubber inlet pH versus predicted
and measured SOz removal - TCA with limestone -
3 7
Shawnee plant.
2-37
-------
L/G,l/inJ
6 7
10
100
90
80
CM
s?
70
60
50
40
( I I i
EFFECTIVE LIQUOR Kg1"1" CONCENTRATION
FACTORIAL TESTS
O 7,000-10,000 ppm
D 3,500-5,500 ppm
& 0-500 ppm
SCRUBBER GAS VELOCITY-3.17 m/s (10.4 ft/se&i
TOTAL HEIGHT OF SPHERES-38.1 cm (15.0 1n.)
SCRUBBER INLET pH=5.4-5.6
INLET S02 CONCENTRATION=2,200-2,800 ppm
LIQUOR Ci" CONCENTRATION-6,000-16,000 ppm
20 30 40 50 60
L/6 RATIO, GAL/MCF
70
80
Figure 2.2.1-5.
L/G ratio versus percent S02 removal at various
magnesium ion concentrations TCA with limestone
Shawnee plant.38
2-38
-------
L/G
100
95
90
85
80
70
60
50
40
30
20
10
0.382; SULFUR COAL
10 15 20
L/G (gpm/1000 SCFM)
25
30
Figure 2.2.1-6. S02 removal versus L/G ratio,
170-MW horizontal module, lime sorbent,
Mohave plant.
[L/G's are for 1-stage of a 4-stage scrubber.
Total L/G's are 4 times the indicated value.
2-39
-------
100-
99-
98-
o
Si 97-
o
01
< 96-
O
ui
DC
CM 95-
UJ
O 94-
93-
92-
91-
90-
TCA 4 STAGES
LIMESTONE
INLETS02 = 200PPM
GAS RATE = 450.000 SCFM
(Redrawn from Reference 5)
60
LiQUID TO GAS RATIO, gal/Mcf
Figure 2.2.1-7.
Effect of liquid-to-gas ratio on SOz removal efficiency
with low sulfur coal at the Mohave power station. If0
2-40
-------
100
90
80
ee
(SI
£70
60
50
40.
D
SLURRY FLOW RAVE FOR
FACTORIAL TESTS
O 6.2 l/sm2(30 gal/min - ft2)
4.66 1/sm, (22.5)
11 1/sirT (15)
OJ
SCRUBBER INLET pH - 7.9 - 8.1
SO, INLET CONCENTRATION = 2,500 - 3,500 ppm
EFFECTIVE LIQUOR Mg CONCENTRATION = 0 ppm
LIQUOR Ci'CONCENTRATION = 8,000 - 13,000 ppm
I
789
SPRAY TOWER GAS VELOCITY, ft/s
10
11
Fieure 2.2.1-8. Gas velocity and slurry flow rate versus
& _.,i _ .1 ri *~* ™,-.. ^i-,-r»-iT ci T»T~ «m
predicted and measured S02 removal, spray
tower with lime, Shawnee plant.
2-41
-------
needed to optimize S02 removal, avoid scaling, and prevent low utilization
of excess reagent.
Lime is much more reactive than limestone. In most cases, complete dissolu-
tion of lime occurs within 20-30 seconds of 'the time that the lime is intro-
duced into the hold tank. Since lime is not a buffered alkali, higher pHs
are maintained in lime systems compared to limestone systems. However,
the combination of reduced reaction time and increased pH can lead to Ca(OH)a
solids entering the scrubber in the feed slurry. This should be avoided
since the dissolution of these lime solids will occur in the low pH environ-
ment of the scrubber. This, in turn, will result in locally high pHs and
high relative saturations of calcium sulfite and sulfate. These conditions
can lead to scale deposition on scrubber internals.
If the hold tank pH in a lime system is maintained below about 8, lime utili-
zations in the 95 percent range are possible. "t2 Lime systems operating in
the optimal pH range for high lime utilization can operate subsaturated
with'respect to calcium sulfite in the scrubber. In fact, since calcium
sulfite dissolution is normally noted in lime scrubbers, it can be concluded
that calcium sulfite scaling is not likely to be a problem, except during
upset conditions.
Lowering a limestone system's slurry pH will increase the limestone dissolu-
tion rate and improve the utilization efficiency, but the resulting pH drop
can lead to decreased S02 removal efficiencies. "*3' "*"* An explanation for this
is found by considering the dissolution step for CaC03.
CaC03(s) j CaC03(aq) 5 Ca4^^ + C07(aq) (2.2.1-.11)
The carbonate ion can then be reacted as follows:
COl + H+ £ HCOl (2.2.1-12)
HCOl + H+ £ H2C03 (2.2.1-13)
2-42
-------
The slurry pH in a limestone system remains buffered according to these
equations. This accounts for the fact that the slurry pH is relatively
insensitive to changes in the amount of dissolved limestone (and therefore
dissolved alkalinity). Because large changes in liquid phase alkalinity can
occur over fairly narrow pH shifts, S02 removal is generally a very sensitive
function of the system pH in a liquid film limiting situation.
This buffering capacity of limestone also means that to achieve higher
scrubber inlet pH's, increasingly larger amounts of limestone must be added
to the hold tank. This results in larger amounts of limestone remaining
undissolved in the hold tank effluent and higher limestone solids concen-
trations in the scrubber feed liquor. The amount of solid CaCOa available
for reaction in the scrubber can have a major effect on S02 removal. As the
concentration of CaCOa increases, generally, the surface area also increases.
This can lead to increased dissolution. From this standpoint, an increase
in scrubber pH in a limestone system will increase the removal for two
reasons: 1) because it increases the amount of dissolved liquid phase
alkalinity and 2) because it increases the amount of limestone solids in the
circulating slurry. Attention must be given to the amount of CaCOa which
dissolves in the scrubber since locally high relative saturations in the
vicinity of dissolving limestone particles can lead to scaling.
The effect of inlet pH on S02 removal in two different Shawnee lime scrubbers
is shown in Figures 2.2.1-9 and 2.2.1-10. (Note that the effect of L/G
discussed above can also be seen from these data.) Typical control points
for a lime system are in the 8-9 range. Limestone systems, on the other
hand, are normally operated at lower pH with typical control points in the
5-6 range. Scrubber inlet pH versus S02 removal are available for various
magnesium ion concentrations (Figure 2.2.1-11) and various L/G ratios
(Figure 2.2.1-12) for the Shawnee TCA scrubber. Both figures show increased
S02 removals with increasing pH over the pH 5-6 range.
Inlet S02 concentration - For fixed design and operating conditions, greater
S02 removal efficiencies are achieved at lower inlet S02 concentrations.
This occurs because the actual quantity of S02 removed per volume of gas is
2-43
-------
100
90-
80-
CM
3
8
Q.
60
50
40
LIQUID TO GAS RATIO
FOR FACTORIAL TESTS
O 9.1 l/mj (68 gal/1000
70-
l/m
j
(51)
ftj)
^
O 4.6 1/
SO, INLET CONCENTRATION * 2,500 - 3,000 ppm
EFFECTIVE LIQUOR (^CONCENTRATION - 0 ppm
LIQUOR CiTCONCENTRATION « 8,000 - 13,000 ppm
7 8
SCRUBBER INLET pH
10
Figure 2.2.1-9.
Scrubber inlet pH and liquid-to-gas ratio
versus predicted and measured S02 removal,
spray tower with lime, Shawnee plant. "*s
2-44
-------
100
90
80
a:
° 70
60
50
40
LIQUID TO GAS RATIO FOR
FACTORIAL TESTS
O 8.0 1/m3 (60 gal/1000 ft3)
A 6.0 1/nC (45 gal/1000 ft3)
D 4.0 1/m3 (30 gal/1000 ftj)
^
"SCRUBBER GAS VELOCITY =3.17 m/S(10.4 ft/sec
INLET SO, CONCENTRATION = 2,200 - 2,800 ppm
TOTAL HEIGHT OF SPHERES =38.1 mm (15.0 in)
EFFECTIVE LIQUOR (^CONCENTRATION » 0 ppm
I I I
7 '8
SCRUBBER INLET pH
10
Figure 2.2.1-10. Scrubber inlet pH and liquid-to-gas ratio
Versus predicted and measured
TCA with lime, Shawnee plant.
"*6
removal,
2-45
-------
100
90
80
70
60
50
40
EFFECTIVE LIQUOR Ngn CONCENTRATION
FACTORIAL TESTS
O 7,000-10.000 ppm
D3.500-5,500 ppm
A 0-600 ppm
SCRUBBER GAS VELOCITY-3.17 m/s,{10.4 ft/secT
LIQUID - TO - GAS RATIO- 6 t/n3 (45 gal/mcf)
TOTAL HEIGHT OF SPHERES- 38.1 cm. (15.0 In.)
.INLET S02 CONCENTRATION-2,300-2,700 ppm
LIQUOR CI* CONCENTRATION-12,000-16,000 ppm
5.0 5.2 5.4 5.6
SCRUBBER INLET pN
5.8
6.0
Figure 2.2.1-11.
Scrubber inlet pH versus percent SOz removal
at various magnesium ion concentrations
TCA with limestone - Shawnee plant. *7
2-46
-------
100
90
80
70
60
50
40
30
LIQUID - TO - GAS RATIO
FACTORIAL TESTS
O 8 t/m (60 gal/mcf)
D 6 I/ml (45 gal/mcf)
0 4 l/mj (30 gal/mcf)
SCRUBBER GAS VELOCITY»10.4 ft/sec
TOTAL HEIGHT OF SPHERES=15.0. in.
EFFECTIVE LIQUOR Mgtf CONCENTRATIONS ppm
INLET S02 CONCENTRATION=2,300-2,700 ppm
LIQUOR CI" CONCENTRATION=5,000-7,000 ppm
4.9 5.1 5.3 5.5 5.7
SCRUBBER INLET pH
5.9
6.1
Figure 2.2.1-12. Scrubber inlet pH versus SOa removal for three
L/G ratios -TCA unit with limestone - Shawnee plant.
2-47
-------
less, thereby reducing the load on the absorbent liquor. This is evidenced
by several sets of limestone scrubbing data from Shawnee obtained with the
TCA and spray tower scrubbers (Figure 2.2.1-13).
It was at one time thought that as inlet S02 concentrations decreased, the
driving force for absorption would decrease and high removal efficiencies
would be more difficult. It was felt that the equilibrium S02 back pressure
over the bulk liquor would approach the SOa gas concentration which would
limit absorption. However, in actual practice this is not the case since
S02 back pressure over fresh lime and limestone slurries is very low (about
1 ppm), and scrubbers can be designed so that the low SOa concentration
scrubbed gas exiting the scrubber is contacted by the freshest slurry.51For
these reasons, very high removal efficiencies have been achieved from flue
gas that averaged only 200 ppm at the inlet as shown in Figure 2.2.1-14.
Reaction tank design - The size of reaction tank can have a major impact
on the amount of limestone that dissolves in the tank. This, in turn,
affects the amount of dissolved alkalinity in the scrubber feed stream.
A given increase in volume will yield a greater percentage increase in
limestone utilization in a system with a shorter residence time tank than
it will in a system with a longer reaction time tank. Also of major impor-
tance is the particle size of the limestone. A decrease in particle size
will result in an increase in the limestone dissolution rate..52' 53
The operating conditions of the scrubbing system can also have a significant
impact on the effects of hold tank size changes. If the scrubber inlet pH is
lowered, the limestone dissolution rate will be enhanced and shorter residence
time tanks can be employed. At Shawnee, limestone utilization was not affected
by hold tank residence time changes when the scrubber inlet pH was maintained
below 5.8. At higher pH's, a 6-minute residence time tank yielded signifi-
cantly lower utilizations than did 12- and 20-minute residence time tanks.5k
Hold tank configuration has also been shown to have an effect on both lime-
stone utilization and SOa removal. Borgwardt's results indicate that plug
2-48
-------
100
O EPAPILOTTCA
SPHERE HEIGHT - 7 INCHES/BED. 3 BEDS
LIQUID TO - GAS RATIO = 85 gal/Mcf
TCA GAS VELOCITY - 7.5 ft/sec
TVA PILOT SPRAY TOWER
LIQUID TO GAS RATIO = 85 gal/Mcf
LIMESTONE SCRUBBING
70
1.000 2.000 3.000
INLET SO2CONC., ppm
4.000
5.000
Figure 2.2.1-13. Effect of inlet S02 concentration on S02 removal
efficiency for fixed design and operating conditions
2-49
-------
100-
99-
98-
Si 97H
o
< 96-
O
ui
EC
tM 95-J
UI
O 94-
ui
Q.
93-
92-
91-
90-
TCA 4 STAGES
LIMESTONE
INLETS02-200PPM
GAS RATE •= 450.000 SCFM
20 40
LIQUID TO GAS RATIO. gat/Mcf
60
Figure 2.2.1-14. Effect of liquid-to-gas ratio on SOa removal efficiency
with low sulfur coal at the Mohave power station.
50
2-50
-------
flow reaction tank designs can yield significant improvements in limestone
utilization and S02 removal.55 (A plug flow design is one that allows the
reacting stream to flow through the reactor such that there is no backmixing.
A plug flow situation can begin to be approximated by a number of mixed tanks
in series.) At Shawnee, at a constant limestone addition rate, the S02 removal
increased from 70% to 79% by changing the reaction tank from a single stirred
tank to three tanks of equal total volume in series. This plug flow effect
apparently drives the additive dissolution reaction further toward completion
and makes more liquid phase alkalinity available for reaction with absorbed S02 .
Effects of soluble species - One of the more important factors in determining
liquid phase alkalinity is the distribution of the dissolved ions. Some of
the species occurring in lime /limes tone systems are soluble and exit the
_i_ j _ i_ _
system only as dissolved solids. Important examples are Na , Mg , and C£ .
These soluble ions can enter' the system as NaO or MgO in the ash or HC£ in
the flue gas. They can also enter or be added to the system with the sorbent.
The soluble species determine to a large extent the levels of total calcium
_i _ i_ _ =
and total sulfur which remain in solution since the Ca , SOs, and
_| _ |_ I
species are constrained by solubility relationships. If Mg and Na ion
concentrations exceed the C& ion concentration (on a normality basis) , there
will be more soluble cations in solution than anions. To satisfy electrical
neutrality requirements in this situation, larger amounts of basic anions
such as SO 3, and SOi* will be required to remain in solution. Because the
I [ __ =
amounts of Ca , SOs, and S0i+ in solution are controlled by solubility con-
i _ i j
straints, solutions which are high in Mg and Na will generally contain
large amounts of sulfur and relatively little calcium. On the other hand,
__ i _j _ i
if C£ concentrations exceed Na and Mg concentrations, there will tend to
I i
be a higher relative concentration of Ca ions in solution.
Addition of relatively small amounts of soluble magnesium species (less than
1 percent by weight) has been shown to be beneficial to S02 removal effi-
ciency. Since the magnesium compounds are much more soluble than the
calcium-based reactants, the S02 can react rapidly with the active liquid-
phase magnesium species, thus making the limestone dissolution rate much
2-51
-------
less of a determining factor in the S02 removal efficiency. This is illus-
trated in Figure 2.2.1-15, and is substantiated by data from both pilot and
full-scale systems. Magnesium addition was tested at two different inlet S02
concentrations (medium and high) on Combustion Engineering's 0.56 m3/s
(1200 cfm) pilot plant using lime as the scrubbing agent. These results are
presented in Figure 2.2.1-16. Some data from a small limestone system
(Shawnee TCA) are shown in Figure 2.2.1-17.
95
»«
5 90
LU
I—•
u
tr
LL.
LU
S 80
,75
70
'0 1000 2000 3000 4000 5000
MAGNESIUM CONCENTRATION, ppm
1 INLET SO, CONCENTRATION - 2000 ppm
2 INLET SO, CONCENTRATION - 3000 ppm
Figure 2.2.1-17- Effect of magnesium on S02 removal efficiency.
58
Test results from two full-scale systems have also been reported.' At
Paddy's Run, magnesium addition increased S02 removal efficiency from 83 per-
cent (average) to 99.7-99.9 percent; the inlet S02 concentrations ranged
from 2150-2230 ppm. At Phillips Station, which is also a lime system, mag-
nesium addition resulted in S02 removal increasing from 50 percent to 83
percent. The source of magnesium was lime containing up to 10 percent Mg.
It should be pointed out that magnesium must be continually added to the
system to make up for that lost in the adherent sludge liquor and that tied
up with chloride ions present in the system.
Number of contact stages - Another way to increase S02 removal is to increase
the number of contact stages in the FGD system. This provides a means to
contact the S02-laden gas with a fresh, unsaturated scrubbing slurry after
2-52
-------
0.20
0.10
0.05
B °-02
5
< o.oi
o
LU
>
S 0.005
0.002
0.001,
C1
1 I I ' I |l
0.10 MOLES/X,
"I
1 1 I I I I I |
CaSO.. sat. =0.3
CaSO, sat. =1.0
Prn = 0.10 atm
C02
~ T = 50'C
0.01
I
0.01 0.02 0.05 0.10 0.2 0.5
Mg CONCENTRATION (MOLES/c)
1.0
Figure 2.2.1-15. Dissolved alkalinity generated
by addition of MgO.
5S
2-53
-------
100
90 •
50 ••
40 ••
30
SCRUBBER INLET pH -
FACTORIAL TESTS
O pH- 5.7-6.0
D pH-5.4-5.6
A pH-5.1-5.3
SCRUBBER GAS VELOCITY - 10.4 ft/ioc
LIQUID - TO - GAS RATIO - 45 gal/Mcf
TOTAL HEIGHT OF SPHERES = 0 in.
INLET S02 CONCENTRATION - 2,300-2,700 ppm
LIQUOR Cl~ CONCENTRATION = 5,000-14,000 ppm
-I-
-f-
-4-
2,000 4,000 6,000 8,000 10,000
EFFECTIVE LIQUOR MAGNESIUM CONCENTRATION, ppm
12,000
Figure 2.2.1-16. Effect of Magnesium on S02 removal efficiency
TCA (no spheres) with limestone.57
2-54
-------
the first stage of SOa removal is complete. This effect tends to level off
as the number of additional stages is increased, as shown by some Mohave lime
scrubbing data obtained on the horizontal module shown in Figure 2.2.1-18.
It was also reported that a similar effect was measured in the TCA scrubber,
at the Shawnee test facility when two different scrubbers were operated in
series. sl The results, said to be generally true for both lime and limestone
systems, are shown in Figure 2.2.1-19.
Ash removal - An important design consideration for coal-fired system appli-
cations is whether or not to remove particulates upstream of the scrubber.
The current trend within the utility industry is installation of dry parti-
culate collection equipment upstream of the FGD system. With low-efficiency
precipitators (90-95 percent removal) or mechanical collectors, the particu-
late emission control costs may be considerably reduced, but the scrubber
design must be capable of removing residual particulate.
Some scrubber types (venturi or mobile-bed) can effectively control both par-
ticulate and S02, in some cases with a two-stage arrangement. Although the
capital cost may be minimized and the ash may contribute alkalinity to the
system, there are several significant disadvantages associated with removing
particulates in the FGD scrubber.63
1. The extent to which the sludge can be dewatered by addition of
dry fly ash is reduced. The importance of this factor depends
on the sludge disposal method to be used.
2. There is a general consensus that ash causes erosion in the
scrubber; on the other hand, some degree of erosivity may be
desirable to keep the internal surfaces free of scale and
deposits.
3. To avoid incidences of exceeding particulate emission regulations,
by-passing the scrubber would be questionable. As the reliability
of lime/limestone systems continues to increase, this factor
diminishes in importance.
4. Fly ash cannot be marketed unless collected dry upstream of the
scrubber.
5. Particulate scrubbing results in an increased pressure drop which
in turn increases power consumption. This is felt not only as an
operating cost item, but also as an increased requirement for
2-55
-------
100
80
60
CM
O
40
20
CONDITIONS
GAS FLOW =212 m3/S (450,000 SCFM)
L/6 =2.7 1/nT (20 GPM/1000 SCFM)
INLET S02 = 220 ppm
2 3
NUMBER OF STAGES
Figure 2.2.1-18.
170 MW horizontal SOz removal versus
number of stages, Mohave plant.60
2-56
-------
100
.90*
95
Ul
K
u
oc
90
85
30
40 50
PERCENT SO2 REMOVAL FOR FIRST SCRUBBER
60
Figure 2.2.1-19.
absorption efficiency for
two scrubbers in series.62
2-57
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overall boiler capacity. On the other hand, burning low-sulfur
coal requires a relatively high power requirement for the preci-
pitator anyway, thus making combined particulate/S02 scrubbing
more suitable, for western coals.
Oxidation and scale prevention - Scaling in the scrubber can be a function of
several operations in the scrubbing loop. Low L/G ratios .may cause scaling
by allowing the calcium sulfate supersaturation in the scrubber liquor to
exceed the critical level. This can lead to precipitation of calcium sulfate
on scrubber internals. This phenomenon would occur if the hold tank were
undersized, not allowing sufficient residence time for precipitation and
desupersaturation of calcium sulfite and sulfate and dissolution of absorbent.
The rate of scaling is sensitive to the supersaturation of calcium sulfate.
If supersaturation is unchecked, calcium sulfate dihydrate crystals start to
crystallize on equipment surfaces, forming a scale. Calcium sulfate super-
saturation levels of up to 1.3 can be tolerated before scaling occurs.
Calcium sulfate formation results from the oxidation of calcium sulfite. For
a given system the oxidation rate will depend on the relative concentrations
of 02 to S02. For example, some systems applied to low sulfur western coals
have shown greater than 90 percent oxidation whereas the eastern coals exhi-
bit lower rates which range from near zero to 40 percent.61* Consequently, in
predicting oxidation, it has been observed that this rate is a strong func-
tion of the oxygen content of the flue gas. This observation must be con-
sidered when designing for systems with relatively high 02/S02 ratios, such
as are encountered in industrial boilers, as well as those where this ratio
might vary. A variation in Oz/SOa ratio would be observed in systems which
encounter large changes in excess air in the boiler operation. Therefore, a
system which must accommodate significant changes in flue gas concentration
should be designed for longer residence times in the hold tank and higher L/G
ratios.
Water balance - Another important design consideration is the water balance
within the system and its role in the overall plant water management system.
The following factors impact the design of the water system:
-------
Evaporation loss in the scrubber
• Adherent liquor discharged with the sludge
• Pump seals
• Mist eliminator wash
• Pond considerations
Water loss via evaporation in the scrubber, generally 0.4-0.5 tons/ton of
coal fired, allows for fresh water addition to the system. Fresh water input
is generally used for mist eliminator washing, slurry preparation, and some-
times pump seals. During periods of low boiler load, however, evaporation
loss drops considerably. Fresh water makeup requirements are not reduced
to the same degree, thereby creating a problem in maintaining the proper
water balance.
Adherent liquor discharged with the sludge constitutes the other main water
loss from the scrubbing system. The quantity varies considerably from site
to site depending on the type of dewatering used, and the amount and type of
solids discharged ('i.e., sulfur-containing species and ash). Water loss can
range from 100 Ib per ton of coal burned for a case using highly efficient
dewatering methods to as much as % ton or more of water per ton of coal.8S
To avoid erosion of pump seals in slurry systems, relatively clean water is
needed. This requirement can be met either by fresh water or clarified re-
cycle water. Another fresh water requirement is for mist eliminator wash,
although in some cases a mix of recycled clarified overflow with fresh water
has been successfully used.
Several features of pond design and operation are also important to the over-
all water balance. For example, rainfall surface runoff, evaporation loss
and leaching make design of completely closed loop operations site specific.
2-59
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Gas Reheat - In general, industrial boiler FGD processes do not use stack
reheat systems whereas utitity FGD applications do.66'67 This discussion
briefly summarizes the reasons for applying reheat, and addresses potential
reheat mehtods.
Flue gases are normally discharged from a utility boiler at a temp-
erature of greater than 120°C (250°F). This temperature is selected to re-
main above the t^SOij dew point to reduce corrosion and permit use of less ex-
pensive grades of material to be used for fans, ducts, and stacks. Industrial
boilers generally have less efficient heat recovery systems and, consequently
discharge flue gas at a temperature of 175-200°C (350-400°F).68 When a
wet scrubber is inserted between the boiler outlet and stack, the flue gas
exiting the scrubber is saturated with water and cooled to the adiobatic
saturation temperature of about 50°C (125°F)- Discharging the cool wet gas
directly to the atmosphere can result in the following problems:
• Condensation of water vapor and sulfur oxides, which can result in
the acidic water corrosion of downstream ducts, fans, and stack
lining.
• Impaired plume rise and, hence, poorer dispersion of residual
pollutants for a given stack height.
Deposition of scrubber residue on downstream fan blades, resulting
in imbalance.
A visible plume as water vapor condenses.
' Stack rain, or mist droplets, that can settle around the power
station.
To prevent these undesirable effects of wet scrubbers, the exit gas can be
reheated to a higher temperature before being discharged.
Several approaches have been developed to reheat scrubber gases
The basic differences among these approaches are in the energy sources and
the method of transferring energy to the gas. Reheat methods currently in
use include:
2-60
-------
• Direct inline reheat - using steam or hot water
exchangers
• Direct combustion reheat - using gas or oil in
inline burners or external combustion chambers
• Indirect hot air reheat - using steam to heat air to
mix with the scrubbed gas
• Bypass reheat - bypassing a portion of the untreated
hot flue gas to mix with the scrubbed gas
Although the majority of utility FGD installations use reheat, there are
several installations which operate under "wet-stack" conditions as do the
majority of industrial FGD systems. Operating without reheat minimizes the
energy consumption of an FGD system.69 Current reheat practice is to provide
about 30°C (50°F) of heat to overcome problems associated with discharging
wet gas to the atmosphere. This practice results in an energy penalty of
about 2 percent of the net heat input to the boiler. Although reheat has
the potential for providing significant benefits, it should not be consi-
dered as a necessity, but as one of a number of approaches for consideration in
optimizing sulfur dioxide absorption systems..
Sludge dewatering and disposal - The sludge dewatering process step is used
to concentrate the solids for ease of handling and disposal and to lower
transportation costs. The clear liquor is usually recycled back to the pro-
cess for reuse. The methods currently used in lime/limestone sludge dewater-
ing on commercial-sized units are thickening and vacuum filtration. Centri-
fugation was also tested, but filtration has been the selected method for
full-scale applications in most cases. In addition to these methods, interim
ponding has been used as a dewatering procedure. °
Clarifiers or thickeners are commonly used as a primary dewatering device for
sludges containing low solids content (10-15 percent solids). They can typi-
cally achieve a solids content of 30-35 percent. Depending on the dewatering
properties of the sludge and the ultimate disposal plans, further steps may
be taken to increase the solids content using vacuum filters, which achieve
50-75 percent solids, depending on the individual system. Because of the
2-61
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physical nature of sulfate crystals compared to sulfite, dewatering
is generally improved by a higher sulfate/sulfite ratio. A sludge from
a forced oxidized system, which contains mostly calcium sulfate can achieve
about 85 percent solids using either centrifugation or filtration.
The dewatering method selected is heavily influenced by the type of disposal
to be used. While the basic options are wet (i.e., ponding) or dry (i.e.,
landfill), the composition of the sludge and the availability of dry fly ash
to supplement dewatering are also important factors.
Erosion and Corrosion. - Early FGD installations experienced many shutdowns
due to corrosion and erosion. While significant advances have been made in
solving materials problems, some material failures are still reported and
studies are still underway to investigate the metallurgical aspects of FGD.
In recent reports, several guidelines for selection of materials for FGD
71,72
systems were laid out based on experience to date. The preferred
material for all wet surfaces, including the FGD module itself as well as
inlet and outlet ducts, is 316L stainless steel. Its superior performance
with respect to resistance to corrosion, erosion and scaling may be due to
its molybdenum content. In comparative metallurgical tests, 316L has proven
better than carbon steel, low alloy steels, type 304, and type 304L. There
are some systems, however, that incorporate such alloys as Hastelloy C-276,
Hastelloy G, Inconel 625, Incoloy 825, 317L stainless steel, 904L stainless
steel, and Jessop JS700. These materials have been used in wet/dry high
temperature, high chloride environments, such as in presaturators.
It is important that pumps, lines, and recycle tanks in wet slurry systems
be protected from erosion. Materials successfully used in these key elements
include synthetic and natural rubber coatings.
Lining materials for the absorbers, exhaust ducts and stacks have also been
tested. Although successful applications have been reported, some liner
2-62
-------
failures have been due to improper application, instability at high tempera-
tures, inconvenience of repair, and cost-related factors. Materials success-
fully used to date in this type of application include resins, ceramics,
polyesters, polyvinyls, polyurethanes, carboline and Gunite.73
2) Fuel variations - The design and operation of lime/limestone scrub-
bing systems are affected by variations in the fuel characteristics
listed below:
• Sulfur content
• Chloride content
Ash alkalinity
• HHV
The effects of these variables on the design of lime/limestone systems are
much like the effects on other wet FGD systems. The following discussion
describes the importance of these fuel variables on the design of lime/
limestone wet scrubbing systems.
Sulfur content - The sulfur content of the fuel determines the removal rate
of S02 which is necessary to comply with a given environmental regulation.
The design of virtually every piece of equipment is affected when attempting
to achieve a set emission rate with fuels of varying sulfur content. Increas-
ing the fuel sulfur content affects the design of the absorber in that higher
liquor circulation rates are required for higher sulfur fuels which results
in bigger pumps and piping, and an increased scrubber pressure drop with
increased system energy requirements. In addition, larger alkali storage,
preparation, and feed equipment is required to handle the increased alkali
necessary to react with the larger amounts of sulfur. Finally, larger solids
separation and waste sludge disposal equipment will be needed to handle the
increased load. In summary, FGD systems designed for high sulfur fuels will
be more complex and costly than systems designed for low sulfur fuels.
2-63
-------
Chloride content - During combustion the chlorine content of the fuel forms
gaseous HC1 which will be readily absorbed in aqueous FGD systems. Chlorine
is important both for corrosion reasons, and because it will react with the
alkaline sorbent thus limiting the-862 removal ability of the system. The
presence of chlorides in the scrubber liquors provides the potential for
stress corrosion which can result in the use of high-alloy equipment wherever
protective coatings are not applicable.
Dissolved chlorides will also react with active alkalis to form inactive
chloride salts, and consequently a high chloride coal will require a signi-
ficantly larger liquor circulation rate than will a coal with a lower
chlorine content to achieve the same amount of S02 removal. The use of
alkali additives, such as MgO, may be used to increase the dissolved alka-
linity which will permit S02 removal without an increased liquor circulation
rate. Alternatively, chlorine can be removed ahead of the SOz scrubber in a
relatively small prescrubber.
Ash alkalinity - Another fuel variation that can affect the design for wet
FGD systems is the ash alkalinity. As previously discussed, a highly alka-
line ash can significantly decrease the quantity of sorbent required for S02
removal. Commercial scale utility systems are in operation today which are
based on a combined sorbent-alkaline ash design. The ash alkalinity will
influence the design consideration of whether or not to remove particulate
material ahead of the scrubber.
HHV - A boiler using a low heating value fuel must fire at a higher burn rate
to generate the same amount of power as for a high heating value fuel. The
low heating value fuel will produce a larger volume of flue gas per unit of
power than the high heating value fuel which will result in the FGD system
being of a larger size to accomodate the greater gas volumes. In addition,
if both fuels had the same sulfur content, the low heating value fuel would
result in greater SOz emissions per unit of power.
2-64
-------
3) Ambient variations - FGD system operations are essentially indepen-
dent of ambient variations. However, as with all wet systems, extreme cold
can adversely affect the operation of a lime/limestone FGD process. Problems
with line freezing and subsequent unit shutdowns were experienced at several
operating facilities in the East during the harsh winter of 1977- These
problems can be solved by heat tracing lines or by locating the FGD
system inside buildings. Also, climatological factors may effect visible
plume formation and should be taken into consideration on reheat decisions.
E. Retrofits—Existing boiler operating parameters such as flue gas
temperature and oxygen content will definitely affect the design of a lime/
limestone system. There are no specific conditions other than equipment
space requirements, however, which would actually prevent application of a
lime/limestone flue gas cleaning process.
Major process equipment for a lime/limestone system includes the scrub-
ber where contact between the flue gas and slurry is promoted, a mist elim-
inator for removal of entrained liquid from the S02-lean gas, a reaction tank
where additional lime/limestone dissolution and solid product precipitation
occurs, and a solids separation device to reduce the amount of liquid waste
leaving the process. Flue gas reheat may also be used on some installations.
In addition to the equipment listed above, several other items are needed.
Storage bins for the alkaline additive are needed. Slurry tanks and slurry
pumps are also required. A waste sludge handling area is required for both
the lime and limestone systems. Depending on whether on-site solids disposal
is planned, equipment in this area could include a clarifier or thickener,
vacuum filters or sludee fixation equipment. Space may also be needed for
sludge storage.
Space required for the scrubbing section of this process is the major
concern for retrofit since this equipment must be placed adjacent to the
powerhouse and stack. Process equipment outside of the scrubber area is of
less concern to the retrofit problem since it can be located on the peripher-
al areas of the plant. By employing good design practices and placing the
2-65
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hold tanks under scrubbers, space requirements for the scrubbing section
can be minimized. However, for the large industrial boiler application con-
sidered in this evaluation, 200 X 105 Btu/hr, a scrubber diameter of only
3M (10 ft) would be sufficient to provide a gas velocity of 3M/S (10 ft/sec).
This relatively small scrubber size does not appear to be large enough to
limit its applicability.
2.2.1.2 System Performance—
A. Emission reductions—A discussion of design and operating
parameters for lime and limestone FGD systems was presented in Section
2.2.I.I.D. Data for that discussion came primarily from EPA's 10 MW test
facility located at TVA's Shawnee steam plant. In this section, results
of process testing recently completed on the Bahco Process installed at
Rickenbacker Air Force Base (RAFB) will be presented to illustrate the
performance of a lime/limestone FGD system applied to industrial boilers.
In the Bahco Process, flue gas is passed through a mechanical collector
to remove coarse particulate matter before entering the booster fan. The
booster fan forces flue gas into the first stage of the scrubber where it is
vigorously mixed with scrubbing slurry in an inverted venturi. In this
stage, flue gas is cooled to its adiabatic saturation temperature and SOa
and particulates are scrubbed from the gas. This partially scrubbed gas
rises to the second stage where it is contacted with slurry containing fresh
lime to complete the required SOz and particulate removal. Gas from the
second stage enters a cyclonic mist eliminator where entrained slurry
droplets are separated from the gas by centrifugal force.
Pebble lime from a storage silo is slaked and added directly to the
slurry in the lime dissolving tank. The resulting fresh lime mixture is
pumped to the second stage (upper) venturi to treat the flue gas stream.
The slurry flows by gravity from the second stage to the first stage where
it contacts hot flue gas entering the scrubber. This countercurrent flow
arrangement results in high SOa removal and efficient reagent usage.
2-66
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Spent slurry flows by gravity from the first stage of the scrubber
to the dissolving tank. Part of the spent stream leaving this stage is
diverted to the thickener where the slurry is concentrated to 35 to 40%
solids. Overflow from the thickener returns to the dissolving tank and
the underflow is pumped to a Hypalon-lined sludge pond near the Heat Plant.
The slurry and gas streams described above are illustrated in Figure
2.2.1-20.
The Bahco system was designed to treat up to 108,000 acfm of flue gas
generated at the peak winter load of approximately 200 X 10 Btu/hr. The
system, which must operate over a relatively narrow range of gas flow,
35,000 to 50,000 scfm, has an essentially unlimited turndown capability for
handling flue gas by mixing air with the flue gas at low boiler loads. This
allows the system to handle seasonal load variations from 20 to 200 X 10
Btu/hr, S02 concentrations from 200 to 2000 ppm, and partiuclate loadings
up to 2 gr/scfd. In addition, the scrubbing system has coped with 100
percent load increases occurring in as little as an hour's time. Table
2.2.1-9 shows scrubber operating conditions during average load conditions
for the RAFB installation. 7k
TABLE 2.2.1- 9. TYPICAL BAHCO OPERATING CONDITIONS76
Flue Gas 64,000 acfm @ 380°F (37,500 scfm)
S02 Concentration 1390 ppm
1st Stage AP 10 in W.C.
2nd Stage AP 8 in W.C.
Lime-SOa Stoichiometry 0.876
S02 Removal 87.6% (0.615 lb/106 Btu)
Lime Utilization 100%
Particulate Emission 0.16 lb/106 Btu
2-67
-------
-THICKENER
ho
I
oo
LIME OR
LIMESTONE
TRUCK
REAGENT SYSTEM
MODULE
MAKEUP
WATER
OVERFLOW
TO
LIME
DISSOLVING
TANK
MIST
ELIMINATOR
REAGENT
STORAGE
R-C/BAHCO
SCRUBBER
LEVEL
TANK
BOOSTER
FAN
THICKENER
OVERFLOW
REAGENT
FEEDER
&SLAKER
LEVEL
TANK
SLUDGE
TO POND
BY-PASS
MAKE UP STACK
FLUE GAS
FROM HEAT
4 PLANT
MECHANICAL
COLLECTOR
TO
FLY ASH
DISPOSAL
UNLOADING"
STATION
•LIME
2nd STAGE PUMP
MILL PUMP
DISSOLVING TANK
Figure 2.2.1-20 R-C/Bahco scrubber system flow diagram75
-------
During the test period, the system operated satisfactorily at slurry
solids from 2 to 25 percent by weight. Particulate emissions were above the
required level of 69 ng/J (0.16 lb/10s Btu) when substantial quantities of
soot were being formed in the heat plant. Sulfur dioxide removal efficiency
was well above the required level, and levels in the 95 percent range were
attainable. The following discussion summarizes the SOz removal results
obtained when operating with lime and limestone, the particulate removal
ability of the system, and the properties of the by-product sludge.
Figure 2.2.1-21 shows a nearly linear relationship exists between lime/
S02 stoichiometry and S02 removal efficiency. Lime utilization approached
100 percent in the range of stoichiometry from 0.3 to 0.9 moles lime/mole
S02 and dropped gradually to 90-95 percent as SO2 removal approached 100
percent. Over the range of liquid circulation rates studied, 1500-2400 GPM,
no effect on S02 removal was observed. Results of these lime tests indicate
that S02 removal is controlled by the lime-S02 stoichiometry. The fact that
SO2 removal when using lime is not adversely affected by changes in operating
variables illustrates that good gas-liquid contact was maintained over the
entire operating range of the system. 7
Results of process testing using limestone indicated that slurry circu-
lation rate, as well as limestone-S02 stoichiometry, are important in
determining the SO2 removal efficiency. Figure 2.2.1-22 shows the positive
effect on SOz removal and limestone utilization of increasing the slurry rate
from 2000 to 2600 gpm. This figure also shows that limestone utilization
varied from 75 to 90 percent at lower SO 2 removal levels but decreased sig-
nificantly above 80 percent SO2 removal. The scatter shown in the data on
Figure 2.2.1-22 were reported to be within the uncertainty limits of the
data, and project scope limitations precluded running additional verifi-
cation tests to investigate this further.79
The Bahco system at RAFB was designed with extra fan capacity in order
to help carry out the particulate test program. Venturi pressure drops
2-69
-------
100
BO
I
a:
CM
O
CO
60
40
20
HAFB CODE REQUIREMENT
UME
UTUZATON .
100%, /
'
V
O SCREENING TEStS
A VERIFICATION TESTS
0.2 0.4 0.6 0.8 1.0
LIME STOICHIOMETHY. MOLES, LIME/SO2
1.2
Figure 2.2.1-21.
removal efficiency as a function
of lime stoichiometry.78'
2-70
-------
100
90
80
6 7°
^ 60
u
I"
LU
3"
30
20
10
.2
J I
I I
.8 1.0 1.2 1.4 1.6
LJMESTONE/SOt STOICHIOMETRY
(LB4WUE CACOVLB MOLE SOi)
Figure 2.2.1-22. SOa removal efficiency as a function
of limestone/SO2 stoichiometry and
slurry pumping rate.80
2-71
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were increased to nearly double their design values of 7 in. W.C. for the
particulate tests. Data from the particulate control tests are plotted in
Figure 2.2.1-23 to show the effect of the combined pressure drop of the two
venturies in the Bahco scrubber on particulate emissions. Below approxi-
mately 18 in. HzO total pressure drop, particulate emissions increase
rapidly.82
A series of scrubber sludge characterization tests were carried out to
compare the lime and limestone sludges. In general, the limestone sludges
settled much more rapidly than the lime sludges as shown in Figure 2.2.1-24.
In both cases, the concentration of settled sludges were not affected by the
addition of flocculant. However, the limestone sludges, due to their higher
gypsum contents, produced a settled layer with a higher solids content (58
8 3
weight percent) than the lime sludges (43-45 weight percent). Results of
centrifugation and filtration tests as shown in Figure 2.2.1-25 also shows
better dewatering results for the limestone sludges.85
B. Reliability—Reliability of lime/limestone FGD systems for industrial
boiler applications is difficult to assess since there are only two installed
systems and only one of those has been operational over a long period of
time. That system is the Bahco system located at Rickenbacker Air Force
Base.
Scrubber performance at the RAFB facility has generally been quite good
except perhaps for the early stages of operation in which several start-up
problems resulted in significant amounts of downtime. From the time period
of November 1976 through December 1978, however, the Bahco system illustrated
that an industrial boiler can operate with high reliability as it operated
95 percent or more of the time except for the months of January-March 1978.
During those three months, system downtime was caused by a severe blizzard
which resulted in the freeze-up of several lines. In February, 1978, the
Governor of Ohio temporarily shelved pollution contorl regulations and the
system was not operated until April at which time it was reported to perform
with an operability index of 100 percent.87
2-72
-------
0.60
0.5O-
I
£ 0.40
0.30
£ 0.20
£
0.10
GUARANTEE EMISSION RATE
A
I
I
I
J_
10 14 18 22 26- 30
TOTAL SCRUBBER PRESSURE DROP - IN. W.C.
Figure 2.2.1-23. The effect of scrubber pressure drop
on particulate emission rates.
2-73
-------
2000
X
0 PPM FLOCCULANT
2 PPM
5 PPM
IB 24 30
SETTUNQ TIME - HOURS
36
42
48
Figure 2.2.1-24. A comparison of lime and limestone
slurry settling rates.
2-74
-------
70
65
60
55
50
D RUNS La-Lt
O RUNSLSi-LS4
O RUNSLSs-LSt
37 WT. % LIMESTONE SLURRY FEED
38 WT. % LIME SLURRY FEED
23 WT.% LIMESTONE SLURRY FEED
10
CENTRIFUGE OPERATION
TIME-MINUTES
15
20
Figure 2.2.1-25. The effect of operating time and slurry feed
concentration on centrifuge cake density.
86
2-75
-------
During the early stages of operation of this unit, several start-up
problems were experienced that resulted in considerable amounts of downtime.
During this period which included 11,024 hours, there were 4,830 hours of
downtime. Most of this downtime, 2,766 hours, resulted from booster fan
problems which have been rectified. An additional 1,088 hours were lost due
to other auxiliary equipment problems. However, of the total of 3,854 hours lost
due to equipment problems, 1,035 hours were due to delays resulting from a
lack of spare parts. Heat plant outages and minor system modifications re-
sulted in an additional 507 hours of downtime. Table 2.2.1-10 summarizes the
downtime during the test period. Once these problems were cleared up, the
system has operated very well for over two years as discussed above.
TABLE 2.2.1-10. SUMMARY OF DOWNTIME DURING THE BAHCO SYSTEM TESTING 87
Category Downtime Hrs.
Fan 2,766
Auxiliary Equipment 1,088
Heat Plant Outages 388
Modifications 119
Other Items 469
4,830
In addition to good performance levels in the United States, Japanese
FGD installations have also demonstrated high SO2 removal efficiencies and
high reliabilities. Recent reports on Japanese installations have documented
system reliabilities of greater than 95 percent. There are several reasons
for the good performances being demonstrated in Japan which are discussed
briefly below. 88>89
Generous design - The design of scrubber systems in Japan appear to have
generally been approached with a more conservative philosophy than that in
the U.S. Specifications provided by users to the scrubber system supplier
do not appear to be more detailed than those supplied in the U.S. However,
2-76
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the user expects and, in fact, demands that the scrubber system supplied
perform with a reliability compatible with that of the steam generating
plant. For example, Electric Power Development Company (EPDC) requires
the scrubber system supplier to correct at his expense any process/equipment
problems that occur within a year after acceptance of the system. This
philosophy has generally resulted in scrubber systems which have required
less subsequent modifications and maintenance as a result of fewer operating
problems than their U.S. counterparts. The fact that one supplier, Mitsubishi
Heavy Industries (MHI), has provided about half of the lime/limestone scrub-
bing systems in Japan has undoubtedly enhanced the reliability obtained
-------
Separate operating crews - The Japanese suppliers and users of FGD
systems recognize and accept the .fact that scrubbers basically involve chemi-
cal processes requiring carefully controlled operation by personnel specifi-
cally trained for this purpose. Many utilities including EPDC contract
with subsidiary companies to supply scrubber system operating and
maintenance services. These specially trained personnel are not rotated into
the power plant for other duties as is generally the case with U.S. utili-
ties.
Government surveillance - Japan employs a stringent continuous monitoring
and enforcement program to insure that utility and industrial sources are in
continuous compliance with environmental regulations. Each prefectural
government operates an environmental research center (subsidized by the
central government), most of which are directly linked via telemetry systems
to automatic monitoring stations located at major emission sources and key
ambient sites. The existence of these monitoring systems likely has been
instrumental in assuring that emission sources remain in constant compliance
since violations would result in fines and/or forced shutdown of the
source.
It appears that although the Japanese systems in general pperate more
reliabily than U.S. systems, the differences between U.S. and Japanese
systems are not great. Consequently, it is reasonable to expect the relia-
bility of U.S. systems to increase to Japanese levels in the coming years.
C. Impacts on boiler—The impacts of a lime/limestone FGD system on
boiler operations should be small. The most important impacts are: l)power
consumption for running the BCD system's pumps and fans, and 2) possible
boiler load reduction during FGD system outages if no bypass is permitted.
The amount of power consumed has been estimated to be about 2-4 percent of
the net power input to an industrial boiler (see Section 5). Boiler load
reduction due to FGD system outages will be the other major impact. However,
increasing FGD reliability as demonstrated by the RAFB system will help to
2-78
-------
reduce these outages to a minimum.
D. Additional maintenance requirements—In all likelihood, additional
maintenance will be required to operate a lime/limestone FGD system over that
required to operate an industrial boiler alone. The Bahco system located
at Rickenbacker A.F.B. has operated with relatively low maintenance require-
ments once the initial mechanical problems were resolved. Part of the
success of this system, however, is probably due to the preventative
maintenance program that R.A.F.B. has established which requires one oper-
ator/shift which is similar to the Japanese operational plan.
The areas or components in the system that should be given the most
attention vary somewhat depending on the design. For most systems they will
include the nozzles, headers, strainers, scrubber internals, pump packing
and impellers, mist eliminators, fans downstream of the scrubber, agitators,
valves, and slurry lines.
2.2.2 Double Alkali Process
The double alkali process encompasses some of the features of lime/
limestone wet scrubbing in that a calcium sulfite/sulfate sludge is produced
for disposal. While double alkali processes are considered throwaway, a
regeneration step is employed to regenerate the active alkali for the S02
sorption. The sodium/calcium double alkali process uses a soluble sodium-
based alkali for S02 sorption and a calcium-based alkali to regenerate the
active sodium solution. Although there are several other types of double
alkali processes which have been investigated (e.g. , ammonia and potassium
as the soluble alkali), the sodium/calcium system is the most advanced.
The following discussion is limited to this type of system.
2-79
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2.2.2.1 System Description —
A. Process description — The double, or dual, alkali process uses a
soluble sodium-based alkali (NaOH, Na2S03, Na2C03 or NaHCOs) to absorb S02
from flue gas in a scrubber. The sulfur oxide-rich effluent liquor is
reacted with a calcium-based alkali to precipitate calcium sulfite, calcium
sulfate, and mixed crystals for disposal. Additionally this reaction regen-
erates the sodium-based alkali for recycle to the scrubber. The processes
developed in the U.S. use lime as the calcium alkali, but other processes
developed in Japan and still in the development stage in the U.S. use lime-
stone.
The process can be divided into three principal areas: absorption,
regeneration, and solids separation. A simplified flow diagram is given in
Figure 2.2.2-1. The principal chemical reactions for a lime regeneration
system are described by the following equations.
Absorption
2 NaOH + S02 + Na2S03 + H20
(2.2.2-1)
Na2S03 + S02 + H20 -> 2NaHS03
(2.2.2-2)
Regeneration
Ca(OH)2 + 2NaHS03 -> Na2S03 + CaS03-%H20 + 3/2 H20
Ca(OH)2 + Na2S03 + H20 ->• 2NaOH + CaS03-%H20
Ca(OH)2 + Na2SOit + 2H20 + 2NaOH + CaS(H'2H20
An important side reaction that also occurs is oxidation:
(2.2.2-3)
(2.2.2-4)
(2.2.2-5)
Na2S03 + %02
(2.2.2-6)
2-80
-------
I
CO
SCRUBBED GAS
SCRUBBER.
•-GAS TO STACK
FLUE GASZr
X
SCRUBBER
EFFLUENT
LIME H20
1 i
LIME
SLAKER
^
REACTOR
SCRUBBER FEED
THICKENER
WASH
WATER
VACUUM
FILTER
£ "
I O
LU O
e^ *O
/I
O
CM
HOLDING TANK
WASTE
CALCIUM
SALTS
Figure 2.2.2-1. Simplified Flow Diagram for Sodium Double-Alkali Process.
9J
-------
In the absorber, SC>2 is removed from the flue gas by reaction with NaOH
and Na2S03s according to Equations 2.2.2-1 and 2.2.2-2. The concentration of
active sodium (sodium associated with anions involved in SOa absorption
reactions, namely sulfite, bisulfite, hydroxide, and carbonate/bicarbonate)
may be either concentrated (>.15M) or dilute (<.15M) depending on the system
design as described in Section D. Because oxygen is present in the flue gas,
oxidation also occurs in the system, according to Equation 2.2.2-6. The
scrubbed flue gas may then be reheated before exiting the stack.
Most of the scrubber effluent is recycled back to the scrubber, but a
slipstream is withdrawn and reacted with slaked lime in the regeneration
reactor according to Reactions 2.2.2-3 and 2.2.2-4. In the process designed
by FMC the regeneration reaction occurs in a continuous stirred tank reactor,
but in the CEA/ADL design it occurs in a specially designed two-stage reactor
system. In the CEA/ADL design, Reactions 2.2.2-3 and 2.2.2-4 both occur; the
regenerated solution has a pH of about 11 to 12.5. In the FMC design, how-
ever, only Reaction 2.2.2-3 occurs. No free hydroxide is formed, and the pH
of the regenerated solution is only about 8.5. The sodium sulfate formed by
oxidation of sodium sulfite does not react with lime as readily as the sodium
sulfite. The presence of sulfate in the absorber is undesirable in that it
converts active sodium to an inactive form, much the same effect as chlorides.
Methods of removing sulfates a.re discussed in Section D, Factors Affecting
Performance. As a key design feature currently under study by process
developers, this topic is also addressed in Section B.2, Development Status,
Recent Improvements.9 2
The regenerated slurry stream, which contains calcium sulfite/sulfate
solids, is sent to a thickener where the solids are concentrated. The clear
solution overflow from the thickener is stored in a hold tank for return to
the absorber, and the underflow containing the calcium sulfite/sulfate
solids is further concentrated in a vacuum filter to about 50 percent solids
or more. The solids are washed, generally with 1 to 2 displacement washes,
to recover sodium salts and then sent to a pond or landfill for disposal.
9 3
The filtrate and wash water are recycled to the thickener.
2-82
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B. Development status—
1) Commercial applications - Several vendors in the U.S. have developed
double alkali systems. These include Arthur D. Little and Combustion Equip-
ment Associated (CEA/ADL), FMC Corporation, General Motors, Zurn Industries,
Neptune Airpol, and Buell/Envirotech. A list of operational and planned double
alkali systsms in the U.S. is given in Table 2.2.2-1. In Japan, the double
alkali processes developed by Showa Denko and Kureha-Kawasaki (which use lime-
stone as a regenerant) have been applied to about 2450 MW of capacity on oil-
fired utility and industrial boilers and sulfuric acid plants. The significant
facilities in operation are listed in Table 2.2.2-2. In general, the double
alkali process uses technology common to other FGD processes, and the
development of double alkali technology benefits from advancements in
technology for other soluble sodium scrubbing systems and for lime/lime-
stone systems.
2) Recent improvements - Process improvements and ongoing research
have provided potential process alternatives in areas of sulfate removal,
forced oxidation and regeneration with limestone. The objectives and current
status of each of these areas are outlined below.
Sulfate removal - Various sulfate removal techniques have been investigated
to regenerate active alkali from inactive sodium sulfate. While recent pro-
cess improvements are presented here, a more basic discussion may be found in
Section 2.2.I.D.I, Factors Affecting Performance - Design. In dilute systems,
with active alkali concentrations less than 0.15 molar, sulfates can be
removed by precipitating calcium sulfate with the addition of lime, according
to Reaction 2.2.2-5. In a dilute system, however, because of the lower con-
centration of alkali in the scrubbing liquor, large volumes of scrubbing
liquor must be circulated through the absorber to achieve good S02 removal.
2-83
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TABLE 2,2.2-1. FULL SCALE DOUBLE-ALKALI SYSTEMS IN THE U=So91"95
NJ
I
00
System Operator
& Location
FMC
Modesto, CA
General Motors
Parma, OH
Caterpillar Tractor Co.
Joliet, 1L
Firestone
Pottstown, PA
Gulf Power Co.
Scholz Plant, Sneads; FL
(Southern Company Services)
Caterpillar Tractor Co.
Mossville, IL
Caterpillar Tractor Co.
Morton, IL
Caterpillar Tractor Co.
East Peoria, IL
Caterpillar Tractor Co.
Maple ton, IL
Louisville Gas & Electric
Cane Run No. 6
Central Illinois Public
Service, Newton No. 1
Southern Indiana Gas &
Electric, A. B. Brown No. 1
ARCO Polymers
Monaca , PA
Grissom Air Force Base
Bunker Hill, IN
Chansler Oil
Bakersfield, CA
System
Application
Reduction kiln
Coal-fired
industrial boilers
Coal-fired
industrial boilers
Coal-fired
industrial boilers
(demonstration)
Coal-fired
utility boiler
(prototype)
Coal-fired
industrial boiler
Coal-fired
industrial boiler
Coal-fired
industrial boilers
Coal-fired
industrial boiler
Coal-fired
utility boiler
Coal-fired
utility boiler
Coal-fired
utility boiler
Coal-fired
industrial boilers
Coal-fired
industrial boilers
Oil-fired
industrial boiler
Vendor or
Developer
FMC
General
Motors
Zurn
Industries
FMC
CEA/ADL
FMC
Zurn
Industries
FMC
FMC
CEA/ADL
Buell/
Envirotech
FMC
FMC
Neptune/
Airpol
FMC
Size
(MW, Equivalent)
10 (gas rate)
30 (regen.)
64
30
4
20
70
19
100
65
277
575
250
118
15
35
Active
Alkali
Cone.
Dilute
Dilute
Cone.
Cone.
Cone.
Dilute
Cone.
Cone.
Cone.
Cone.
Cone.
Cone.
Cone.
Cone.
Start-up
Date
Dec. 1971
March 1974
Sept. 1974
Jan. 1975
Feb. 1975a
Oct. 1975
Jan. 1978
April 1978
(Jan. 1979)
(Dec. 1978) b
(Nov. 1979 )b
(April 1979)
b
(June 1980)
(Sept. 1979)b
(Mar. 1979)b
a. System ceased operation in July 1976
b. Projected Start-up date
c. Site based on • conversion of 200O SCFM per MW equlvale
-------
TABLE 2.2.2-2. SUMMARY OF SIGNIFICANT OPERATING FULL SCALE SODIUM/CALCIUM
I
CO
ALKALI SYSTEMS IN JAPAN
96.97
System Operator
& Location
Showa Denko KK
Chiba
Tohoku Electric
Shinsendai Sta.
Nippon Mining,
Sagonoseki
Toho Zinc,
Annaka
Showa Pet. Chem.
Kawasaki
Kanegafuchi Chera.
Takasago
Toly Plastic
Fuji
Kyowa Pet. Chem.
Yokkaichi
Kinuura Utility
Nagoya
Daishuwa Paper
Fuji
Sikoku Electric Power
Anan
Sikoku Electric Power
Sakaide
Kyushu Electric Power
Buzen
System
Application
Oil-fired elec.
power boiler
Oil-fired
utility boiler
IhSOii plant
H2SO» plant
Oil fired
industrial boiler
Oil-fired
industrial boiler
Oil-fired
industrial boiler
Oil-fired
industrial boiler
Oil-fired
industrial boiler
Oil-fired
boiler
Oil-fired
utility boiler
Oil-fired
utility boiler
2 oil-fired
utility boilers
Vendor or
Developer
Showa Denko
Kawasaki/
Kureha
Showa Denko/
Ebara
Showa Denko/
Ebara
Showa Denko
Showa Denko
Showa Denko/
Ebara
Showa Denko/
Ebara
Tsukishima
Tsukishima
Kawasaki/
Kureha
Kawasaki
Kureha
Kawasaki/
Kureha
Size
(MW, Equivalent)
150
150
37
43
62
93
65
46
63
85
450
450
(2-450)
900
Active
Alkali
Cone .
Cone .
Cone.
—
—
Cone .
Cone.
Cone.
Cone .
Cone .
Cone.
Cone.
Cone.
Calcium
Sources
Limestone
Limestone
Limestone
Limestone
Limestone
Limestone
Limestone
Limestone
Lime
Lime
Limestone
Limestone
Limestone
Start-up
Date
June 1.973
Jan. 1974
1973
1974
1974
1974
1974
1974
1974
1974
1975
1975
1977
-------
These large volumes must also be circulated through the regeneration and
solid/liquid separation equipment. Capital and operation costs for S02
removal would thus be higher for dilute systems.
Sulfate can also be removed by precipitation as a mixed ("double") salt with
calcium sulfite as shown in Equation 2.2.2-7.
xNaHS03 + (x+y)Ca(OH)2 + (z-x)H20 •*• (x+2y)NaOH
(2.2.2-7)
This occurs in concentrated active sodium systems. In this mode, sulfate
removal cannot be accomplished by precipitation as gypsum since the high
sulfite levels prevent the soluble calcium concentration from reaching a
n O
level required to exceed the gypsum solubility product. However, calcium
sulfate is precipitated along with calcium sulfite in the regeneration step
even though the mother liquor is unsaturated with respect to gypsum. The
mixed crystal which results is in effect a solid solution of the two salts.
This phenomenon also takes place in lime/limestone wet scrubbing systems.
Another sulfate removal method uses sulfuric acid to dissolve calcium sul-
fite, increasing the concentration of calcium ions in solution enough to
exceed the solubility product of calcium sulfate. The method requires more
HaSOi* than the stoichiometric amount indicated by Reaction 2.2.2-8, however,
and also increases calcium consumption. Thus, it is economically unfavorable
in cases of high oxidation or where the gypsum is not marketable. This is
the sulfate removal method used in most Japanese double alkali processes.98
Na2S04 + 2CaS03'%H20 + H2SOlf + 3H20 ->• 2NaHS03 + 2CaSOif-2H20 (2.2.2-8)
Forced oxidation - The quality of the solid wastes produced can be improved
by forced oxidation of the calcium sulfite to calcium sulfate (gypsum) .
2-86
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Sludges containing a high ratio of calcium sulfate to calcium sulfite are
less thixotropic than sludges having a higher proportion of calcium sulfite,
and can also be settled, filtered, and dewatered more easily. In addition,
calcium sulfate or gypsum has a higher compressive or load-bearing strength
and is suitable for landfill disposal. In the Buell-Envirotech process,
spent scrubbing liquor is oxidized to further increase the sulfate concen-
tration before it is contacted with lime, thus producing a sludge with higher
q q
gypsum content. In Japan, forced oxidation has been used to produce
gypsum solid wastes suitable, for wallboard or cement production, although
this is not a likely market in the U.S. A liquor purge stream may be neces-
sary to prevent the build-up of soluble species (i.e. , chloride and magnesium)
in systems using forced oxidation since little water is discharged with the
solid wastes. This purge stream, however, would be no larger than the amount
of water discarded with conventional sludges.
Regeneration with limestone - The double alkali processes developed in the
U.S. generally use lime for regeneration, but the use of limestone has been
investigated in laboratory tests by A. D. Little. °° The full-scale pro-
cesses developed in Japan utilize limestone for regeneration. In these
processes, flue gas is scrubbed with a solution of Na2SOs. A bleed stream
from the scrubber is neutralized with limestone, to regenerate the Na2SOs and
precipitate calcium sulfite. The calcium sulfite is reslurried, acidified
with sulfuric acid to reduce the pH and dissolve calcium sulfite, and oxi-
-dized with air to produce gypsum. The gypsum is dewatered to 5-10 percent
moisture in a centrifuge. A small slipstream of scrubber effluent is reacted
with calcium sulfite and sulfuric acid to remove sulfate. This combination
of processing steps makes the process relatively complex and expensive. High
levels of oxidation in the scrubber can be tolerated but are uneconomical
because of the expense of sulfuric acid addition and the additional limestone
required for neutralization.
Magnesium enters as an impurity with the limestone and tends to delay the
reaction of limestone with sodium bisulfite in the regeneration step. Lime-
stone is a less expensive raw material than lime but is less reactive so
that larger reactors and a more complex reactor system are required. In
2-87
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addition, limestone occurs in many forms which vary in reactivity. Process
control may thus be more difficult with limestone than with lime. More
solid wastes would also be produced with limestone. Nevertheless, the use
of limestone may be an attractive alternative if it can be accommodated with
a reasonable degree of increase in process complexity, control requirements,
and capital costs.
C. Applicability—Double alkali systems may be somewhat easier to
retrofit than lime/limestone systems since, because of the lower liquid-to-
gas ratio, smaller equipment is used to contact the flue gas. However, a
large amount of land is still required to dispose of the calcium sulfite
waste sludge. Future use of double alkali systems on industrial boilers
is foreseen since these systems have already been applied successfully to
small industrial boilers.102
One disadvantage of the double alkali process is the production of large
quantities of solid waste. The waste consists primarily of calcium sulfite
and generally contains from 30 to 50 weight percent water. Because of high
concentration of soluble species in the scrubbing solution, the wastes will
also contain soluble solids (such as NaaSOa, NaaSCK, and NaCl as well as the
relatively insoluble calcium salts. The wastes are washed to recover the
sodium, and thus the soluble solids content can be reduced to less than 1
weight percent.103 Even with washing, however, the soluble solids content
of double alkali waste solids and the total dissolved solids content of the
adherent liquor will probably be somewhat higher than those typically present
in lime/limestone wastes.
D. Factors Affecting Performance—
1) Design and Operating Considerations - Potential operating problem
areas in double alkali systems that should be considered in system design an
operation include oxidation, scale formation, chloride buildup, and sludge
quality.
Oxidation - A potential problem in double alkali systems results from the
oxidation of sulfite to sulfate. The sulfate thus formed does not take part
2-88
-------
in SO 2 removal reactions, and is termed "inactive". If the sulfate is not
removed from the system, deterioration of S02 removal capability or precipita-
tion of sodium sulfate can result.
Theoretically, sulfates can be removed by precipitation as CaSOit-2H20,
according to Equation 2.2.2-5. This reaction, however, does not occur to a
significant extent in a solution with a high concentration of sulfite ions.
The dissolved calcium ions can bond with either the sulfite ions and precipi-
tate as calcium sulfite or with the sulfate ions and precipitate as calcium
sulfate. Since calcium sulfite is less soluble than calcium sulfate, it
will precipitate first, leaving the sulfate ions in solution.
If the concentration of sulfite and bisulfite is kept low (active sodium
concentration less than 0.15 molar), enough sulfite will be oxidized to
sulfate so that there will be more sulfate ions than sulfite ions in solu-
tion. If a sufficient excess of sulfate ions is present, the calcium sulfate
will be precipitated. A double alkali system in which calcium sulfate pre-
cipitates in this manner is classified as a "dilute" system. A system with
» Q ft.
a higher concentration of sulfite and bisulfite is a "concentrated" system.
Because of the lower active alkali concentration in the dilute system,
larger volumes of scrubber liquor must be circulated than in a concentrated
system to achieve the same level of S02 removal. The dissolved calcium con-
centration will also be higher than that in a concentrated system since
calcium sulfate is more soluble than calcium sulfite. Therefore, the dis-
solved calcium level must build up to a higher level before calcium sulfate
is precipitated. The higher calcium level can lead to gypsum scaling in the
scrubber, as discussed in the next section. Operation in the dilute mode
does allow for easy removal of sulfates, however, and is thus used for appli-
cations where the oxidation rate is expected to be high (for example, in
systems with low S02 and high 02 concentrations).
Under certain conditions, sulfate is coprecipitated with calcium sulfite,
forming a mixed crystal. The amount of sulfate removed in this manner
depends on the concentration of sulfate in the reactor liquor. The ratio of
2-89
-------
to CaSOs in the reactor solids appears to be roughly proportional to
the ratio of the concentration of SO^ to SOl in the reactor liquor.
Scaling - Gypsum and calcium carbonate scaling have occurred in double alkali
systems but can be prevented by proper design and operation. Gypsum scaling
occurs when sulfate ions formed by oxidation react with calcium ions in the
scrubbing liquor, forming gypsum scale. Gypsum scaling is generally not a
problem in concentrated systems, since the high concentration of sulfite
keeps the calcium ion concentration low. In dilute systems NaaCOs or COa
may be added to the liquor to form CaC03 which will precipitate to reduce the
calcium ion concentration before the liquor is returned to the scrubber.
Calcium carbonate scaling can occur in high pH scrubbing liquor when COa
absorbed from the flue gas reacts with dissolved calcium ions. This carbo-
nate scaling can be eliminated by pH control in the scrubber. At a scrubber
liquor pH below 9 the carbonate/bicarbonate equilibrium tends to limit the
concentration of free carbonate ion and thus prevent precipitation of cal-
cium carbonate. If the pH of the regenerated solution is above 9, it should
be mixed with the lower pH recycled scrubber liquor before being fed to the
absorber to avoid localized high pH areas that could result in carbonate
scaling.105
Chloride build-up - Another possible problem is the buildup of chlorides in
the system. Chlorides are absorbed from the flue gas, and the only mechanism
for them to leave the system is in the liquor contained with the solid wastes.
The wastes are washed to recover sodium, however, and this washing also
recovers chlorides that are then recycled to the absorber. In addition to
decreasing the concentration of active alkali in the absorber, high levels
of chlorides also can result in stress corrosion.
One possible solution that has been proposed is to use a prescrubber to
remove chlorides before the double alkali system. This method has been pro-
posed by Buell-Envirotech in their High Chloride process. The use of a
prescrubber with a separate liquor loop, however, could cause water balance
problems in the system. Since the evaporation loss would occur in the
2-90
-------
prescrubber, the only water loss from the double alkali system itself would
be the water occluded with the solid waste. This small water loss would not
allow addition of enough water for the normal cake washing (more than one
displacement wash), demister washing, pump seals, and lime slaking.106
Solids quality - Poor settling solids can result from several causes and
create problems in thickener and filter or centrifuge operation. ADL has
reported that solids settling deteriorates with high calcium sulfate concen-
trations in the solids. Normally calcium sulfate solids can be dewatered
more easily than calcium sulfite. In a concentrated double alkali system,
however, the sulfate is incorporated with the calcium sulfite in a mixed
crystal rather than as distinct CaSO^*2H20. High levels of sulfate in the
mixed crystal appear to inhibit solids settling. High concentrations of
magnesium also inhibit solids settling. In addition, improper control and
excess lime feed can make solids settling more difficult.
2) Fuel variations - The importance of fuel characteristics to dual
alkali systems is chiefly determined by the following factors:
• Sulfur content
Chloride content
• Ash alkalinity
HHV
The effects of these parameters are similar to effects on other FGD systems
as discussed in Section 2.2.I.I.D.2. One fuel element which is especially
important in dual alkali systems is the potential for sulfite oxidation.
Combustion of low-sulfur coals produces conditions which promote oxidation.
since the ratio of oxygen to sulfur dioxide in the flue gas is higher than
in high sulfur coal applications. Therefore, the sulfur content is a major
factor in the selection of design mode and operability of dual alkali sys-
tems; is.e., concentrated systems are more suited to application where low
oxidation (high sulfur fuel) is expected. l °8 Although not a fuel variable, high
excess oxygen levels which arc commonly found in inrHtntrial boilers will also
cause high oxidation rates and may affect the design of the FGD system.
2-91
-------
Another fuel variable of special importance in dual alkali scrubbing is
the chloride content. This species is difficult to purge from the system,
but if allowed to build up will significantly reduce the S02 absorption capa-
city of the alkali and/or cause materials problems as discussed previously.
3) Ambient variations - FGD systems are essentially independent of am-
bient variations. However, as with all wet systems, extreme cold can adversely
affect the operation but this effect can be designed for and eliminated by
using heat traced lines or by enclosing the FGD system in a building to pre-
vent freezing.
2.2.2.2 System Performance —
A. Emission reductions — Double alkali systems have demonstrated S02
removal efficiencies up to 99 percent and have been used to treat gases
with inlet SOa concentrations as high as 3800 and as low as 250 ppm.
Table 2.2.2-3 summarized some of the design features and performance
characteristics of operational dual alkali systems.
PEDCo recently summarized some guidelines for achieving high SOz removal
efficiencies from dual alkali and sodium carbonate systems based on available
data. Results of their study recommended the following design and operat-
ing practices:
Upstream flue gas treatment - Prescrubbing with a separate water
recirculating system for particulate removal and chloride control
for high chloride coal (>0.04 weight percent Cl in the coal).
Scrubber type - Use of a two-stage tray or packed tower absorber
--L/G 1.3-2.7 £/m5 (10-20 gal/1000 acf)
—absorber pressure drop 15-30 cm (6-12 in.) of water
pH - recycle liquor pH in 6.0 - 7.0 range
Reactor tank residence time -
—In lime regeneration system, approximately 10 minutes with
lime utilization ^90 percent
—In limestone regeneration, approximately 30 minutes with
reactant utilization of 75-85 percent.
Soda ash (sodium carbonate) addition - to effect carbonate
softening and sodium ion makeup
2-92
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TABLE 2.2.2-3. PERFORMANCE OF DUAL ALKALI FGD FACILITIES IN THE U.S. 109
Company/Utility '.
Plant /Unit;
Location '•
Size of System ;
Fuel
Characteristics '•
Active Alkali
Mode :
Startup I
Status ;
S02 inlet '.
S02 outlet:
S02 removal
efficiency :
General Motors
Chevrolet
Parma, Ohio
124 m'/s
(262,000 acfm)
32MW
Coal
1.5-3.0% sulfur
Dilute
March 1974
Operational
800-1300 ppm
20-130 ppm
90-99 percent
Caterpillar Tractor Co.
Joliet Plant
Joliet, Illinois
48.8 m'/s
(103,500 acfm)
18MW
Coal
4% sulfur
Dilute
September 1974
Operational
2300 ppm
115-350 ppm
85-95 percent
Firestone Tire and
Rubber Co.
Pottstovm Plant
Pottstown, Penn.
6.6 m'/s
(14,000 acfm)
3MW
Coal
3% sulfur
Concentrated
January 1975
Operational
1000 ppm
100 ppm
90 percent
Caterpillar Tractor Co.
Mossville Plant
Mossville, Illinois
113 m'/s
(240,000 acfm)
57MW
Coal
3.2% sulfur
Concentrated
October 1975
Operational
—
—
90 percent
Utah Power and
Light Co.
Gadsby Station/Unit 3
Gadsby, Utah
1.2 m'/s
(2500 acfm)
M).6MW
Coal
0.4% sulfur
Dilute and
Concentrated
1971
Terminated 1973
250-1500 ppm
15-40 ppm
90 percent
Gulf Power Co.
Scholz, Unit 1
Chattahoochee,
Florida
35 m'/s
(75,000 acfm)
20MW
Coal
3-5% sulfur
Concentrated
February 1975
Terminated July
1976
1800-3800 ppm
(dry)
—
95 percent
-------
Some actual test data which document the above guidelines are presented
in this section. These specific results are from different test facilities
operating under given sets of test conditions, and therefore, may not be
generally applicable to all dual alkali FGD systems.
Table 2.2.2-4 summarizes the key operating parameters and results of
three test periods for the EPA's sponsored tests at General Motors Parma
facility and Figure 2.2.2-2 shows the flow diagram and normal operating con-
ditions. The original literature should be consulted for more details.
In general, it was found that despite low S02 concentrations in the flue gas
(800-1300 ppm by vol.) caused by high excess air, S02 removals in the 90 per-
cent range were achieved. This exceeded the average statutory requirement
of 60-70 percent S02 removal for the coals burned (1.5-3.3 percent S;
11,000-13,500 Btu/lb). Poor S02 removal (20-85 percent) in one operating
mode was attributed to lower pH in the scrubber which resulted from a modifi-
cation to the process feed system. Also, calcium sulfite plugging occurred
during this operating mode. This mode was, therefore, abandoned for further
tests.114
The effects of scrubber feed pH and L/G were investigated during pilot
testing of Envirotech's double alkali system at the 85 M3/min (3000 acfm)
unit at Gadsby Station.115'115 The scrubber used for this work was a poly-
sphere unit having perforated plastic balls supported on trays for the actual
scrubber packing. Figures 2.2.2-3 and 2.2.2-4 present the results. In
general, it was found that S02 removals decreased dramatically below pH 6,
and that S02 removal could be maximized by using multiple trays and by raising
the pressure drop across the scrubber. Increasing the liquid flow within a
certain range for a given gas flow also increased S02 removal.119
Performance results have also been reported for the CEA/ADL tests on
pilot and prototype levels.120'121'122 The Scholz prototype unit demonstrated
high S02 removal capabilities of the dual alkali system when operating in a
concentrated mode (0.2-0.4 M active sodium). In over 15 months of operation
the average S02 removal was 95.5 percent (inlet S02 800-2800 ppmv) using both
the venturi and absorber. With the venturi alone slightly lower removals
2-94
-------
TABLE 2.2.2-4. KEY OPERATING PARAMETERS AND RESULTS FOR INTENSIVE TESTING AT PARMA
112
I
VO
Ln
8/19/74-9/13/74
S02 removal 90-95% (rpt. by GM)
Recycle pH
Scrubber feed location
Lime stoichiometry mole Ca/mole S in cake
Soda ash stoichiometry mole Naa/mole SO.,
A
removed
Solubles loss mole Naa/mole S in cake
Oxidation mole CaSOit/mole CaSO in cake
1 I
Ca in scrubber, ppm
Scaling?
Na = Total, M in scrubber feed
Na = Active, M in scrubber feed
Cake washing?
Total solids, wt.%
Soluble solids, wt.% of wet solids
Insoluble solids, wt.% of wet solids
Soluble solids/total solids
5.5-6
Top tray
1.90
_*
0.11
60%
305
Yes(CaC03)
0.58
0.087
No
47.1
2.4
44.7
0.05
Operating period
2/17/75-3/14/75
65%
7-7.5
Recycle tank
1.65
0.08
0.12
47%
465
Yes (CaSO 3)
0.66
0.13
No
41.4
2.4
39.0
0.06
4/19/76-5/14/76
90%
6.0
Recycle line
1.32
0.12
0.03
83%
490
No
0.96t
0.12
Yes
56.0
1.0
55.0
0.02
*No data taken on S02 removal
Includes %0.18M Na+ as NaCl
-------
Flue Gjs
ho
I
L/G = 2.7t/m3
(20 Gal/1000 cu. Ft.)
800—1300 ppm
vv) Weight
S02 Concentration
s ) Sample Point
Sludge
Blend
Tank
Flue Gas From
Boiler No. 1 Only
(Typ.)
02 =* 10%
15-20% Solids
Filtrate
pH= 12.2-12.5
Soda Ash C02
0]
Solids Contact
Reactor #2
Water
0
Surge
Tank
Figure 2.2.2-2. Data collection points and normal operating conditions.
I I 3
-------
100T
90
Q
UJ
?? sol
HI
cc
«
O
70
60
o o
L/G 2.5 4/m3
AP - 4 in. H20
Two-stage absorber
345678
SCRUBBER EFFLUENT pH
Figure 2.2.2-3.
removal versus scrubber effluent pH for the
Envirotech/Gadsby pilot plant with a two-stage
absorber-117
2-97
-------
100
JL/ACTUAL m°
2 3 4
95
90
CM
O
l/l
85
80
75
Fuel - coal
pH - 5.8-7.1
Single stage absorber
10 20 30 40
L/G - Gal/1000 acf
Figure 2.2.2-4.
removal versus L/G ratio for the
Envirotech/Gadsby pilot plant with a
single stage polysphere absorber.118
2-98
-------
were obtained, averaging 90.7 percent with SO inlet of 700-1300 ppmv. The
chief contorl parameter was pH, with little effect on SO removal found due
to highly fluctuating gas flows.
The effects of pH are shown in Figure 2.2.2-5 and 2.2.2-6 for two coal
types. The two curves illustrate the range of data for each of the opera-
ting configurations evaluated. The operating pH of the scrubber system
can be adjusted by changing the sorbent feed rate and/or pH of the regen-
erated liquor. In concentrated systems the liquid flow requirements for
SOz removal are low, as illustrated in Figure 2.2.2-7 which shows the
required stoichiometry to achieve various SO- removal levels over a range
of inlet concentrations. To achieve 95 percent removals, L/Gs on the order
of 3.3 £/m3 (25 gal/lOOOacf) in the venturi and 0.7-0.9 £/m3 (5-7 gal/lOOOacf)
in the tray tower were required.
The laboratory and pilot plant work were performed on both dilute and
concentrated modes. Results regarding S02 removal and oxidation which have
recently been published125 were based on higher scrubber temperatures
(60-65°C (140-150°F)) than those normally encountered' in conventional boiler
flue gas applications (50-54°C (120-130°F)). This would have the effect of
decreasing S02 removal efficiency (due to elevated SOa partial pressures for
a given solution) and increasing oxidiation rates. The results, however, are
still very useful in generating design data for larger systems. The effects
of scrubber feed stoichiometry on S02 removal for various S02 inlet concentra-
tions are presented in Figure 2.2.2-7. In any range of S02 concentration,
increasing stoichiometry increased the S02 removal, while a higher stoichio-
metry was required to achieve a given S02 removal in the lower inlet S02
range than in the higher inlet S02 range.
B. Availability/reliability—Since 1973 double alkali systems have
been started up on three large utility boilers and several industrial boilers
in Japan. Although detailed operating details are not available, few
2-99
-------
250
200-•
150--
100--
50--
4.0
•n,.tS02 Venturi AP ("H20)
(ppmv) 4.5-7 11-12.5
2400 ± 200
2000 ± 200
low AP
High AP
4.5
5.0 5.5
Scrubber Bleed Liquor pH
6.0
6.5
Figure 2.2.2-5.
°
CEA/ADL Dual Alkali Process - SOa removal as a
function of pH - high-sulfur coal.123
2-100
-------
500
t-o
I
400
-- 300
Q.
£.
CM
O
CO
« 200
+-*
3
O
100
4.5
yVenturi plus
trays
Operational Configuration
• Venturi + 2 Trays
o Venturi + 2 Trays
A Venturi (No Feed to Trays)
Venturi
AP, (in. H20)
5-7
8-11
8-11
Inlet SO2 = 1050 — 1250 ppmv
Venturi with
no trays
Active
Na+, (M)
0.25-0.4
0.15-0.3
0.15-0.3
5.0
5.5 6.0
Scrubber Bleed Liquor pH
6.5
7.0
Figure 2.2.2-6. CEA/ADL Dual Alkali Process - S02 removal as a function of pH -
low/medium-sulfur coal.124
-------
100
95
90
85
80
75
70
0.7
Scrubber Operating Conditions
Blri-il Liquoi Tcmpri.iUiri' 140°F 150°F
Activ,.' [Na' I 0.3 -0.5 M
Total Dissolved Solids 5-15 wt"..
Sulfili! Oxidation 250 ppm S02 Equivalent
Vrrmiii: P 1 1 14 in. H20
LG 15 - 17 cial'Macf satuiatnd
O
high SO2
inlet -»•
O
0.8
0.9
Range of SO2 Inlet Levels
D 500-700 ppm
A 800-1.000 ppm
02,250-2,650 ppm
1.0
1.1
1.2
1.3
1.4
1.5
Scrubber Feed Stoichiometry (mols Ma capacity/Tools SO2 inlet)
Figure 2.2.2-7.
Effect of feed Stoichiometry on removal efficiency in
the Venturi/2-tray tower absorber for the EPA-ADL
double-alkali pilot program.126
2-102
-------
problems have been encountered.12 Scaling and corrosion, while not rare in
any FGD operation, are reportedly controlled so that they do not hinder
overall plant operation.128 Operabilities for the Shikoku Electric Sakaide
plant and the Showa Denko Ichihara plant were reported to be 95 and 98 per-
cent, respectively, with corresponding SOz removal efficiencies of 95 and
90-95 percent.129 Operability is defined as the ratio of the FGD system's
operating hours to the scheduled operating hours of the gas source, normally
11 months continuous operation with a one month shutdown for utility boilers.
Since there are few double alkali systems with long operating histories
in the U.S., it is difficult to predict the long-term reliability of a system
for a given industrial boiler application. The largest operational double
alkali system in the U.S. is operated by Caterpillar Tractor in Mossville,
Illinois. This system, which was designed by FMC, removes the full load of
fly ash from the flue gas. Several problems have been experienced including
plugging of the demister (which has occurred several times), difficulty in
controlling the liquor flow rate to the absorber, difficulty in dewatering
fly ash, and erosion problems.130 Most of these problems, however, are
apparently due to the large amounts of fly ash in the system. Thus, a system
installed with upstream particulate removal should be more reliable than the
Caterpillar installation.
The 3 MW demonstration unit designed by FMC for Firestone Tire and
Rubber, Pottstown, Pennsylvania, had an overall reliability (based on boiler
operation) of about 95 percent when firing with oil (93.5 percent the first
year) and 88 percent when firing with coal. Figure 2.2.2-8 illustrates the
reliability of this system over the last four years. Nothing was spared
in this small installation, including pumps, which accounted for 15 percent
of the total system downtime the first year. Most of the remaining down-
time the first year was due to thickener pluggage and problems with the
cake conveyor, fan, lime feeder, and spray nozzles, with less downtime
due to instrumentation and control valves. During the period of coal
2-103
-------
O
-P-
100 -
90 -
80 -
g 70 H
o
V-i
0)
CM
— 60 H
2 5° •
cfl
•H
H
0)
« 40 -
•M
W
30 -
20 -
10 -
1st
2nd 3rd
1975
4th I 1st
2nd 3rd 4th I 1st 2nd 3rd
1976 1977
Time in Quarters of Years
4th I 1st
2nd 3rd
1978
4th
Figure 2.2.2-8. Performance trends for the Firestone double alkali system
1 31
-------
firing, most of the downtime was due to fly ash, which the system was not
designed to handle. Erosion caused by fly ash has ruined the fan and
damaged the venturi nozzles. Firestone has been satisfied with the relia-
bility demonstrated by the system. They are investingating the possibility
of converting their boilers to coal, using the double alkali process to
132
control 862 emissions.
The oldest full-scale double alkali system in the U.S., at FMC in Modesto,
California, treats gas from reduction kilns. The gas has a high SO2 content,
up to 8000 ppm SO 2. The system reportedly has had an availability over
95 percent since it started up in December 1971. In 1977 the scrubbing system
reportedly did not cause any kiln downtime (100 percent operability). The
scrubber (a packed tower) is flushed with water for cleaning during normal
plant maintenance shutdowns, which occur for several days about 3 times a
year. The gas to be cleaned contains calcium, which has led to scaling in
the scrubber packing. The packing has been changed twice and cleaned with
acid four times since 1971, but no scaling has been noted in the year since
the recirculation rate was increased. 33
The purpose of the 20 MW prototype unit at Scholz, designed by CEA/ADL,
was to characterize and evaluate the performance of the double alkali pro-
cess. It was not intended as a demonstration unit nor a test of the ultimate
reliability possible from a full-scale application. Nevertheless, the over-
all operative record of the system over the 17 months of operation from
initial startup through completion of the test program was impressive,
including delays in receiving replacement parts and in awaiting repair of
the boiler air preheater. Most of the downtime occurred between operating
periods; during the operating periods the availability averaged about 90
percent. The various parameters actually achieved by this unit are shown
in Table 2.2.2-5.
As reported by Rush and Edwards the process operability (ease of
operation) was excellent in all respects. The system was successfully
2-105
-------
TABLE 2.2.2-5,
CEA/ADL DUAL ALKALI PROCESS VIABILITY PARAMETERS
(February 1975-July 1976)
1 3 if
Parameter
Availability*
Reliability^
OperabilityT
§
Utilization Factor
Parameter 5
Value
%
78
80
70
58
95
Total time
Numerator
9,679
7,128
7,128
7,128
9,679
(hours)
Denominator
12,376
8,911
10,172
12,376
10,172
*Availability: Hours the FGD system was available for operation (whether
operated or not) divided by hours in the period, expressed as a percentage.
Reliability: Hours the FGD system was operated, divided by the hours the
FGD system was called upon to operate, expressed as a percentage.
TOperability: Hours the FGD system was operated divided by boiler operating
hours available to the process in this period-, expressed as a percentage.
§
Utilization Factor: Hours that the FGD system operated divided by total
hours in this period, expressed as a percentage.
Parameter 5: Hours the FGD system was available for operation (whether
operated or not) divided by the hours the boiler was available to the pro-
cess, expressed as a percentage.
2-106
-------
operated over a range of widely fluctuating inlet flue gas sulfur dioxide
levels (1.6 to 4 wt% sulfur fuel), oxygen concentrations, and flow rates
with little or no change in the sulfur dioxide removal efficiency, waste
cake properties, or lime utilization. It demonstrated an ability for con-
tinuous operation with large and frequent variations in pH in both the
scrubber circuit (4 to 7) and the regeneration/solids dewatering section
(6.5 to 13).
The most impressive aspect of system operation was its resistance
to short-term upsets. Low soluble calcium levels throughout the
system, even during most upset conditions, resulted in a low potential
for calcium/sulfur salt precipitation, particularly in the scrubber cir-
cuit. In addition, operation for over 4500 hours without mist eliminator
washing and without any mist eliminator scaling confirmed the viability of
system operation without a mist eliminator wash.
The dual alkali process does have a limitation in that there is a minimum
level of sulfur content in the fuel below which it cannot be successfully
operated in the concentrated mode due to the potential for gypsum scaling.
This lower limit is a function of the rate of oxidation in the system and is
therefore dependent not only on the sulfur content in the fuel, but also the
level of oxygen in the flue gas. For typical pulverized-coal-fired boilers, the
design of a concentrated-mode dual alkali system for use on flue gases produced
from the combustion of coal containing between 1 and 2 wt% sulfur will have to
be considered carefully on a case-by-case basis. For fuels containing less
than 1 wt% sulfur, a concentrated-mode dual alkali system cannot be operated
at the excess air levels that are typical for pulverized-coal-fired boilers
without an intentional purge of sodium sulfate, either in the filter cake
or as a separate purge stream. However, above 2 wt% sulfur in the fuel,
and in many cases for coal containing between 1 and 2 wt% sulfur, the oper-
ation of the system is excellent. In fact, overall system operation improves
significantly as higher sulfur content fuel is utilized.
2-107
-------
It is difficult to generalize with respect to the long-term reliability
potential of any mechanical/chemical process. However, the results at
Scholz indicate that the overall performance of a properly designed and
operated dual alkali system can be superior to that of direct lime and
limestone systems because:
• As discussed above, the system is highly
resistant to upset. The potential for
calcium/sulfur salt precipitation (scaling)
is eliminated by the high sulfite levels
in solution, except in extreme upset
conditions.
• The handling of slurries in the absorption
section is completely eliminated.
* The most important control parameter—pH—
has a wide, acceptable range of operation,
and unlike lime and limestone systems,
has no effect on scaling up to values of
12.5 to 13 in the reactor system.
The General Motors' double alkali system at Parma, Ohio, was started up
in February 1974. The dilute process was the subject of an EPA-sponsored
test program from August 1974 to May 1976, during which significant improve-
ments in both process and mechanical performance were observed. Test objec-
tives included determination of SOa removal capability, process reliability,
sulfate control, waste characteristics, degree of closed loop operation, lime
and other chemicals utilization, and material balances. The test program
involved three one month intensive tests and 18 months of lower level nonin-
tensive testing. Numerous modes of operation were investigated, necessitating
shutdowns for equipment changes and modifications. There were, however,
several shutdowns due to equipment or other operational malfunctions as
described in detail in the literature.137 The overall results and conclu-
sions regarding system reliability were that, in part due to its then develop-
mental status, it was judged "not yet proven over an extended period to be a
commercially viable process for SOz control." However, significant
improvements observed during the last four week test period (April-May 1976)
were thought to indicate that the system had capability for long-term
2-108
-------
reliability. The total scrubber availability to the boilers over the whole
test program was 77.9 percent, excluding four long-term planned shutdowns
139
for system modifications.
During the two year period since the end of the EPA study the Parma
FGD system has continued to operate, experiencing varying operability indexes
(1.25-100 percent). Figure 2.2.2-9 presents operability data from the
General Motors Parma system over the last four years. Shutdowns have been
for annual scheduled overhauls, severe winter conditions, low boiler loads
during summer months, and in some cases mechanical problems. The most
frequent problem was continued solids buildup in the mix tank. To resolve
this problem the mix tank agitators have been totally redesigned and re-
placed by high-carbon steel, propeller types.1 ^
C. Impacts on boiler—The major impacts of a dual alkali FGD system
on boiler operations are similar to those of a lime/limestone system; 'i.e.,
1) power consumption for running the FGD system's pumps and fans and
2) possible boiler load reduction during FGD system outages if no bypass is
permitted. The power consumption for the CEA/ADL prototype unit at Scholz
was reported to be 2.5-3.0 percent of the unit's generating capacity with the
system operating at full fan and at full venturi and absorber liquid recir-
culation capacity. Correcting for the excess fan and pump capacity, the
power consumed by the equipment actually required in the application (tray
tower at an L/G of 0.7-1.3 £/m3 or 5-10 gal/1000 ft3) is approximately 1-1.5
percent of the design generating capacity. lltl Although an industrial boiler
generally produces steam instead of electricity, the proportion of boiler
product usage by the dual alkali system should be about the same assuming
that steam can be used to drive the pumps and fans.
Boiler load reduction due to FGD system outages will be the other major
impact. The relatively high system reliabilities exhibited by existing
double alkali processes will help to reduce these outages to a minimum.
2-109
-------
100 -\
90-
80-
S 70-
!-)
0)
PM
4-1
•H
60-
•H
.0
2 50
0)
a.
o
e
01
4-1
cn
40 -
30-
20-
10 -
, , n . , , r
1st 2nd 3rd 4th I 1st 2nd 3rd
1976 1977
Time in Quarters of Years
1st 2nd
3rd 4th
1975
4th
1st
T
2nd 3rd
1978
T
4th
Figure 2.2.2-9. Performance trends for the GM Parma double alkali system
1 to
-------
D. Additional maintenance requirements—Operation of a dual alkali
system increases the total maintenance requirement of an industrial boiler
facility. Manpower requirements for the GM double alkali process is reported
to be 1.4 men/shift for direct operation and 0.5 man/shift for maintenance.
FMC claims that under most conditions one boiler house operator is sufficient
to operate the FGD system. In comparison to direct lime/limestone scrubbing,
the maintenance requirement should be less because the scaling potential is
minimized and a. solution rather than a slurry is used for scrubbing. 1 lf3
2.2.3 Wellman-Lord Sulfite Scrubbing Process
2.2.3.1 System Description—
A. System—The Wellman-Lord Sulfite Scrubbing Process marketed by
Davy Powergas is based on the ability of sodium sulfite solution to absorb
S02, thus forming a solution of sodium bisulfite which can be thermally
regenerated. It is a regenerable process and is presently being commercially
employed on a large scale. A concentrated stream of SOa is produced which
can be processed to elemental sulfur, sulfuric acid, or liquid SOa. A by-
product purge of sodium sulfate is produced as the result of sulfite oxidation.
Antioxidants have been used to reduce the sulfate purge rate, but have been
found to be uneconomical.
A simplified process flow sheet appears in Figure 2.2.3-1. The Wellman-
Lord Process consists of five basic processing steps:
1) Gas Pretreatment
2) Absorption
3) Purge Treatment
4) Regeneration
5) SC>2 Conversion
2-111
-------
AOSURBEK
Figure 2.2.3-1. Process flow diagram Wellman-Lord Process,
-------
No unusual or unique equipment is used in any of these areas with the possible
exception of the S02 conversion step which is licensed technology. The gas
pretreatment and absorption sections are essentially the same as those found
in most aqueous scrubbing systems.
1) Gas pretreatment - The flue gas to be treated is taken downstream
of the electrostatic precipitator. This gas, which is at a temperature of
about 300°F, is passed through a venturi or tray type prescrubber where it
is cooled to around 130°F and humidified. The venturi scrubber is preferred
because it removes 70 to 80 percent of any remaining fly ash and 95 to 99
percent of the chlorides. A tray-type prescrubber is satisfactory for cool-
ing and humidifying low pressure drop, but is less efficient than a venturi
for fly ash and chlorides removal. Humidification of flue gas in the pre-
scrubber prevents evaporation of excessive amounts of water in the absorber.
A well-designed prescrubber can remove up to 99 percent of all chlorides
145
in the flue gas. This should help maintain a low level of chloride in the
scrubbing liquor, and reduce the potential for stress corrosion. Fly ash and
other solids collected by the prescrubber are pumped to the ash disposal pond
as about 5 percent slurry. When absorption of HC£, 862, and SOs causes the
slurry water to become too acidic, the slurry is neutralized with lime before
it is pumped to disposal.
2) Absorption - Humidified gas from the prescrubber is passed through
the absorption tower where the 862 level is reduced by at least 90 percent.
The cleaned gas may be reheated by heat exchange with high-pressure steam and
exhausted to the atmosphere. There are various alternatives to this method
of reheating the gas, but use of steam allows coal to be used indirectly
rather than premium fuels such as oil or natural gas.
2-113
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Absorption of S02 proceeds according to Equation 2.2.3-1.
Na2S03 + S02 + H20 -> 2NaHS03 (2.2.3-1)
Makeup sodium carbonate also reacts with SO 2 in the absorber to form
sodium sulfite by Equation 2.2.3-2.
Na2C03 + S02 -> Na2S03 + C02 (2.2.3-2)
A very important side reaction, which will be discussed in detail later, is
the oxidation of sulfite to sulfate by oxygen in the flue gas as in Equation
2.2.3-3.
Na2S03 + %02 -> Na2S0lt (2.2.3-3)
Some sodium sulfate is also formed by absorption of S03 from the flue gas:
2Na2S03 + S03 + H20 ->• 2NaHS03 + Na2S0lf (2.2.3-4)
Davy Powergas offers two types of absorption units, a packed tower for
small volume applications and a valve tray tower for large volume applications.
The valve tray unit is generally built in a square configuration and includes
three to five trays depending on the inlet S02 concentration and the degree
of desulfurization. Because of the large capacity of concentrated sodium
sulfite solution to absorb S02, the feed liquor flow is fairly small and re-
circulation is practiced on each stage to maintain good hydraulic character-
istics over the trays. With the recirculation, the L/G ratio is kept at about
3 gal/mscf, per tray. Superficial gas velocity in the absorber is about 10
feet per second.
2-114
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Scaling and plugging problems in the absorber are virtually eliminated
because:
the prescrubber removes solids,
the scrubber solution is clear, and
the absorption product, sodium bisulfite, is more soluble
than sodium sulfite.
3) Purge Treatment - About 15 percent of the absorber product liquor is
drawn off to the purge treatment area for separation of sodium sulfate. Feed
liquid is cooled by heat exchange with treated liquid. In a chiller-
crystallizer the purge is further cooled to about 32°F and a mixture of
sodium sulfate and sulfite is crystallized out. The slurry is put through
a centrifuge to produce a 40 percent solids cake and the cake is dried by
steam. A small liquid purge stream is also drawn off the evaporators in the
regeneration section and added to the cake to remove some of the thiosulfate
formed in the system. The crystalline product is a mixture of anhydrous
sodium sulfate (70 percent) and sodium sulfite (30 percent), plus small amounts
of thiosulfates, pyrosulfites, and chlorides. The vent gases from the dryer
are scrubbed to remove dust and recycled to the main flue gas stream. The
centrate is heated by passing through the feed cooler and is returned to the
product liquor stream entering the evaporator loop. Refrigeration for the
chiller-crystallizer is provided by an ethylene glycol system.
4) Regeneration - The regeneration section consists of a set of double
effect, forced circulation evaporators, condensers, a condensate stripper,
and a dump dissolving tank. Regeneration of sodium sulfite is performed by
simply reversing the absorption reaction through addition of heat as shown
in reaction 2.2.3-5.
2NaHS03 ~> Na2S03 + H20 + S02 (2.2.3-5)
2-115
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However, higher temperatures also increase the formation of thiosulfate by
the following disproportionation reactions.
6NaHS03 -»• ZNazSCH + Na2S203 + 2S02 + 3H20 (2.2.3-6)
2NaHS03 + 2Na2S03 -* 2Na2S04 + Na2S203 + H20 (2.2.3-7)
As depicted in Figure 2.2.3-1, the combined stream of absorber product liquor
and purge centrate is split between the two evaporator effects, each of
which operates under a vacuum, with 55 percent going to the first effect and
45 percent going to the second effect. The first effect operates at 200°F
and is heated with low pressure steam by way of an external shell and tube
exchanger. The S02 and water vapor driven off in the overhead from the first
effect are used to heat the second effect which operates at about 170°F. In
each effect the undissolved solids content of the recirculated liquor is
about 45 percent, primarily sodium sulfite. Operating at such a high solids
concentration eliminates the need to centrifuge the product stream going to
the dissolving tank. The regeneration reaction is limited by the equilibrium
concentration of sulfite ion in solution. Fortunately, since sodium sulfite
is less soluble than sodium bisulfite, it is continuously removed from
solution by crystallization, driving the reaction forward.
The S02 and water vapor from the evaporators is subjected to partial
condensation to remove most of the water and thus concentrate the S02 . The
condensate contains several hundred ppm of dissolved S02 and is steam strip-
ped to lower these values. Stripped condensate is returned to the dump
dissolving tank along with a small amount of makeup water and sodium carbo-
nate. This mixture is then agitated with the sodium sulfite slurry from the
evaporators to provide absorber feed. The S02 stream exiting the condenser
contains only 5 to 10 percent water. It is compressed and transferred to the
S02 conversion section.
2-116
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5) SOz Conversion - One of three product options may be selected for
conversion of gaseous S02 to useful chemicals. Sulfur dioxide may be:
compressed and liquified for sale,
catalytically oxidized for production of sulfuric acid
(Contact Process), and
reduced to elemental sulfur.
These options are listed in order of increasing processing cost.
The questions of what to do with the SOa product stream will have to be
answered separately for each individual site. But for production of sulfur
to be a viable alternative, a process able to use a reductant other than
methane will often be necessary. Current Wellman-Lord sulfur producing
systems employ the Allied Chemical sulfur production unit which uses methane
as the reductant. In some isolated cases, methane may be readily available
but in general the most reliable sources of reducing agent will be coal or
coke. Coal might be used in solid form as in Foster Wheeler's RESOX process
or as H2/CO producer gas. Unfortunately, this whole area is a source of
uncertainty because to date, only the Allied Chemical process using methane
has been demonstrated on a large scale. The use of reducing gas from coal
gasification in the Allied Process is still in the relatively early develop-
mental stage.
B. Status of development—The Wellman-Lord Process was first conceived
in 1966 and a pilot plant using potassium sulfite rather than sodium sulfite
was operated in 1967. The problems encountered with scaling and high steam
consumption on this original pilot induced a switch to the sodium system.
The potassium bisulfite product is less soluble than potassium sulfite, while
sodium bisulfite is more soluble than sodium sulfite. Thus, in a sodium sys-
2-117
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tern, as S02 is absorbed and water evaporates, the saturation point will not
be exceeded because the product of absorption, sodium bisulfite, is more
soluble than the reactant, sodium sulfite.
The first sodium sulfite system was installed for Olin Corporation in
1970 to treat a 45,000 scfm stream of acid plant tail gas. Acid plant tail
gas differs from industrial boiler flue gas in that it generally contains
a higher SO2 concentration (up to 10,000 ppm) and a lower oxygen concentra-
tion. Despite some initial difficulties, the plant operated successfully.
Subsequently, the process was applied to an oil-fired industrial boiler of
the Japan Synthetic Rubber Company in 1971. Again, the plant was operated
successfully after some initial difficulty. This unit has since been able
to achieve better than 90 percent removal of SO2 with an onstream factor of
better than 97 percent.
Wellman-Lord systems currently account for about 4 percent of the total
USA operating and planned utility FGD capacityt^7 and about 10 percent of
1 M- 8
total Japanese capacity. Although the fuel base is not specified in these
statistics, it is safe to say that the U.S. systems are predominately
coal-fired, and the Japanese, oil-f ired.1 "*9
Tables 2.2.3-1 and 2.2.3-2 summarize the operating Wellman-Lord systems
in the U.S. and Japan, and Table 2.2.3-3 summarizes the systems planned
for U.S. operation. ' '
/
Table 2.2.3-4 gives worldwide distribution of Wellman-Lord units. Of
these 31 installations, two of the most significant are the unit in operation
since 1973 for Chubu Electric Power's Nagoya Station and the unit at Northern
Indiana Public Service Company's Dean H. Mitchell Station. The Chubu plant
is significant because it is the largest Wellman-Lord system with a fairly
long operating history. The application is a 220 MW, oil-fired peaking
plant. This Wellman-Lord Unit has been highly successful, achieving 90
percent removal of S02 consistently with a high on-stream factor. The
2-118
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TABLE 2.2.3-1. SUMMARY OF OPERATING WELLMAN-LORD SYSTEMS IN THE U.S.150
Company /location
Olin Chemical*
Paulsboro, New Jersey
Std. Oil of California
El Segundo, California
Allied Chemical
Calumet, Illinois
Olin Chemical
Curtis Bay, Maryland
K>
1 Std. Oil of California
j^ Richmond, California
St. Oil of California
El Segundo, California
Northern Indiana Public
Service
Gary , Ind iana
Public Service Co. of
New 'lexico
Waterflow, New Mexico
Public Service Co. of
New Mexico
Waterflow, New Mexico
Completion
date
July 1970
September 1972
November 1972
May 1973
August 1974
January 1975
December 1976**
November 1978
November 1978
Feed gas Gas f
origin 1000 Nm'/hr
Sulfuric acid 76
plant
Claus plants 51
Sulfuric acid 51
plants
Sulfuric acid 133
plants
Claus plant 51
Claus plant 51
Coal-fired 527
boiler
(115 MW)
Coal-fired
boiler
(314 MW)
Coal-fired —
boiler
(306 MW)
Design
.ow S02 concentration,
(scfm) ppm
(45,000) in 6,000
out 500
(30,240) in 10,000
out 250
(29,850) in 2,700
out 250
(78,046) in 4,000
out 250
(30,000) in 10,000
out 250
(30,000) in 10,000
out 250
(310,000) in 2,200
out 200
90% removal
— 90% removal
Disposition
of S02
Recycle to acid
Recycle to Claus
Recycle to acid
Recycle to acid
Recycle to Claus
Recycle to Claus
Elemental sulfur
Elemental sulfur
Elemental sulfur
plant
plant
plants
plants
plant
plant
Plant operation suspended as of January 1, 1976.
integrated operation of the plant was in December , 1976; however , the Wellman-Lord sys tern
began operations in July 1976, before coniple tion of the Allied sulfur recovery unit .
-------
TABLE 2.2.3-2. SUMMARY OF OPERATING WELLMAN-LORD SYSTEMS IN JAPAN
151
Company /location
Japan Synthetic Rubber
Chiba
Toa Nenryo ,
Kawasaki
Chubu Electric Power,
Nagoya
Sumitomo Chemical,
Sodeguara
Japan Synthetic Rubber
Yokkaichi
Kashlma Oil,
Kashima
Toa Nenroyo ,
1 Hatsushima
| — >
fo Toyo Rayon,
° Nagoya
Japan National Railway
Kawasaki
Mitsubishi Chemical,
Mizushima
Kurashiki Rayon,
Okayama
Fuji Film,
Fu j inomiya
Shin Daikyowa,
Yokkaichi
Sumitomo Chemical,
Niihama
Mitsubishi Chemical,
Mizushima
Mitsubishi Chemical,
Kurosaki
Tohoku Electric Power,
Niigata
Completion
date
August 1971
August 1971
May 1973
November
1973
December
1973
February
1974
October
1974
December
1974
September
1974
April 1975
July 1975
1974
December
1975
February
1976
August 1976
September
1976
March 1977
0Gas flow,
Feed gas origin
Oil-fired boilers
Glaus plants
Oil-fired boiler
(220 MW)
Oil-fired boiler
Steam boiler
(2% S fuel oil)
Glaus plants
Glaus plant
Oil-fired boiler
Steam boiler
(200 MW-3% S
fuel oil)
Oil-fired boiler
Oil-fired boiler
Oil-fired boiler
Oil-fired boiler
Oil-fired boiler
Oil-fired boiler
Oil-fired boiler
Oil-fired boiler
1000 mJ/hr
200
67
620
360
450
30
17
330
700
628
410
150
400
155
628
530
380
(scfm)
(124,000)
(41,000)
(390,000)
(225,000)
(280,000)
(20,200)
(10,000)
(218,000)
(435,000)
(373,000)
(248,000)
(89,000)
(253,000)
(91,000)
(390,000)
(330,000)
(236,000)
SO- concentration,
ppm Disposition of SO.
in 2,100
in 6,500
out 200
in 1,600
out 150
in 1,550
out 250
in 1,000
out 100
in 11,000
out 200
in 18,580
out 250
in 1,500
out 150
in 1,500
out 45
in 1,500
out 150
in 1,500
out 150
in 1,300
out 125
in 1,500
out 150
in 1,600
out 130
in 1,300
out 130
in 1,500
out 75
in 1,000
out 100
Sulfuric acid
Recycle to Glaus
Sulfuric acid
Sulfuric acid
Sulfuric acid
Recycle to Glaus
Recycle to Glaus
Sulfuric acid
Sulfuric acid
Sulfuric acid
Sulfuric acid
Liquid SO
Sulfuric acid
Sulfuric acid
Sulfuric acid
Sulfuric acid
Sulfuric acid
plants
plants
plant
-------
Table 2.2.3-3. SUMMARY OF WELLMAN-LORD SYSTEMS
PLANNED IN THE U. S.
152
System Operator
and Location
Size
(MWe)
New or
Retrofit
Startup Disposition
Mo / Yr of S02
Public Service Co. of 468
New Mexico
San Juan #3
Public Service Co. of 472
New Mexico
San Juan //4
Delaware Power & Light 180
Delaware City, 1,2,+ 3
New
New
January '81 Acid
January '82 Acid
Retrofit April '80
2-121
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nature of the source is such that the FGD system has had to follow rapid
load swings from 35 to 105 percent of plant nameplate capacity, thus its
flexibility and turndown capabilities are well proven. The flexibility is
1 c q
obtained mainly by the provision of large solution surge storage.
TABLE 2.2.3-4. DISTRIBUTION OF WELLMAN-LOKD SULFUR DIOXIDE REMOVAL PLANTS
(Operational and Planned)1 51f'1 55
Country
USA
Japan
W. Germany
TOTAL
Coal-fired
Industrial
-
-
-
0
boilers
Utility
6
-
2
8
Oil-fired
Industrial
-
11
-
11
boilers
Utility
-
3
-
3
Glaus /Acid
plant
6
3
-
9
Total
12
17-
2
31
The NIPSCO plant is important because it is the first large-scale instal-
lation in the U.S. and the first full-scale demonstration of the process on a
power plant burning high sulfur coal. The Wellman-Lord unit is installed on
the Number 11 boiler at the Mitchell station. The nameplate rating of the
boiler is 115 MW and it burned 2.9 percent sulfur coal during its demonstration.
Another installation of significance is the unit to go on-line at New
Mexico Public Service's San Juan Station. This unit is designed to treat
flue gas from a 314 MW boiler burning low sulfur New Mexico coal. The
Allied process will be used to produce sulfur from the concentrated S02.
Currently, New Mexico Public Service Co. is planning to install Wellman-Lord
systems on an additional 1200 MW of electrical generating capacity with
planned start-up date from July 1978 to January 1982.156
Current areas of concern are improvement in evaporator and purge treat-
ment design. The fractional crystallization method of purge concentration
was developed by Davy Powergas' two Japanese licensees. Some effort has
also^ been spent to devise an economical means of converting the sodium sul-
tate to a form such as sodium carbonate or caustic soda which could be returned
2-122
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to the system in order to reduce the sodium makeup requirements. So far this
work has not been completed. Ionics' SULFOMAT electrolytic cell and various
techniques from the paper industry have shown that sodium sulfate conversion
or reduction is possible, but the economics are still uncertain and some
technical problems remain.
The use of an antioxidant has been evaluted by the process developer in
order to reduce the quantity of sodium sulfate formed. However, the chemical
compounds which were able to substantially decrease the oxidation rate turned
out to be so expensive that the cost of using the antioxidants was greater
than the cost of simply replacing the sodium sulfite which was converted to
sodium sulfate. Therefore, use of antioxidants has been dropped in favor of
various operational techniques to minimize oxidation by reducing oxygen
transfer.
Data for selection of the materials of construction is plentiful due to
the large amount of previous experience. The materials which will be
employed are generally carbon steel lined with fiberglass reinforced polyester
(FRP), 304, 316 and 316L stainless steel, all of which are fairly common.
The only area where material problems are still encountered is the prescrubber.
There the chlorides and ash may make the use of Hastelloy G necessary.
C. Applicability of Wellman-Lord (W-L)—Worldwide, Wellman-Lord has been
applied to industrial and power boilers as well as to Glaus sulfur units and
acid manufacturing tail gases. These applications range from about 17 to
1100 MW (30,000 to 2,510,000 scfm).157 All industrial boiler applications
have been on oil-fired boilers in Japan.
Approximately 2/3 of U.S. industrial boilers are of 60,000 Ib/hr (VLO MW)
steam capacity or less. 158 A particular concern in applying W-L to
small coal-fired boilers is that these boilers are often stoker-fired, and
may have no provision for over-firing. When the operator must burn under-
sized coal containing a high percentage of fine particles, as sometimes
delivered, incomplete combustion occurs. To compensate for these factors,
the boiler operators use high percentages of excess air. One FGD vendor
2-123
-------
measured ^240% excess air in one flue gas test of a coal-fired industrial
boiler. High excess air leads to excessive sulfite oxidation in the absorber
liquor (Reaction 2.2.3-3). An increase in the sodium sulfate purge,
of course, increases both disposal and soda ash makeup costs.
Since the W-L process is a clear liquor process, it offers a flexibility
not easily achieved in slurry processes. Due to the high SOz loading of the
scrubbing liquor the absorber and regenerator can be decoupled by storing
absorption liquor such that a centralized regenerator may serve several
absorber units.
D. Factors affecting performance—
1) Design and Operating Considerations - A design area of major concern
in the Wellman-Lord process is sulfite oxidation. Other areas of concern are
the evaporator-crystallizer, and the availability/cost of reducing agents.
The oxidation rate of the sulfite scrubbing solution is a function of
several factors: impurities in the solution, recirculation rate, temperature
and oxygen content of the flue gas, contact efficiency, pH, and solution
strength.
It is probable that certain components of fly ash such as iron and man-
ganese catalyze the oxidation and thus increase the total sulfate formation.
Despite the upstream electrostatic precipitator and venturi prescrubber, some
small amounts of fly ash will be picked up by the scrubbing liquor. The
liquor is routinely filtered as it leaves the absorber to prevent build-up of
these solids. Davy Powergas has noted about 25 percent increase in oxidation
for oil-fired boiler applications over acid plant tail-gas applications where
fly ash is not present. This information was obtained from data on Wellman-
Lord units now in commercial operation. Short term tests on coal-fired
boiler flue gas, however, indicated that fly ash from coal does not increase
the oxidation rate over that obtained in oil-fired boiler applications. Long-
term data on oxidation in applications with fly ash from coal-fired boilers
has been obtained from the 115 MW demonstration unit at Northern Indiana
Public Service Company's Dean H. Mitchell Station, where acceptance tests
2-124
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159
were completed September 14, 1977. Long-term tests are currently under
way. Unpublished results showed no significant effects of fly ash upon the
overall system oxidation rate over that obtained from oil-fired operations.
The liquor recirculation rate on each stage of the absorber is governed
by the need to maintain adequate flow for good hydraulic characteristics.
If liquid flow across the trays is too low, poor distribution of liquor occurs
and stagnant areas will result. Any recirculation above that needed for this
reason merely increases the oxidation rate without increasing SOa removal.
The reason for this is that SC>2 transfer from gas to liquid is controlled by
gas film resistance while oxygen transfer is controlled by liquid film resis-
tance. Increasing liquid recirculation decreases liquid film resistance to
oxygen transfer.
The absorption temperature is fixed at the adiabatic saturation temper-
ature of the gas, about 125°F. Oxygen content of the flue gas is determined
by the amount of excess air fired. Flue gas N02 may be a significant contri-
butor to oxidation but the major cause is direct oxidation of sulfite to
sulfate by oxygen from the gas.
Oxygen transfer is impeded by salts in solution so the fresh scrubbing
liquor is originally saturated with sodium sulfite. This also minimizes the
amount of solution to be circulated. The pH range of 7 to 5.5 is controlled
by the sulfite-bisulfite equilibrium. Contact efficiency is a function of
tray design which is set by the level of S02 removal.
As a result of the oxidation problem, several approaches have been tried
to reduce the amount of purge and the amount of makeup sodium carbonate.
The most successful method so far has been the development of a fractional
crystallization technique to increase the concentration of sodium sulfate in
the purged solids to about 70 percent. This step was demonstrated in the
U.S. for the first time at NIPSCO. It has been used three times previously
in Japan. Further improvements in sulfate concentration are still possible.
2-125
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2) Operating procedures - Operation of small coal-fired industrial
boilers may deviate from large boiler operation because of several factors:
• Small companies will probably not have the luxury of long-term
coal contracts for coal of consistent physical and combustion
characteristics.
• Coal feed systems are more likely to be stoker type with
accompanying fuel: air ratio control-problems.
• Some boilers may not have overfiring capability.
These factors all call for incremental excess air over what a utility boiler
would use. Air rates xrould also be set high to assure combustion for:
• the finest particle size coal,
• the highest HHV coal, and
• the highest fuel feed rate excursion at a given set point.
High excess air provides driving force to increase the sulfite-to-sulfate
oxidation rate. As previously discussed this leads to increases soda ash
makeup and solids disposal costs.
3) Performance vs. maintenance - The Wellman-Lord system uses a
clear liquor for absorption and consequently does not have the deposition
problems of lime/limestone slurry absorption. A Wellman-Lord representative
estimates maintenance at 3.5 percent of fixed capital which, if accurate,
is moderate.J 6 2
4) Fuel variations - As discussed under Operating Procedures, the
boiler operator will ensure that the boiler does not smoke or develop a
reducing zone around boiler tubes by resorting to increased excess air.
This protects against short-term variations in fuel:air ratio, but leads to
increased sulfite oxidation.
2-126
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The main fuel characteristics that will affect Wellman-Lord operations
are sulfur content and chlorine content. The effects of varying these fuel
characteristics on this system are similar to the effects on other wet FGD
systems and have been previously discussed in Section 2.2.1.1.
E. Retrofits—Other than problems associated with operating a Wellman-
Lord system in a relatively high oxygen atmosphere, space limitations will be
the major concern in retrofitting the system to industrial boilers. However,
a Wellman-Lord system may be easily decoupled by adding intermediate liquor
storage between the absorption and regeneration section, thus limiting the
required space around the boiler. Wellman-Lord also affords the possibility
of one regeneration section to economically serve several small boiler-S02
absorber systems if the boilers are located a reasonable distance from one
another.
2.2.3.2 Wellman-Lord FGD System Performance—
A. Emission reduction—In September, 1977, the Wellman-Lord system
installed at NIPSCO completed acceptance testing. Since that time, the
system has been undergoing a year of extensive demonstration testing.
Results of those tests have not yet been reported, but will be forthcoming
in early 1979, Currently, the system is undergoing a second year of tests.
Although results of the extensive testing at NIPSCO are not yet available,
both Japanese and U.S. sources point to S02 removals of 90 percent and
system operabilities of greater than 95 percent. 3
Test results are, however, available from the acceptance testing con-
ducted in 1976 and 1977- These are relatively short-term monitoring results,
but they do illustrate the systems' S02 removal ability. In addition, EPA
has recently completed a program to collect long term continous monitoring
data from this system. Results of that monitoring indicated an average SO
removal of 89 percent throughout the monitoring period. 61*
2-127
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Figures 2.2.3-2 through 2.2.3-4 illustrate three periods of sustained
operations prior to the acceptance tests. As can be seen from the figures,
the system effectively removed SCL from the flue gas for the first and
second run. Poor SCL removals during the third run resulted from poor
quality feed solution caused by mechanical problems in the soda ash feed
system and evaporation area, and low feed rates to the absorber.165
3,000
2,000 h
Q.
IN
8
7,000 I-
KUN DURA TION, days
THE S03 CONCENTRA TION CUR VES HA VE BEEN
EXTRAPOLA TED THROUGH DA YS 3 AND 5 BECAUSE
OF INOPERA Tl VE INS TRUMEN TA TION.
Figure 2.2.3-2. Inlet and outlet S02 concentrations during run no. 1.
i ss
2-128
-------
2,000
7,500
7,000
500
200
0
5 70
RUN DURA 7VO/V, rfuyj
15
Figure 2.2.3-3. Inlet and outlet S02 concentrations during run no. 2.
166
3,000
2,000 -
7,000 -
5 70
RUN DURA TION, days
THE POOR S02 RECOVERIES DURING THIS PERIOD
RESULTED FROM POOR QUALITY SOLUTION CAUSED
BY MECHANICAL PROBLEMS IN THE SODA ASH FEED
SYSTEM AND EVAPORA T/ON AREA, AND LOW FEED
RA TES TO THE ABSORBER WHILE BALANCING TANK
INVENTORIES.
75
Figure 2.2.3-4. Inlet and outlet SOz concentrations during run no. 3.
1 6 7
2-129
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The actual acceptance tests were divided into two phases. Phase I was
a 12 day test at an average flue gas flow of 320,000 acfm, and Phase II was
an 83 hour test at an average flue gas flow of 390,000 acfm. Results of the
Phase I tests are shown in Figure 2.2.3-5. This figure shows that the S02
removal of the system was in excess of 90 percent (the minimum acceptable
level) throughout the test period. In addition, particulate emissions from
the absorber remained below the Federal NSPS of 0.1 lb/106 Btu throughout
the test period.
92
£ 91
o
90
CM
O
00
89
MINIMUM ACCEPTABLE
TEST DURATION, days
10
11
12
Figure 2.2.3-5. S02 removal efficiencies during 12-day test.
IBB
During the Phase II, 83 hour test, the FGD system was also required to
achieve an S02 removal of 90 percent. Under the more stringent Phase II
operating conditions, the S02 removal efficiency averaged 91 percent.
Particulate emissions during these tests were also below the Federal NSPS
limit.169
2-130
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B. Impact of Wellman-Lord performance on boiler performance—Major
effects of a Wellman-Lord system on a boiler's performance would be:
1) boiler derating due to the parasitic energy required to run the pumps and
fans and to regenerate the sodium sulfite sorbent, and 2) boiler load reduc-
tions due to Wellman-Lord downtimes, assuming no bypass of flue gas is
allowed. Energy consumption for a 500 MW rated W-L system was reported to
be 32 MW, not including SC>2 conversion. 17° This amounts to a derating of
6.4 percent. Calculations were performed to estimate the energy requirements
for industrial boiler application (see Chapter 5) which showed the process to
consume from 3 to 8 percent of the net heat input to the boiler, depending
mainly upon the amount of S02 removal.
C. Additional maintenance requirements—Wellman estimates maintenance
cost at 3.5% of fixed capital. Since the W-L process is a clear liquor
process, it should not be expected to have excessive maintenance problems
in the absorber section. It is, however, more mechanically complex than a
once-through system such as lime/limestone slurry FGD.
2.2.4 Magnesia Slurry Absorption Process
2.2.4.1 System Description—
The Magnesia Slurry Absorption Process uses magnesium hydroxide to
absorb S02 in a wet scrubber. Magnesium sulfite is the predominant species
formed by the reaction with 502 in the scrubber according to Equation 2.2.4-1
Mg(OH)2 + S02 + MgS03 + H20 (2.2.4-1)
Reaction 2.2.4-2 occurs to a lesser extent.
MgS03 + H20 + S02 -> Mg(HS03)2 (2.2.4-2)
2-131
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The aqueous slurry is centrifuged and the cake is then dried with fuel oil
combustion gas to remove free and bound moisture. The magnesium oxide is
regenerated in a calciner by thermal decomposition of magnesium sulfite
according to Equation 2.2.4-3.
MgS03 + MgO + S02 (2.2.4-3)
Hot combustion gases are again used to heat the magnesium sulfite in the
calciner and to remove the 862. The SOa gas stream may be used to produce
sulfuric acid or elemental sulfur.
Three magnesia-based wet scrubbing processes have been developed since
the early 1930's: 1) A basic (high pH) MgS03/Mg(OH)2 slurry process,
2) an Mn02 activated absorbent slurry system, and 3) an MgSOs acidic clear
liquor process. The basic slurry process is the most advanced system and,
therefore, will be discussed in this chapter.
A. Process description—The basic magnesia scrubbing process can be
divided into four major process areas: SOz absorption, MgSOa/MgSOi* separa-
tion and drying, MgO regeneration and SOz recovery, and SOz conversion.
Figure 2.2.4-1 is a simplified flow diagram for the process.172
1) S02 absorption - Absorption of S02 takes place after the flue gas
is treated for particulate removal in a wet scrubber or electrostatic pre-
cipitator. A separate system ahead of the 862 scrubber is used for particu-
late removal for two reasons. First, it eliminates some components of the
fly ash such as vanadium and iron compounds which can catalyze the oxidation
of MgS03 to MgSOif. Also, there is no easy way to remove fly ash from the
circulating scrubbing slurry.
2-132
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REIIEATER
SO} ABSORBER
FLUE QAS •
N3
OJ
OJ
EFFLUENT HOLD TANK
^~"
TO STACK
COKE
CALCINED
Figure 2.2.4-1. Process Flow Diagram for the Magnesia Slurry Absorption Process.
-------
An aqueous slurry of magnesium hydroxide and magnesium sulfite (pH range
6.5 to 8.5) is used to absorb the S02 according to Equations 2.2.4-1 and
2.2.4-2. The MgS03 is formed as a crystalline solid in slurry. A bleed
stream is sent to a centrifuge as a first step for MgO recovery. Makeup
water, recycle MgO, and makeup MgO are added to the recycle slurry to main-
. . 173
tain a constant slurry composition.
Sulfite oxidation gives rise to sulfates in the system by the following
reaction:
MgS03 + %02 + MgSOij (2.2.4-4)
The test facility at PEPCO-Dickerson reported sulfate concentrations to be
about 7 mole percent for the anhydrous solids transferred to the regeneration
section.171* The sulfite and sulfate solids precipitate as hydrated crystals
as illustrated in the following equations:
MgS03 + 6H20 -»• MgS03«6H20 (2.2.4-5)
MgS03 + 3H20 •> MgS03-3H20 (2.2.4-6)
MgSO.* + 7H20 -* MgSOi»-7H20 (2.2.4-7)
MgS03*6H20 is the preferred form because of its large, easily separable crys-
talline form. Laboratory work indicates that it can be preferentially formed,
given proper design and operating conditions.175 The bisulfite in the spent
scrubbing liquor is reacted with magnesium hydroxide which is formed by
slaking the fresh and recycle magnesium oxide.
Mg(HS03)2 + Mg(OH)2 + 4H20 -> 2(MgS03'3H20) (2.2.4-8)
MgO + H20 •*• Mg(OH)2 (2.2.4-9)
2) MgS03/MgSOi( separation and drying - After absorption of S02 in the
scrubber, a portion of the slurry from the main scrubber circulation loop is
2-134
-------
removed as a 10-15 weight percent slurry and sent to a thickener. The
thickener is an optional piece of equipment that has the potential for
increasing the surge capacity in the recovery portion of the plant and for
improving centrifuge operation.176
A stainless steel, solid bowl centrifuge has been used to recover a wet
cake of MgSOs/MgSCK hydrate crystals. Satisfactory centrifuge operation is
necessary so that solids are removed at a sufficient rate to maintain control
of recycle solids concentration. Wet crystals are discharged from the cen-
trifuge through a vertical chute into a screw feeder which provides a seal
and a continuous flow of wet solids into a rotary fluid-bed dryer. The
rotary kiln type dryer is presently used in the three U.S. magnesia scrub-
bing demonstration units. Combustion gas from an oil burner, which can be
tempered by a sidestream of stack gas, is used to dry the crystals. The
dried MgSOs/MgSOit is discharged from the dryer and conveyed to a calciner for
Mgo regeneration and S02 recovery.
3) MgO regeneration and SOz recovery - Dried MgSOs/MgSCK solids are
heated in an oil-fired rotary kiln or fluidized-bed reactor until decomposed.
The main decomposition reaction is shown in Equation 2.2.4-3. The MgSOi* is
also reduced in the calciner using carbon as a reducing agent.
+ hC -> MgO + S02 + %C02 (2.2.4-10)
Two installations have used a rotary kiln to regenerate the MgO. High
dust losses in the rotary kiln require the use of a hot cyclone and venturi
scrubber to recover all of the MgO. If a fluidized-bed reactor were used
most of the MgO formed would go overhead with the S02 and combustion gases
and separation equipment would also be required.
2-135
-------
The optimum calcining temperature in the reactor is set by the fact that
it must be high enough to decompose all of the MgS03/MgSOi+ solids without
"dead burning" the MgO. "Dead burned" MgO is that which has been melted into a
refractory like material and is chemically unreactive and not effective for
further S02 removal. Operating temperatures in the 815 °C (1500°F) range
have been used for rotary calciners in this service.
4) SOa conversion - After dust removal, the sulfur dioxide rich gas
from the calciner is piped to either a sulfur or sulfuric acid production
unit. This concentrated SOa stream is actually not as well suited for sulfur
production as it is for acid production. This is because of the oxygen in
the concentrated S02 , which is introduced by the excess air used for combus-
tion of fuel oil in the calciner. This oxygen will consume additional reduc-
tant than what is required for reduction of SOa to sulfur, thus increasing
sulfur production costs. However, the SOa stream is well suited for acid
production, as it is at approximately 38°C (100°F) , is saturated with water,
and contains 8-10 percent
B. Status of development — The magnesia slurry scrubbing process has
been shown to be feasible on a full-scale size. Three retrofit units in the
U.S. of the 95-150 MW size have demonstrated greater than 90 percent SOa
removal on both oil-fired and coal-fired systems. * A list of the
operating and planned magnesia scrubbing units is shown in Table 2.2.4-1.
MgO units make up about 5 percent of total U.S. FGD capacity. None of the
U.S. applications of MgO technology are on industrial boilers.179
The magnesia scrubbing process has been used on a commercial scale at
three locations in Japan, and constitutes about 1 percent of total FGD
i R n
capacity. A summary of the status of the three Japanese installations is
shown in Table 2.2.4-2. SOa removals of 90 to 99 percent have been demon-
strated for the Japanese units. The Mitsui unit uses a special cross-flow
type absorber which they have developed. The scrubber consists of an empty
chamber with two rotating shafts with many spoons. The scrubber has little
2-136
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TABLE 2.2.4-1. OPERATING AND PLANNED MAGNESIA
U.S. POWER PLANTS AS OF AUGUST
SCRUBBING UNITS ON
1978 181> 182> 183
I
M
to
Utility Company, New or
Power Station Retrofit
Boston Edison, Retrofit
Mystic No. 6
Potomac Electric Retrofit
and Power,
Dickerson No. 3
Philadelphia Retrofit
Electric Co.
Eddystone No. 1A
Philadelphia Retrofit
Electric Co.. ,
Eddystone No. IB
Philadelphia Retrofit
Electric Co. ,
Eddystone No. 2
Philadelphia Retrofit
Electric Co. ,
Cromby
* Reference 181 given long term
Reference jgg gives long term
Size of
FGD Unit
(MW)
150
95
120
240
336
150
S02 removal
SOa removal
Fuel,
Process Sulfur Content,
Vendor (%)
Chemico Oil, 2.5
Chemico Coal, 2.0
United Coal, 2.5
Engineers
United Coal, 2.5
Engineers
United Coal, 2.4
Engineers
United
Engineers
of 77-87%.
of 77-83%.
TABLE 2.2.4-2. OPERATING MAGNESIA SCRUBBING UNITS ON
Type of
Company Location Plant
Onahama Smelter Copper
Onahama, Japan Smelter
Mitsui Mining H2SO.,
Hibi, Japan
Idemitsu Kosan Glaus
Chiba, Japan Unit and
Boiler
Size of
FGD Unit
(MW)
28
24
162
Flue Gas
Process S02 Content
Vendor (ppm)
Onahama- 15,000-25,000
Tsukishima
Mitsui- 1,500-2,000
Mining
Chemico-
Mitsui
SO?
Recovery
(%) Start-Up Date, Status
90* Start-up in April, 1972; Test program com-
pleted in June, 1974. Not currently opera-
tional.
90 Start-up in September, 1973; Test program
completed in September, 1975. Not currently
operational.
96-99 Start-up in September, 1975; Currently
operational.
Start-up due in June, 1980; Letter of
intent signed.
Start-up due in June, 1980; Considering
FGD system.
Start-up due in June, 1980; Considering
FGD system.
JAPANESE POWER PLANTS AS OF AUGUST 1978 181*
S02
Recovery By-Product
(%) (tons/day) Start-Up Date, Status
99.5 H2SOI,-240 Start-up in December, 1972;
Currently operational.
90 H2SO,,-18 Start-up in October, 197];
Currently operational.
95 Sulfur-70 Start-up in 1974; Currently
operational.
-------
possibility of scaling because of the simple structure, but the size of the
scrubber is limited to the treatment of no more than 1700 Nm /min (60,000
scfm) or about 29 MW.
The magnesium oxide FGD system at Boston Edison has been able to demon-
strate 80 percent availability during sustained operation. The Potomac
Electric Dickerson station was only able to operate at 64 percent availabil-
ity for their SO2 recovery system during their best month.
C. Applicability of the magnesia slurry FGD process (MgO)—Magnesia
slurry has been used in Japan for oil-fired power plant flue gas cleanup and
on industrial tail gases.185 In the U.S. the process has been applied to
both oil- and coal-fired utility boilers but has not been used on an indus-
trial boiler.186 The relative complexity of this process may limit applica-
tion to larger industrial boilers or to multiple-boiler situations where
multiple absorbers and a common regenerator might be installed.
D. Factors affecting performance—
1) Design and operating considerations - It is not entirely clear
which design factors have most influenced the contrasting performance of
magnesia FGD systems as reported. Generally, high SOz removals are achieva-
ble because magnesium sulfite is relatively soluble. The amount of soluble
alkalinity available in magnesia scrubbing systems is more than that available
in lime/limestone systems but less than for sodium-based scrubbing systems.
Therefore, magnesia systems can operate at a lower L/G than lime/limestone
systems but at a higher L/G than sodium-based systems. An L/G of 33 gal/1000
acf was used at the Boston Edison facility using a venturi scrubber.187
The major design variables that can be used to regulate the scrubbing
operation are the recycle pH and percent solids. A higher pH gives a
higher S02 recovery. The pH can be controlled by MgO addition while the
amount of solids in the slurry can be controlled by adjusting the rate of
bleed from the recycle slurry.
2-138
-------
The process sequence for magnesia slurry scrubbing has been fairly
standard for the U.S. installations that have already been operated. A cen-
trifuge and rotary dryer have been used in all three U.S. demonstration plants.
Two of the installations used a rotary calciner while Philadelphia Electric
uses a fluidized bed unit to calcine the MgSOa/MgSOif product. A thickener
is an optional piece of equipment that has been used to concentrate the
slurry prior to the centrifuge. The use of a thickener gives the prospect
for improved centrifuge operation.
Particulate control is carried out separately from the SOz scrubbing
system. High dust removal efficiencies are required to minimize the amount
of contaminants that enter the system. Small amounts of fly ash will still
enter the circulating slurry system since particulate control devices are
not 100 percent effective in removing fly ash.
Contaminant control is required to reduce the amount of impurities that
will accumulate in closed-loop operation. Both soluble and insoluble impuri-
ties are of concern. Insoluble impurities come primarily from the fly ash
in the flue gas and coke used in the calciner. These insoluble impurities
are controlled by high efficiency particulate removal equipment upstream of
the S02 absorber and by the use of coke with a low ash content.
Soluble contaminants enter the system from the makeup water, makeup
MgO, fly ash, and calciner coke. These soluble species must be purged from
the system. Chloride attack of scrubber internals can occur when carbon
steel or stainless steel are used as absorber construction materials ; conse-
quently a prescrubber can be used for chloride removal, Glass reinforced
polyester resins have been successfully used to prevent corrosion in the scrubber.
Heat recovery from the dryer and calciner off gas is desirable but is
complicated by the entrainment of MgSOs or MgO fines in the combustion gases.
The dryer off gas at 200°C (400°F) and the calciner off gas at 870°C (1600°F)
offer a useful source of energy savings for the process.
2-139
-------
2) Operating procedures - Numerous operating problems occurred early
in the program at Boston Edison. Most of the problems were of a materials
handling nature resulting from the characteristics of the solids generated
in the scrubber loop. It was found that trihydrate crystals with an average
size of 10-15 microns were formed in the absorber instead of the larger
hexahydrate which had been formed in the pilot plant. As a result, the
centrifuge cake contained as much as 25 percent unbound moisture. This led
to problems of solids adhesion to the dryer drum.
These problems were solved by several operating and design modifications.
Among these modifications were changing the dryer to function as a granulator
and adding hammers to loosen any material which tended to adhere to the dryer
shell. The granulator material discharge was screened and sent through lump
breakers to eliminate oversize agglomerated granules of the magnesium sulfite.
The dryer off gas was sent to the 862 absorber to prevent high dust losses.
This subsequently caused the loss of 8°C (15°F) of reheat of the saturated
flue gas expected from the dryer off gas.
Other process problems occurred in the calcining system. A rotary kiln
has been used for these operations. The formation of the very fine trihydrate
crystals in the oil-fired power plant application also resulted in dusting
problems in the rotary calciner. The facility at Essex Chemical resolved
the dusting problem in the calciner by the use of a cyclone followed by a
venturi scrubber to remove all of the MgO fines from the gas. Leakage of air
into the calciner was a problem since high oxygen levels interfere with the
reduction of MgSOi,. Installation of new seals on the rotary calciner cor-
rected the problem.
In addition, high calciner temperatures can cause sintering or "dead
burning" of the regenerated MgO which will result in unreactivity of the
product for reuse in the scrubber. High reactivity magnesia is favored by
low calciner temperature and by increased amounts of carbon in the calciner
2-140
-------
feed. Recycle MgO reactivity was improved by the correct calciner operating
conditions, by pulverizing the calcined MgO, and by heating the MgO slurry
tank. More recent tests indicate that continuous stable operations are
possible with high SOz removal efficiencies and that continuous use of
recycle MgO has only a slight effect on the system.
The chemical and mechanical performance of the scrubber was excellent
at Boston Edison. No internal plugging was encountered and the polyester
lining of the scrubber was in sound condition after two years of intermittent
operation.l88
Erosion and corrosion were experienced in the carbon steel recirculating
slurry piping. The use of rubber-lined pumps, valves, and piping in certain
areas of the system is considered to be a practical solution to this problem.
The slurry recirculating pumps in the system have withstood corrosion using
316 stainless steel impellers.189
Potomac Electric's coal-fired Dickerson Station FGD system achieved
only 78-83 percent SOa removal and 48 percent availability at 75 percent
design over a five month test period.190 The maximum Dickerson stream time
without shutdown from equipment malfunction was 120 hours.191 Although
corrosion, erosion, leaks, centrifuge outages, etc. were mentioned as con-
tributing to low reliability, no figures were given to indicate outage time
as a function of the type of equipment failure, except for leaks. A report
of Dickerson operations revealed during the test period there were nine major
outages, three of which were leaks, and that there were forty-seven leaks
observed during four months of record keeping, sixteen of which induced shut-
downs. Plant construction with off-spec pipe, fittings, and rubber linings
was blamed for many equipment failures. 9 2
2-141
-------
3) Performance vs. maintenance - Maintenance problems at the coal-
fired Dickerson plant appear to. have stemmed primarily from erosion/corrosion
These problems were in part blamed on construction with off -specif ication
pipe, fittings and insufficient rubber piping liners. Over 40 leaks occurred
in approximately 5 months of testing.
4) Fuel variations - The long-term problem of acquiring coal of con-
sistent quality applies to owners of all small-scale industrial boilers,
who will probably not be able to write the same long-term contract that a
large utility company might. The day-to-day impact is that higher excess
air may be introduced to the boiler than would be used in a utility boiler-.
As a result, the flue gas will have a lower SOa concentration, and the
absorber will be larger- relative to the amount of S02 to be removed, for
a given heat release, coal sulfur content, and SOz removal efficiency.
2.2.4.2 Magnesia Slurry FGD System Performance —
A. Emission reductions — Only limited data are available from the full-
scale systems that show the interrelationship of SOa removal to the various
process variables. Operation of these units did, however, show that 90 per-
cent control of S02 was achievable. Part of the reason that data are not
readily available from these systems is the relatively low system avail-
abilities obtained at these installations.
For a given L/G the effect of inlet S02 concentration on S02 removal
efficiency was relatively minimal at high pressure drops at the 155 MW
Mystic facility, as shown in Figure 2.2.4-2. At lower pressure drops across
the venturi absorber significant effects were noted over the 400-2000 ppm
S0£ range. The relationship between S02 removal efficiency and pressure drop
is more directly illustrated in Figure 2.2.4-3.
In 1970 Babcock and Wilcox evaluated S02 absorption responses to various
process variables using both a floating bed absorber and a venturi scrubber/
absorber. Devitt, e+. o.l . , report that in order to achieve a 90 percent S02
removal efficiency, the pH of the scrubbing slurry should be maintained
2-142
-------
TOO
90 -
80 -
70
C£
CvJ
o
-p-
to
60
50
40 "—
200
I
400 600
AP = 30.48 G/SQ.CH (12 IN.)
AP'= 15.24 G/SQ.CH (6
AP = 10.16 G/SQ.CM (4 IN.J
AP = 7.62 G/SQ.CM. (3 IN.)
VENTURI ABSORBER
AT L/G = 40 gal/1000 acf
=5.3 H/m3
I I
800 1000 1200
INLET S02 - ppm
1400
1600
1800
2000
Figure 2,2.4-2.
Effect of inlet SOz concentration and venturi pressure drop on S02 removal
for the Mystic venturi absorber.197
-------
I
I—'
90
80
m 70
_j
<;
>
UJ
OL
CNJ
S 60
50
40
T
INLET S02 = 1000 ppm
INLET S02 700 ppm
INLET S02 400 ppm
VENTURI ABSORBER
AT L/G = 40 gal/1000 acf
= 5,3 ft/m3
j I i
5.08 10.16 15.24 20.32 25.40 30.48 35.56 40.64 45.72
PRESSURE DROP - GRAMS/SQUARE CENTIMETER
Figure 2.2.4-5. Effect of pressure drop on 862 removal for the Mystic venturi absorber.191
-------
between 6.0 and 7.5.195 This is based on data reported by Semrau shown in
Figure 2.2.4-4. A 227 kg/hr (500 Ib/hr) coal feed rate was used for these
tests. Results showed that 862 absorption efficiency was a function of the
bisulfite concentration and pH of the recycle liquid for a given absorber
L/G. Additional work, which was not particularly successful, attempted to
rationalize the mass transfer coefficient determined from this work to
fundamentally developed values.196
B. Impacts on boiler—The major impacts of a magnesium oxide FGD system
on boiler operations are similar to those of other wet FGD systems; i.e.,
1) power consumption for running the system's pumps and fans, and 2) possible
boiler load reduction during FGD system outages if no bypass is permitted.
The power consumption for a 500 MW magnesium oxide unit was reported to be
about 1.8 percent of the unit generating capacity. This consumption was
for the power to operate the systems pumps and fans and excluded other energy
requirements such as fuel oil for drying and calcining the magnesium sulfite
solids.
Boiler derating due to load reductions, however, is more difficult to
evaluate based on U.S. operating experience. Past U.S. experiences with
magnesium oxide systems indicate that this impact may be more significant
than for other FGD systems; however, Japanese magnesium oxide systems have
operated with high reliabilities.
C. Additional maintenance requirements—Reported estimates of mainten-
ance costs for the magnesium oxide process vary from four to seven percent
of fixed capital.200 This is higher than maintenance costs reported for
Wellman-Lord (3.5 percent). Experience to date in the U.S. points to
problems of corrosion and erosion of metal parts as the most clearly defined
maintenance problem affecting system performance.
2-145
-------
o
CD
CQ
CO
o
oo
100
99
98
97
96
95
94
93
92
91
90
89
88
87
86
85
I
I
1
1
6 7
pH OF SCRUBBING SLURRY
8
Figure 2.2.4-4. The effect of pH on S02 scrubbing efficiency.
199
2-146
-------
2.2.5 Sodium Scrubbing
The sodium scrubbing process is capable of achieving high S02 removal
efficiencies over a wide range of inlet S02 concentrations. The process
consumes a premium chemical and produces a soluble waste salt which under
current practice, is normally discharged to a lined evaporation pond for
drying .
2.2.5.1 System Description —
A. System — Sodium scrubbing processes currently being used for flue
gas desulfurization (FGD) on utility and industrial boilers employ a
wet scrubbing solution of NaOH, Na2C03 or NaHC03 to absorb S02 from the flue
gas. The operation of the wet scrubber is characterized by a low liquid-
to-gas ratio (L/G) and a clear scrubbing liquid because of the high
solubility of sodium salts. The absorption reactions which take place
in the scrubber are:
2NaOH(Jl) + S02(g) -> Na2S03 (A) + H-) (2.2.5-1)
Na2C03(£) + S02(g) -> Na2S03W + C02(g) (2.2.5-2)
2NaHC02 + S02(g) -> Na2S03(^) + H20 + 2C02(g) (2.2.5-3)
Na2S03 + S02(g) + H20 •* 2NaHS03 (2.2.5-4)
Simultaneously some sodium sulfite reacts with the oxygen in the flug
gas to produce sodium sulfate:
Na2S03W + Js02(g) + NaaSOifOO (2.2.5-5)
2-147
-------
The scrubber effluent solution thus consists of a mixture of Na2S03, NaHS03,
and Na2S04. Scrubber effluent from this nonregenerable process is then
treated in a variety of ways. If fresh caustic or soda ash is added for
makeup the majority of the effluent is recycled to the scrubber with a slip-
stream going to wastewater treatment and disposal. If a process waste stream
is used for scrubbing, the effluent is used on a once-through basis and then
disposed. Pulping operations can make use of the sulfite-bisulfite solution
in their digesters. Figure 2.2.5-1 presents a simplified process diagram
for a sodium scrubbing system.
Sodium scrubbing encompasses two categories; systems that use chemical
addition for reagent makeup and systems that use a process waste stream for
scrubbing. Chemical addition can be differentiated according to reagent,
either Na2C03 or NaOH. Some FGD systems that are located at a plant which
produces an alkaline waste'stream use the waste stream in their scrubbing
process. Soda ash plants which use end liquor and pulping operations which
produce a caustic waste stream are examples. Table 2.2.5-1 gives a summary
of the capacities of sodium systems that are operational or under construc-
tion in the United States.
TABLE 2.2.5-1. SUMMARY OF SODIUM FGD PROCESSES
ON U.S. COAL-FIRED INDUSTRIAL AND UTILITY BOILERS
201
System Category
System Category
Operational
Construction
Planning
Totals
Industrial Boilers
No. Units SCFM
102
19
4
125
4,145,000
656,000
286,000
5,087,000
Utility Boilers
MWe
2073
328
143
2544
No. Units
3
1
_
4
MWe
375
509
-
884
2-148
-------
I
M
_!>.
Co
Stack
Fresh Sorbent
Waste to
Treatment
Figure 2.2.5-1. Simplified Flow Diagram Sodium Scrubbing System.
-------
Even though the sodium scrubbing process is one of the least complex
FGD processes both chemically and mechanically, some accessory equipment is
required. Particulate control may or may not be used upstream of the absorber
depending on the absorber type. Flue gas reheat may be necessary, and facil-
ities for waste disposal are required.
If fly ash is not removed prior to S02 removal, a venturi scrubber may
be used for both particulate and SOz removal. The FGD system must then be
designed for slurry handling and must deal with special disposal considera-
tions. If fly ash is removed prior to S02 removal, the absorber can be
either a packed tower or a tray tower.
Solid reagent handling requirements may vary from site to site, but some
generalities can be made. Storage, usually silos, is required. Transport
from the silo to a mixing tank can be done by conveyor. A mixing tank is
needed to dissolve the reagent. The solution is sometimes then pumped to a
clarifier where insoluble impurities settle out. The clarified liquor is
then ready for use in the scrubber.
Wastes from wet sodium processes contain sodium sulfite, sodium sulfate,
sodium carbonate, sodium hydroxide, and some inerts. Disposal of this
material presents problems because it is highly water soluble. Several
disposal possibilities have been studied, but more research is required.
The disposal options studied include:
Developing a market for the resultant salt cake.
Short term storage of the waste with eventual recovery of
sodium products.
Insolubilizing the waste by forming complex salts.
Returning waste to mines.
Permanent storage of waste in lined and covered pits.
2-150
-------
At present, the trend for disposal practices is toward wastewater
treatment and holding ponds for evaporation. The specifics of the
202
wastewater treatment from users are not reported. However, a general
scheme for the treatment can be outlined from several sources. The
liquor must be clarified to remove fly ash and solids. The clarified
liquor may then be aerated to oxidize sulfite to sulfate to reduce
chemical oxygen demand. Solution pH is then adjusted and the liquor
discharged to sewers or rivers.
B. Development status—Sodium scrubbing systems are considered commer-
cialized technology. These systems have been developed in both the U.S. and
Japan. The first full-scale application of sodium FGD in the U.S. was at a
General Motors plant in St. Louis, Missouri. Two systems were installed on
coal-fired boilers in 1972. In the U.S. sodium carbonate scrubbing is
commercially practiced on 375 MW of generating power at the Nevada Power Reid
Gardner Station. The Reid Gardner Station has operated 250 MW of the system
since April 1974 and 125 MW of the system since July 1976. The system has
operated well without any major operating problems. FMC's soda ash plant
in Green River, Wyoming has had two scrubbers on two .100 MW (equivalent) units
since May 1976. Table 2.2.5-2 gives a summary of FGD processes on U.S. coal-
fired utility boilers. This table shows that in the utility sector, sodium
carbonate scrubbing is used on 375 MW from a possible 14,420 MW of operational
systems (about 3 percent). However, Table 2.2.5-3 shows that on industrial
boilers, sodium scrubbing is the preferred FGD process. Use of sodium
scrubbing on industrial boilers is because of several attractive features
of the system:
• It can use a process waste stream as the SCL sorbent.
• It uses a clear liquor rather than slurry which lessens
the potential for plugging and scaling.
• It is more tolerant of changes in boiler load conditions.
2-151
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Many of the sodium systems on industrial boilers were the result of adding
NaOH to the circulating liquor in particular scrubbers for pH control. It
was found that adding NaOH for pH control not only aided in minimizing corro-
sion, it also increased S0? removal to the point that the particulate scrubber
assumed a dual role of both particulate and SCL removal. Table 2.2.5-4
presents a summary of operating sodium scrubbing systems applied to U.S.
industrial boilers.
9 n 3
TABLE 2.2.5-2. SUMMARY OF UTILITY BOILER FGD PROCESSES
Percent of Capacity by Process Type
Process Operational Construction Total Market
Lime/Limestone
Sodium Carbonate
Magnesium Oxide
Double Alkali
Wellman-Lord
91
3
1
2
3
89
3
0
5
3
90
3
-
4
3
TABLE 2.2.5-3 SUMMARY OF INDUSTRIAL BOILER FGD PROCESSES 2° "
Percent of Capacity by Process Type
Process Operational Construction Total Market
Sodium Scrubbing
Double Alkali
Lime/Limestone
Others
76
11
3
10
54
35
-
11
72
16
2
10
2-152
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TABLE 2.2.5-4. PERFORMANCE DATA FOR OPERATING SODIUM SCRUBBING SYSTEMS205
S3
1
h-1
Ln
Installation/ location
Alyeska Pipeline
Valdez, Alaska
American Thread
Martin, NC
Belridge Oil
McKittrick, CA
Canton Textiles
Canton, GA
Chevron
Bakersfield, CA
FMC
Green River, HY
General Motors
Dayton, OH
General Motors
Pontiac, MI
General Motors
St. Louis, MO
General Motors
Tonawanda, NY
Georgia Pacific
Orosett, AK
Getty Oil
Bakersfield, CA
Great Southern
Cedar Springs, GA
ITT Rayonier
Fernandina, FL
Kerr-McGee
Trona, CA
Mead Paperboard
Stevenson, AL
Mobil Oil
San Ardo, CA
Nekoosa Papers
Ashdown, AK
Northern Ohio Sugar
Freemont, OH
St. Regis Paper
Cantonment, FL
Texaco
San Ardo, CA
Texasgulf
Granger, WY
(1) C=coal
0=oil
B=bark
Sorbent
NaOH
Caustic waste
NaOH
Caustic waste
Na2C03
Na2C03
NaOH
NaOH
NaOH
NaOH
Caustic waste
NaaCOa
NaOH
Na2C03
Na2C03
Na2C03/NaOH
NaOH
NaOH
NaOH
NaOH
Na2C03
N
Fuel Start-up
Type %S Date I
0 <0.1 6/77
C 1-1.5 1973
0 1.1 6/78
C 0.8 6/74
0 1.1 7/78
C 1 5/76
C 0.7-2.0 9/74
C 0.8 4/76
C 3.2 1972
C 1.2 6/75
B,C,0 1.5-2 7/75
0 1.1 6/77-12/78
BCD 1-2 1975
B,0 2-2.5 1975
0 0.5-5 6/78
0 1.5-3 1975
0 2-2.5 1974
C 1-1.5 2/76
C 1 10/75
B,0 <1 1973
0 1.7 11/73
C 0.7 9/76
lo.of
FGD
1
2
2
1
3
2
2
2
2
4
1
6
2
2
2
1
28
2
2
1
32
2
S02
150
500
500
500
700
800
1.430/10'BTU
-
2000
1«/106 BTU
500
600
1000
1200
-
1500
1300
600
-
-
1000
860
Percent
96
70
QO
yu
70
90
95
86
—
90
90
80
90-96
85-90
80-85
98
95
90
90
-
80-90
73
90
oxidation/dilution
pond
pond /waste treatment
pond /waste treatment
pond
clarify/adjust pH/
to sewer
combine with ash/
landfill
oxidize /neutralize/
discharge
combine with ash/
landfill
to city sewers
pond
to paper process
pond
to paper process
pond
waste treatment
pond
clarification/
aeration
pond /wells/softening
and resuse
pond
-------
C. Applicability—Sorbent costs and disposal of the waste liquor are
the major limitations to this process. Because of the sorbent cost, Na2C03
is $60/ton f.o.b., and about $90/ton delivered a distance of 1,000 miles,206
applications of this process may become centered near large raw material
sources which are in the Western part of the U.S. However, a significant
number of industrial installations produce a sodium based waste stream (e.g.
paper mills) that can be used as the sorbent such that the process may con-
tinue to be applied throughout the U.S.
Waste liquor disposal is the other major limiting factor with
regard to the application of this technology. The majority of sodium scrub-
bing systems in use today are located in the California oil fields, where the
aqueous wastes are disposed of in evaporation ponds. If wastes from this
system cannot be treated in existing waste water treating facilities or used
as a process make-up stream, costs associated with achieving a zero discharge
water system will more than likely limit the system's application due to
economic reasons.
Sodium scrubbing, because it is simple both chemically and mechanically,
can be applied to boilers of varying size and type. At present, sodium sys-
tems are being employed on industrial boilers ranging in size from about 4
to 100 MWe with satisfactory results. The process has been applied to both
stoker-fed and pulverized coal boilers. The FGD system at the G.M.C. Delco
Moraine plant in Dayton, Ohio controls S02 emissions from two stoker-fed,
coal-fired boilers. The S02 removal efficiency is reported to be 85 percent
Of|C
and the average operability of the system is 95 percent.
D. Factors affecting performance—
1) Design and operating considerations^ - One of the major factors to
consider in system design is whether or hot to remove fly ash prior to SOz
removal. If the fly ash is removed then the absorber can be either a packed
tower or a tray tower. If fly ash is not removed a venturi scrubber may be
used for both particulate and SOz removal but this penalizes the process by
requiring a higher absorber pressure drop. The fly ash contributes to solids
buildup at the wet-dry interface and causes erosion of pipes, pumps,
2-154
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spray nozzles and scrubber internals. The dramatic effect of pressure drop
can be seen in Figure 2.2.5-2 representing Shawnee data for the prototype
venturi scrubber.
Because the sodium alkali is very reactive the design L/G ratios can be
low. In 1972 a series of tests were carried out at the Shawnee Test Facility
using sodium carbonate. The effects of L/G are shown in Figure 2.2.5-3 for a
marble bed scrubber, and for a TCA operating with no internals (i.e. , as a
spray tower) in Figure 2.2.5-4. Efficiencies of 99 percent were realized in
210
a TCA with a normal three-bed configuration.
2) Fuel Variations - Variations in the fuel characteristics will affect
the design of sodium scrubbing systems in the same manner they affect the
design of the other wet FGD systems. The most important fuel characteristics
with regard to FGD systems design are: sulfur content, chloride content, and
alkalinity, and fuel heating value. Effects of these variables on process
design are discussed in Section 2.2.1.1.
3) Ambient Variations^ - FGD systems are essentially independent of
ambient variations. However, as with all wet systems, extreme cold can
adversely affect the operation of an FGD process. Cold weather effects can,
of course, be accounted for in the system design by providing heat traced
lines or by enclosing the FGD system within a building.
4) ^Maintenance Problems - Problems encountered with sodium systems
should be generally low. The system is simple to operate and does not
require a complex sorbent regeneration process. Using a clear solution
rather than a slurry minimizes potential plugging or scaling problems.
Some operating problems have, however, been reported in the EPA Indus-
trial Boiler FGD Survey by system users. Most of the problems experienced
to date concern corrosion of the scrubber and pH monitoring equipment, and
erosion in systems that remove both SOa and flyash. A potential cure for
these problems is the use of more exotic grades of materials as is commonly
done in the utility industry.
2-155
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100
95 --
90 --
O
Z
LU
o 85
L_
li.
<
O
80 --
c-g
O
CO
ai
O 75
c:
70
65
60
O
SCRUBBER INLET LIQUOR pH = 9.5
SCRUBBER LIQUOR SODIUM CONC. = 1.0 wt %
SO2 INLET CONC. = 600-3,300 ppm
(Air/SO2 & Flue Gst)
LIQUID TO-GAS RATIO 8-50^l/Mcf
THROAT GAS VELOCITY = 41-105 ft/sec
PLUG OPENING = 40-80 percent
468
PflESSURE DROP, in H2O
10
12
Figure 2.2.5-2.
Effect of pressure drop on S02 removal efficiency
venturi with sodium carbonate (10 MW size).207
2-156
-------
100
o
2 removal efficiency.
20
2-157
-------
100
O
LU
o
u_
u.
LU
cc
Cvi
O
in
H
2
LU
O
cc
UJ
c.
95
90 •-
85 -~
80 --
75 --
70 --
65
20
D
SCRUBBER INLET LIQUOR pH = 9.5
SCRUBBER LIQUOR SODIUM CONC. = 0.5 wt %
SO2 INLET CONC. = 900 ppm (Air/SO2)
D GAS VELOCITY = 6.2 ft/iec
O GAS VELOCITY = 9.2 ft/sec
A GAS VELOCITY = 12 ft/sec
40 60 80 100
LIQUID-TO-GAS RATIO, gal/Mcf
120
Figure 2.2.5-4. Effect of liquid-to-gas ratio on S02 removal efficiency
TCA (no spheres) with sodium carbonate (10 MW size).
209
2-158
-------
E' Retrofits—Sodium scrubbers should be well suited for retrofits.
The system can be easily modularized and a relatively small amount of equip-
ment is needed for the process. Effects on the boiler should be minimal.
Retrofits may be limited by the availability of raw scrubbing materials and
the cost of waste disposal.
2.2.5.2 System Performance—
A. Emission reduction^—Although there are many sodium scrubbing systems
in operation today, there is a lack of performance data for these systems
with regard to their S02 removal capabilities. System operators, have,
however, reported some data on process operations in the EPA Industrial
Boiler FGD Survey. These S02 removal data, which are summarized in Table
2.2.5-4, indicate the S02 removal capabilities for the currently operating
sodium scrubbing systems to be about 90 percent.
B. Availability/Reliability;—Overall reliability of sodium scrubbing
systems applied to industrial boilers has generally been quite high. Although
no quantitative reliability/availability/operability indices are available
from the EPA Industrial Boiler Survey prior to the 4th quarter of 1978,
the majority of user responses indicated that the systems have been opera-
ting well with no problems being experienced, and hence little boiler down-
time has been attributed to scrubber reliability problems.
An examination of the 4th quarter 1978 data shows that of the 22 indus-
trial boiler installations that have operating sodium scrubbing systems, 15
reported quantitative reliability or operability indices that ranged from
89 to 100 percent with an average of 97.8. Of the 15 responses, 9 reported
a 100 percent reliability/operability and all but two reported reliabilities
of greater than 95 percent.2
Of the seven installations that did not report quantitative reliability
indices, two reported that the FGD system had no problems, two reported
erosion/corrosion problems, one was down for reconstruction, one was having
2-159
-------
mechanical problems with pump packings, and one system had no reported
comments.
C. Impacts on boiler—Major effects of a sodium system on a boiler
performance would be 1) boiler derating due to energy required to
run pumps, fans, and if necessary, flue gas reheat, and 2) boiler load reduc-
tion due to scrubber downtime assuming no bypass of flue gas. The energy
penalty for sodium scrubbing systems has been estimated to be about 2 percent
of the net heat input to the boiler (see Section 5). Boiler load reduction
due to scrubber downtime should not be significant due to the high reliabil-
ity of sodium systems.
i
D. Additional maintenance requirement^—Since sodium systems are a clear
liquor process excessive maintenance problems are not expected in the absorber
section. Manpower needs have been reported to vary from 0.25-1.0 man/shift
for the smaller units (10-40 MW) to 3 men per unit for the larger systems
(125 MW).212'213
2.2.6 Processes Under Development
At present, there are some 100 FGD systems in various stages of develop-
ment. This includes systems in the very early stages of development to
systems whose development efforts have ceased. A description and evaluation
of all these systems is outside the scope of this project. However, some of
these developing FGD systems may prove to be advantageous for near-term
application to industrial boilers. Consequently, this section presents brief
evaluations of the following FGD technologies that have not yet been operated
in a commercial system but that have the potential for future commercial
applications:
1) Dry Scrubbing
2) Citrate/Phosphate Buffered Absorption
3) Bergbau-Forsehung/Foster Wheeler
2-160
-------
4) Atomics International Aqueous Carbonate
5) Shell/UOP
6) Chiyoda 121
2.2.6.1 Dry Scrubbing—
Several types of dry scrubbing processes are currently under development
by EPA and FGD process vendors. Dry scrubbing systems that appear to be
particularly applicable to industrial boilers include spray drying of a lime
or sodium sorbent, and firing of a pellitized limestone/coal mixture in
a stoker-fired boiler. This evaluation will focus on spray drying control
systems since they are currently being installed on two commercial sized in-
dustrial boiler applications. The pellitized limestone/coal scrubbing
process will only be briefly discussed due to its early development status.
In this process, the limestone/coal pellets are fired as ordinary fuel
in stoker boilers. The S02 formed during combustion reacts with the lime-
stone present in the fuel pellets to form calcium sulfite and calcium sulfate
salts. The majority of calcium salts remain in the ash bed and are discharged
from the boiler along with the bottom ash. This system does produce an
increase in the boiler's particulate emissions which may affect the design
of the fly ash control equipment.
Preliminary results from laboratory testing at EPA indicated that S02
removal efficiencies of 70-80 percent may be achievable with this type of
control technology. These encouraging results were obtained on an Illinois
No. 6 coal (about 3.5 percent sulfur) using fuel pellets with a calcium/sulfur
ratio of about 7- Additional work has been conducted to optimize the struc-
tural properties of the pellets to obtain a fuel with the same handling
characteristics as raw coal. This work has resulted in pellets with a
calcium/sulfur ratio of about 3. Tests are currently being conducted to
evaluate the combustion and 502 removal ability of the improved pellets.
2-161
-------
Preliminary economic estimates have been prepared which indicate this
technology will cost about $15/ton of coal. This cost compares very well
with the annual costs estimated for other industrial boiler FGD systems
which range from about $23 to $27 per ton of coal burned assuming a
200 x 10s Btu/hr boiler. This control technology, when developed, has the
potential for providing a low cost S02 removal option that may be rather
easily applied to both new and existing industrial boilers.
A. System description — In a spray drying process, a slurry of soda
ash or lime is used to remove S02 from boiler flue gas. The spray dryer
product is a dry mixture of sodium or calcium salts and unreacted sorbent
which can be collected with flue gas fly ash for disposal.
Flue gas enters the spray dryer at air preheater exit temperatures,
generally between 250 and 300 °F. In the spray dryer the gas is contacted
by high speed centrifugal atomizers and driven outward in cross flow to
the flue gas. Flue gas velocity in the dryer vessel is on the order of
5 ft/sec. The dryer vessel is sized for approximately five to ten seconds
residence time.
In the scrubber, gaseous S02 is sorbed into the liquor where it reacts
with sorbent material to form sulfite salts, as indicated in the following
Reactions :
S02 (g) + Na2C03(aq) •* Na2S03(aq) + C02 (g) (2.2.6-1)
S02 (g) + CaO(s) + %H20 -> CaS03«%H20 (2.2.6-2)
In addition to these primary reactions, sulfate salts will be produced by
the following reactions :
Na2S03(aq+s) +%02(g) -> Na2SOit (aq+s) (2.2.6-3)
S03(g) + Na2CO,(s) + Na2SO.f(s) + C02(g) (2.2.6-4)
S02(o) + CaO(s) + %02(g) + 2H20 -> CaSO^ • 2H20(s) (2.2.6-5)
2-162
-------
A typical product mixture formed by these reactions when using a sodium
sorbent is approximately 60 weight percent sodium sulfite, 20 percent sulfate,
and 20 percent excess carbonate.2lk Expected composition of the waste mater-
ial from a lime based spray drying system is about 55 percent calcium sulfate,
30 percent sulfite, and 15 percent CaC03 and lime inerts. This estimate is
for a system that recycles a.portion of the collected sorbent fly ash mixture
in order to improve the overall lime utilization.215
Liquid to gas (L/G) ratios for spray drying are typically 0.2 to 0.3
gal/1,000 scf. This low liquid rate is insufficient to saturate the gas.
Gas exit temperatures are typically in the 150 to 200°F range to provide a
safe margin against water condensation. Spent reactant is entrained in the
flue gas as dry particulate material.216
S02-clean flue gas exits the spray dryer and is routed to a normal
particulate collection device such as an ESP or baghouse, where spent reac-
tant and fly ash are removed for disposal. Systems using a baghouse for
particulate removal report additional S02 sorption occurring in the baghouse
as will be discussed later. Care must be taken to maintain flue gas temper-
ature well above saturation at this point to avoid condensation on the rolids
collection device surfaces.
Accessory equipment consists of liquor preparation and dry waste dis-
posal facilities. In general, liquor preparation facilities include dry
storage, a liquor mix tank, and associated liquor tankage and pumps.
Facilities for handling the collected spray dryer product and fly ash and
transporting them to the ultimate disposal site are similar to those normally
associated with baghouse or ESP collection devices. A generalized process
flow diagram for a typical spray drying scheme is shown in Figure 2.2.6-1.
B. Development status—Spray drying technology for removing S02 from
boiler flue gas is still under development, although spray dryers have been
used in various industrial applications for many years. Development and
2-163
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Economizer
Boiler
NJ
I
Air Preheater
Combustion
Air
Sodium Carbonate
Solution
Spray
Dryer
Particulate
Collection
I.D. Fan
Figure 2.2.6-1. Simplified Flow Diagram for Spray Drying Process,
-------
pilot plant installations have been conducted in the United States by Atomics
International/Wheelabrator Frye, Joy Manufacturing Company/Niro Atomizer,
Babcock and Wilcox, and Carborundum. Additional process vendors which have
announced dry SD2 control system development efforts include Buell/Envirotech,
Ecolaire, Micropul, and Research Cottrell. Other process vendors may also
be pursuing dry scrubbing systems due to the widespread interests on the part
of potential utility and industrial users. Major U.S. development efforts
have included pilot-scale testing at Southern California Edison's Mohave
Station, Basin Electric's Leland Olds Station, and Ottertail Power's Hoot
Lake Station. In addition, it has been reported that Koyo Iron Works is
9 1 fl
undergoing development of a sodium based spray dryer system in Japan.
To date, although there are no spray drying S02 control systems currently
operating on either an industrial or utility application, five contracts have
been awarded for construction of commereial sized systems ranging in size
from 22,000 SCFM (about 11 MWe) up to about 575 MW. Two of these systems
are industrial boiler applications. Table 2.2.6-1 summarizes the status of
these five systems. These activities indicate a significant amount of inter-
est in this rapidly developing control technology.
TABLE 2.2.6-1 SUMMARY OF COMMERCIAL DRY SCRUBBING APPLICATIONS219'220'221
' System/Location
Strathmore Paper Co.
Woronoco. MA
Celanese
Cumberland, MD
Basin Electric Co.
Antelope Valley #1
Basic Electric Co.
Laramie ?/3
Otter Tail Power
Coyote Unit #1
Vendor Sorbent Size Start-up Date
Mikropul
Wheelabrator-Frye/
Rockwell Int .
Joy/NIRO
Babcock & Wilcox
Wheelabrator-Frye/
Rockwell Int.
Lime 22,000 SCFM
(11 MWe)
Lime 50,000 SCFM
(25 MWe)
Lime 440 MWe
Lime 575 MWe
Na2C03 410 MWe
5/79
12/79
4/82
4/82
Spring '81
2-165
-------
C. Design considerations—Based upon pilot unit test results, 90 per-
cent S02 removal can be achieved using either lime or sodium based sorbents.
Stoichiometric ratios of 2.3-3.0 were required for lime operations whereas
stoichiometric ratios of only 1.0-1.2 were required to achieve the same S02
removal for sodium operations. It has also been reported that 90 percent
S02 removal may be achieved with a stoichiometric lime requirement of 1.3-1.7
O O p
by recycling some of the unreacted sorbent. These reported values are for
relatively low sulfur coal operations. It still remains to be demonstrated
whether or not a lime based spray dryer system will be able to achieve high
S02 removal efficiencies when applied to a high sulfur coal. A sodium based
system should be able to achieve high S02 removals on high sulfur coals due
to its higher reactivity.
Process control of the spray dryer feed solution is basically simple and
straightforward. However, the amount of scrubbing solution added to the
spray dryer must be carefully controlled as too much will result in conden-
sation in downstream equipment and not enough will prevent attaining the
required S02 removal efficiency.
The primary reaction of S02 with sorbent material in the spray dryer
appears to occur in the aqueous phase. S02 removal and sorbent utilization
are, therefore, enhanced by increasing the exposure time of liquid droplets
to flue gas S02. This liquid phase residence time in the spray dryer can be
lengthened by increasing the ratio of liquid to flue gas (L/G) entering the
scrubber. An increase in L/G increases aqueous phase residence time not
only because the mass quantity of water to be evaporated is larger, but
because adiabatic evaporation of the increased moisture content lowers the
flue gas outlet temperature. As the flue gas temperature decreases toward
its adiabatic saturation temperature, the driving force for evaporation also
decreases. Higher L/G ratios also result in better gas-liquid contact
because higher liquid feed ratios result in more spray droplets.
2-166
-------
For a given flue gas inlet moisture content and temperature, an upper
limit for L/G ratios is set by the temperature requirement for the outlet
flue gas. Flue gas from the spray dryer must be approximately 30-50°F above
its adiabatic saturation temperature in order to avoid condensation on par-
ticulate removal device surfaces and to avoid plume opacity due to condensa-
tion.
For a given inlet moisture content, a higher flue gas inlet temperature
will result in the capability of evaporating more water. For example, a flue
gas entering a spray dryer at 250°F with a humidity of 0.06 Ib water per Ib
dry gas would have an adiabatic saturation temperature of approximately 124°F.
Adiabatic evaporation to within 50°F of saturation would result in an outlet
humidity of approximately 0.08 Ib H20/lb dry gas. This humidity increase
across the spray dryer corresponds to an L/G ratio of 0.17 gal/1000 scf.
Increasing the inlet temperature to 290°F at the same 0.06 Ib H20/lb dry gas
inlet humidity raises the adiabatic saturation temperature to approximately
127°F, and raises the humidity at a 50°F approach to saturation to approxi-
mately .09 Ib H20/lb dry gas. This raises the maximum L/G ratio possible
to 0.26 gal/1000 scf.
For this reason, industrial boilers appear to be particularly well suited
for spray drying applications since they typically do not have as sophisti-
cated of a flue gas heat recovery system as utility boilers and tend to
have hotter flue gas exit temperatures. For the cases considered, the in-
dustrial boiler flue gas temperatures ranged 350 to 400°F instead of 250-300°F
which is typically found on utility boilers.
Since spray dryer FGD systems have significant amounts of unreacted sor-
bents in the exit gas, the potential exists for further reaction with flue
gas SOz in downstream equipment. Since the temperature and gas/solid rela-
tive velocities are low in the outlet from a spray dryer, no significant
reaction of S02 with unreacted sorbent can be expected in the subsequent duct
work. However, collection of the solid material in a baghouse offers a high
gas/solid relative velocity as well as a high solid residence which are both
conducive to SOz-sorbent reactions.
2-167
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Data reported from various test installations appears to vary somewhat
with regard to the amount of S02 absorption occurring in a baghouse particu-
late collection device. Results of one pilot unit indicated that up to 20
percent of the system's total SOz removal occurred in the baghouse when sodium
was the sorbent and up to 10 percent occurred when lime was the sorbent.
Results from another pilot system which was investigating the performance of
lime as the sorbent indicated that virtually no S02 absorption was observed
223
in the baghouse.
Design considerations directly related to the performance of a spray
dryer-based FGD system are the operation and maintenance of the system. The
successful operation of a spray dryer FGD system is greatly affected by the
process control instrumentation. Unlike wet scrubbing systems in which SOa
removal can be controlled by some gross recycle liquor physical parameter
such as pH, this FGD system must be controlled by a feedback loop from con-
tinuous monitoring of outlet S02 concentration and temperature. Increases in
spray dryer outlet temperatures from set point would result in increased
liquor flow rates to the scrubber and vice versa. Increases or decreases in
system outlet SC>2 concentrations would result in a corresponding increase or
decrease in sorbent feed to the scrubbing liquor mix tank. In such a control
scheme, an undetected S02 monitor drift or error could result in failure to
meet control requirements or waste of sorbent, depending upon the direction
of error. Error in outlet temperature measurement could result in excessive
L/G ratios and consequent condensation on downstream particulate control
equipment.
During one of the pilot plant installations, a spray dryer was deliber-
ately operated under upset conditions to examine the effect of condensation
on baghouse performance. After the system had returned to normal operating
conditions for one hour it was inspected and no changes were found on the
baghouse and bags when compared to previous inspections. This was because
the exit gases at normal temperatures dried the wet deposits to the point
that they were removed in the normal bag cleaning cycle. This upset condi-
2-168
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tion was for a short duration (15 minutes) so it is still unclear what the
9 p h.
effects of a long term upset will be.
A spray dryer-based FGD system offers maintenance and operation advan-
tages over other systems in 1) the high turndown ratios which can be achieved
and 2) the anticipated low maintenance requirements and high reliability of
the major process equipment. Turndown of the spray dryer system should be
relatively simple and high turndown ratios should be possible. As gas rate
decreases, liquid feed rate to the scrubber is decreased proportionally.
Turndown ratios of up to 4 to 1 have been demonstrated in a single scrubber
with a single atomizer with no decrease in S02 removal efficiency. The bene-
fit of increased gas residence time appears to maintain S02 removal efficien-
cies at lower flow rates. In larger installations where multiple spray
dryer vessels would be employed, dampers could be employed to remove whole
vessels from service to increase the possible turndown ratio. The major
limit to system turndown is the onset of condensation in downstream equipment
due to greater heat losses at the significantly lower flue gas flow rate.225
The spray dryer scrubber could offer some advantage over wet scrubbers
in maintenance and reliability of process equipment. Because of the much
lower L/G ratio and the relatively small head requirements, pump maintenance
requirements for this system should be significantly lower. Scrubber vessel
problems such as scaling or erosion should be avoided, since wet particles
should never come into contact with vessel walls. While some deposition of
dry material on scrubber walls has been observed in demonstration tests,
occasional mechanical rapping has been found to adequately control buildup.226
Since the energy for atomization comes from a spinning disc rather than
from a nozzle constriction, atomizer erosion problems should be minimal.
However, due to their high speed of operation (up to 18,000 rpm) it can be
expected that the atomizers will require frequent periodic maintenance. Due
to the simplicity of removal of a single atomizer and the fact that multiple
2-169
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atomizer arrangements can be employed in spray dryers in a large installation,
failure or removal from service, of one atomizer should have a minimal effect
on overall system performance. In October 1975 tests on a 7-foot test scrub-
ber by Bowen Engineering, the equipment supplier for the Atomics International
process, indicated that removal efficiency in a three atomizer spray dryer
could be maintained when removing flow from one atomizer by adding this flow
227
to the other two atomizers.
D. Applicability to industrial boilers—There are no apparent technical
constraints in applying spray drying technology to small industrial boilers.
In fact, due to the relative simplicity of the system and its claimed high
reliability, this technology may prove to be a very desirable option for
small industrial applications.
Possible concerns with applying this technology to industrial boilers
are in the following areas: 1) the relatively large diameters of the spray
drying vessels, 2) the possible detrimental effects on downstream particulate
collection equipment from system upsets, and 3) lack of data concerning the
system's S02 removal ability for high sulfur eastern and midwestern applica-
tions. Another concern, which is not unique to this system, is the problem
of solid waste disposal.
Condensation effects on downstream particulate control equipment and
system performance on various coals under actual operating conditions are
questions that will be resolved after additional operating experience with
the system is gained. Industrial boiler spray drying systems at Strathmore
Paper Co. and Celanese which are expected to be operational in late 1979 will
be the first commercial applications of the process,
E. Summary—The major advantages claimed for using a spray drying
system include: 1) less equipment requirements than a dry system, 2) produc-
tion of a dry rather than wet waste material, 3) reported high reliability
of system components, 4) lower system costs, and 5) reduced energy requirements
2-170
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The major unresolved issue appears to be the effect of auxiliary equipment
such as control instrumentation on overall system reliability. Although the
major process equipment has proven to be quite reliable in other applications,
it remains to be seen how precisely the spray drying system can be controlled
to avoid condensation in subsequent particulate collection equipment under low
temperature, high moisture boiler flue gas conditions.
2.2.6.2. Citrate Buffered Absorption —
The citrate flue gas desulfurization process is a regenerable process
which removes SOa from flue gases by absorption in an aqueous sodium citrate
solution. The absorbed S02 is converted to elemental sulfur in subsequent pro-
cessing steps. The U.S. Bureau of Mines pioneered the development of the
Citrate Process with developmental work also being done by Arthur G. McKee
and Company, Peabody Engineering, and Pfizer, Inc.
A. System description228 — Flue gas is quenched and prescrubbed with
water to remove HC1, SOs, and residual fly ash. S02 is removed by a buffered
aqueous solution in an impingement tray or packed bed absorber. The flue gas
may be reheated to 175°F before discharge to the atmosphere. The primary
reactions occurring in the absorber are as follows:
S02 + H20 -> HSOs + H+ (2.2.6-6)
HCit" + H+ -> H2Cit~ (2.2.6-7)
The citrate acts as a buffer maintaining a pH of 3-5 throughout the
system. The regeneration reaction requires a pH in this range.
Absorber liquor is reacted with H2S in several reactors placed in series
to form elemental sulfur by the following overall reaction:
+ 2H2S -> 3S + 2H20 (2.2.6-8)
The sulfur is concentrated and separated from the regenerated liquor.
H2S for the regeneration reactors can be produced by reacting reducing gas
with two-thirds of the product sulfur from the process.
2-171
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Some sulfite oxidation occurs and results in the formation of sodium
sulfate which must be removed from the system by crystallization as Glaubers
salt (Na2SO^*10H20). The amount of sulfite oxidation in the systems has been
reported to vary from one to four percent. The low oxidation rates are
attributable to the presence of thiosulfate ions which are formed in the
regeneration step.
Figure 2.2.6-2 presents a simplified flow diagram of this process which
highlights the interaction of the various processing areas. This process is
made up of many components and is consequently relatively complex.
Process equipment used in addition to the prescrubber and absorber are:
regeneration reactors, sulfur flotation and melting equipment, sulfate purge
crystallization equipment, and H2S generation equipment. Except for the re-
generation reactors and H2S generation equipment, the process operations have
been well demonstrated in other applications. Integration of all equipment
and demonstration of system reliability and ability to follow boiler opera-
tions will be demonstrated on an industrial boiler (64 MW) located at St.
Joe Minerals.
229
B. Development status —The Citrate Process was pioneered by the
Bureau of Mines with developmental work also done by Arthur G. McKee, Pea-
body Engineering, and Pfizer, Inc. The Bureau of Mines operated a 1,000
scfm pilot unit on a lead sintering furnace tail gas for about two years.
Pfizer, Peabody, McKee treated 2,000 scfm from a coal-fired industrial
boiler for about seven months in 1974. A Bureau of Mines/EPA co-funded
demonstration unit on 120,000 scfm of flue gas from a coal-fired industrial
boiler is scheduled to start up in mid 1979. This unit is at St. Joe
Minerals zinc smelter in Pennsylvania.
The primary technical hurdles that remain for the Citrate Process are
H2S generation from a hydrogen source other than methane, integration of an
H2S generation unit with the process, and a more complete understanding of
the chemistry of the regeneration reaction. The process has been totally
integrated except for producing H2S from a coal based reducing gas. HZS
was produced from methane in the pilot plant installations.
2-172
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TO STACK
SULFATE TO
DISPOSAL
REGENEBATED LIQUOR
FLUE GAS
SULFUR
MELTING
PRODUCT
SULFUR
Figure 2.2.6-2. Simplified Plow Diagram for Citrate/Phosphate Process.
-------
Currently, no data can be obtained for this process except for the
earlier pilot plant test runs. However, the EPA demonstration of this
system at St. Joe Minerals should provide data to answer questions concerning
process chemistry and operability.
C. Design considerations — Major design factors affecting the
absorption portion of this process are flue gas temperature, liquor cir-
culation rate, and pH. Process developers attempt to optimize these factors
to achieve a desired SOa removal (up to 99 percent has been reported from
pilot units) while producing as concentrated a liquor as possible. An efflu-
ent liquor highly concentrated with 862 is desirable since a high liquor
loading produces less solution for regeneration which results in a smaller
regeneration unit.
Regeneration reactor design is based on both reaction rate and mass
transfer considerations. One developer reports that reaction kinetics are
limiting; the mass transfer of H2S into the solution occurs more rapidly
than reactions involving intermediates formed in the system. Another vendor
reports that the regeneration reaction is mass transfer-limited in the first
reactor where the ratio of HaS concentration to aqueous SOz concentration is
lowest, and kinetics-limited in the other two reactors where the ratio of
H2S concentration to aqueous S02 concentration is greater.230
The concentration of hydrogen sulfide in the gas fed to the regeneration
reactors is another design consideration since it directly affects reactor
size. Vendor specifications vary from a 39 percent to a 96 percent nominal
H2S gas stream.
Small bubbles of HjS are desirable in the regeneration reactors to pro-
vide more liquid-gas interface. Pilot plants have used an open pipe or
sparger generating large bubbles with a high tip-speed agitator shearing the
large bubbles into tiny bubbles. The regeneration section will precipitate
sulfur and, therefore, must be designed to minimize plugging problems.
2-174
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An HaS stream produced on site from reducing gas or hydrogen and sulfur
will contain some elemental sulfur vapors. Potential condensation and solidi-
fication of these vapors upon contact with the liquor must be considered in
determining the method of introducing H2S gas to the regeneration reactors.231
Potential plugging at the entrance to the regeneration reactors could
be reduced by several methods. Two parallel heat exchangers operated in a
cyclic mode could be used. One would cool the HaS gas and condense the sul-
fur vapors while steam was injected in the other to vaporize condensed
sulfur. Alternately, a water quench could remove the sulfur and saturate
the HaS gas streams producing a sour water containing hydrogen sulfide.
Flotation with either kerosene or air has been examined for use in
separating the precipitated sulfur from the regenerated liquor. The
flotation unit should produce an essentially sulfur-free citrate/phosphate
solution for recycle, and a sulfur stream containing as little liquor as
practicable. Kerosene flotation yields a powdery damp sulfur of about 50
percent solids, while air flotation yields a pumpable slurry of about 7-15
percent solids. The sulfur slurry produced with air flotation contains more
liquor than that produced by kerosene flotation. As a result, systems using
air flotation will have more solution exposed to the sulfur melting and de-
canting steps. A potential for buffer and thiosulfate degradation exists at
the operating temperature of about 280°F in the melting and decanting steps. 32
Kerosene flotation may affect product sulfur quality and may result in sig-
nificant kerosene losses. An open vessel was used at the pilot plant with
kerosene losses of about 90 Ib/net ton of sulfur produced. A vapor recovery
system on an enclosed vessel might reduce these losses to a more acceptable
level.233
The sulfur melter is simply a steam heat exchanger that raises the tem-
perature of the sulfur slurry out of the flotation cell to about 280°F. The
equipment transporting the sulfur/liquor mixture to the melter and the melter
itself must be designed for the appropriate slurry service and to minimize
any potential for plugging. The decanter is a pressurized gravity flow
2-175
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separation unit. Molten sulfur is continuously or intermittently taken off
the bottom while clear liquor is recycled to the absorber. All of the sulfur
transfer lines must, of course, be heat traced to prevent solidification.
The sulfate purge section should present no unique design problems. A
crystallizer, refrigeration unit, and Glauber's Salt (Na2S(V 10H20) separator
are used to remove the sulfates from the citrate/phosphate solution. The by-
product is then dried to produce anhydrous Na2SOi(. The unit may be operated
intermittently to control the sulfate concentration in the liquor. It should
be noted that commercial sulfate removal units have not been operated at any
of the pilot plants using the Citrate/Phosphate Process.
Hydrogen sulfide (H2S) generation design considerations in the buffered
absorption processes are different from most FGD processes that require H2S
because H2S is produced externally by reacting elemental sulfur with hydrogen
or reducing gas (H2/CO). A gas stream with a high H2S concentration is de-
sirable for use in the regeneration section. A higher H2S partial pressure
results in a higher mass transfer rate and reduced reactor size. The H2S
concentration will, however, be determined mainly by the composition of the
reducing gas or hydrogen fed to the H2S generation unit rather than the pro-
cess selected for H2S generation.234
Plugging of the system by condensed sulfur has been one of the major
operational problems with the H2S generation facility. This unit must be
designed so that condensation of sulfur vapors occurs only in knockout ves-
sels included in the design.
The ability to turn down the unit or store H2S is important for a work-
able process design. The H2S generation unit must be able to follow the load
swings of the boiler since hydrogen sulfide requirements are directly propor-
tional to the quantity of S02 absorbed. Turn down capability has not been
fully developed or demonstrated at any of the pilot plants.235 The small
amount of residual H2S leaving the regeneration reactors will be collected
and combusted in the boiler or an incinerator.236
2-176
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D. Applicability to industrial boilers—Overall system complexity
would appear to limit the applicability of this system to small industrial
boilers. Since all of the liquor effluent from the absorber is routed to
the regeneration section, this process is not well suited for a decoupled
centralized regeneration.
The major concern with applying this system to an industrial boiler is
with the sulfur reduction/HzS generation portion of the process. Unless a
stream of HZ or CE^ were available, the reductant would have to be produced
from coal gasification which is still relatively undeveloped and would add
to the overall process complexity. This process does, however, become
rather attractive for applications to facilities with existing waste H2S
streams such as oil refineries. For these applications the HzS stream could
be introduced directly into the regeneration reactors, which would eliminate
process problems associated with integrating an HzS generator into the FGD
operations. 237
E. Summary—The unique features of the process are the reaction of HaS
gas with absorbed SOa in the regeneration step, the separation of precipi-
tated sulfur by flotation, and the external generation of HaS from elemental
sulfur and reducing gas. Advantages of this process are production of sul-
fur as a by-product, and simple absorber operation which should result in
good S02 removal. The major disadvantage of this process for application to
industrial boilers is its overall complexity especially in regard to the t^S
generation section. The demonstration of this process at St. Joe Minerals
will provide data to answer questions on process control, operability, and
reliability.
2.2.6.3 Bergbau-Forschung/Foster Wheeler Process—
Flue gas SOz is adsorbed in a moving bed of char to form HzSCK in the
Bergbau-Forschung/Foster Wheeler (BF/FW) system. Saturated char is then
heated to yield gaseous S02 and regenerated char. The SC>2 stream is reduced
2-177
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in the presence of anthracite coal to elemental sulfur. The adsorption
and regeneration steps were developed by Bergbau-Forschung in Germany while
the reduction step is proprietary technology of Foster Wheeler.
A. System description — The Bergbau-Forschung/Foster Wheeler process
uses the concept of dry carbon adsorption of S02 and thermal regeneration of
the carbon. A flow diagram for the process is presented in Figure 2.2.6-3.
Flue gas from the boiler passes through dust collection equipment to remove
the bulk of the particulate matter. The gas enters the adsorber at tempera-
tures between 250-300°F. The flue gas passes through one or two beds of
activated char where S02, SOs, oxygen, and water vapor are adsorbed on the
char. HC£ is reported to pass through the adsorber without reacting the char.
The following reactions take place in the presence of activated char:
S°2(g)+%°2(g) 3(g) (2.2.6-9)
S03(g) + H20( } H2SOn (2.2.6-10)
The sulfuric acid formed remains in the pores of the char. The spent char is
conveyed to a regeneration vessel where the char is mixed with hot sand at a
mixture temperature of about 1200°F. The following reactions take place in
the regeneration step.
H2S(K( j (2.2.6-11)
H2S04(g) H20(g)+S03(g) (2.2.6-12)
2S03 + Cf , C02f . + 2S02, v (2.2.6-13)
(s) •* (g) (g)
The carbon in Reaction 2.2.6-13 is supplied by consuming part of the char
adsorbent. The char is recycled to the adsorber and the sand is recycled
to a sand heater where it is reheated to about 1500°F.238>239
2-173
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FLUE GAS
—i
v-D
FLUE GAS
TO STACK
f FUEL
SAND
Figure 2.2.6-3. Simplified Flow Diagram for Bergbau-Forschung Process.
-------
The off gas from the regenerator contains about 20 volume percent
and can be used to make sulfuric acid or sulfur. If the SOa was to be used
to make sulfuric acid the gas would first be cooled and cleaned before being
sent to the acid plant. If sulfur was to be pro-duced the gas could be sent
to the proprietary Foster Wheeler RESOX process. The gas would then be taken
directly from the regenerator and reacted with crushed anthracite coal at
1100-1500°F to form elemental sulfur. Foster Wheeler generally includes the
KESOX process in its Bergbau-Forschung/Foster Wheeler system designs.240
This process is characterized by its relative mechanical complexity as
compared to wet FGD processes. It is a dry process that adsorbs S02 on solid
carbon particles which must be transported to the regeneration processing
area and then back to the adsorption area. Current designs use a series of
conveyors and bucket elevators for this transfer. Since the regenerated char
will be at approximately 1200°F this mechanical equipment must be designed
for high temperature service.
Major equipment items are: 1) an adsorber which handles a moving bed
of char, 2) a thermal regenerator which uses hot inert sand to raise the
char temperature to about 1200°F for regeneration, 3) a char-sand separator,
4) a fluid bed sand heater to heat the sand to 1500°F, and 5) a RESOX
sulfur production system which uses coal to directly reduce the regeneration
off gases to sulfur. All of this equipment is solids handling equipment
which in general, presents more operating and maintenance problems than
gas-liquid handling equipment.
B. Development status—The BF/FW Process has been thoroughly and
successfully tested at the bench, pilot, and prototype level in West Germany.
In 1963, research and development began using a 20 scfm bench-scale system
that included adsorption and regeneration steps. A process for the manufac-
ture of char was developed concurrently in a facility which could produce
approximately 5000 tons per year. As a result of the bench-scale work a
pilot plant operating on a 1,750 scfm slip stream of flue gas from a 35 MW
2-180
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pulverized coal-fired boiler was operated from 1968-1970. A prototype 45 MW
unit was installed by Bergbau Forschung in 1974 at the Kellerman Power Plant
in Lunen, West Germany. The completely integrated process operates on an
88,000 scfm slip stream from a 350 MW coal-fired peaking service boiler.
Off gas from the regenerator is treated in a modified Glaus unit (BAMAG) to
produce elemental sulfur. Testing at the prototype unit, which is sponsored
by the West German government, is scheduled to end in late 1978.
The first fully integrated application of the BF/FW Process using RESOX
was demonstrated by FW on a 47.50 MW coal-fired boiler at Gulf Power's Scholz
Stream Plant. The unit consisted of a 20 MW adsorption section and a 27.5 MW
reduction section (RESOX). The facility operated intermittently from August
1975 to May 1976. The unit experienced only limited success due primarily
to mechanical problems or failures caused by Foster Wheeler in efforts to
reduce costs. The Bergbau Forschung system at Lunen, West Germany, is the
most successful system to date and is more representative of the process than
the system at Scholz.2"*1
C. Design considerations—The adsorber in the BF/FW Process is a unique
gas-solid contacting device. The adsorber must be designed for gas flow as
well as char flow. It is very important that both materials flow with even
distributions and velocities throughout the adsorber. In the Bergbau design,
a set of louvers perform several functions that influence the distribution
and velocity of the gas and the char as shown in Figure 2.2.6-4. As a result,
the most important aspect of the adsorber is the louver design.
The primary design criteria for the louvers are the angle of inclination
at the gas inlet and outlet, the vertical spacing between the louvers, and
the width of the louver measured in the direction of gas flow. These criteria
have a direct bearing on:
distribution of the char weight on the structural support columns,
the uniform downward flow of char to insure plug flow in the beds,
2-181
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REGENERATED CHAR IN
ro
I
oo
r-o
FLUE GAS IN
~\
1
1
t
£
/
1
I
FLUE GAS OUT
SATURATED
J_ J_ J_ _[ CHAR OUT
Figure 2.2.6-4. FW/BF Dry Adsorption Process Adsorber Module Detail.
-------
the mixing action of fly ash with the char to prevent buildup
at the inlet face of the char bed, and
the minimization of entrainment of fines in the exit gas.
Design considerations for the char beds are the char depth, the number
of modules placed in series to form a bed, and the number of beds in parallel.
These criteria influence the degree of S02 removal, gas side pressure drop,
and gas residence time. The char beds have been modularized for commercial
application. The depth of the char bed is then set by the module choice.
The number of modules in a bed and the number of parallel beds is determined
by the SOa concentration in the flue gas and the removal efficiency desired.
The modules in a bed are not separated by partitions and, therefore, are more
conceptual than real. Consequently, the char can undergo horizontal movement
(cross flow) with a possible increase in char attrition.2^
The operating temperature of the absorber is primarily a function of the
flue gas temperature and the temperature of the recycled char pellets from
the regeneration section. The design operating temperature of the adsorber
is approximately 280°F. Dilution air is used to cool the flue gas to this
operating temperature. The char out of the regenerator must be adequately
and uniformly cooled by a water quench prior to recycle to the adsorber.
Solids handling and control is another important consideration in the
BF/FW adsorption section. Regenerated char must be uniformly distributed
over an entire adsorber bed in a full-scale design. Solids handling equip-
ment must be designed for the operating temperatures of the adsorption section.
Control of the adsorber is accomplished by adjusting the char flow rate
and the amount of dilution air added to the flue gas. The char flow rate is
not critical as long as enough char is fed to the adsorber to achieve the
required degree of removal. Addition of dilution air should be carefully
controlled to keep the flue gas temperature above the acid dew point. Also,
a water quench might be an alternative cooling method for a commercial
installat ion.
2-183
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The regeneration reactor design is based on a Bergbau-Forschung unit
that has been in operation over five years. The primary design factors are
assurance of complete mixing, adequate residence time, and a noncombustible
atmosphere. Conductive heat transfer using sand at about 1500°F as an inert
heat transfer medium occurs to raise the char temperature to about 1200°F.
The flow of the char and sand through the regenerator is gravity-induced.
Char out of the adsorption section is transported by bucket elevator to a
surge tank above the reactor to supply the necessary head to flow through
the regeneration section. The fluid bed sand heater is likewise elevated
to provide a head for gravity flow of the sand. The fluidized level is
maintained by sand leaving through a refractory-lined overflow pipe which
empties into the regenerator.
Complete and uniform mixing of the char and sand is very important. The
mixing of two hot solids of widely differing particle size is difficult.
Operating problems were reported at early pilot-scale installations, but BF
has since improved the regenerator design.
No excess oxygen should be permitted to enter the regeneration vessel as
this would increase the char consumption above that necessary for the chemis-
try of the regeneration process. For this reason, a reducing atmosphere with
a slightly positive pressure is maintained in the vessel. Minimum resistance
to upward flow of the evolved gases is also necessary. A means for the gas
to move upward without having to find its way through the char-sand mixture
may be required. Otherwise, evolution of gas may continue until after the
char leaves the regenerator, resulting in "burps" in the char-sand separ-
2 t it, 2 4 5
ator. '
The char cooler must quench the char pellets uniformly to prevent
localized concentrations of hot char being recycled to the adsorber. The
exit temperature of the char from the cooler should also be low enough so as
not to initiate hot spots in the adsorber. The water quench of the hot char
also must not produce thermal strains that result in breakage of the char
pellets. BF feels that the char is strong enough to resist thermal stresses
satisfactorily as long as air leakage-into the unit is prevented.
2-184
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Design considerations in the RESOX process for reduction of S02 are
primarily concerned with the reaction of SC>2 gas with anthracite coal to
produce elemental sulfur. Some factors that might influence this reaction
are feed gas composition, pressure, temperature, residence time, gas distri-
bution, and stoichiometry of carbon to S02.
Foster Wheeler reports that the pressure and gas distribution are not
critical design considerations. The operating pressure is atmospheric.
Proper gas distribution is accomplished by injection of the S02-rich gas at
various ports around the lower vessel circumference. The stoichiometry of
carbon to S02 is likewise not a sensitive parameter as long as an excess of
carbon is present. Presently FW uses a carbon to S02 stoichiometry of about
2:1. About half of the RESOX coal feed leaves the bottom of the reactor as
part of the RESOX ash stream. An economic incentive may exist to recover
24-6 2. *t 7
the heating value of the RESOX ash by burning it in the boiler. '
D. Applicability to industrial boilers 2tf8' 2tf9—The overall mechanical
complexity of this system would appear to limit its applicability to large
industrial boilers. In addition, the process employs rather large adsorbers
and solids handling equipment, and unless extensive conveying systems are to
be used, it will be necessary to locate both the adsorption and regeneration
sections near the stack. The sulfur or sulfuric acid production unit can be
located more remotely from the stack.
Another process characteristic that will limit the application of this
process is that its long-term reliability is undetermined and may be a prob-
lem. In general, the process appears to be more applicable to large boilers
than small boilers because of the large amount of regeneration equipment
required. It does not appear cost effective for a small boiler application
to use a complicated process requiring large amounts of equipment.
A potential process advantage is the ability to remove NOX from the
flue gas. Reported NO removal ability of the Bergbau system has varied
2-185
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from zero to near seventy percent. Recent data taken at the Scholz pilot
unit has indicated that average NOX removals of about twenty percent were
being achieved.250 At this time, however, the ability of the BF/FW process
for NO removal has not been conclusively demonstrated.
x
Another attractive feature of this process is its ability to produce
sulfur by using coal as the reductant. However, the process currently
requires anthracite coal which has a limited availability. Current develop-
ment efforts are underway by Foster Wheeler to evaluate methods of using
bituminous coals as the reductant.
E. Summary—The Bergbau-Forschung/Foster Wheeler Process is a dry
adsorption process' which uses activated char pellets to adsorb SOz from flue
gas. The char is thermally regenerated, and the SOa in the off gas is
reduced in the presence of anthracite coal to elemental sulfur by a unit.
The unique features of the process are the louvered wall design of the
adsorber and the RESOX unit. The louvers influence the distribution and
velocity of the flue gas and the char. RESOX is a proprietary process for
the reduction of gaseous SOa to elemental sulfur. The major problems lie
in the mechanical reliability of equipment and the availability of design
and operating data for the RESOX unit.
2.2.6.4 Atomics International Aqueous Carbonate Process—
The Aqueous Carbonate Process (ACP) as developed by Atomics Interna-
tional uses an aqueous sodium carbonate (Na2C03)solution to sorb sulfur
dioxide (SOg) from power plant flue gas. The dry scrubber product is treated
to regenerate the scrubbing solution and to produce elemental sulfur. The
technology for the regeneration and sulfur production steps is based prin-
cipally upon established practice in the pulp and paper and the chemical
industries.
A. System description—The ACP is divided into five major processing
areas. These are gas cleaning, reduction, quenching and filtration, carbona-
tion, and sulfur production. Figure 2.2.6-5 presents a simplified flow
2-186
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FLUE GAS
I
M
Oo
ELEMENTAL
SULFUR
ASH &
COKE
CLEAN GAS
TO STACK
PARTICULATE
REMOVAL
SULFUR
PRODUCTION
Na2C03 SOLUTION
— •>-
-n
\j
S02
SCRUBBING
ENTRAINED SOLIDS
REDUCER OFF-GAS
__
H2S RICH
GAS
H2O
f
p APRH
H2O
f
NATION
Na2S RICH QUENCH &
SOLUTION FILTRATION
PRODUCT
COLLECTION
co
-------
diagram for this process illustrating the interaction of the various pro-
cessing areas.
The S02 scrubbing and product collection equipment are combined in the
gas cleaning subsystem. This subsystem employs a spray dryer for S02 scrub-
bing, and a bank of cyclones in series with an electrostatic precipitator
for product collection and final particulate removal from the gas stream.
Before the gas enters the scrubber, the ash content must be reduced to 1.0
gr/scf or less to limit the size of the ash removal equipment in the regenera-
tion section. The following reactions occur in the scrubber.
S02 + Na2C03 ->• Na2S03 + C02 (2.2.6-14)
Na2S03 + ig02 -> Na2SOi, (2.2.6-15)
S03 + Na2C03 ->- Na2S0lt + C02 (2.2.6-16)
Reaction 2.2.6-14 is the primary reaction for S02 removal.
Sulfur dioxide-clean flue gas exits the spray dryer and is routed to a
bank of product collection cyclones, where the majority of the dry particles
are removed. Final particulate removal is accomplished in high efficiency
electrostatic precipitators or baghouses in which particulate emissions are
limited to 0.01 grain/scf or less. Spent reactants from both the cyclones
and precipitator or baghouse are collected and sent to the reduction sec-
tion. 251
The dry product collected in the gas cleaning system is stored in a silo
and then conveyed pneumatically to the reducer vessel. This vessel contains
a pool of molten salts at temperatures between 1700 and 1900°F. Carbon is
injected in the form of petroleum coke (or coal). Combustion air is bubbled
through the melt from injection nozzles in the vessel walls.252
In the molten salt pool, the following reactions take place:
2-138
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Na2S03W + 3/2C(s) -»• Na2S(£) + 3/2C02 (g) (2.2.6-17)
NaaSCMA) + 2C(s) + Na2S(£) + 2C02(g) (2.2.6-18)
C(s) + 02(g) -* C02(g) (2.2.6-19)
Reactions 2.2.6-17 and 2.2.6-18 are endothermic. Reaction 2.2.6-19 is exo-
thermic, however, and provides heat for both the endothermic reactions and
system heat losses. The mechanism for reaction 2.2.6-19 is complex, in-
volving sequential oxidation-reduction of the sulfur-containing salts as
well as direct oxidation of carbon.253 ' 2Stf
The C02-rich off gas from the reducer is sent to the carbonation sec-
tion after passing through a recuperator, waste heat boiler, and gas cooling
tower. Reducer melt is continuously withdrawn and directed to the quench/
dissolver vessel.
The sodium sulfide melt from the reducer is dispersed into fine drop-
lets by steam shatter jets and dissolved in a green liquor solution which
is near its boiling point. Insoluble material, mostly fly ash and reacted
coke, is filtered out at this point using a rotary drum vacuum filter.
Sodium is recovered from the ash filter cake using a simple washing technique.
Both the quench and filtration operations are considered proven technology in
the pulp and paper industry.
After the quench-filtration step, green liquor is contacted with C02-
rich reducer off gas in a series of carbonation towers. The technology
for this process step has been developed in the pulp and paper industry and
proven processes are available. AI has developed their own carbonation
scheme by modifying existing technology. The process reacts the C02-rich
reducer off gas and green liquor from the ash filter, producing concentrated
Na2COa solution for recycle to the gas cleaning subsystem plus an H2S-rich
Glaus plant feed gas. A vent gas stream is also produced and sent to the
Glaus plant incinerator, and some additional ash is removed and sent to
disposal.
2-189
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The HaS produced in the carbonation section is directed to a Glaus plant
for conversion to elemental sulfur. Glaus technology is commercially avail-
able but has not been tested for specific application with this desulfuriza-
tion process. The Glaus plant tail gas and precarbonator off gas are com-
busted and returned to the spray dryer for treatment. AI proposes the use of
an incinerator to burn these gases; however, they may be returned to the
255 .256
boiler for combustion.
B. Development status—Development of the AGP has been underway since
1972. All of the process steps have been tested on a 1000 scfm scale or are
considered proven technology by Atomics International. The spray dryer has.
been tested on pilot units at Bowen Engineering's North Branch, New Jersey
facility and at Southern California Edison's Mohave Station. Early bench-
scale work on the molten salt regeneration unit was performed at AI test
laboratories in 1973. Based on this work, pilot scale tests were conducted
with a feed rate equivalent to 1.25 MW. Bench-scale tests have also been
conducted to determine the feasibility of using coal as the carbon source
in the reducer. A fully integrated AGP system has never been operated,
u 257
however.
This process has been selected by EPA for demonstration at the 100 MW
scale at Niagrara Mohawk Power Corporation's Charles R. Huntley Station.
This will be the first full-scale demonstration of a regenerable process that
can produce sulfur using coal as the reductant. A preliminary design and
cost estimate have been completed and are under review. The current schedule
for this process demonstration calls for plant startup to occur in late 1979
or early I960.258
950 9 K 0 ? R 1
C. Design considerations ' ' —Since this process uses a spray
dryer for S02 absorption, the design considerations for the gas cleaning
section of this process will be similar to those for the sodium based spray
dryer process. Consequently, the reader is referred to Section 2.2.6.1 for
a discussion of gas cleaning design considerations.
2-190
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Sodium sulfite, sulfate, and unreacted carbonate are collected from the
gas cleaning process area and sent to the spent absorbent storage hopper
which has at least two days storage capacity. This is done in an attempt to
decouple the operations of the gas cleaning and regeneration processing
areas.
Operation of the molten salt reactor is the key design feature for this
process. From a preliminary engineering viewpoint the design basis appears
sound, as AI has identified potential problem areas associated with this
reactor and have taken appropriate control measures. Historically, however,
the operation of molten salt beds under reducing atmospheres has been parti-
cularly difficult to control, and it is likely that any unanticipated process
problems would occur in this area.
Spent reactant, a source of carbon (petroleum coke or coal), and air
are fed to the reducer to accomplish the reduction of sulfite and sulfate to
sodium sulfide. Pilot plant data indicate that a steady-state reduction of
95 percent can be achieved with about 4 wt percent excess coke maintained in
the reducer melt. Approximately one-half of the feed carbon is consumed
chemically by the reduction reactions. The other one-half is burned to
supply the heat needed for endothermic reduction to Na2S, to melt the in-
coming salts, and to offset thermal losses and preserve the 1700 to 1900°F
operating temperature. Typical reducer off gas contains less than 1 per-
cent 02 and 35 percent carbon oxides with a CC>2:CO ratio of 10 to 1 or higher.
Heat is recovered from the off gas in a recuperator and waste heat boiler
before the gas is cooled for use in the carbonation step.
The reducer can be continuously run at rated capacity or can be turned
down to a standby condition corresponding to 5 percent of the design feed
rate. To compensate for vessel heat losses and maintain operating tempera-
ture during standby operation, approximately 400 Ib/hr of carbon is required
in excess of that corresponding to the normal absorbent-to-carbon ratio. It
will be important to avoid wide swings in melt temperature since appreciable
2-191
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spalling of refractory lining was noted in pilot tests under these conditions
For typical boiler load cycles or downtime, the preferred operating mode for
the reducer is either 100 percent capacity or standby.
Because of its volatility, NaC£ will be the predominant salt in the
reducer off gas. It is likely that some of these salts will condense on the
heat exchange surfaces of the recuperator and waste heat boiler and result in
serious corrosion problems. AI has noted such deposits in pilot scale tests
but claims, however, that the material can be easily removed with periodic
washing or soot blowers. Most of the NaC£-rich particulates will be scrubbed
from the gas in the gas cooling tower. A portion of the recirculating cool-
ing tower liquid stream will be used as a chloride purge for the ACP-
The technology for quench and filtration of the reducer melt is con-
sidered standard practice in the pulp and paper industry. The design of
equipment in this processing area borrows heavily from this technology.
The quench-dissolver is a thick-walled carbon steel vessel similar to
those used in the paper industry. Some corrosion of the carbon steel is
expected but the walls will be 5/8 inch thick to allow for it. Carbon steel
is the preferred construction material for this application because it mini-
mizes the possibility of embrittlement and cracking at the quench operating
temperatures (around 200°F).
Two factors that are key to avoiding explosive conditions in the quench-
dissolver are:
1) the melt must be broken up and dispersed into fine droplets, and
2) the green liquor should be maintained at or near its boiling
temperature to avoid pressure excursions.
To assure operation in this mode, AI has specified that two completely re-
dundant steam shatter systems be provided. One will use process steam gen-
erated in the waste heat boiler and the other will use low pressure steam
2-192
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generated in the boiler. Also, all pumps in the circulating liquor loop
will be spared to insure an adequate water supply to the quench tank.
Filtration of fly ash and other solids from process liquor has caused
problems in other alkali scrubbing systems. This situation is complicated
in the ACP by the effect of high temperature reduction on the filtering
characteristics of ash and unreacted coke in green liquor. Tests at Al's
Molten Salt Test Reactor facility have shown that the fly ash and fuel ash
present in the melt do not dissolve but do interact with other melt consti-
tuents to significantly alter their physical and chemical properties upon
quenching. Laboratory tests have shown that it is possible to filter the
ash and coke from simulated green liquor and concentrate the solids to a
40-60 wt percent solids filter cake. A prototype test loop has been pre-
pared to demonstrate the rotary drum filter and other pieces of green liquor
handling equipment.
Like the quench and filtration steps, the technology for carbonation is
well developed by virtue of its use in the pulp and paper industry. Several
commercial carbonation processes are available and can be applied to the ACP
regeneration scheme. Most of these processes feature high energy and steam
consumption, however, which has led AI to develop their own carbonation
process. This process uses a series of sieve-tray columns to react the C02
from the reducer off gas with the sodium sulfide in the green liquor from
the quench step. The diameters of the columns are set by the superficial
velocity of the carbonation gas, in this case between 1 and 2 ft/sec. The
heights of the columns are dictated by the mass transfer characteristics of
the sieve trays. The key design criteria are HaS and COz uptake.
Production of elemental sulfur from the HaS in the crystallizer off gas
is standard Glaus technology. Because of the low HaS concentration (30-35
percent) of this stream, however, a three-stage Glaus plant is required for
97 percent sulfur recovery. Other carbonation schemes produce a more con-
centrated feed stream and thus require a smaller Glaus plant, but AI claims
the overall economics favor their process concept.
2-193
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D. Applicability to industrial boilers262'263—Overall chemical and
mechanical complexity will limit the application of this system to small
industrial boilers. As with the spray dryer throwaway system, space require-
ments may also limit its application; however, the regeneration section can
be easily decoupled from the gas cleaning section and located in a peripheral
plant access or completely off site.
One of the major process concerns of this system is operation of the
high temperature reducing section. Since this section of the process has
only been operated at the 1.25 MW scale, and then not in fully integrated
operations, questions remain as to how the overall system will operate when
coupled to a coal-fired boiler.
One of the main process advantages is that it provides a mechanism to
produce sulfur without requiring a reducing gas. This advantage may not
overcome the process disadvantage of increased complexity for applications
to small industrial boilers.
E. Summary—The major advantage of the Aqueous Carbonate Process is its
ability to produce sulfur with a solid carbonaceous reducing agent, eliminat-
ing the need for a reducing gas that must be produced from natural gas or by
coal gasification. The decoupled nature of the process makes the gascleaning
subsystem the sole operational interface with the flue gas source, and allows
regeneration to be carried out at a distance from the scrubber. The need
for reheat has been avoided in many applications by operating the spray
dryer scrubber at L/G ratios too low to saturate the flue gas. A key to
the success of the ACP is the successful operation of the molten salt reduc-
tion step which produces a very corrosive melt. Other potential problem
areas of the process include the quench and solids filtration steps. Although
the individual process steps have been demonstrated at the pilot plant level,
integration of process components has not yet been demonstrated at any level
of development.
2-194
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2.2.6.5 Shell Flue Gas Desulfurization Process-
The Shell Flue Gas Desulfurization (SFGD) Process which is licensed by
UOP's Air Correction Division is a dry metal oxide process. It is based on
the ability of copper oxide to react with S02 in a flue gas and to be regen-
erated by hydrogen reducing gas. Although the Shell process is sometimes
termed an adsorption process, the removal mechanism is the reaction of S02
with copper oxide to form copper sulfate according to Equation 2.2.6-20.
CuO + %02 + S02 + CuS04 (2.2.6-20)
The process developers refer to this reaction as acceptance and to copper
oxide as an acceptor. Copper is regenerated and S02 evolved by passing
reducing gas, preferably hydrogen, over the copper sulfate. The end product
i
of the process can be sulfur, S02, or sulfuric acid depending on the avail-
able markets and other economic considerations. Figure 2.2.6-6 is a
simplified flow diagram of the system.
The unique feature of the process is a set of specially designed,
parallel passage, fixed-bed reactors which use a copper oxide-on-alumina
acceptor for S02 removal. The reactor configuration was specifically
designed by Shell to maintain good catalyst stability and minimize pressure
drop. The reactor cross section is proportional to the gas flow while the
depth of the reactor determines the removal capacity. The SFGD Process can
achieve 90 percent removal of S02 irrespective of inlet S02 concentration.
The process also has demonstrated NO removal from the addition of NHs which
X
reduces NO over CuO or CuSOtj according to the following reaction.
2NH3 + 2ND + ^02 •+ 2N2 + 3H20 (2.2.6-21)
A. Process description261>> 26 5' 26 6—The system operates at about 750°F in
the acceptance and regeneration steps. Flue gas is routed from a hot electro-
static precipitator through the Shell/UOP system before going to the air
preheater.
2-195
-------
r-o
I
Flue Gas
Naphtha,
Methane.,
Coal
Air Preheater
Acceptance
Regeneration
Reactors
Waste Heat
Recovery
S02 Recovery
and
Concentration
•H20
Hydrogen
Production
S02
SOa
Conversion
'Tail Gas
Sulfur
Figure 2.2.6-6. Block Flow Diagram - Shell Flue Gas Desulfurization Process.
-------
The acceptance reaction depicted in Equation 2.2.6-20 and the regeneration
reactions in Equations 2.2.6-22 and 2.2.6-23 are exothermic. Since acceptance
and regeneration occur in the same temperature range, there is no energy
lost in heating or cooling the beds.
When the majority of the copper in an acceptor has been converted to
copper sulfate and S02 breakthrough begins to occur, the acceptor vessel is
blocked off from the flue gas stream and purged in preparation for sorbent
regeneration. The spent sorbent is regenerated by passing a diluted hydrogen
containing reducing gas through the vessel. S02 is reevolved according to
Equation 2.2.6-22. Swing reactor acceptors allow continuous flue gas treat-
ment.
2H2 •+ Cu + S02 + 2H20 (2.2.6-22)
Additional hydrogen is consumed by the following side reactions. First,
any CuO which was unused during S02 adsorption is reduced back to copper by
Reaction 2.2.6-23.
CuO H- H2 -> Cu + H20 (2.2.6-23)
After regeneration the acceptor is again purged and returned to flue gas
service where any copper sulfide is reoxidized to copper sulfate by oxygen
in the flue gas .
Cu2S + 02 •+ CuSOit + CuO (2.2.6-24)
Current SFGD designs use a combination of gas holder and a gas compressor
for recovering the concentrating S02 from the regeneration gases. The gas
holder is used to dampen flows from the cyclic absorption section to the gas
compressor. The gas is compressed so that water can be condensed and sepa-
rated from the recovered S02 before it is fed to the Glaus unit.
2-197
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The SOa conversion section presents several options. The concentrated
S02 stream can be processed by Allied Chemical, RESOX, or BAMAG process
technology, or a variation on one of these, to produce sulfur. However,
sulfuric acid and liquid SOa production are options which appear, in general,
to be more economical in terms of use of energy and reducing media, and are
more proven technically.
Hydrogen is needed to regenerate Cu from the spent sorbent according
to Reaction 2.2.6-22. UOP has generally proposed steam-methane or steam-
naphtha reforming as a means of supplying needed reducing gas. These sys-
tems provide fairly high concentrations of Ha compared to air-blown coal
gasification which produces a gas highly diluted by nitrogen. A possible
problem with using these systems is the decreasing availability of both
methane and naphtha.
B. Development status—Shell has gone through several phases of testing
but has yet to build a completely integrated unit applied to a coal-fired
boiler. Extensive bench-scale testing to determine a suitable metal oxide
acceptor, to select a reactor configuration, and to collect mass transfer
data preceded the installation of a 400 to 600 scfm pilot unit at Shell's
Pernis Refinery. The Pernis unit accumulated 20,000 hours of operating time
during which various types of copper on alumina acceptors, regeneration
agents, and reactor internal construction materials were examined. The
flue gas used for the tests came from a high sulfur, heavy oil-fired heater.
The stability of the acceptor when subjected to a flue gas environment was
verified and the proposed reactor configuration was successfully demonstrated.
Information from Pernis was used to scale up and build a commercial-
scale unit which went onstream in mid 1973 at the SYS refinery in Japan.
The unit there consists of two acceptors operating on flue gas from an
oil-fired boiler equivalent to about 40 MM of electric generating capacity
2-198
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(75,000 scfm). Only the acceptors and water gathering SOz recovery sys-
tem were installed and tested because refinery hydrogen and excess Glaus
plant capacity were available. The longest period of continuous operation
has been about 2 months. The fairly successful SYS installation confirmed
Shell's computer design methods for the reactor and their ability to scale
up the system. Demonstration of both the viability of automatic sequencing
controls and the large flue gas valves was an additional result.
The next step in Shell's development program was the testing of a
reactor using coal-fired flue gas without particulate removal. This experi-
ment took place at a Dutch utility station. The resistance of the reactor
configuration to fouling by dust was examined and information was generated
on erosion/corrosion tendencies of various materials of construction.
Actual operation of the acceptance and regeneration steps in a coal-
fired flue gas environment has been tested on a 0.6 MW unit at Tampa Elec-
tric' s Big Bend Station. Results to date indicate no increase in pressure
drop over time and stable catalyst activity. The system was operated with
r) r a
high fly ash loadings without noticeable effect on the acceptors.
The other parts of the system, S02 concentration and sulfur production, are
considered to be standard technology.
The availability of design information appears to be good but more
confidence could be placed in the system's overall reliability if a fully
integrated system were in operation. A fully integrated system could
answer the troubling question of process operability under varying loads.
C. Design considerations269'278—Two major design considerations with
the Shell/UOP process are: 1) the difficulty of integrating the cyclic be-
havior of the fixed-bed adsorption scheme with the variable SC>2 load from the
power plant, and 2) the relatively inflexible operation of the hydrogen pro-
duction facility and the Glaus plant, if used. The fixed-bed acceptors them-
selves are not affected by the variations in S02 loading but the problem
arises rather in maintaining efficient control of a widely varying situation.
2-199
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Because unused copper oxide acceptor consumes H2 duriag regeneration and
because H2 consumption is an important contributor to operating costs, accep-
tor operations should be optimized to minimize hydrogen consumption.
The capability of the H2 plant to respond to frequently varying loads
will be an important point, along with its economic viability and raw material
supply, in deciding what type of hydrogen production facility to use. Also,
sulfur producing plants such as Claus plants are rather inflexible at present.
Two alternatives might be to produce sulfur from SOa using less expensive or
more readily available reductants, or to instead produce sulfuric acid. Acid
plants are more amenable to variable feeds, but good temperature control is
still essential.
D. Applicability to industrial boilers—As with other regenerable sys-
tems, the application of this system to industrial boilers will be limited due
to its overall complexity. In addition, this system requires hydrogen for the
regeneration step. Hydrogen is either produced from raw materials whose
future supply is uncertain (methane and naphtha) or from a coal gasification
process which would add to the overall process complexity. In addition,
space requirements for this process are more of a concern that with other
processes due to the uncertain ability to decouple the acceptance and S02
971
recovery sections of the process.
Retrofit applications will be limited due to the requirement of flue
gas temperatures of about 750°F for the acceptance section of the process.
Although there are no technical impediments to retrofitting this system it
would require extensive duct work or reheating the flue gas to about 750°F.
E. Summary—The Shell Flue Gas Desulfurization Process has as its
principal advantages the ability to remove NO simultaneously with SOa, and
X
the ability to produce a choice of three potential by-products: Sulfur, sul-
furic acid, and liquid S02. The process is relatively complex, consumes hydro-
gen, and requires high temperature absorber operations. Process applications
appear to be best suited to installations where NC> removal must be achieved.
X
2-200
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2.2.6.6 Chiyoda Thoroughbred 121 Process—
A. System description—The Chiyoda Thoroughbred 121 Process (CT-121)
uses a j et bubbling reactor to absorb, oxidize, and neutralize SC>2 from flue
gas. The reactor is a single vessel, consisting of flue gas inlet and outlet,
air inlet, limestone slurry inlet, and gypsum slurry outlet. Air and mechani-
cal agitation are also provided. A simplified flow diagram for this process
is shown in Figure 2.2.6-7.
Hot flue gases are first routed through a prescrubber for gas cooling
and particulate removal purposes. Chiyoda has reported that this prescrub-
bing step may be incorporated into the reactor.272
Flue gas is sparged into the reactor through an array of vertical
spargers to generate a bubbling or froth layer. S02 is absorbed in the
froth layer to form calcium sulfite which is oxidized to calcium sulfate.
The froth layer provides good mass transfer for this S02 absorption reaction
to occur. Cleaned flue gas is demisted and exhausted to the atmosphere.
Although the chemistry of this process is similar to that of conven-
tional limestone scrubbing processes, it differs in that the sorbed S02 is
oxidized to sulfate (gypsum), leaving only trace amounts of sulfite. Addi-
tionally, all chemical and process steps are carried out in one vessel.
The overall reaction for the system is described by the following
equation:
S02 + CaC03 + %02 + 2H20 -* CaSOit'2H20 + C02 (2.2.6-25)
Crystallized gypsum is discharged from the reactor as a slurry of 5 to
20 weight percent. Solids are separated and the liquor is returned to the
reactor. The gypsum by-product is of a high quality and may be dewatered
to produce a 90 percent dry product and landfilled. As an alternative,
2-201
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O
NJ
WATER.
FLUE GAS,
GAS
BLOWER
PRESCRUBBER
ABSORPTION/
OXIDATION
REACTOR
AIR
-PURGE
MIST
ELIMINATOR
TO STACK
LIMESTONE
GYPSUM
Figure 2.2.6-7. Simplified Flow Diagram for Chiyoda Thoroughbred 121 Process.
-------
there is a possibility that the gypsum may be sold to the wallboard or
273
Portland cement industry.
B. Development Status—Development of the CT-121 Process was initiated
in 1975 in an effort to reduce the cost and complexity of the commercial
CT-101 Process. Tests were initiated at the bench and laboratory scale and
finally at the 650 scfm level. A conceptual design of a 125,000 scfm plant
has been completed and plans are now under way to convert the 23 MW CT-101
system at Southern Company's Scholz Power Station to a CTrl21 process con-
figuration for further testing.
C. Design Considerations—The scaling potential of this system is of
concern in that calcium sulfate is formed in the reaction vessel. However,
this system maintains a gypsum crystal concentration of 10 to 20 weight per-
cent in the reactor which provides area for crystal growth. Consequently,
gypsum should precipitate on the crystal surfaces and not form scale on the
11 275
reactor walls.
SC"2 removals of about 90 percent were achieved during the pilot plant
testing. Higher removal efficiencies can probably be obtained, but would
be at the cost of higher pressure drops and capital costs since it would be
necessary to increase contact time by using larger scrubbers with more
severe breakdown of bubbles by increased sparger pressure drop.
Control of the CT-121 process should be accomplished by monitoring pH
levels in the absorber. The process should be rather easily controlled and
rather insensitive to process upsets. However, ability of the process to
follow load could be limited by sparger performance. As gas velocity through
the sparger decreases, pressure drop decreases and bubble size increases as
a result. This decreases mass transfer and S02 removal. It may be necessary
to use multiple absorbers or a manifold system to distribute gas into any of
a number of spargers in the absorber.276
2-203
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D. Application to industrial boilers'—Application of the CT-121 process
to industrial boilers may be limited by space requirements for the scrubber
and land area for product disposal. If the gypsum can be sold, by-product
disposal problems would be eliminated. Marketability of gypsum will be
extremely dependent upon local requirements for wallboard and, in general,
it is expected that the product will have to be landfilled.
The process could potentially be decoupled, with the use of storage
ponds for the slurry during centrifuge down time. This would disrupt normal
process operations and increase land requirements. Ability of the process
to follow load may be limited as discussed previously. Due to its relative
mechanical and chemical simplicity, the process should be fairly acceptable
2. 7 7
for industrial boiler applications.
E. Summary—The Chiyoda Thoroughbred 121 Process is a developing pro-
cess that seems attractive for several reasons. Operation of the system
should be relatively simple; capital costs and operating costs are poten-
tially low. The process uses a plentiful raw material and produces a product
that can be sold or landfilled. It is, however, undeveloped and there are
many uncertainties as to its overall viability and costs. The upcoming
demonstration of the plant at the Scholz Power Station will resolve many of
these uncertainties.
2.3 CONTROLS FOR OIL-FIRED BOILERS
The major differences between oil- and coal-fired boilers are attribu-
table to the differences in fuel properties. Since fuel oils are generally
lower in sulfur and ash, their particulate and S02 emissions will be lower
than for coal for a given sized boiler. The sulfur variability in fuel oil
is also less than for coal which results in more stable FGD operations than
required for coal-fired installations. In addition, it is possible to oper-
ate with as little as 5 to 7 percent excess air in oil-fired boilers whereas
excess air requirements for coal-fired boilers may vary from 15 to 60 percent,
2-204
-------
depending on the method of firing. In general, these differences will not
effect FGD process design considerations as discussed in this chapter, but
will result in smaller FGD systems with correspondingly lower costs for the
oil-fired FGD systems.
In the United States, there are currently 73 FGD systems being used to
control SOa emissions from oil-fired boilers with an additional 19 systems
under construction. As shown in Table 2.3-1, all of these systems,
except for one, are applied to oil-fired steam generating units which burn
crude oil as a fuel. The non-oil field steam generating system is applied
to an oil-fired boiler at a paper mill. Ninety-one of these FGD units are
sodium throwaway systems and one is a double-alkali system. The actual
designs of the systems are very similar to coal-fired FGD system designs
except that the oil-fired systems have lower flue gas flow rates for a given
boiler size because less excess air is used for combustion. Consequently
the oil-fired FGD systems are smaller than coal-fired FGD systems would be
for the same boiler size. Consequently, the discussions of FGD process
design considerations in Section 2.2 for coal-fired boiler installations
are applicable to oil-fired FGD systems.
2-205
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TABLE 2.3-1. SUMMARY OF UNITED STATES
OIL-FIRED INDUSTRIAL BOILER INSTALLATIONS
279
S02 Removal
Total Size
Installation (SCFM)
Alyeska Pipeline
Belridge Oil
Chevron
Getty Oil
Mead Paperboard
Mobil Oil
Texaco
*Belridge Oil
*Chansler Oil
*Getty Oil
*Mobil Oil
*Sun
*Texaco
Total tn Operation 1,
Total in Construction
TOTAL - 2 ,
50,000
24,000
248,000
567,000
100,000
125,000
380,000
12,000
70,000
396,000
80,000
12,000
117,000
544,000
687,000
231,000
Number of
FGD Units
1
2
3
6
1
28
32
1
1
4
7
3
3
73
19
92
Inlet
(ppm)
150
500
700
600
1500
1500
1000
500
710
600
500
700
1000
Percent
Removal
96
90
90
90-96
95
90
73
90
96
90
85
85
95
Type FGD
Sodium
Sodium
Sodium
Sodium
Sodium
Sodium
Sodium
Sodium
Double Alkali
Sodium
Sodium
Sodium
Sodium
*Installations in construction phase
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REFERENCES
1. Stern, R.D., et- al. Interagency Flue Gas Desulfurization Evaluation.
Revised Draft Report, 2 Vols. DCN 77-200-187-16-10. Austin, TX.
Radian Corporation. November 1977. Page 46.
2. Tuttle, J. t:t al.. EPA Industrial Boiler FGD Survey: Fourth Quarter
1978. Final Report. EPA Contract No. 68-02-2603, Task 45. EPA-600/
7-79-067a. Cincinnati OH. PEDCo Environmental Inc. November 1978.
3. Melia, M., et al. EPA Utility FGD Survey: June-July 1978. EPA
Contract No. 68-02-2603, Task 24, EPA 600-7-78-051d. Cincinnati, OH.
PEDCo Environmental, Inc. November 1978.
4. Communication with Mr. G.B. Blue, Hartford Steam Boiler Inspector.
October 17, 1978.
5. Communication with George Christman, Babcock & Wilcox Power Generation
Group. October 18, 1978.
6. Leivo, C.C. Flue Gas Desulfurization Systems: Design and Operating
Considerations. Final Report, 2 Vols. EPA Contract No. 68-02-2616,
Task 2, EPA 600/7-78-030a,b. San Francisco, CA. Bechtel Corporation.
March 1978. pp. 3-17.
7. Biedell, E.L., et al. EPA Evaluation of Bahco Industrial Boiler Scrub-
ber System at Rickenbacker AFB. Report No. EPA-600/7-78-115. June
1978.
8. Radian Corporation. By-Product/Waste Disposal for Flue Gas Cleaning
Processes, EPRI RP786-2, Task 1, Report Review and Assessment of the
Existing Data Base. Draft Report. DCN 77-200-169-01. EPRI RP786-2,
Task 1. Austin, TX. March 1977.
9. Duvel, W.A., Jr., et al. State-of-the-Art of FGD Sludge Fixation.
Final Report. EPRI FP-671, EPRI RP786-1. Beaver, PA. Michael Baker,
Jr., Inc. January 1978.
10. Jones, B.F., III., K. Schwitzgebel, and C.M. Thompson. Evaluation of
the Physical Stability and Leachability of Flue Gas Cleaning Wastes.
Draft Report. DCN 77-200-169-09. EPRI Research Project 786-2, Task 2.
Austin, TX. Radian Corporation. December 1977.
2-207
-------
11. Lunt, R.R., et al. An Evaluation of the Disposal of Flue Gas Desulfuri-
zation Waste in Mines and the Ocean. Initial Assessment. EPA 600/7-
77-051, EPA Contract No. 68-03-2334, PB 269 270. Cambridge, MA.
Arthur D. Little, Inc. May 1977.
12. Duvel, et at.., op.clt., p. 2-5.
13. Stern, et al., op. olt. , Vol. II., Appendix D.
14. Gregory, N., et al. EPA Utility FGD Survey; February-March 1978.
EPA Contract No. 68-01-4147, Task 3, EPA 600/7-78-051b. Cincinnati,
OH. PEDCo Environmental, Inc. June 1978. p. 59
15. Ibid.
16. Ibid.
17. Ibid.
18. Tuttle, J., Op. olt.
19. Biedell, et al., op-, alt.
20. Ando, Jumpei. "Status of SOa and NOX Removal Systems in Japan."
Presented at the Fourth Symposium on Flue Gas Desulfurization. Holly-
wood, FL. November 1977. p. 2.
21. Ando, J. S02 Abatement for Stationary Sources in Japan. EPA Contract
No. 68-02-2161. EPA 600/7-78-210. Tokyo, Japan. November 1978.
22. Ando, "Status of S02... p. 4.
23. Head, H. N., et al. "Recent Results from EPA's Lime/Limestone Scrubbing
Programs - Adipic Acid as a Scrubber Additive." Presented at the
Flue Gas Desulfurization Symposium, Las Vegas, Nevada. March 1979.
24. Rochelle, G.T., and C.J. King. "The Effect of Additives on Mass Transfer
in CaCOs or CaO Slurry Scrubbing of SO from Waste Gases." Industrial
Engineering Chemistry, Fundamentals Vol. 16, No. 1. 1977. pp. 67-75.
25. Ibid.
26. Head, H. N., et al., op. Git.
27. Ibid.
28. Stern, et al., op. olt., pp. D-86f.
29. Ibid., p. D-87.
2-208
-------
30. Ibid., p. D-88.
31. Stern, et a'l., op.ait., p. D-84.
32. U.S. Bureau of Mines. Minerals Yearbook, 1975, Vol. 1, Metals, Miner-
als, and Fuels. Washington, DC, GPO. 1977.
33. Stern, et al. op.ait., p. D-138.
34. Hargrove, O.W. and R.S. Merrill. Summary of the Effects of Important
Chemical Variables Upon the Performance of Lime/Limestone Wet Scrub-
bing Systems. Final Report. DCN 77-200-126-08. EPRI Research Pro-
ject 630-3, Task 1. Austin, TX. Radian Corporation. December 1977.
p. 2-16.
35. Devitt, T., et al. Flue Gas Desulfurization System Capabilities for
Coal-Fired Steam Generators. Final Report, 2 Vols. EPA Contract No.
68-02-2603, Task 1, EPA 600/7-78-032a,b. Cincinnati, OH. PEDCo
Environmental, Inc. March 1978. p. 3-43.
36. ibid. , p. 3-45.
37. Ibid. , p. 3-100.
38. Ibid., p. 3-104.
39. Ibid. , p. 3-46.
40. Leivo, op.ait., p. 4-6.
41. Devitt, et al. , op.ait. , p. 3-58.
42. Hargrove, op.oit.
43. Kim, Yong K., Melvin E. Deming, and John D. Hatfield. "Dissolution of
Limestone in Simulated Slurries for the Removal of Sulfur Dioxide from
Stack Gases." ACS, Div. Env. Chem., Prepr. 13(2), 33-38 (1973).
44. Epstein, M., et al. "Results of Mist Elimination and Alkali Utiliza-
tion Testing at the EPA Alkali Scrubbing Test Facility." In Proc. ,
Symposium on Flue Gas Desulfurization, New Orleans, March 1976, Vol. 1
Research Triangle Park, NC. EPA, Office of Research and Development.
1976.
45. Devitt, et al., op.cit., p. 3-52.
46. Ibid. , p. 3-49.
47. Ibid. , p. 3-105.
2-209
-------
4-8. Ibid., P. 3-102.
49- Leivo, op.ait., p. 4-4.
50, Ibid. , p. 4-3.
51- Ibid. , p. 4-6.
52. Potts, J.M., A.V. Slack and J.D. Hatfield. "Removal of Sulfur Dioxide
From Stack Gases by Scrubbing with Limestone Slurry: Small-Scale
Studies at TVA." Presented at the 2nd International Lime/Limestone
Wet Scrubbing Symposium, New Orleans. November 1971.
53 . Kim, op. ait.
54. Hargrove, op.ait.
55 . Borgwardt, Robert H. "Increasing Limestone Utilization in FGD
Scrubbers." Presented at the 68th Annual AIChE Meeting, Los Angeles.
November 1975.
56. Ibid. , p. 3-55.
57. Leivo, op.oit. , p. 4-17.
58. Devitt, et at. „ op.oit., 3-53.
59. Devitt, et al., op.ait., p. 3-41.
60. Ibid. , p. 3-56.
61. Ibid. , p. 3-55.
62. Leivo, op.cit. , p. 4-19.
63. Slack, A.V. "Lime-Limestone Scrubbing: Design Considerations."
CEP 74(2), 71(1978).
64. Ottmers, D.M., Jr., et al. Evaluation of Regenerable Flue Gas Desul-
furization Processes, Revised Report, 2 Vols. EPRI RP535-1, EPRI
FP-272. Austin, TX. Radian Corporation. July 1976. p. 338.
65. Slack, op. ait.
66. Tuttle, J., e-bal.t op. ait.
67. Melia, M. , et al., op. ait.
68. Tuttle, J., et al., op. ait.
2-210
-------
69. Environmental Protection Agency, Office of Air Quality Planning and
Standards. Electric Utility Steam Generating Units - Background
Information for Proposed SO Emission Standards. EPA-450/2-78-007a.
Research Triangle Park, North Carolina. July 1978.
70. Radian, By-Product/Waste Disposal...
71. Devitt, et al.3 op.cit. , p. 4-5,6.
72. Laseke, Bernard A. and Timothy W. Devitt. "Status of Flue Gas Desul-
furization Systems in the United States." Presented at the EPA
Symposium. Hollywood, FL. November 1977.
73. Radian Corporation. Factors Affecting Ability to Retrofit Flue Gas
Desulfurization Systems. EPA Contract No. 68-02-0046, EPA 450/3-74-015.
Austin, TX. December 1973,
74. Biedell, op. cit., p. 22.
75. Biedell, op. oit-, p. 24.
76. Ibid.
77. Ibid., p. 13,64.
78. Ibid., P. 66.
79. Ibid., p. 74.
80. Ibid., p. 70.
81. Ibid., p. 80.
82. Ibid., p. 78.
83. Ibid,, p. 104.
84. Ibid., p. 106.
85. Ibid., P- 103,108.
86. Ibid., p. 17.
87. Ibid , p. 17.
88- Maxwell, M.A., H.W. Elder and T.M. Morasky. Sulfur Oxides Control
Technology in Japan. Research Triangle Park, NC, Muscle Shoals, AL.,
and Palo Alto, CA. EPA, Tennessee Valley Authority and EPRI. June
1978. p. 20.
2-211
-------
89. Slack, A.V. Technology for Power Plant Emission Control Survey of
Developments in Japan. August 1978. Sheffield, AL. SAS Corporation.
October 1978. p. 1.
90. Dennison, L.L. and J.C. Dickerman. Trip Report. Meeting at Ricken-
backer Air Force Base, 31 October 1978. Austin, TX. Radian Corpora-
tion. November 1978.
91. Curtis, Robert E., William R. Menzies, and Teresa G. Sipes. Flue Gas
Desulfurization Process Evaluation for Southland Paper Mills, Inc.
DCN 78-200-225-01. Southland Project No. 20-77-LLP-077. Reviewed by
Douglas H. Brown and James C. Dickerman. Austin, TX. Radian Corpora-
tion. February 1978. p. B-3.
92. Ibid.
93 LaMantia, C.R., et al. Dual Alkali Test and Evaluation Program. 3 Vols,
Final Report. EPA 600/7-77-050a-c. Cambridge, MA. Arthur D. Little,
Inc. May 1977.
94. Tuttle, J., et al., op. cit.
95. Melia, M., et al.3 op. ait.
96. Kaplan, Norman. "Introduction to Double Alkali Flue Gas Desulfurization
Technology." Presented at the EPA Flue Gas Desulfurization Symposium.
New Orleans. March 1976.
97- Ando, "Status of Flue Gas...
98. Stern, et al., op. eit. p. D-140.
99. Kaplan, Norman. "An Overview of Double Alkali Processes for Flue Gas
Desulfurization." Presented at the Flue Gas Desulfurization Symposium.
Atlanta, GA. November 1974. p. 26.
100. Oberholtzer, J.E., et al. Laboratory Study of Limestone Regeneration
in Dual Alkali Systems. Final Report. EPA Contract No. 68-02-1332,
Task 26. EPA 600/7-77-074. Cambridge, MA. Arthur D. Little, Inc.
July 1977.
101. Stern, et al.., op.oit. , p. D-141.
102. Tuttle, J., et al., op.oit.
103. Interess, Edward. Evaluation of the General Motors Double Alkali S02
Control System. Final Report. EPA Contract No. 68-02-1332, Task 3,
EPA 600/7-77-005. Cambridge, MA. Arthur D. Little, Inc. January 1977.
2-212
-------
104. Curtis, Menzies, and Sipes, op.dt., p. D-Bl
105, Ibid. , p. B-7-
106. Ibid. , p. B-10.
107. Ibid. , p. B-ll.
108. Kaplan, "Introduction...
109- Devitt, et al. 3 op.ait. , pp. 3-130, 3-150.
110- Ibid. , p. 386.
111. Interess, op.dt.
112. Ibid. , p. 2.
113. Ibid. , p. 13.
114. Ibid. , pp. 1-27.
115. Cornell, C.F. and D.A. Dahlstrom. "Performance Results an a 2500 ACEM
Double-Alkali Plant." Presented at the 66th Annual AIChE Meeting.
Philadelphia, PA. November 1973.
116. Cornell, C.F. and D.A. Dahlstrom. "Sulfur Dioxide Removal in a Double-
Alkali Plant." CEP 69(12), 47(1973).
117,. Ibid.
118 . Ibid.
119 . Ibid.
120. Rush, Randall E. and Reed A. Edwards. Operational Experience with
Three 20 MW Prototype Flue Gas Desulfurization Processes at Gulf Power
Company's Scholz Electric Generating Station. Summary Report.
Southern Company Services, Inc. 1977.
121. Rush, Randall E. and Reed A. Edwards. Evaluation of Three 20 MW Proto-
type Flue Gas Desulfurization Processes. Final Report, 3 Vols.
EPR1-FP-713-SY, EPRI RP536-1. Birmingham, AL. Southern Company
Services, Inc. March 1978.
122. LaMantia, et al. 3 op.dt.
2-213
-------
123. Rush and Edwards, Operational Experience... p. 371.
124. Ibid. , P. 370.
125. Lamantia, et al., op.oit. , Vol. II.
126 . Ibid.
127. Ando, "Status of Flue Gas...
128. Ando, "Status of S02...
129 . Ibid.
130. Kirschner, Gwen, (Caterpillar Tractor Co>, Massville, IL>. Private
Communication with Teresa Sipes, 4 January 1978.
131. Tuttle, J., et al., op. oit.
132. Tunison, Dave (Technical Superintendent, FMC Plant, Modesto, CA).
Private Communication with Teresa Sipes, 11 January 1978.
133. Ibid.
134. Rush and Edwards, Evaluation of Three... p. 9.
135. Ibid.
136. Ibid.
137. Interess, op.oit.
138. Ibid. , p. 5
139. Tuttle, J.,etal., op. oit.
140. Tuttle, et al. f op.oit.
141. Devitt, op.cit.
142. Choi, P.S., et al. SOa Reduction in Non-Utility Combustion Sources-
Technical and Economic Comparison of Alternatives. Final Report.
EPA 600/2-75-073, Task 13. Columbus, OH. Battelle-Columbus Labora-
tories. October 1975.
143. FMC Corporation. Environmental Equipment Division, Capabilities
Statement, Sulfur Dioxide Control Systems. Itasca, IL. March 1976.
144. Ottmers, et al.^ op.oit., p. 247.
2-214
-------
145. Ibid., p. 248.
146. Ibid., p. 251.
147. Melia, M., et al., op. cit.
148. Ando, J., "S02 Abatement..."
149. Pedroso, R.I. "An Update of the Wellman-Lord Flue Gas Desulfurization
Process." In Proceedings, Symposium on Flue Gas Desulfurization,
New Orleans, March 1976. 2 Vols. EPA, 1976. pp. 719ff.
150. Devitt, op. ait., p. 3-275.
151. Devitt, op. cit., p. 3-292.
152. Melia, M., et al.3 op. cit.
153. Ottmers, et al., op. cit. p. 260.
154. Melia, op. cit*3 P- 77.
155. Ayer, Franklin A. Comp., Flue Gas Desulfurization. Hollywood, FL.
November 1977. Symposium Proceedings, 2 Vols. EPA Contract No.
68-02-2612, Task 38, EPA 600/7-78-058a. Research Triangle Park, NC.
Research Triangle Institute. March 1978.
156. Laseke, et al.„ op.cit.
157. Ibid.
158. Putnam, A.A., E.L. Kropp, and R.E. Barrett. Evaluation of National
Boiler Inventory. Final Report. EPA Contract No. 67-02-1223, Task 31.
Columbus, OH. Battelle Columbus Labs. October 1975.
159. Boyer, Howard A. and Roberto I. Pedroso. "Sulfur Recovered From SO2
Emissions at NIPSCO's Dean H. Mitchell Station." Presented at the
Fourth Symposium on Flue Gas Desulfurization, Hollywood, FL. November
1977.
160. Ibid.
161. Ottmers, et al. J op. cit. , p. 258.
162. Pedroso, op. cit.
163. Boyer, op. cit.
2-215
-------
164. Kelly, W.E., et at. Air Pollution Emission Test, Second Interim Report,
Vol. 1, Consinuous Sulfur Dioxide Monitoring at Steam Generators.
EPA/EMB Report No. 77SPP23B. Research Triangle Park, NC. EPA, Office
of Air Quality Planning and Standards, Emission Measurement Branch.
March 1979.
165. Sedman Paper.
166. Link, F. William and Wade H. Ponder. "Status Report on the Wellman-
Lord/Allied Chemical Flue Gas Desulfurization Plant at Northern Indiana
Public Service Company's Dean H. Mitchell Station." Presented at the
EPA Symposium, Hollywood, FL., November 1977.
167. Ibid.
168. Ibid.
169. Ibid.
170. Ibid.
171. Boyer, op.ait.
172. Ottmers, et at., op.ait., p. 310.
173. Ibid., p. 313.
174. Sommerer, Diane K. Magnesia FGD Process Testing on a Coal-Fired Power
Plant. EPA Contract No. 68-02-1401, Tasks No. 1, 10, 24, and 25, EPA
660/2-77-165. Stanford, CT. York Research Corporation. August 1977.
175. Lowell, Philip S. and Frank B. Meserole. "Crystallization of MgSO.,.
3H?0 and MgSO^-6H?0 in the Magnesium Oxide Flue Gas Desulfurization
'resented at the AIChE 83rd National Meeting, Houston,
TX. March 1977.
176. Ottmers, et al., op.cit.
177- Koehler, George and James A. Burns. The Magnesia Scrubbing Process as
Applied to an Oil-Fired Power Plant. Final Report. EPA-600/2-75-057,
EPA Contract No. CPA 70-114. New York Chemical Construction Corp.
October 1975.
178. Sommerer, op.ait.
179. Tuttle, Patkar, and Gregory, op.oit.
180. Ibid.
181. Ando, "Status of Flue Gas...
182. Koehler and Burns, op.cit.
2-216
-------
183. Kelly, W.E., et al. Air Pollution Emission Test, First Interim Report.
Vol. 1, Continuous Sulfur Dioxide Monitoring at Steam Generators.
EPA Contract No. 78-02-2818, Work Assignment No. 2, EPA/EMB Report No.
77SPP23A. Research Triangle Park, NC. EPA, Office of Air Quality
Planning and Standards, Emission Measurement Branch. August 1978.
184. Sommerer, op. oit.
185. Ando, "Status of Flue Gas...
186. Ibid.
187. Tuttle, Patkar, and Gregory, op. oit.
188. Koehler and Burns,
189. Ibid.
190. Ibid., p. 100.
191. Sommerer, op. oit.
192. Ibid.
193. Ibid.
194. Ando, "Status of Flue Gas...
195. Sommerer, op. oit.
196. Devitt, et al., op oit., p. 3-236.
197- Downs, W. and A.J. Kubasco. Magnesia Base Wet Scrubbing of Pulverized
Coal Generated Flue Gas—Pilot Demonstration, PB 198 074. Alliance, OH
Babcock and Wilcox Co. 1970.
198. Devitt, et al. , op.oit. , p. 3-207-
199. Ibid., p. 3-208.
200. Ibid. , p. 3-237.
201. Choi, op.oit.
202. Tuttle, Patkar, and Gregory, op.oit.
203. Ibid.
204. Melia, op. oit., p. 79.
2-217
-------
205 Tuttle, J., et dl,. EPA Industrial Boiler FGD Survey: Third Quarter
1978. Final Report. EPA Contract No. 68-02-2603, Task 36. EPA 600/
7-78-052c. Cincinnati, OH. PEDCo Environmental, Inc. November 1978.
206. Ibid.
207. Ibid.
208. Leivo, op.ait.3 p. 4-25.
209. ibid., P- 4-27.
210. Ibid., p. 4-26.
211. Ibid., p. 4-24.
212. Tuttle, J., et dl., op. ait.
213- Energy and Environmental Analysis, Inc. Survey of the Application of
Flue Gas Desulfurization Technology in the Industrial Sector. FEA
Contract No. CO-05-60469, FEA/G-77-304. Arlington, VA. December 1976.
214. Fujimoto, op.ait.
215. Ottmers, et dl., op.ait. , p. 98.
216. Janssen, K.E., and R.L. Erickson. "Basin Electric's Involvement with
Dry Flue Gas Desulfurization." Paper No. 4M. Presented at the Fifth
Symposium on Flue Gas Desulfurization. Las Vegas, Nevada. March 1979.
217- Morre, K.H., R.D. Oldenkamp, and M.P- Schreyer. "Dry FGD and Parti-
culate Control Systems." Prepared for the Fifth Symposium on Flue
Gas Desulfuruzation. Las Vegas, Nevada. March 1979.
218. Davis, R.A., J.A. Meyler, and K.E. Gude. "Dry SO-j Scrubbing at
Antelope Valley Station." Presented at the American Power Conference.
April 1979.
219. Canada, Ministry of the Environment, Sulphur Dioxide Removal From Stack
Gases, A Review of Available Methods. Report No. ARB-TDA-02-75.
Ontario. March 1975.
220. Tuttle, J., et dl.-3 op. ait.
221. Wilkens, John. "S02 and Particulate Control for the Antelope Valley
Station, An Evaluation." Presented at the Western Precipitation
Seminar. Tamarron, Durango, CO. May 1978.
2-218
-------
222. Felsvang, K. "Results of Pilot Plant Operations for S02 Absorption."
Presented at the JOY, Western Precipitation Division Seminar. Durango,
CO. May 1978.
223. Kaplan, S.M., and Karsten Felsuang. 'Spray Dryer Absorption of SOa
From Industrial Boiler Flue Gas.' Presented at the 86th National
AIChE Meeting. Houston, Texas. April 1S79.
224. Dustin, Donald F. Report of Coyote Pilot Plant Test Program, Test
Report. Canoga Park, CA. Rockwell International, Atomics International
Division. November 1977.
225. Ibid.
226. Ottmers, et al. _, op.ait., p. 111.
227. Dustin, op.cit.
228. Ottmers, et at., op.cit., p. 120.
229. ibid. , pp. 159-165.
230. Stern, et al. , op.cit., p. D-46, D-47.
231. Rochelle, Gary T. "Process Alternatives for Stack Gas Desulfurization
With Steam Regeneration to Produce S02-" Presented at the EPA Sympo-
sium of Flue Gas Desulfurization. Hollywood, FL. November 1977.
232. Madenburg, R.S., D.M. Paulsruke, and C.H. Sherman. "The Citrate Pro-
cess for Sulfur Dioxide Emission Abatement-Process Instrumentation and
Control." Paper No. 75-658. Presented at the ISA Industry Oriented
Conference and Exhibit. Milwaukee, WS. October 1975.
233. McKinney, W.A., et al. "Pilot Plant Testing of the Citrate Process
for S02 Emission Control." Presented at the Flue Gas Desulfurization
Symposium. Atlanta, GA. November 1974.
234. Ottmers, et al., op.cit.. p. 175.
235. Rochelle, op.cit. «s
236. Ottmers, et al., op.cit., p. 178.
237- Madenburg, R.S. and R.A. Kurey. "Citrate Process Demonstration Plant -
A Progress Report." Presented at the EPA FGD Symposium. Hollywood,
FL. November 1977.
238. Stern, et al., op.cit. , p. D-50.
239. Ottmers, et al., op.cit., p. 61.
2-219
-------
240. Holighaus, Rolf. "Desulfurization Technology in the Federal Republic
of Germany." Presented at the EPA Flue Gas Desulfurization Symposium.
Hollywood, FL. November 1977.
241. Rush and Edwards, Operational Experience...
242. Stern, et al. , op.oit. , p. D-27 .
243. Ottmers, et at., op.oit., pp. 72-76.
244. Rush and Edwards, Evaluation of Three — p. 5-12.
245, Ottmers, et al., op.oit., pp. 76-79.
246. Rush and Edwards, Evaluation of Three —
247- Ottmers, et al., op.oit., pp. 79-82.
248. Rush and Edwards, Evaluation of Three... pp. 5-19, 5-23.
249. Stern, et al., op.oit., p. D-28.
250. Ottmers, et al., op.oit., pp. 90-92.
251. Rush and Edwards, Evaluation of Three...
252. Ottmers, et al. , op.oit., p. 99.
253. Weiss, Lawrence H. "Evaluating Sulfur-Producing FGD Processes."
ACS, Division Fuel Chem. , Prepr. 21(7), 125-33(1976).
254. Binns, Donald R. and Robert G. Aldrich. "Design of the 100 MW Atomics
International Aqueous Carbonate Process Regenerative FGD Demonstration
Plant." Presented at the Fourth Symposium on Flue Gas Desulfurization.
Hollywood, FL. November 1977.
255. Ottmers, et al. , op.oit., p. 100.
256. Binns and Aldrich, op. ait.
" 257. Weiss, op.oit.
258. Stern, et al. , op.oit., p. 15.
259. Binns and Aldrich, op.oit.
26°' FlueWSi *nte™ati°nal> Atomi" International Division. Advanced
Flue Gas Desulfunzation Process, Vol. 1, Statement of Qualications
Response to EPA Solicitation DU-75-2001. Undated. ^cations.
2-220
-------
261. Binns and Aldrich, op.oit.
262. Ottmers, et al. , op.oit., pp. 107-119.
263. Stern, et al.., op.oit., p. D17, 20.
264. Ottmers, et al., op.oit. , pp. 127-129.
265. Dautzenberg, F.M., J.E. Nader, and A.J.J. vanGinnekan. "Shell's Flue
Gas Desulfurization Process." Chemical Engineering Progress 67(8),
86-91(1971).
266. Ottmers, et al., op.oit., pp. 27-35.
267- Universal Oil Products Company. Shell Flue Gas Desulfurization Process.
Des Plaines, IL. April 1974.
268. Ottmers, et al., op.eit., pp. 45-46.
269. Arneson, Allen D., Frans M. Nooy, and Jack B. Pohlenz. "The Shell FGD
Process Pilot Plant Experience at Tampa Electric." Presented at the
Fourth Symposium on Flue Gas Desulfurization. Hollywood, FL. November
1977.
270. Ottmers, et at., op.eit., pp. 41-44.
271. Arneson, Nooy, and Pohlenz, op.oit.
272. stern, et al., op.oit., pp. D117-120.
273. Chiyoda International Corporation. Introduction to Chiyoda Thoroughbred
121 Flue Gas Desulfurization Process. Seattle, WA. December 1976.
27.4. Idemura, H.T. Kanai and H. Yanagioka. "Jet Bubbling Flue Gas Desul-
furization." CEP 1978 (Feb.), 46.
275. Stern, et al., op.oit., p. D40.
276. Idemura, Kanai, and Yanagioka, op.ait.
277. Clasen, D.D. and H. Idemura. "Limestone/Gypsum Jet Bubbling Scrubbing
System." Presented at the EPA Symposium on Flue Gas Desulfurization.
Hollywood, FL. November 1977.
278. Stern, et al., op.eit., pp. D41-42.
279. Tuttle, J., et al.^ op.oit.
2-221
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SECTION 3
CANDIDATES FOR BEST SYSTEMS OF S02 EMISSION REDUCTION
In Section 2 of this report, eleven FGD processes were described and
discussed with regard to their application to industrial boilers. These
processes included both commercially available and developing technologies.
In this section, these eleven processes are compared against each other in
order to select those that appear to be the best candidates for industrial
boiler application. The processes selected as best candidates will then
be evaluated in detail with regard to their energy, environmental, and
economic impacts. Results of the detailed evaluations will be used to
compare the candidate processes for application to different sized boilers,
with different coal types and S02 removal levels.
The processes selected as best candidates for application to industrial
boilers are:
Sodium Scrubbing
Double Alkali
Lime/Limestone
Spray Drying
• Wellman-Lord
The first three processes are all currently used to control S02 emis-
sions from industrial boilers throughout the United States. These processes
are the ones of choice as evidenced by the fact that 118 of the 132 operating
industrial boiler FGD systems are of these process types. The remaining
operating industrial boiler FGD systems use ammonia process waste waters as
a sorbent and are predominately found in sugar processing plants. Lime/
3-1
-------
limestone, double alkali, and sodium scrubbing processes also appear to be
the processes of choice for future installations as evidenced by the fact
that 36 out of 39 systems in the planning or construction stages are of
these process types.
The spray drying process was selected as candidate technology for
industrial boiler applications due to its potential for widespread
use as evidenced by the large amount of interest expressed in this
rapidly developing process. Presently, there are no full-scale spray
drying FGD systems in operation; however, orders have been placed for
five commercial spray drying systems, two of which are industrial boiler
applications. Data on spray drying is available only from pilot scale
units. This data will be used in later sections to compare the energy,
environmental, and economic impacts of this process versus the other
candidate processes.
The Wellman-Lord process was selected to compare the impacts of a
regenerable FGD process against the other candidate processes which are all
throwaway systems. There are currently no regenerable systems in operation
on small industrial boilers, however, a demonstration of the Citrate Process
is scheduled for operation in the near future. The Wellman-Lord process was
selected as a candidate process over the Citrate Process primarily because of
its increased development status and availability of data.
In order to select the candidate processes, a set of screening
criteria was first developed that would insure consistency in evaluating
all the processes. Next, a tabular comparison of the FGD processes was
made for each screening criteria. Finally, the FGD processes judged to
be the best candidates for industrial boiler applications were selected.
The following sections describe the criteria that were developed for this
study and present tabular comparisons of the FGD systems.
3-2
-------
3.1 CRITERIA FOR SELECTION OF BEST SO CONTROL SYSTEMS
As an aid to developing the screening criteria, several FGD process
vendors and operators of industrial boiler FGD systems now in use or being
planned were contacted to determine what factors they considered to be
important in the application of FGD processes to their industrial boilers.
The people contacted were asked why they selected the process they did; if
they had been able to successfully operate it; and if they would use the
same process if additional capacity were needed. Results of these contacts
indicated that proven performance, process simplicity, and economics were
the key factors for selecting a particular FGD process. Most FGD operators
also indicated that they were satisfied with their process choices and
seemed content with improving existing operations rather than switching
2
processes.
Evaluation or screening criteria were then established to provide an
objective and consistent means of comparing the processes and to insure that
the same factors were considered for each process. The screening criteria
were then applied to each process and the results were compared and used to
select the processes that appeared to be best suited for near-term industrial
boiler applications. In some cases, engineering judgment was applied to
rate the processes due to their relatively early development stage and lack
of accompanying operating data. The screening criteria are listed in
Table 3.1-1 and discussed in the following sections as they relate to the
application of FGD systems to industrial boilers.
3-3
-------
TABLE 3.1-1. FLUE GAS DESULFURIZATION SCREENING CRITERIA
1. Status of Development
• Overall Process Development
• Availability of Data
2. Performance
• SO2 Removal
• Reliability
3. Applicability
• Simplicity
• Flexibility
• Controllability
• By-Product Marketability
4. Economic Considerations
• Capital Investment Costs
• Operating Costs
5. Energy Considerations
• Liquid Pumping Requirements
• System Pressure Drop
• Regeneration Energy
• Requirement for Reducing Gas
6. Environmental Considerations
• Multipollutant Control
• Secondary Pollutant Emissions
1) Development Status - These criteria considered both the overall
process development and the availability of data. Development
status is important because the more developed process will in
general have more operating data and will provide a sounder data
base for evaluating process economics.
2) Performance - Performance was interpreted to mean how well a
system has demonstrated its desulfurization ability. Factors
that are implicit in this definition are: S02 removal ability
and system reliability. Particulate removal is often considered
a performance criteria for FGD systems. However, for this process
screening, particulate removal was assumed to occur upstream of
the FGD process.
3-4
-------
S02 removal ability is self-explanatory. Three removal efficien-
cies of 75, 85, and 90 percent were considered. Reliability is
an important performance criteria since it can impact process
capital and operating costs. Systems with poor reliability may
require extra modules to maintain acceptable operations.
3) Applicability - The applicability of an FGD process to industrial
boilers can be assessed by considering such factors as process
simplicity, flexibility, controllability, and by-product market-
ability. Process simplicity is especially important to small
industrial boiler applications, where FGD system capital and
operating costs may become a signficant portion of the boiler
expenses. In general, the less complex the process the more
applicable it is.
Process flexibility was defined to be the ability to separate
process steps within a system (decoupling). A process that can
be easily decoupled can maintain operations in one section of
the process while another section undergoes- repairs. Process
decoupling may also permit the concept of a centralized regen-
eration center. This concept may significantly reduce the
overall costs of a regenerable system. However, for the
concept to be viable, several FGD systems must be located in
a close proximity.
Controllability refers to the relative ease of operating the FGD
process and will have an effect on process operating costs. In
general, the more easily controlled processes have lower operating
costs which is a significant benefit particularly to small
installations.
By-product marketability will also affect process applicability
since available land-fill disposal space is often difficult to
3-5
-------
find. However, in some areas of the country, land-fill may be
selected for economic reasons, whereas in other land-limited
areas, it may be more economical to use processes which make
saleable products such as gypsum, sulfuric acid, sulfur or
liquified S02.
4) Economic Considerations - FGD process costs are very important
for small industrial boiler applications and will probably be
the major factor in selecting one FGD process over another.
Both capital and annualized costs must be evaluated for each
FGD system. For this preliminary screening, however, only
rough process cost estimates could be made for the eleven FGD
systems being evaluated. More detailed cost estimates of the
selected candidate processes will be presented in Section 4.
5) Energy Considerations - Criteria to evaluate process energy require-
ments considered process features as liquid pumping requirements,
system pressure drop, the relative amount of energy required for
regeneration, and the requirement for reducing gas. Liquid pumping
requirements and system pressure drops indicate the relative amounts
of energy required for operating a system's pumps and fans.
Regeneration energy requirements were based on an examination of
the relative energy intensiveness of the unit operations used in
each of the regenerable processes. The requirement for a reducing
gas must be considered a process liability since reducing gases
are produced either from reforming premium fuels such as naphtha or
methane, or coal gasification. Either method of producing a reduc-
ing gas will add to process costs and complexity, in addition to
increasing its energy consumption. Section 5 of this report will
evaluate the energy usage of the candidate processes selected for
industrial boiler application in detail.
3-6
-------
6) Environmental Considerations^ - Criteria to evaluate environmental
impacts considered two main items: 1) multipollutant control,
and 2) secondary pollutant emissions. Multipollutant control
capability is important since it has the potential of reducing the
overall costs of an emission control system. Secondary pollutant
emissions are important since it is undesirable for an FGD system
to remove S02 at the expense of producing a secondary air, liquid
or solid emission. Section 6 of this report will evaluate the
environmental impact of the selected candidate processes in more
detail.
3.2 SELECTION OF BEST CONTROL SYSTEMS
Each of the emission control techniques discussed in Section 2 can be
designed to achieve any of the three emission control levels. The impact of
increased S02 removal efficiency from 75 to 90 percent is predominately an
economic one and will differ from system to system. In general, there will
be less of an economic impact associated with the increased level of S02
control for the highly alkaline, clear liquor, sodium-based systems than
there will be for the lime/limestone systems. This impact will be evaluated
by performing economic sensitivity studies which will be presented in
Section 4 for the selected candidate systems.
As discussed in Section 2, the impact of fuel type when going from
coal- to oil-fired operations should not have an adverse affect on FGD sys-
tems. This fact is illustrated by several Japanese systems which have shown
good operability on both coal- and oil-fired boilers. Consequently, this
evaluation will consider only coal-fired operations.
All of the emission control techniques discussed in Section 2 will be
considered in this section as possible candidates for the best control system
for industrial boiler applications. The systems will be compared and dis-
cussed with regard to the criteria discussed in Section 3.1.
3-7
-------
3.2.1 Development Status
One of the most important criterion for selection of candidate control
systems was status of development. Table 3.2-1 shows the current status of
development for the S02 control systems being considered. Of the opera-
tional processes, it is immediately apparent that the lime/limestone, double
alkali, and sodium throwaway processes are the most highly developed and most
widely used systems on industrial and utility boilers in the U.S. Wellman-Lord
has recently undergone a demonstration on a coal-fired utility boiler with
two additional utility units also under construction. The magnesium oxide
process is currently operational at the 120 MW scale with three units in
the planning stages.
With regard to systems in the planning or construction phases, the
lime/limestone, double alkali, and sodium'throwaway processes are again the
predominate processes. The spray drying process has five systems in planning
or construction phases which illustrates the large amount of interest in
this rapidly developing technology. The remaining processes, Citrate,
Bergbau Forschung/Foster Wheeler, Atomics International, Shell, and Chiyoda
121 are all currently in the demonstration stage and have no announced plans
for commercial systems to be built in the near future.
3.2.2 Performance
Comments on S02 removal and reliability are highlighted for each of the
candidate processes in Table 3.2-2. Early lime/limestone systems experienced
many reliability problems, but better understanding of the process chemistry
and improved process design has led to an improved reliability for these
systems. Generally, clear liquor processes such as double alkali and sodium
scrubbing have shown reliable operations. The first Wellman-Lord system,
also a clear liquor system, applied to a coal-fired boiler in the U.S. has
reported high system reliability during the first testing period. The
magnesium oxide process has demonstrated greater than 90 percent SOz removal
but has not been operated continuously for longer than eight days.
3-*
-------
TABLE 3.2-1. OVERALL STATUS OF DEVELOPMENT OF CANDIDATES FOR BEST SYSTEMS OF S02 REDUCTION
Process
Operational
Industrial Utility
No. No.
Planned/Construction
Industrial Utility
No. No.
Comments
u>
I
VO
Lime/Limestone
Double Alkali
Wellman-Lord
Magnesium Oxide
Sodium Scrubbing
Spray Drying
Citrate/Phosphate
Bergbau Forschung/
Foster Wheeler
Atmocs International
Aqueous Carbonate
Shell/UOP
Chiyoda 121
34
63
14
102
11
2
1
3
0
0
23
4
3
1
0
0
0
0
0
0
0
0
0
0
1
0
0
0
0
0
0
1
0
1
Total industrial capacity approxi-
mately 489,000 scfm. Total
utility capacity approximately
43,000 MWe.
Total industrial capacity approxi-
mately 1,445,000 scfm. Total
utility capacity approximately
1100 MWe-
Total utility capacity approxi-
mately 1850 MW£.
Total utility capacity approxi-
mately 846 MWe«
Total industrial capacity approxi-
mately 5,081,000 scfm. Total
utility-capacity approximately
874 MW£i
Total industrial capacity approx-
imately 73,000 scfm. Total
utility capacity approximately
1300 MWe.
Total industrial capacity
approximately 142,000 scfm.
No active units in U.S.
100 MW demonstration unit planned
for operation in 1980.
.6 MWepilot unit in operation
Capacity approximately 23 MWe,
-------
TABLE 3.2-2. PERFORMANCE OF CANDIDATES FOR BEST SYSTEMS FOR S02 REDUCTION
Process
Comments
Lime/Limestone
Double Alkali
Lime/limestone systems have demonstrated greater
than 90 percent S02 removal.1* In the past relia-
bility has been a problem, but there is a trend
toward better reliability.5
Very high removals (>90 percent) and reliabilities
have been demonstrated by double alkali systems on
industrial boilers in the U.S. High oxygen levels
in the flue gas will lead to high sulfite oxidation,
which results in sodium sulfate which is less
active.
Wellman-Lord
Magnesium Oxide
Sodium Scrubbing
The first application of the Wellman-Lord process to
a coal-fired boiler in the U.S. has produced very
good results. S02 removals greater than 90 percent
are reported. Reliability was high during the
'testing period. High levels of oxygen in the flue
gas can promote sulfite oxidation which produces
sodium sulfate which must be purged from the system.
S02 removal efficiencies of greater than 90 percent
have been demonstrated.8 Overall process relia-
bility is difficult to evaluate since the longest
continuous period of operation to date has been
eight days.9
SO removal levels of greater than 90 percent are
reported.
excellent.
System reliability is generally
Spray Drying
Citrate/Phosphate
Pilot unit test results have shown greater than 90
percent removal for low sulfur coal operations.11
Reliability should be very good although only pilot
scale units have been operated to date. A careful
balance between flue gas condition and dryer oper-
ation is necessary to prevent condensation in
downstream equipment.
Pilot plant operations have reported up to 99 per-
cent S02 removal.12 Operation of the 64 MWe unit at
St. Joe Minerals will provide more information con-
cerning performance and reliability.
3-10
-------
TABLE 3.2-2. (Continued)
Process Comments
Bergbau/Forschung/ Operation of the BF/FW process in the U.S. at Gulf
Foster Wheeler Power's Scholz Steam Plant experienced only limited
success. Many mechanical problems were experienced
which adversely affected process performance. The
system at Liinen, Germany has reported much more
successful operations.
Atomics International A fully-integrated Aqueous Carbonate Process system
Aqueous Carbonate has never been operated. A 100 MW demonstration
unit is planned for operation in 1980 which will
help to answer performance and reliability questions.
Shell/UOP Shell has yet to build a completely integrated unit
applied to a coal-fired boiler. Questions about the
system's performance and reliability remain.
Chiyoda 121 Conversion of the 23 MWe CT-101 system at the Scholz
Power Station to a CT-121 process is underway. S02
removals of about 90 percent were achieved during
the pilot plant testing. ^ The process should be
rather insensitive to flue gas variations, in fact,
excess oxygen should be beneficial.
3-11
-------
Three processes (double alkali, Wellman-Lord, and spray drying) are
sensitive to flue gas conditions. Double alkali and Wellman-Lord are ad-
versely affected by high oxygen'levels in the flue gas since any oxidation
products formed are less active thus requiring additional alkali and regen-
eration gases, respectively. Spray drying is sensitive to flue gas temper-
ature, so the sorbent feed rate must be adjusted accordingly to avoid
downstream condensation.
3.2.3 Applicability
Table 3.2-3 presents a summary of the applicability of the candidate FGD
processes to industrial boilers. The throwaway processes (lime/limestone,
double alkali, sodium scrubbing, spray drying and Chiyoda 121) are charac-
terized by lower space requirements in the area near the boiler, fewer pro-
cess steps, and no marketable by-product. These processes will, however,
require land for solid waste disposal. The regenerable processes (Wellman-
Lord, MgO, Citrate, Bergbau, Atomics International, and Shell/UOP) require
more space near the boiler than the throwaway processes, have more process
steps, and produce marketable by-products. In addition, they are generally
more complex with sensitive controls. All the processes can be decoupled
to some extent.
3.2.4 Economic Considerations
Preliminary cost estimates for 200xl06 Btu/hr industrial sized FGD
systems, which are reported in Table 3.2-4, were developed by directly
scaling down utility FGD system cost estimates by process size. The
utility costs were based on a series of cost estimates prepared by
TVA for an Interagency Evaluation of FGD systems, and as such are
developed on a consistent bases.
3-12
-------
TABLE 3.2-3. APPLICABILITY OF CANDIDATE SYSTEMS TO INDUSTRIAL BOILER
OJ
i
Process
Lime/Limestone
Double Alkali
Wellman-Lord
Magnesium Oxide
Sodium Scrubbing
Spray Drying
Citrate/Phosphate
Bergbau-Forschung/
Foster Wheeler
AI Aqueous Carbonate
Shell/UOP
Chiyoda 121
Space
requirements
Medium
Medium
High
High
Low
Low
High
High
High
High
Low
Process
simplicity
Moderate
Moderate
Complex
Complex
Simple
Simple
Complex
Complex
Complex
Complex
Simple
By-product
Sludge/Gypsum
Sludge/ Gyp sum
Sulfur/Acid
Acid
High TDS Liquid
Dry Sodium/
Calcium Salts
Sulfur
Sulfur
Sulfur
Sulfur/Acid
Sludge/Gypsum
-------
TABLE 3.2-4. PRELIMINARY FGD SYSTEM COST ESTIMATES
(200 X 106 BTU/hr Sized Systems)16
Process
Lime /Limes tone
Double Alkali
Wellman-Lord
Magnesium Oxide
Sodium Scrubbing
Spray Drying
Citrate /Phosphate
BF/FW
Atomics International-ACP
Shell/UOP
Chiyoda 121
Preliminary
capital
cost
(106 $)
1.9
2.0
2.3
2.1
1.8
1.8
2.9
3.2
2.2
3.8
*
Relative
Economic
Rank
1
1
2
2
1
1
3
3
2
3
1
*No capital cost data was available; however, the vendor claims capital costs
should be about 60 percent of the capital costs of a limestone system.
3-14
-------
Results of these estimates indicate that, in general, the mechanically
simple throwaway processes will be less expensive than the more complex
regenerable processes. This result is expected when one considers that users
can be expected to choose an FGD process on the basis of cost, and essen-
tially all the existing and planned FGD systems on U.S. utility and industrial
boilers are of the throwaway type.
Of the throwaway processes, the sodium scrubbing and spray drying
processes are shown to be the least expensive followed by the lime/limestone,
and double alkali processes. As noted in Table 3.2-4 the cost for the
Chiyoda 121 system was not available except for a statement by the vendor
which claimed that its capital costs would be 60 .percent of a limestone
system. If these costs do become verified upon further process develop-
ment, the Chiyoda 121 process will become the least expensive.
Of the regenerable processes, the magnesium oxide, Wellman-Lord, and
Atomics International process are shown to be the least expensive, with the
Citrate, Bergbau, and Shell/UOP processes being somewhat more costly. The
regenerable processes, except for the Wellman-Lord and magnesium oxide pro-
cesses, are in relatively early development stages and consequently, their
estimated costs are more speculative. A more in-depth cost analysis of the
processes selected as being best suited for industrial boiler applications
will be presented in Section 4 of this report.
3.2.5 Energy Considerations
A summary of the estimated energy requirements for the candidate systems
is shown in Table 3.2-5. Results of these estimates indicate that the
throwaway processes are less energy intensive than the regenerable processes.
This is largely due to the energy required by the regenerable processes for
regeneration of the SOz sorbent and for producing the by-product sulfur or
sulfuric acid.
3-15
-------
TABLE 3.2-5. FGD SYSTEM ENERGY REQUIREMENTS T 8' 1 9
(200 X 106 BTU/hr system, 3.5%S coal, 90% S02 removal)
Liquid
Pumping
Process Requirements
(L/G)
Lime/ Limes tone
Double Alkali
Wellman-Lord
Magnesium Oxide
Sodium Scrubbing
Spray Drying
Citrate/Phosphate
w Bergbau-Forschung/
*-• Foster Wheeler
Atomics International
Aqueous Carbonate
Shell/UOP
Chiyoda-121
High
Low
Low
Medium
Low
Low
Low
None
Low
None
Medium
System AP
Medium
Low
Low
Medium
Low
Low
Medium
High
Low
Medium
Medium
Regeneration
Energy
Required
No
No
High
Med ium
No
No
Low
Medium
Low
Medium
No
Reducing
Gas
No
No
Yes
No
No
No
Yes
No
No
Yes
No
Overall
Energy Requirements
(10b Btu/hr)
6.4
5.6
16.8
10.7
5.6
2.6
11.8
17.9
8.7
9.2
6.7
* Ranking
Medium
Low-Medium
High
Medium-High
Low-Medium
Low
Medium-High
High
Medium-High
Medium-High
Medium
* Energy requirement estimates include energy for stack gas reheat of 3 X 10 BTU/hr except for
the Spray Drying, Atomics International, Bergbau-Forschung, and Shell/UOP Processes.
-------
This preliminary analysis indicates that the spray drying process will
be the most energy efficient process, to be followed by the sodium scrubbing
and double alkali processes, and the limestone and Chiyoda 121 processes.
The low energy requirements of the spray drying system are due to the fact
that no reheat energy is required. If reheat requirements were not included
in these estimates, the overall energy required by the sodium, dual alkali,
and spray drying processes would be essentially the same.
Of the regenerable processes the Atomics International, Shell/UOP, and
magnesium oxide processes are the least energy intensive processes, and the
Bergbau and Wellman-Lord processes are the highest energy users. Although
the Wellman-Lord process has a relatively high rate of energy use, most of
the energy consumed is in the form of low pressure steam which in many cases
can be supplied from a waste steam source. A more detailed study of FGD pro-
cess energy requirements is presented in Section 5 for the systems selected
as best candidates for industrial boiler applications.
3.2.6 Environmental Considerations
The major environmental impacts considered for this process screening as
shown in Table 3.2-6 were multipollutant control and secondary pollutant
emissions. With regard to multipollutant control, all of the FGD processes
under consideration have the potential for some particulate (fly ash)
removal. However, removal of fly ash in the SC>2 scrubber of a regenerable
system may lead to contamination of the by-product sulfur or acid and may
also adversely affect the process chemistry. Consequently, particulate
removal should occur upstream of regenerable systems. In general, the
throwaway systems can remove both fly ash and SC>2 simultaneously in the gas-
liquid contractor since a throwaway sludge is produced. Upstream particu-
late removal will probably, however, be required for the Chiyoda 121 process
if a marketable quality product gypsum is desired.
3-17
-------
TABLE 3.2-6. FGD SYSTEM ENVIRONMENTAL IMPACTS
t_o
I
Multipollutant control
Process
Lime/Limestone
Double Alkali
Wellman-Lord
Magnesium Oxide
Sodium Scrubbing
Spray Drying
Citrate/Phosphate
Particulate
removed
Yes
Yes
No
No
Yes
Yes
No
Bergbau-Forschung/ No
Atomics International No
Aqueous Carbonate
Shell/UOP
Chiyoda 121
No
No
NOX
removal
No
No
No
No
No
No
No
10-40%
No
Yes
No
Secondary pollutants generated
Air
None
None
None
Possible par-
ticulate emis-
sions from
calcining
None
None
Possible H2S
fugitive
emissions
Possible en-
trained carbon
fines in exit gas
Possible H2S
fugitive
emissions
Possible NH3
emissions
None
Solid
Sludge (CaS03, CaS04)
Sludge (CaS03, CaSOj
Relatively small
Na2SOit purge
None
None
Dry Sodium/
Calcium Salts
Relatively small purge
Na2S04
None
Fly ash filter cake
Attrited copper oxide
Gypsum solids
Liquid
Possible leaching
from waste solids
Possible leaching
from solids
Possible Chloride
purge
Small purge strean
High TDS waste-
water stream
None
None anticipated
Small purge streai
Small purge streai
Small purge strea
Possible leaching
from solids
-------
In general, no appreciable NO removal will be experienced by the FGD
systems except for the Shell/UOP process which has demonstrated up to 80
percent NO removal with ammonia addition. Up to 70 percent NO removal was
X X
also reported to occur on the carbon of the Bergbau process. However, addi-
tional testing has not confirmed the early results and this matter is still
under investigation.
As far as secondary pollutants are concerned, none of the processes
should have significant air emissions. Significant quantities of solid
wastes will be generated for the limestone, dual alkali, and spray drying
processes. The limestone and dual alkali processes will produce a calcium
sludge containing from 30-70 percent solids (rough calculations indicate a
58.6 MW (200 X 106 Btu/hr) system may produce about 2 tons per hour) and
the spray drying process will produce a dry sodium or calcium salt waste
material. Although the Chiyoda 121 process can produce a marketable quality
gypsum, disposal of this material may be necessary if a market cannot be
found. The amount of gypsum produced should be similar to the solids from
the limestone process. The regenerable processes should not have major
solid emission streams except perhaps for a Naj SO^ purge from the Wellman-
Lord process if high oxidation occurs. No major liquid effluent streams
are expected from these FGD processes except for the sodium scrubbing pro-
cess whose disposal product is a high TDS aqueous stream that is composed
mainly of dissolved sodium salts.
3-19
-------
REFERENCES
1. Tuttle, J., et al. EPA Industrial Boiler FGD Survey: Fourth Quarter
1978. Final Report. EPA Contract No. 68-02-2603 Task 45, EPA 600/
7-79-067a. Cincinnati, OH. PEDCo Environmental, Inc. November 1978.
2. Phillips, W.R. and J.C. Martinez. Private Communication with variuos
vendors and operators. September-October, 1978.
3. Tuttle, J., et al., op.cit.
4. Biedell, E.L., et al. EPA Evaluation of Bahco Industrial Boiler Scrubber
System at Rickenbacker AFB. Report No. EPA-600/7-78-115. June 1978.
5. Laseke, Bernard A. and Timothy W. Devitt. "Status of Flue Gas Desulfuri-
zation Systems in the United States." Presented at the EPA Symposium,
Hollywood, FL. November 1977.
6. Interess, Edward. Evaluation of the General Motors Double Alkali S02
Control System. Final Report. EPA Contract No. 68-02-1332, Task 3,
EPA 600/7-77-005. Cambridge, MA. Arthur D. Little, Inc. January 1977.
7. Boyer, Howard A. and Roberto I. Pedroso. "Sulfur Recovered from S02
Emissions at NIPSCO's Dean H. Mitchell Station." Presented at the Fourth
Symposium on Flue Gas Desulfurization, Hollywood, FL. November 1977.
8. Devitt, T., et al. Flu-e Gas Desulfurization System Capabilities for
Coal-Fired Steam Generators. Final Report, 2 Vols. EPA Contract No.
68-02-2603, Task 1, EPA 600/7-78-032a,b. Cincinnati, OH. PEDCo
Environmental, Inc. March 1978. p. 3-207.
9. Sommerer, Diane K. Magnesia FGD Process Testing on a Coal-Fired Power
Plant. EPA Contract No. 68-02-1401, Tasks No. 1, 10, 24, and 25,
EPA 660/2-77-165. Stanford, CT. York Research Corporation. August 1977.
10. Tuttle, J., A. Patkar and N. Gregory. EPA Industrial Boiler FGD Survey:
First Quarter 1978. EPA Contract No. 68-02-2603, Task 4, EPA 600/7-78-
052a. Cincinnati, OH. PEDCo Environmental, Inc. March 1978.
11. Felsvang, K. "Resulting of Pilot Plant Operations for S02 Absorption."
Presented at the JOY, Western Precipitation Division Seminar, Durango,
CO. May 1978.
3-20
-------
12. Ottmers, D.M., Jr., et ul. Evaluation of Regenerable Flue Gas Desul-
furization Processes, Revised Report, 2 Vols. EPRI RP535-1, EPRI
FP-272. Austin, TX. Radian Corporation. July 1976. pp. 69-70.
13. Stern, R.D., et at. Interagency Flue Gas Desulfurization Evaluation.
Revised Draft Report, 2 Vols. DCN 77-200-187-16-10. Austin, TX.
Radian Corporation. November 1977. V II, App. D.
14. Clasen, D.D. and H. Tdemura. "Limestone/Gypsum Jet Bubbling Scrubbing
System." Presented at the EPA Symposium on Flue Gas Desulfurization,
Hollywood, FL. November 1977.
15. Stern, et al. 3 op.cit.., App. D.
16. Ibid.
17. Stern, et al., op.cit.
18. Ottmers, et al. _, op.cit,., pp. 69-70.
19. Stern, et al. , op.cit., App. D, VII.
20. Ottmers, et al. _, op.cit.
3-21
-------
SECTION 4
COST ANALYSIS OF CANDIDATES FOR BEST EMISSION CONTROL SYSTEM
Five flue gas desulfurization (FGD) processes were selected in Section 3
as being the best candidates for application to small industrial boilers. The
processes selected were: Lime/Limestone, Sodium Scrubbing, Double-Alkali,
Spray Drying, and Wellman-Lord. In this section the costs of applying each of
these systems to the-Standard sized boilers will be considered. In addition,
the-cost impacts of achieving various SOa control levels, and of treating flue
gases of fuels with different sulfur contents will be addressed. Costs
presented in this section are estimates of installing a single FGD system on
each of the model boilers.
In order to minimize the number of cases and yet cover the range of
expected costs to be incurred, cost estimates will not be presented for the
lime process. Lime process costs have been evaluated by TVA and were
found to generally be within 10 percent of limestone process costs. In
addition, limestone is becoming the sorbent of choice for most new installa-
tions due primarily to the high calcining costs involved in producing lime
from raw mined limestone.2
The general approach used in developing the process costs consisted of
four main steps. First, a series of material and energy balance calculations
were performed for each process, to establish process strearr, flow rates and
energy requirements as functions of unit size, SOz removal, and the amount of
4-1
-------
sulfur in the coal. The energy requirements for each process are presented and
discussed in Section 5 and the material balance results are presented in
Appendix A. Second, each of the FGD processes were divided into a number of
process areas, or modules, which represented separate processing areas.
Third, equipment sizes were then developed for each process module based on the
results of the material and energy balances. Finally, capital cost estimates
were prepared by contacting process equipment vendors for price quotations in
the size range for the standard industrial boilers of this study. All of the
capital costs for each process area are reported in the form of direct capital
costs which include all materials and labor installation costs. Except for the
spray drying process, particulate control equipment costs were not included in
this study.
In order to assess the reasonableness of the cost estimates calculated for
this study, a series of graphs were prepared to illustrate the range of calcu-
lated installed costs versus site specific installed process costs reported in
the EPA Industrial Boiler FGD Survey.3 A range of calculated costs for a .given
process size is presented because of the cost differences attributable to coal
sulfur content and S02 removal percentage.
Figure 4.0-1 presents the comparison of the limestone total capital costs,
Figure 4.0-2 the dual alkali costs, and Figure 4.0-3 the sodium scrubbing
costs. No comparison can be made for the Wellman-Lord process since there are
no industrial boiler applications of this process in the United States.
There are currently two spray drying processes under construction on
industrial boilers, but only one of those systems has reported costs published
in the Industrial Boiler FGD Survey. That system treats 22,000 scfm of flue gas
and has reported costs of $1.4 million. Spray drying costs calculated for this
study on a 2.3 percent sulfur coal had estimated capital costs of $1.1 million
for a 19,000 scfm system and $2.7 million for an 86,000 scfm system. Interpola-
ting to the size of the reported spray drying system results in an estimated
capital cost of $1.2 million which compares well with the reported cost for
this process.
4-2
-------
2500
2000
CD
J-l
cfl
1500
a
-------
A
3.5% S Coal-
0.6% S Coal'
i
I
;90% SO 2 Removal
0 Calculated Costs
Reported Costs
_ I
1.4
(50)
2.8
(100)
4.2
(150)
5.6
(200)
7.0
(250)
Process Size 103 Nm /min (103scfn)
Figure 4.0-2. Comparison of calculated and reported total capital investment
for the double-alkali process.
4-4
-------
I
Ul
A
3.5% S Coal,
\
\
'90% S02 Removal
0.6% S Coal
A
2.8
(100)
5.6
(200)
Process Size 103
8.4
(300)
Nm3/min (103 scfm)
© Calculated Costs
Reported Costs
11.2
(400)
14.0
(500)
FIGURE 4.0-3. Comparison of Calculated and Reported Total Capital Investment
for the Sodium Scrubbing Process (excludes
wastewater treating costs) .
-------
The costs calculated for this study were for a general case and do not
consider site specific factors that produce the range in reported costs shown
in Figures 4.0-1 through 4.0-3. Although the process costs calculated for
this study may not predict the actual installed costs for a given site specific
case, they are representative of the actual installed costs reported by the
process users as the Figures show.
4.1 CONTRIBUTORS TO CONTROL COSTS AND COST BASES
Annualized process costs consist of three components:
annualized capital charges,
direct operating and maintenance costs, and
indirect operating costs.
The annualized capital charges component includes both direct and indirect
costs. Direct costs consist of the cost of equipment and auxiliaries as well
as the cost of installation. Direct capital costs were developed for each of
the FGD processes in the following manner.
As mentioned previously, for purposes of estimating process capital
costs, each of the five processes under consideration was divided into several
process areas or modules which were individually costed. Their results were
summed in order to develop overall process costs. This approach was used to
permit comparisons of the costs of various process areas from one process to
another. The major process areas defined for this evaluation and the process
variables used to size the equipment in each process area are briefly described
in the following sections as they impact the overall process costs.
The raw material handling and preparation process areas contain such
equipment items as storage silos, conveyors, screw feeders, mix tanks, agi-
tators, and pumps. This equipment is basically solid handling equipment and
costs of the solids handling equipment were estimated to vary with the raw
4-6
-------
material feed rate for each process. Pumping costs varied with liquid flow
rate through the process area. The raw material feed rates were calculated
for each case and are presented in Appendix A. Feed rates were found to vary
with boiler size, inlet S02, and S02 removal efficiency.
The S02 scrubbing and fans process areas are made up predominately of
gas/liquid contacting and gas handling equipment. Accordingly, their costs
were estimated to vary with the flue gas flow rate. Flue gas flow rates were
calculated for each case and are presented in Appendix A. The major factor
affecting flue gas flow is boiler capacity. In addition, the SOz scrubbing
process area contained agitated tanks and pumps. Costs of these equipment
items varied according to the liquid flow rate through the scrubber which was
found to vary with boiler size, inlet SO2 , and SO2 removal efficiency.
The sizes of the regeneration, solids separation, purge treatment, and
sulfur production process areas will vary with the amount of SO 2 being processed
for each case. Accordingly, costs for these process areas were estimated to
vary with the amount of SO 2 removed from the flue gas. The amount of SO 2
removed was calculated for each case and was found to vary with unit size, SO 2
inlet concentration, and SO2 removal level.
The wastewater treatment processing area for the sodium scrubbing process
consisting of oxidation and pH neutralization was assumed to be associated
with the boiler or plant in question. Therefore, wastewater treatment appears
as an operating cost only. Similarly, solids disposal for limestone and double
alkali processes was assumed to be contracted out for offsite disposal and it,
too, appears only as an operating cost.5 The discharge rates were calculated
for each of the cases and are presented in Appendix A. Unit size, S02 inlet
concentration, and percent removal all affect the wastewater and sludge
discharge rates.
Indirect capital costs are those not attributable to specific equipment
itea^'listed and include:
4-7
-------
engineering costs,
construction and field expenses,
contractors' fee,
allowance for funds used during construction
initial charge of chemicals and catalysts,
start-up and performance test,
contingency costs, and
working capital.
Often the indirect cost items are estimated as a percentage of the direct
capital costs or of the annual direct operating costs. Specific percentage
used in this study are indicated in Table 4.1-1. Engineering costs were
calculated only for the larger size processes and were assumed to be a constant
value for the smaller process cases. This was done to reflect a more accurate
impact of indirect costs for the small size FGD processes.
TABLE 4.1-1. INDIRECT CAPITAL COST FACTORS6
Item
Amount
Engineering
Construction and Field Expense
Contractor Fee
Start-up
Performance Tests
Contingency
Land
Working Capital
10% of Installed cost of 58.6 MWt
(200xl05 Btu/hr), eastern coal, 90%
removal case for each process. Con-
stant value for smaller process cases.
10% of installed cost of 118 MWt(400x10
Btu/hr), cases.
10% of installed cost
10% of installed cost
2% of installed cost
1% of installed cost or $2000
20% of total direct and indirect costs
0.084% of total turnkey costs
25% of total direct operating costs
-------
Annualized capital charges can be calculated via several methods, including
1) utility financing, 2) private investor financing (discounted cash flow
method), and 3) use of capital recovery factors. Capital charges can also
include local taxes, insurance costs, and general and administrative (G&A)
costs. By direction of EPA, the capital recovery method was used for this
analysis. The capital recovery factor was calculated from Equation 4-1:
i.,—1 (4~1}
where: i = interest (10%)
n = number of year (15)
for this case the capital recovery factor is 0.13147. When G&A expenses of 4
percent are added, the annual capital recovery factor becomes 17 percent of
the total capital investment.
Direct operating and maintenance (O&M) costs include:
utilities
raw materials,
operating labor,
maintenance and repairs,
fuel,
by-product credits, and
waste disposal.
Unit costs used to calculate the direct O&M costs are summarized in Table 4.1.2,
Transportation costs for the raw materials were not considered in this study.
The operating labor requirements were based upon one operator per shift
and the supervision labor requirements were based upon 0.15 men/Shift.7
Maintenance costs were based upon estimates reported by TVA.8
4-9
-------
TABLE 4.1-2. VALUES USED FOR ANNUAL COST ITEMS9
Item
Direct Labor
Supervision
Maintenance Labor
Maintenance Material
Electricity
Steam
Process Water
Methane
Lime
Limestone
Soda Ash
Solids Disposal
'" »i4
$ /man-hour
$ /man-hour
$/year
$/year
mills/kWh
$/GJ
$/m3
$/GJ
$/kg
$/kg
$/kg
$/kg
Cost
12.02
15.63
0.04(TDC)
0.04(TDC)
25.8
1.84
0.04
2.05
0.0385
0.0143
0.0991
0.044
Utility and fuel requirements were calculated for each case in a series
of material and energy balances. Results of these calculations showed the
variation in process energy requirements as a function of size, coal sulfur
level, and S02 removal and are reported in detail in Section 5, Energy Impact
of Candidates for Best Emission Control Systems. These calculations also
serve as the basis for the energy costs reported in this section.
Results of the raw material and energy balance calculations also serve as
the basis of the raw material cost estimates (see Appendix A). There are four
basic raw materials used by the FGD processes under consideration: limestone,
lime, sodium carbonate, and water. Other materials used by these systems such
as steam and natural gas were costed according to energy equivalency. The
amount of raw material used by each process was found to vary with both process
size and the amount of SO2 removed from the flue gas.
The third component of annualized costs is indirect operating costs.
These costs include both payroll and plant overhead charges. In this study,
payroll overhead was taken as 30 percent of the annual operating and super-
vision labor costs and plant overhead as 26 percent of the total annual labor
and maintenance costs.
4-10
-------
TABLE 4.2-2. SODIUM THROWAWAY PROCESS COST SUMMARY
Capital Cost
Boiler size
and type
8.8 MW
(30 x 106Btu/hr)
Underfeed Stoker
22 MW
(75 x 106Btu/hr)
Chaingrate Stoker
44 MW
(150 x 106Btu/hr)
Spreader Stoker
58.6 MW
t
(200 x 106Btu/hr)
Pulverized Coal
118 MW
t
(400 x 106Btu/hr)
Pulverized Coal
Coal
type
Eastern
3.5% S
Western
0.6% S
Eastern
3.5% S
Eastern
2.3% S
Western
0.6% S
Eastern
3.5% S
Western
0.6% S
Eastern
3.5% S
Western
0.6% S
Eastern
3.5% S
Eastern
2.3% S
Western
0.6% S
S°2 ,
removal
level
90
85
75
90
85
75
90
75
90
90
75
90
75
90
75
90
85
75
90
85
75
90
90
90
Direct
capital
costs
(103$)
222
222
219
188
187
187
369
359
346
294
291
553
538
452
447
615
607
599
499
497
494
937
856
766
Total
capital
investment
(103$)
457
456
449
394
392
392
705
685
657
559
554
1028
993
810
800
1151
1129
1108
886
882
875
1783
1588
1355
% Increase
over
„ Annualized costs uncontrolled
(10J$)
374
371
363
303
302
302
547
515
476
377
374
811
752
489
479
947
902
865
529
525
519
1532
1197
760
($/J/S)
0.043
0.042
0.041
0.035
0.035
0.035
0.025
0.023
0.022
0.017
0.017
0.018
0.017
0.011
0.011
0.016
0.015
0.015
0.009
0.009
0.009
0.013
0.010
0.006
($/MBtu/hr)
12,466
12,366
12,100
10,100
10,066
10,066
7,293
6,866
6,347
5,026
4,986
5,406
5,013
3,260
3,193
4,735
4,510
4,325
2,645
2,625
2,595
3,830
2,993
1,900
boilers
34
34
33
28
28
28
28
26
24
19
19
26
25
16
16
21
20
19
12
12
12
19
IS
10
4-13
-------
TABLE 4.2-3. DOUBLE -ALKALI PROCESS COST SUMMARY
Capital Cost
Boiler size
and type
8.8 MW
(30 x I06Btu/hr)
Underfeed Stoker
22 MW
(75 x 106Btu/hr)
Chaingrate Stoker
58.6 MW
(200 x IOsBtu/hr)
Pulverized Coal
118 MW
(400 x Iff Btu/hr)
Pulverized Coal
Coal
type
Eastern
3.5Z S
Western
0.6Z S
Eastern
3.5Z S
Eastern
2.3* S
Western
0.6Z S
Eastern
3.5Z S
Western
0.6Z S
Eastern
3.5Z S
Eastern
2.3Z S
Western
0.6Z S
SOi
removal
level
90
90
90
90
90
90
90
90
90
90
Direct Total
capital capital
costs investment
(10S$) (10S$)
369
315
522
489
442
775
641
1,083
1,000
912
699
603
960
889
799
1,422
1,115
2,105
"l,819
1,587
: Annualized coe
U0'$)
437
350
625
519
425
1,053
574
1,778
1,247
800
(S/J/S)
0.050
0.040
0.028
0.024
0.019
0.018
0.010
0.015
0.011
0.007
Z Increase
over
its uncontrolled
($/MBtu/hr)
14,566
11,666
8,333
6,920
5,666
5,265
2,870
4,445
3,117
2,000
boilers
41
32
32
28
20
24
13
23
16
10
-------
TABLE 4.2-4. SPRAY DRYING PROCESS COST SUMMARY
I
I—'
Ui
Boiler size
and type
17.6 MW
(60 x l66Btu/hr)
Chaingrate Stoker
22 MW
(75 x 106Btu/hr)
Chaingrate Stoker
44 MW
(150 x 106Btu/hr)
Spreader Stoker
58.6 MW
(200 x I06Btu/hr
Pulverized Coal
118 MW
(400 xt!06Btu/hr)
Pulverized Coal
Coal Sorbent
type type
Western
0.6% S Lime
Eastern Lime
2.3% S
Western Sodium
0.6% S
Lime
Western Sodium
0.6% S
Eastern Lime
2.3% S
Western Lime
0.6% S
S02
Removal
level
75
75
90
75
50
90
75
50
75
70
70
Capital
Cost
Direct Total
capital capital
costs investment
(103$) (103$)
450
582
845
834
826
865
850
839
917
1573
1501
828
1035
1450
1420
1401
1460
1431
1408
1555
2682
2503
% Increase
over un-
Annualized costs controlled
(ioj$)
432
520
695
644
607
630
592
565
718
1242
946
($/J/S)
0.025
0.024
0.016
0.015
0.014
0.014
0.013
0.013
0.012
0.011
0.008
($/MBtu/hr)
7200
6933
4633
4293
4047
4200
3947
3767
3590
3105
2365
boilers
28
22
21
19
20
19
18
16
16
12
-------
TABLE 4.2-5. WELLMAN-LORD PROCESS COST SUMMARY
Boiler size
and type
8.8 MW
(30 x ioSBtu/hr)
.P. Underfeed Stoker
i— '
o>
22 MW
(75 x 10GBtu/hr)
Chaingrate Stoker
58.6 MW
(200 x IOBBtu/hr)
Pulverized Coal
Coal
type
Eastern
3.5% S
Western
0.6% S
Eastern
3.5% S
Western
0.6% S
Eastern
3.5% S
Western
0.6% S
SO
removal
level
90
90
90
90
90
90
Direct
capital
costs
(103$)
796
370
1420
732
2573
1354
Total
capital
% Increase
over
investment Annualized costs uncontrolled
(103$)
1539
896
2483
1440
4233
2378
(10*$)
558
385
809
529
1289
771
(S/J/S)
0.063
0.044
0.037
0.024
0.022
0.013
($/MBtu/hr)
18,600
12,833
10,786
7,053
6,445
3,855
boilers
52
35
41
24
29
17
-------
With regard to annualized costs, the relative rankings of the
FGD processes remain the same as with the capital costs for all cases
considered. The sodium throwaway process again emerged as the least
costly alternative. It should be noted, however, that part of the low
costs for this process are attributable to the relatively simple water
treating process consisting of oxidation plus pH neutralization. If a
more elaborate water treating scheme were required, possibly to comply
with zero discharge regulations, process costs would increase accordingly
and could even make this process the most costly alternative.
Figures 4.2-1 and 4.2-2 illustrate this graphically by showing the
variation in total capital investment costs and total annualized costs
with unit size for each process applied to high sulfur coal at 90%
removal. Figure 4.2-1 indicates a capital cost increase of 100-200
percent in each of the processes in going from an 8.8 MW (30xl06 Btu/hr)
to a 58.6 MW (200xl06 Btu/hr)size. Figure 4.2-2 shows the annualized
costs for the processes as a function of unit size. As with the capital
costs, almost a 200 percent increase in annualized costs will occur in
going from 8.8 MW to a 58.6 MW size. These figures indicate that the
relative annual cost of the Wellman-Lord process is less than its capital
costs when compared to the other FGD processes. This is due to the
credits received for the production of sulfur as a by product.
Figures 4.2-3 and 4.2-4 show the variation in total capital costs
and total annualized costs with unit size for the candidate processes
applied to low sulfur coal boilers at 75% removal. One can see from
these plots that the sodium throwaway process again has both the lowest
capital and annualized costs for low sulfur coals requiring low SOz
removal levels. It should be noted here that the capital costs of
the spray drying systems include costs for a baghouse as a particulate
collection device which is an integral part of the spray drying
system. The annual costs of the spray drying systems also include the
4-17
-------
4000
en
!-i
cd
O
rH
W
4J
CO
o
a,
as
o
3000
2000
1000
Double Alkali
Sodium Throv/away
29.2
(100)
58,6
(200)
87.9
(300)
118
(400)
Size in MWt(106 Btu/hr)
Figure 4.2-1. FGB capital costs versus unit size.
(3,5% S coal, 90% removal)
4-18
-------
2000
M
a 1500
O
TJ
CO
4J
O
O
0)
N
ctf
(3
1000
500
Dual Alkali
Sodium
Throwaway
29.3
(100)
58.6
(200)
Size in MW (10C
87.9
(300)
Btu/hr)
118
(400)
Figure 4.2-2.
FGD annualized costs versus unit size,
(3.5% S coal, 90% removal)
4-19
-------
O
ro
O
3000 I—
2000
'Spray Drying (lime)
Spray Drying (Sodium)
CD
O
0
cd
4-1
1 -rl
§* 1000
u
Limestone
Sodium Throwaway
29.3
(100.)
58.6
(200)
87.9
(300)
118
(400)
Size in MW (106Btu/hr)
Figure 4.2-3. FGD capital costs versus unit size.
(0.6% S coal, 75% removal)
4-20
-------
1500
W
O
0
T)
0)
N
•H
I
1000
500
— Spray Drying
Limestone
Sodium Throwaway
29.3
(100)
58.6
(200)
87.9
(300)
118
(400)
Size in MWfc (106Btu/hr)
Figure 4.2-4. FGD annualized costs versus unit size,
(0.6% S coal, 75% removal)
4-21
-------
capital charges associated with the baghouse. However, to be consistent with
the other candidate processes which were designed solely for S02 control,
disposal costs were only charged for that portion of the solid waste stream
associated with SOa removal.
Figures 4.2-5 and 4.2-6 illustrate the changes in capital and annualized
costs as a function of coal sulfur content. Figure 4.2-5 illustrates the
capital cost impacts and indicates that for a given FGD process size and SOa
removal level, the capital cost of the FGD processes will increase about 50
percent in going from a 0.6 to 3.5 percent sulfur coal. Also, with regard to
capital costs, the relative ranking of the processes did not change with an
increase in the coal sulfur content. Figure 4.2-6 illustrates the annualized
cost impacts and indicates that annualized costs for the FGD processes will
increase almost 100 percent in going from a 0.6 to a 3.5 percent sulfur coal.
Figures 4.2-7 and 4.2-8 illustrate the change in capital and annualized
costs as a function of SOa removal for the limestone and sodium throwaway
processes applied to the 58.6 MW boiler burning 3.5 percent sulfur coal.
Since the double-alkali and Wellman-Lord processes were only considered for
the 90 percent removal cases, their cost variation with SOz removal could not
be assessed. Capital costs of the limestone and sodium processes are shown to
increase about 5-10 percent in going from an SOa removal of 75 to 90 percent.
This low variation is due to the fact that no flue gas by-pass was assumed
for the calculations, and consequently, the gas handling costs remain constant
for a given unit size. Additional cost sensitivity cases will be developed
for flue gas by-pass operations and will be reported in a separate section
of this report. Figure 4.2-8 shows that the annualized costs for both
these processes increase about 10 percent in going from 75 to 90 percent
removal.
4-22
-------
5000
4000
3000
w
M
cd
o
"O
w
o
cd
•H 2000
ex
td
o
1000
Wellman-Lord
Limestone
Double Alkali
Sodium Throwaway
123
Coal Sulfur Content (%)
Figure 4.2-5. FGD capital costs versus coal sulfur content
(58.6 MW, 90% Removal)
4-23
-------
1500 r
1250 -
cfl
I-l
H
O
CD
ra 1000
O
til
N
cd
750
500
Wellman-Lord
Limestone
Double "Alkali
Sodium Throwaway
123
Coal Sulfur Content (%)
Figure 4.2-6.
FGD annualized costs versus coal sulfur content.
(58.6 MW, 90% removal)
4-24
-------
2000 _
1500
M
4J
w
o
0
r-l
cd
o
1000
.500
Limestone
Sodium Throwaway
70 80 90
Percent SO2 Removal
100
Figure 4.2-7. FGD capital costs versus S02 removal,
(58.6 MWfc 3.5% S coal)
4-25
-------
1500
CO
t-t
cd
1250
03
O
O
0)
N
-------
Figures 4.2-9 and 4.2-10 illustrate the change in capital and
annualized costs with SO removal for the limestone, sodium throwaway,
and spray drying processes applied to the 44 MW boiler burning 0.6
percent sulfur coal. One can see from these plots the capital intensive-
ness of the baghouse for the spray drying processes. The limestone and
sodium throwaway process capital costs do not include charges for
particulate control. With regard to annualized operating costs,
however, Figure 4.2-10 illustrates that the relative annualized cost
difference between the spray drying process and the sodium and limestone
processes is not as great. This is due to the lower disposal costs
associated with handling a dry waste product. This figure also illustrates
the relative costs of lime and sodium as sorbents for use in the spray
drying process.
Cost Effectiveness
The cost effectiveness of the various FGD processes was also determined
as part of this study. Cost effectiveness was defined as dollars per kilogram
of removed SOa ($/kg S02) and was calculated by dividing the annualized process
costs by the kilograms of SOz removal in a year assuming a 60 percent load
factor. Results of these calculations are shown graphically in Figures 4.2-11
through 4.2-15.
Results of these figures show that both coal sulfur content and process
size significantly affect the cost effectiveness of an FGD process. For a
given size system, cost effectiveness increases with an increasing coal sulfur
content. For a fixed coal sulfur content, cost effectivenss increases with
increasing process size. Consequently, the most cost effective systems are
those designed for large boilers burning high sulfur coal, and the least cost
effective systems are those designed for the small boilers burning low sulfur
coal. All FGD systems examined here are most cost effective at the stringent
level of control, 90 percent SO2 removal.
4-27
-------
1500
Spray Drying
CSodium and Lime)
Limestone
1000
CO
M
cd
Sodium Throwaway
O
r-l
CO
O
O
500
•H
PJ
I
0 50
60 70 80
Percent S02 Removal
90
100
Figure 4.2-9. FGD Capital Costs versus S02 removal.
(44 MW , 0.6% S coal)
4-28
-------
1000
w
)-i
nj
O
13
CO
4J
tn
O
0)
N
•H
I
500
Spray Drying (Sodium)
Spray Drying (Lime)
Limestone
••Sodium Throwaway
0 50
60 70 80
Percent S02 Removal
90
IDTT
Figure 4.2-10. FGD Annualized Costs versus S02 removal.
(44 MW , 0.6% S coal)
4-29
-------
D
>
O
O
CO
bO
to
03
OJ
§
4-1
O
-------
5 -
•n
0)
1
<&
CN4
o
60
to
to
•H
4-1
O
0)
U-l
m
w
4-1
03
O
@ 90% S02 removal
J.
_L
_L
29.3 58.6
(100) (200)
Size in MW (10
87.9 118
(400) (400)
6 Btu/hr)
Figure 4.2-12. Sodium throwaway process cost effectiveness
4-31
-------
5 -
13
0)
(N
o
00
to
CO
0)
c
4J
0
0)
en
o
o
@ 90% SO2 Removal
0.6% Sulfur Coal
2.3% Sulfur Coal
3.5% Sulfur Coal
29.3
(10)
58.6
(200)
87.9
(30)
118
(400)
Size in MW (106 Btu/hr)
Figure 4.2-13. Double alkali process cost effectiveness.
4-32
-------
Lime Sorbent @ 75% S02 removal
"d
0)
i
0)
O
W
60
cn
CO
fi
a)
o
cu
0.6% Sulfur Coal
to
O
u
2.3% Sulfur Coal
29.3
(100)
58.6
(200)
87.9
(300)
118
(400)
Size in MW (10sBtu/hr)
Figure 4.2-14. Spray Drying Process Cost Effectiveness,
4-33
-------
6r-
@ 90% S02 Removal
0)
o
cvl
O
CQ
CO
01
O
cu
W
4-1
CQ
O
0.6% Sulfur Coal
3.5% Sulfur Coal
(50)
(100)
(150)
58.6
(200)
Size in MWt(106 Btu/hr)
Figure 4.2-15. Wellman-Lord process cost effectiveness.
4-34
-------
Retrofit Applications
Application of FGD systems to existing plants usually entails higher
costs than those for application to similar new plants. Whereas an FGD system
for a new plant can be incorporated into the overall plant design, retrofitting
requires that the system be adapted to the given plant configurations; space
is often limited, and ongoing plant operations further constrain installation
of the system.
Configuration of equipment in the plant sometimes governs the location of
the FGD system. For instance, if the boiler stack is on the roof of the
boiler house, as it is in many older plants, the FGD system may have to be
placed at ground level; this placement could entail long ducting runs from the
absorber to the stack or could require construction of a new stack. At some
plants the stack is situated directly adjacent to the boiler house or particu-
late control device, a placement that often necessitates locating the FGD
system at some distance. At some plants, especially those located in urban
areas, the available space at ground level is inadequate to accommodate the
entire FGD system. In such cases either the FGD scrubber units must be stacked,
one on top of the other, or additionl land must be acquired adjacent to the
plant property.
Terrain of the power plant site also affects the capital cost of the FGD
system by sitework and structural requirements. Hilly terrain requires consid-
erable grading and filling to prepare the site for construction of foundations
and possible additional structural components. Subsurface conditions can
necessitate piling to provide adequate support for the concrete foundations of
the FGD system.
Other capital cost components that can be increased because of space
restrictions are construction labor and expenses, interest charges during
construction (because of longer construction periods), contractor fees and
expenses, and allowances for shakedown. Table 4.2-6 summarizes the capital
cost impacts of several retrofit requirements. 12
4-35
-------
TABLE 4.2-6. TYPICAL INCREASE IN CAPITAL COSTS WITH VARIOUS
RETROFIT REQUIREMENTS 13
Retrofit requirements
Long duct runs 4 - 7
Tight space 1-18
Delayed construction (1 year delay) 5-15
Hilly terrain 0-10
New stack 6-20
16 - 70
4.2.2 Example Calculation
The sodium throwaway FGD process was selected for the sample calculation
due to its present predominance in industrial boiler installations. The
approach used to estimate FGD capital investment costs was as follows. First,
a series of material and energy balance calculations were performed to establish
process flow rates and energy requirements. Second, each of the FGD processes
were divided into a number of process modules which represented separate
processing areas. Equipment sizes were then developed for each process module
using detailed engineering designs based on the results of the material and
energy balances. Finally, capital cost estimates were prepared by contacting
process equipment vendors for price quotations in the size range for the
standard industrial boilers of this study.
All of the capital costs for each process area are reported in the form
of direct capital costs which include all materials and labor installation
costs. Particulate control equipment costs were not included in this study
except for the spray drying process, and there a baghouse was used for the
particulate control device. Table 4.2-7 illustrates the process operating
conditions for the example case design, one of the larger FGD systems considered
in this study.
4-36
-------
TABLE 4.2-7 PROCESS OPERATING CONDITIONS
Operating condition Value
Boiler combustion rate, MW 58.6
(MBtu/hr) (200)
Coal HHV, Btu/lb 11,800
Coal sulfur content, wt% 3.5
Coal ash content, wt% 10.58
Total excess air, % 30
S02 feed rate, Ib/hr 1126.5
S02 recovery, % 90
Flue gas rate; acfm, T°F/psia 74,800; 400/14.7
Year of economics mid-1978
Raw Material Handling
Raw material handling costs will be a function of the NazC03 makeup rate.
The material balance calculations presented in Appendix A showed a NaC03 feed
rate of 1900 Ib/hr for the 58.6 MW industrial boiler. The raw material hand-
ling costs include costs for the following equipment items:
- pneumatic conveyor (sorbent unloading to storage)
- storage silo (2-week capacity to avert shortages in case of
delays in sorbent delivery due to bad weather, strikes, etc.)
- screw feeder (sorbent feed to mixing tank)
- mixing tank (5 minute residence time)
mixer (agitation for dissolution)
2-feed prep pumps (pump feed to circulation tank)
The F.O.B. costs for each of these items were obtained via quotations from
process equipment vendors and are presented in Table 4.2-8.
4-37
-------
TABLE 4.2-8. RAW MATERIAL HANDLING CAPITAL COSTS
Item
Capacity
Mid-1978
F.O.B. Cost
Pneumatic conveyor
Silo
Screw feeder
Mixing tank
Mixer
Feed prep pumps (2)
2 +ph (smallest available) $ 3,400
(1900 rr)(24 ^7) (14 day) = 638,000 Ib 87,000
1,200
1,600
930
11,000
2 +ph (smallest available)
(123 GPM)(5 min) = 615 gal
615 gal
123 GPM ea.
$105,130
SO Scrubbing
The F.O.B. costs (acquired from equipment vendors) for the equipment
items in the SO scrubbing area are shown below in Table 4.2-9. The capacities
are obtained from the material balance presented in Appendix A.
TABLE 4.2-9. S02 SCRUBBING CAPITAL COSTS
Item
Capacity
Mid-1978
F.O.B. Cost
Absorber (scrubber)
Circulation tank
Mixer
Circulation pumps (2)
74,800 acfm
(748 + 123 GPM)(5 min) = 4360 gal
4360 gal
779 GPM ea.
$72,000
5,700
3,100
16,000
$96.800
4-38
-------
Fans
The F.O.B. costs for the fan (including motor) are shown below in Table
4.2-10.
TABLE 4.2-10 FAN CAPITAL COSTS
Item
Fan (including motor)
Capacity
74,800 acfm
Mid-1978 F.O.B. Cost
$33,000
Wastewater Pumps
The F.O.B. costs for the wastewater pumps (for pumping the wastewater to
the boilers' wastewater treating facility) are shown below in Table 4.2-11.
TABLE 4.2-11. WASTEWATER PUMPS CAPITAL COSTS
Item
Capacity
Mid-1978 F.O.B. Cost
Wastewater Pumps (2)
92 GPM ea.
10,000
The F.O.B. costs for each module (process area) are multiplied by an
installation factor to obtain the installed costs (including transporation) and
are summarized below in Table 4.2-12.
TABLE 4.2-12. INSTALLED COSTS
Process Module
Raw Material Handling
SO 2 Scrubbing
Fans
Wastewater Pumps
Total
F.O.B. Cost
$105,130
96,800
33,000
10,400
Installation
Factor l "*
1.69
3.11
2.29
2.49
Installed
Cost (103$)
178
301
76
26
Subtotal 581
4-39
-------
Costs for utilities and services were estimated at 6% of the installed
1 5
costs of all process equipment.
Utilities and services cost = .06 ($581,000)
= $35,000
Installed Cost Summary
For the 58.6 MWt Case:
Process Area Costs (10 $)
Raw Material Handling 178
SO Scrubbing 301
Fans 76
Wastewater Pumps 26
Utilities and Services 35
Total 616
All of the indirect capital costs are based upon a percentage of the
direct costs as discussed in Section 4.1. Indirect costs include such items
as engineering, construction expenses, contractor fees, start-up expenses, and
contingencies. Table 4.2-13 illustrates how these indirect capital costs were
applied to determine the total turnkey costs (TTC) for the sodium throwaway
process.
Land and working capital costs are added to TTC to arrive at Total Capital
Investment (TCI). Land is estimated based upon percentages of process costs
derived from TVA. Working capital is estimated at 25 percent of total direct
operating costs as discussed in Section 4.1 and shown in Table 4.2-13.
Annual process costs for the sodium throwaway process are shown in Table
4.2-14. These costs are based upon a 60 percent stream factor 16(5256 hrs/yr)
@ 100 percent flow. The material balance calculations presented in Appendix A
and the energy balance calculations presented in Section 5 provided the basis
of the fuel and chemical costs. Prices for labor, raw materials, and utilities
were established by PEDCo for consistency17 and are shown on Table 4.1-2 and
on Table 4.2-14.
4-40
-------
TABLE 4.2-13
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Thrnwaway
Boiler Capacity: _58.6 MWt (200 MBtu/hrj
Coal Feedstock: J_, 5% S Eastern
S02 Control Level: 90%
Item
Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling
S02 Scrubbing
Fans
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services
Total Direct Costs (TDC)
Indirect Capital Costs
a
Engineering
Construction and Field Expenses (0.1 TDC)
Construction Fees (0.1 TDC)
Start-up (0.02 TDC)
Performance Test (0.01 TDC)
Total Indirect Cost (TIC)
Contingencies (0.2 (TDC+TIC))
TOTAL TURNKEY COSTS (TTC)
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)
TOTAL CAPITAL INVESTMENT (TCI)
178
301
76
26
39
62
62
62
12
615
204
164
983
0.8
167
1151
a. Engineering Costs
0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOj removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
S0? removal.
b. Reference: 4
c. From Annual Cost Table
4-41
-------
TABLE 4.2-14.
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coa'l Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
58.6
(200 MBtu/hr)
3.5% S Eastern
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials ( . OA TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (. 2 6x( 1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
kW
864
_kg/hr
_kg/hr
_kg/hr
105
21
27
27
20
GJ/hr
27
.9
m3/hr
6
GJ/hr
20
.9
m3/hr
11
kg/hr
450
38
47
$/106 Btu
I/kg SO;
667
85
196
947
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) C&A, property taxes and insurance. Federal and State income
taxes are not included.
4-42
-------
Overhead costs were estimated as percentages of direct labor and main-
tenance costs as discussed in Section 4.I.18 Both payroll and plant overhead
expenses were estimated as shown in Table 4.2-14.
Capital charges of 17 percent of the Total Capital Investment TCI are
based upon a 10 percent return on investment, 15 yr plant life, straight-line
depreciation, and 4 percent G&A expenses. The sum of the direct operating
costs, overhead costs, and capital charges is the total annualized cost. In
the cases presented, all costs, both capital and annual, are expressed in
thousands of dollars. Cost tables for the other cases considered are found in
Appendix B.
4.3 COSTS TO CONTROL OIL-FIRED BOILERS
Both capital and annualized costs of controlling SO2 emissions from the
standard residual oil-fired boiler were estimated using the same cost esti-
mating procedure as described in Section 4.1. Costs were developed for the
limestone process in order to determine the impact of treating flue gas from a
residual oil-fired boiler. For comparison, costs were also developed for a
coal-fired system of the same size treating flue gas from a 3.5 percent sulfur
coal. Table 4.3-1 summarizes the results of these calculations.
TABLE 4.3-1. COST COMPARISON OF OIL- AND COAL-FIRED LIMESTONE FGD COSTS
Boiler size
44 MWt
(150 x 105Btu/hr)
Percent Capital
SOz investment costs
Fuel type removal (10 3 $)
Eastern Coal
3.5% S
Res id Oil
3.0%'S
90
75
, i
90
75
1385
1270
1017
942
Total annualized
costs
(103 $)
974
865
742
683
4-43
-------
The major differences between the oil- and coal-fired systems are that
the oil-fired system processes less gas and removes a small quantity of S02
than the coal-fired system. This is because the oil-fired boiler operates
with 15 percent excess air versus 50 percent for the standard spreader stoker
boiler, and the uncontrolled S02 emissions from the oil-fired boiler are 471
kg/hr versus 846 kg/hr for the coal-fired system. These differences result in
the oil-fired system having smaller raw material feed equipment, smaller gas
handling equipment, and smaller solids handling equipment.
Figure 4.3-1 illustrates the difference in the oil- and coal-fired costs
graphically. Costs of the coal-fired system are about one-third higher for a
given SOa removal. Similar cost differentials should exist for the other
processes.
4-44
-------
200Q
CO
M
d
1500
M
O
O
n)
4J
•H
U
1000
500
Coal-Fired
Oil-Fired
70 80 90
Percent SO2 Removal
100
Figure 4,3-1. Comparison of oil- and coal-fired FGD costs.
(Limestone Process, 44 MW )
4-45
-------
REFERENCES
1. McGlamery, G.G., et al. Detailed Cost Estimates for Advanced Effluent
Desulfurization Processes. Final Report. EPA-600/2-75-006. Muscle
Shoals, AL. TVA. January 1975. pp. 4, 100-113.
2. Weismantel, Guy E. "Limestone and Magnesium: A New S02 Control Team."
Chem. Eng. 1978 (11 Sept), p. 111.
3. Tuttle, J., et al. EPA Industrial Boiler FGD Survey - Second Quarter
1978. EPA Contract No. 68-02-2608, Task 36, EPA-600/7-78-0526. Cincinnati,
Ohio. PEDCo Environmental, Inc. July 1978.
4. Noe, David N. Memorandum to James C. Dickeman, PEDCo Environmental, Inc.
Cincinnati, Ohio. June 13, 1979.
5. Ponder, T.C = Memorandum to John Protapas, OAQPS. PEDCo Environmental,
Inc. Dallas, Texas. May 9, 1979.
6. PEDCo Environmental Specialists, Inc. "Emission Control System Economics."
Section 3.0. Cincinnati, OH. October 1978.
7. Ponder, Thomas C., et al. Simplified Procedures for Estimating Flue Gas
Desulfurization System Costs. EPA 600/2-76-150. EPA Contract No. 68-
02-1321, Task 12. Cincinnati, OH. PEDCo Environmental Specialists,
Inc. June 1976.
8. McGlamery, et al., op. cit., p. 239.
9. PEDCo Environmental Specialists, "Emission Control...". op. cit.
10. Ponder, T.C., "Memorandum...", op. cit.
11. PEDCo Environmental Specialists, "Emission Control...", op. cit.
12. Ponder, Thomas, "Simplified Procedures...". op. cit.
13. Ibid.
14. Guthrie, K.M., "Capital Cost Estimating", Chemical Engineering, March 24,
1969, pp. 114-142.
15. McGlamery, et al., op. cit.
4-46
-------
16. PEDCo Environmental Specialists, Inc. "Cost of New Boilers." Section
4.0. Cincinnati, OH. October 1978.
17. PEDCo Environmental Specialists, Inc., "Emissions Control..."
18. PEDCo Environmental Specialists, Inc., "Cost of..."
4-47
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SECTION 5
ENERGY IMPACT OF CANDIDATES FOR
BEST EMISSION CONTROL SYSTEMS
5.1 INTRODUCTION
This section addresses the energy requirements associated with the con-
trol of S02 emissions from small industrial boilers. Five FGD systems were
selected as being the best candidate S02 control systems and the energy
requirements associated with each are shown in Table 5.1-1. These energy
requirements compare quite well with FGD energy requirements presented in a
recent study prepared by Rubin1 and with energy requirements reported by
FGD system operators at a recent symposium on FGD energy requirements held
Q
at Lehigh University. Both of these sources have reported energy require-
ments of limestone systems with stack gas reheat to vary from about 3 to 3.5
percent of the net heat input to the boiler.
TABLE 5.1-1. RANGE OF FGD SYSTEM ENERGY REQUIREMENTS*
Energy Requirement Energy Requirement
S02 control method Not Including Reheat Including Reheat
Limestone
Sodium Scrubbing
Double Alkali
Spray Drying
Wellman-Lord
0.9-1.8
0.5-0.8
0.5-0.6
0.5-0.8
1.6-6.2
2.4-3.5
2.0-2.6
2.0-2.3
0.5-0.8
3.2-7.9
* Energy Requirements expressed as percent of net heat input to boiler.
The variations in energy requirements for these processes are due to diff-
erent levels of sulfur in the coal, different levels of S02 control, and to
a smaller extent, unit size.
5-1
-------
There are four major energy consumption areas that are used for all of
these processes:
• Raw materials handling and fuel preparation
S02 scrubbing
Fans
Utilities and services
A fifth area of energy consumption, stack gas reheat, was not considered
as part of the industrial boiler FGD systems' energy requirements since
the majority of industrial boiler FGD applications do not reheat the exhaust
gas. To illustrate the impact of reheat, Table 5.1-1 was prepared assuming
reheat of the stack gases to 175°F. This table shows that stack gas reheat
creates an energy penalty of 1.5-1.7 percent of the net heat input to the
boiler and significantly increases the energy consumption of all wet
scrubbing processes. The spray drying process exhausts its stack gas at
175°F and, consequently, no reheat would be required.
In addition to the above energy consumption areas, the sodium scrubbing
process has energy requirements for wastewater disposal and the Wellman-Lord
process requires low pressure steam as part of its regeneration operations and
methane for sulfur production. No energy impact for disposal was charged to the
three other throwaway processes as the waste is a solid and was assumed to be
hauled off-site and disposed of by a contractor. Each of the above process
areas will be discussed briefly below as they relate to FGD energy requirements
The raw material handling and feed preparation operation is designed
to receive, store and prepare makeup reagents for the FGD processes. This
requires storage silos, conveyors, mixers, slurry or solution preparation
tanks and pumps. The double-alkali process uses lime as the regenerant
or precipitation reagent. Facilities for calcining limestone to produce
the lime were not included in the design basis. If these facilities were
included, the energy requirements of the raw material handling and feed
preparation operations would increase and would impact the total energy
5-2
-------
requirements of the double-alkali process with an approximate 25 percent
increase.
The major energy requirement associated with the SC>2 scrubbing process
area is electricity to run the flue gas fans and process recirculation
PUTTIPR - Liquid circulation (L/G) rates for these energy consumption calculations
were based on information provided in Section 2 of this report for each of the
candidate processes. The pressure drop to be overcome for each case was
estimated by using the following empirical relationship developed by TVA.1*
AP = 1.68 + 7.17 x Id"3 (Bed Height) x L/A x V2 (5-1)
where:
AP = Pressure drop (in H20)
Bed Height = Total height of TCA spheres (in)
L = Liquid circulation rate (gpm)
A = Scrubber area (ft2)
V = Gas velocity (ft/sec)
To evaluate the energy requirement for installations using reheat,
stack gas reheat was assumed to occur in indirect steam reheaters. The
minimum stack exit temperature required to prevent sulfuric acid mist
formation and provide plume buoyancy is not well defined. Common practice,
and the design basis for these additional calculations is to reheat the
scrubbed gas to achieve an exit temperature of 353°K (175°F) which requires
^50°F of reheat.5
Utilities and services such as instrument air, lighting, heating,
cooling, etc. are also required for each FGD process. For this analysis the
amount of energy consumed for utilities and services was estimated from
process energy requirements developed for utility boilers.
The limestone, double-alkali, sodium scrubbing, and spray drying FGD
processes produce waste materials for disposal. Ponding is the normal
5-3
-------
on-site method of FGD waste disposal. The processes with a solid phase
waste material were assumed to use a contractor for waste removal and
disposal so no energy penalty was charged. It was assumed that the liquid
wastes from the sodium scrubbing process would be pumped one mile to a
central on-site wastewater treatment plant.
The Wellman-Lord process produces a concentrated S02 stream from regen-
eration of the S02 scrubbing liquor. A set of evaporator/crystallizers are
used to regenerate the S02 scrubbing liquor. This equipment uses low pres-
sure steam as the energy source. Because the regeneration operation produces
a concentrated S02 stream, conversion to elemental sulfur or sulfuric acid
is possible. For this study, a proprietary process of the Allied Chemical
Company which uses methane as a reductant was selected as the basis for con-
verting the S02 to elemental sulfur.6
5.2 ENERGY IMPACT OF CONTROLS FOR COAL-FIRED BOILERS
A series of material balance calculations were performed for each of the
FGD systems to identify stream flow rates, equipment sizes, effluent streams,
and raw material requirements. Results of these calculations are presented
in Appendix A. Process energy requirements were then calculated using the
material balance results and the assumptions and design bases listed in
Table 5.2-1. Process energy requirements calculated for each of the FGD
systems under consideration are shown in Tables 5.2-2 through 5.2-6.
Results of these calculations indicate that the process energy penalties
range from about 0.5 to 2 percent of the gross heat input to the boiler for
the three throwaway processes and from about 2 to 6 percent for the Wellman-
Lord process. The larger energy consumption for the Wellman-Lord process
is due to the steam and methane requirements for the regeneration and S02
reduction portions of the process. Energy penalties associated with the
SOa reduction section of the Wellman-Lord process were estimated from the
results of a previous study.7
5-4
-------
TABLE 5.2-1. MATERIAL AND ENERGY BALANCE ASSUMPTIONS AND DESIGN BASES
en
C71
Process Parameters
L/G Jl/ms,
(gal/103 acf)1
Particulate Removal2
Stoichiometry
(moles sorbent/mole sorbed S02)
Gas Pressure Drop
Pa (in H20)3
Pump Discharge Pressure
Pa (psi)2"
Pumping Height
M (ft)1
(A) Limestone System L/G and AP
Limestone
A '(A)
Sodium
FGD PROCESSES
Spray Dryer
Throwaway Wellman-Lord Double Alkali
1335 (10) 1335
(10) 1335 (10)
Sodium
40 (0.3)
Lime
40 (0.3)
Baghouse Downstream
1.2
A (A)
5227 (15)
6 (20)
99 Percent Upstream
B
100 (8) 100
3484 (10) 3484
6 (20) 6
of FGD System ,
1.0 1.0
(8) 100 (8)
(10) 3484 (10)
(20) 6 (20)
(B) Sodium Throwaway
of Spray
C
75 (6)
3484 (10)
6 (20)
Stoichiometry
Dryer
C
75 (6)
3484 (10)
6 (20)
Percent SOi removal
Coal type 90
Eastern (3.5% S)
L/G 10.7 (80)
AP 214 (17)
Western (0.6% S)
L/C 9.3 (70)
AP 189 (15)
85
8.0 (60)
176 (14)
7.3 (55)
164 (13)
75
5.3 (40)
151 (12)
5.3 (40)
151 (12)
90% Removal
85% Removal
75% Removal
- 1.05
- 0.95
- 0.85
(C) Spray Dryer Stoichiometry
90% Removal
75% Removal
50% Removal
Sodium* Lime5
1.1 2.0
- 0.8 1.2
0.5 0.65
Baaed on process data presented in Section 2.
2Radian assumption.
'Based on TVA empirical relationship.
^Reference 8.
'Reference 9.
-------
TABLE 5.2-2. ENERGY REQUIREMENTS FOR THE LIMESTONE FGD PROCESS
Energy consumed in
Boiler size and type
8.8 KW (30xl05 Btu/hr)
Underfeed stoker
•
22 >fWc(75x!06 Btu/hr)
Chaingrate stoker
44 MWt (ISOxlO5 Btu/hr)
Spreader stocker
58.6 MWt (200xl06 Btu/hr)
Publverized coal
Ui
1
8.8 MW (30xl06 Btu/hr)
Underfeed stoker
22 MK,. (75xl05 Btu/hr)
Underfeed stoker
44 MW,. (150xlOs Btu/hr)
Spreader stoker
58.6 MWt (200x10* Btu/hr
Pulverized coal
Coal type
Eastern
3.5% sulfur
Eastern
3.5% sulfur
Eastern
3.5% sulfur
Eastern
3.5% sulfur
Western
.6% sulfur
Western
.6% sulfur
Western
.67. sulfur
Western
.6% sulfur
S02
removal
level
90
85
75
90
75
90
75
90
85
75
90
85
75
90
75
90
75
90
85
75
Raw material
handling and
preparation
17
16
14
42
35
86.
71,
114.
108.
95.
3,
3.
3.
9.
7.
18.
15.
23.
22.
20.
.0
.1
.0
.7
.7
.1
.4
.2
.3
,4
i 5
,4
,0
.0
6
2
,2
9
,7
.2
Liquid
54,
41
27
136
69
274
140
317
240.
162.
45.
36.
27.
115
66.
230.
132.
267.
210.
154.
.8
.1
.8
.7
.7
.6
.0
.4
.0
.2
.7
.0
.8
.4
.5
.1
.9
.6
.8
.0
control device (kW )
Fan
81.
67.
57.
203,
143,
408,
288
471,
388.
333
72.
63,
58,
183,
146.
364.
292.
423.
367.
339.
.3
.0
.3
.7
.8
.7
.5
.7
.3
.0
.7
.0
.2
.0
,5
.8
.0
.7
.2
.0
Utilities
&
services
1
1
1
3
3
7
7
8
8
8
1
1
1
3
3
7
7
8
8
a
.5
.5
.5
.7
.7
.4
.4
.6
.6
.6
, 5
.5
.5
.8
.8
.6
.6
.8
.8
.8
Total
kWt 106 Btu/hr
154.6
125.7
100.6
386.8
252.9
776.8
507.3
911.9
745.2
599.2
123.4
103.9
90.5
311.2
224.4
620.7
447.7
724.0
609.5
522.0
0
0
0
1
2
1
3
2
2
0
0
0
1
0
2
1
2
2
1
.53
.43
.34
.32
.86
.65
.73
.11
.54
.05
.42
.35
.31
.06
.77
.12
.53
.47
.08
.78
Percent increase
in energy use
over uncontrolled
boiler
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
0
.8
.4
.1
.8
.2
.8
.2 .
.6
.3
.0
.4
.2
.0
.4
.0
.4
.0
.2
.0
.9
-------
TABLE 5.2-3. ENERGY REQUIREMENTS FOR THE SODIUM THROWAWAY FGD PROCESS
Energy consumed in control device (kW ) F-rcent
Boiler size and type
8.8 MWt(30xl08 Btu/hr)
Underfeed stoker
22 MWt (75x10 6 Btu/hr)
Chaingrate stoker
44 MW (150xlOs Btu/hr)
Spreader stoker
58.6 MW (200xlOs Btu/hr)
Pulverised coal
118 MH,. (400x10' Btu/hr)
Ul
1 22 MW,. (75xlOs Btu/hr)
-J f.hainorate stoker
SO 2
removal
Coal type level
Eastern
3.5 sulfur
Eastern
3.57. sulfur
Eastern
3.5% sulfur
Eastern
3.5% sulfur
Eastern
Medium sulfur
2.3X sulfur
90
85
75
90
75
90
75
90
85
75
90
90
Raw material
handling and
preparation Liquid
1.4 6.7
1.3 6.7
1.2 0.7
3.5 17.3
2.9 16.7
7.0 34.6
5.8 33.4
9.3 40.7
8.8 40.1
7.7 39.5
18.6 80.9
2.0 16.2
irtTTTtTes
& Total
Fan
38.2
38.2
38.2
95.8
95.8
192.3
192.3
222.0
222.0
222.0
433.0
95.7
Disposal
23
21.
19
60
49
.3
.7
.5
.3
.0
121.0
97.
160.
145.
130.
4
4
.4
.2
322.7
34.
9
services kW,. 10° Btu/hr
1.6 71.2
1.6 69.5
1.6 67.2
4.0 180.9
4.0 168.4
8.0 362.9
8.0 336.9
9.3 441.7
9.3 425.6
9.3 408 . 7
17.7 872.9
3.9 152.7
0.
0.
0.
0.
0.
1.
1.
1.
1.
1.
2.
.24
.24
.23
.62
.57
.24
15
51
45
39
98
0.52
in energy use
over uncontrolled
boiler
0
0
0
0
0
0
0
0
0
0.
0
0
.8
.8
.8
.8
.8
.8
.8
.3
.7
.7
.7
.7
118 MWt(400xlO° Btu/hr) Medium sulfur
Pulverized coal 2.3% sulfur
75.3 438.1 138.4
8.8 MWe(30xl05 Btu/hr) Western
Underfeed stoker .6% sulfur
90
85
75
0.3
0.3
0.2
5.8 38.7
5.8 38.7
5.8 38.7
5.1
4.5
4.2
1.6
1.6
1.6
51.5 0.18
50 9 0.17
50.5 0.17
0.6
0.6
0.6
22 MWt(75xlOE Btu/hr) Western
Chaingrate stoker .6% sulfur
90
75
14.3 97.7 12.7
14.3 97.7 10.3
4.1
4.1
129.5 0.44
127.0 0.43
0.6
0.6
44 MWt(150xlO£ Btu/hr) Western
Spreader stoker .6% sulfur
90
75
1.5
1.2
28-7 194.7
28.7 194.7
25.4
20.6
8.2
8.2
258,5 0.88
253.4 0.86
0.6
0.6
58.6 MWt(200xlO€ Btu/hr) Western
Pulverized coal .6% sulfur
90
85
75
1.9
1.9
1.7
34.0 226.0
33.7 226.0
33.7 226.0
34.0
30.7
27.4
9.5
9.5
9.5
305.4 1.04
301.8 1.03
298.3 1.02
0.5
0.5
0.5
118 MWt(400xlOs Btu/hr) Western
Pulverized coal .6% sulfur
67.2 445.6
-------
TABLE 5.2-4. ENERGY REQUIREMENTS FOE THE DOUBLE-ALKALI PROCESS
Boiler size and type
8.8 HW (30xlOG Btu/hr)
Underfeed stoker
22 MKt(75xl05 Btu/hr)
Chaingrate stoker
58.6 MKt (200x10* Btu/hr)
Pulverized coal
118 MKt(400xl06 Btu/hr)
Pulverized coal
Coal type
Eastern
3.5% sulfur
Western
.67. sulfur
Eastern
3.5% sulfur
Medium sulfur
2.':% sulfur
Wes-tern
.6% sulfur
Eastern
3.5% sulfur
Western
.6% sulfur
Eastern
3.5% sulfur
Medium sulfur
2.3% sulfur
Western
. 6% sulfur
S02
removal
level
90
90
90
90
90
90
90
90
90
90
Raw material
handling and
preparation
1.9
0.4
4.8
2.7
0.9
12.7
2.4
25.5
14-2
4.8
Energy consumed in
Liquid Fan
10.4 38.2
6.2 38.7
26.1 94.3
20.8 95.7
15.5 97.7
63.2 222.0
36.9 226.0
129.4 433.0
100.6 438.1
73.3 446.6
control device (kW )
Utilities
& .___Tqtal
services kW 10sBtu/hr
1.6 52.1 0.18
1.6 46.9 0.16
3.8 129.5 0.44
3.9 123-1 0.42
4-1 118.2 0.40
9.3 307.2 1.05
9.5 274-8 0.94
17.7 605.6 2.07
17.9 570.8 1.95
18.4 543.1 1.85
Percent increase
in every use over
uncontrolled boiler
0.6
0.5
0.6
0.6
0.5
0.5
0,5
0.5
0.5
0.5
-------
TABLE 5.2-5. SPRAY DRYING ENERGY REQUIREMENTS
Energy consumed in control device.
Coal
Boiler size sulfur
MW (105 Btu/hr) content
17.5
U, "
1 44
44
58.6
22
44
44
44
118
118
(60) 0.6
(150) 0.6
(150) 0.6
(150) 0.6
(200) 0.6
(75) 2.3
(150) 0.6
(150) 0.6
(150) 0.6
(400) 0.6
(400) 2.3
SO:
removal
level (A)
75
90
75
50
75
70
90
75
50
70
70
Type
alkali
used
Sodium
Sodium
Sodium
Sodium
Sodium
Lime
Lime
Lime
Lime
Lime
Lime
Raw materials
handling and
preparation
0.6
1.4
1.2
0.8
1.6
2.0
1.8
1.5
1.0
3.8
10.6
Liquid
pumping
0.2
0.5
0.5
0.5
0.6
0.4
0.5
0.5
0.5
1.8
1.9
Fan
55.6
139.7
139.7
139.7
162.8
70.7
139.7
139.7
139.7
330.5
323.7
71.8
122.3
121.3
120.2
132.9
106.0
121.3
120.0
119.1
246.6
257.2
kW
Utilities
an ser ices
3.2
8.2
8.2
8.2
9.7
3.9
8.2
8.2
8.2
18.4
17.9
P(
Total i,
klJt
131.4
272.1
270.9
269.4
307.6
183.0
271.5
269.9
268.5
601.1
611.3
10" Btu/hr ui
0.45
0.93
0.92
0.92
1.05
0.62
0.93
0.92
0.92
2.05
2.09
;rcent increase
i energy use over
.^controlled boiler
0.7
O.b
0.6
0.6
0.5
0.8
0.6
0.6
0.6
0.5
0.5
-------
TABLE 5.2-6. ENERGY REQUIREMENTS FOR THE WELLMAN-LORD PROCESS
Energy consumed in control device (kW )
S02 Raw materials Utilities
removal handling and ^ Process & Total i
Boiler size and type Coal type level preparation Liquid Fan steam Methane services kW 10°Btu/hr u
8.8 MWt(30xl06Btu/hr) Eastern 90 1.4 11.5 35.2 337 155 1.7 541.3 1.85
Underfeed stoker 3.5% sulfur
1 Western 90 .3 6.8 35.7 70 35 1.7 149.5 0.51
M .6% sulfur
O
22 MWt(75xl06Btu/hr) Eastern 90 3.5 31.2 88.5 844 392 4.3 1363.5 4.65
Chaingrate stoker 3.5% sulfur
Western 90 .7 16.3 90.2 179 82 4.4 372.6 1.27
.6% sulfur
58.6 MWt(200xl06Btu/hr) Eastern 90 9.3 82.9 205.5 2220 1048 9.6 3575.3 12.20
Pulverized coal 3.5% sulfur
Western 90 1.9 42.2 208.6 469 219 10.2 950.9 3.25
.6% sulfur
Percent increase
n energy use over
ncontrolled boiler
6.2
1.7
6.2
1.7
6.1
1.6
Includes electricity for operation of SOz scrubbing pumps and pumps for the SOz reduction process area.
-------
Figure 5.2-1 which is based on the calculations developed for this study
illustrates the effect of coal sulfur content on process energy requirements.
This figure shows a larger energy requirement for a limestone system applied
to a 3.5 percent sulfur coal than for a system applied to an 0.6 percent sul-
fur coal. The energy increase, about 12 percent, is primarily due to the
increased energy requirements of the feed preparation area since more alkali
is required for the higher sulfur coal. In addition, higher liquid pumping
rates and increased system pressure drop for the 3.5 percent sulfur coal case
also contribute to the increased energy requirements. The double alkali and
sodium throwaway systems show even less of a variation in energy consumption
with coal sulfur content. The most noticeable change is with the Wellman-
Lord system which is due primarily to increases in the regeneration and S02
reduction sections of the process.
Figures 5.2-2 and 5.2-3 illustrate the effect of boiler size on process
energy requirements. In all cases, the amount of energy consumed by an
FGD process increases with increasing boiler capacity. If the FGD energy
consumption is plotted as a percent of the heat input to the boiler,
however, a different result is shown. Figures 5.2-4 and 5.2-5 illustrate
that except for the large sized boilers, the boiler size has essentially no
effect on the energy penalties when they are expressed as a percent of
boiler heat input. In this figure, the percent energy penalty for each of
these processes is shown to decrease in going to the large 58.6 MW (200 X
106 Btu/hr) size. The reason for the decrease is the relatively lower fan
energy requirements for the larger sized standard boilers which were used
for the basis of this study because the percent excess air decreases from
50 percent down to 30 percent as specified by the standard boiler character-
izations. rtl Figures 5.2-3 and 5.2-5 also illustrate that the energy require-
ments for a spray drying process are essentially the same as those for a
sodium throwaway process when reheat is required.
5-11
-------
a
o
•H
3
to
(3
O
CJ
bO
M
0)
3000 —
2000 _
1000 _
Wellman-Lord
s tone
Sod iuni '1'lirownwny
I)oul)le-Al k.-il i
% Sulfur in Coal
(58.6 MWfc Boiler Size; 905 S02 Removal)
Figure 5.2-1. Energy Consumption versus coal sulfur content.
5-12
-------
C
w
We11man-Lord
3000-
c 2000-
o
c.
E
1000-
Limestone
Sodium
Throwaway
Dual
Alkali
(30)
22
(75)
44
(150)
58.6
(200)
Boiler Size, MK't (] Oc litu/lir )
(90% SO? Kunioval)
Figure 5.Z-2. Energy consumption versus boiler size:
high sulfur eastern coal
5-13
-------
90% S02 Removal
75% SO, Removal
1000—
I
\->
-P-
C
o
•H
4-1
CX
6
o
a
a)
d
w
Wellman-Lord
500—
Spray Drying
Sodium Thrdwaway
Dual Alkali
(30)
22
(75)
(150)
58.6
(200)
118
(400)
Boiler Size, MW (106Btu/hr)
Figure 5.2—3. Energy Consumptxon versus boiler size:
low sulfur western coal.
-------
0
cs
ID
PL,
c.
3
K
C
O
o
c
w
5—
3—
Wellman-Lord
Limestone
Sodium Throwaway
Dual Alkali
8.8 22
(30) (75)
44
(150)
58.6
(200)
Boiler Size, MW (10G BLu/lir )
Figure 5.2-4,
Percent energy consumption versus boiler size:
high sulfur eastern coal
5-15
-------
I
H
2.5-J
~ 2.0.
J! 1.5.
^ 1.0—1
90% SCU Removal
75% S02 Removal
Wellman-Lord
Limestone
Spray Drying
Jjodium Throwaway
Alkali
3 '.3
(30)
2'°
(75)
(150)
5'8.6
(200)
(400)
Boiler Size, MWt (106Btu/hr)
Figure 5.2-5. Percent energy consumption versus boiler size:
low sulfur western coal.
-------
Figures 5.2-6 and 5.2-7 illustrate how energy requirements vary with
removal. These figures show that the energy consumption for the lime-
stone system increases by about 40-50 percent when going from an S02
removal of 75 to 90. This is due primarily to an increased liquid pumping
rate required to achieve the higher removal efficiency and an increased
system pressure drop. Only a very small change in energy consumption (<10
percent) is shown for the sodium throwaway process since its liquid pumping
rate is low and does not vary appreciably for a change in S02 removal. The
change is equally low for the spray drying process as the fan requirements
do not change with removal level. Also, the alkali slurry atomization energy
requirements did not change appreciably due to the small sized equipment
involved. Calculations for the double-alkali and Wellman-Lord processes
were only developed for the 90 percent removal cases. Consequently, the
effects of varying S02 removal levels on process energy requirements are
not addressed.
5.2.1 Sample Calculations
The process most widely used in industrial boiler applications is the
Sodium Throwaway process; consequently, it was selected for example calcu-
lations. The following information was taken from the material balance
results presented in Appendix A.
Boiler Size - 58.6 MW (200 106 Btu/hr)
Coal Type - 3.5% S Eastern Coal
S02 Removal - 90%
Exit Flue Gas Rate - 1430 Nm /min (50,552 scfm)
Liquid Pumping Requirements - 3767 £/min (994 gpm)
From Table 5.2-1:
Discharge Pump Pressure - 3484 Pa (10 psi)
Pumping Height - 6m (20 ft)
Gas Pressure Drop - 100 Pa (8 in. H20)
5-17
-------
1000
900
800 —
700
Limestone
g 600
•i-i
D.
e
3
cn
c
o
o
500 —
Sodium Throwaway
£ 400
w
300 —
200 _
100 —
75
85
95
Percent S02 Removal
(58.6 >rw Boiler; High Sulfur Kastern Coal)
Figure 5.2-6. Energy consumption versus 502 removal
5-18
-------
700.
Limestone
600-
500—
400—
C-
E
0}
C
o
CJ
°f 300—
c
Spray Drying
-• Sodium Throwaway
200-
100—
50
60
70
I
80
90
Percent S02 Removal
(44 MWt Boiler; Low Sulfur Wostrrn Co;il)
Figure 5.2-7. Energy consumption versus S02 removal
5-19
-------
5.2.1.1 Calculate Raw Material Handling and Preparation —
From a previous study of energy requirements for soda ash handling
and preparation:
Electricity = 24.8 kW/ (kg S02 removed/sec)
S02 removed = 0.127 kg/sec
Electricity =24.8 (0.127) = 3.2 kWg
= 9.3 kWt (.032 106 Btu/hr)
5.2.1.2 Calculate Liquid Pumping Energy —
w n = Ik + t_ zb + |bi
P p gc 2gc
where: W = pump work (ft-lb/sec)
n = pump efficiency [assume 0.6]
Pb = discharge pressure (lb/ft2)
p = density (lb/ft3)
g = gravitational acceleration (ft/sec2)
gc = conversion factor (32.179 ft-lb /lb -sec)
m f
Zb = pumping height (ft)
Vb = fluid velocity (ft/sec) [assume 10]
applying these values:
w n = 110X144)_ _(32.2^ (.20) _, (10)2
P 62.3 + 32.17 + 64734
P lb
m
W = 745-^1
P lb
m
5-20
-------
To calculate the required power :
= m W
B 550
where: P = brake horsepower (Hp)
B
m = mass flow rate (Ib/sec)
nr>/ gal min 8.33 Ib
m = 994 -^T— x x
rn - ^
mm 60 sec gal
m = 138.0 Ib/sec
_ (138.0) (74. 5) _
PB -- 550 - 18.7 Hp
PB
= 40.7 kWt
5.2.1.3 Calculate Fan Energy--
scfm = 50,552 scfm
Pressure drop = 8 in. HaO
Gas density = .08 lb/ft3
Fan efficiency = .6
Exit gas velocity = 60 ft/sec
AP
W n = —
P P
W n = -
p .6
lb/ft2
(8 in.H20) (5.2
(.08 Ib /ft3
m
ft-lb
= 867-lb-JL
m
= m W
P
(550)(1.34)
5-21
-------
coo sc£ mole 29 Ib min
m - 5U,>ZZ x x x
5.2.1.4 Calculate Waste Disposal Requirements
Assume viscosity of waste water equals that of water
y = .0012 Ib/ft-sec
Assume linear fluid velocity in pipe -
v = 10 ft/sec
Average flow rate for all cases examined:
= ™ PPM (ft3) (min)
^ n (7.48 gal.-> (60 sec)
= .087 ftVsec
Cross-sectional area of pipe
.087 ftVsec
10 ft/sec
= .0087 ft2/sec = ~-
D = .093 ft
Reynolds Number—
Re = *—
y
= (.093 ft) (10 ft/sec) (62.3 lb/ft3)
(.0012 Ib/ft-sec)
= 4.84 x 10*
Friction Factor—
f = .0052
Assume wastewater is pumped to onsite treatment facility, one mile from
FGD system—
L = 5280 ft
5-22
-------
Friction Loss—
D 2gc
9(.OQ52) (5280 ft) (10 ft/sec)2
(.093 ft) (2) (32.2 Ib ft )
m
seczlbf
ft-lb
Tb—
m
Pump Work—
P
W n = — + -S- Z, + -r^— + F
P P gc b 2 gc
= 10 (144) . (10)2 ,
62.3 2(32.2)
= 23 + 0 + 2 + 1834
ft-lb
= 1859
lb
1859 m ft'lbJ
Wp = ~ = 3100 -^—J
m
Required Power—
m W
p = 1
B 550
from Appendix A —
flow =92 gpm
= 92 x -- X
mm 60 sec gal
= 13 lb /sec
m
(13 lb /sec) (3100 ft-lb./lb )
_ m _ i m
ft-lb /sec
550
hp
5-23
-------
=73-3 h?
= 54.7 kW
e
= 160.4 kW
5.2.1.5 Calculat_e_Utilities .and Services —
1 2
From a previous study of energy requirements:
Electricity = 0.15 kW/(Nm3/sec)
Inlet flow rate = 45,228 scfm
45,228 ft3
min
492°R
520°R
1 min
60s
(.3048)3m3
1 ft3
Vs
Electricity = .15(21) = .2 kW£
= 9.3 kWfc (.033 106 Btu/hr)
5.2.1.6 Calculation Summary—
kW 106 Btu/hr
Raw materials preparation/handling 9.3 0.03
Liquid pump energy requirements 40.7 0.14
Fan energy requirements 222.0 0.76
Wastewater disposal pumping 160.4 0.55
Utilities and services 9.3 0.03
441.7 1.51
5.2.2 Methods to Reduce Energy Consumption
Energy consumption in nonregenerable FGD systems that do not use reheat
is primarily attributed to two main sources: electricity for driving fans
and electricity for driving pumps. Other significant energy consumers are
the feed preparation and utilities and services. For small industrial
boiler applications electrical power requirements will probably be purchased
5-24
-------
whereas steam requirements may result in boiler derating- Ecth electricity
and steam may have to be purchased for some installations such as those for
boilers producing hot water only. Table 5.2-7 presents a summary of the
relative percentage of each of the energy requirements of each process area
as compared to the overall energy requirement for the FGD process. This
table shows that fans are the largest energy consumer for each of the wet
throwaway processes when stack gas reheat energy requirements are not
included. However, when reheat energy requirements are included, they
become the dominant energy consuming portion of these processes.
TABLE 5.2-7. PERCENTAGE ENERGY CONSUMPTIONS FOR NONREGENERABLE PROCESSES
(58.6 MWt Boiler, 90% Removal, Eastern Coal)
Source of
energy
consumption
Raw materials
handling and
preparation
Liquid pumps
Fans
Disposal pumps
Utilities and
services
TOTAL
Reheat Steam
Limestone
kW
114
317
471
-
8
911
884
1795
.2
.4
.7
.6
.9
.0
.9
Percent
of total
13
35
51
-
1
Double
Alkali
Percent
kW of total
12.7
63.2
222.0
-
9.3
307.2
884.0
1191.2
1
21
72
-
3
Sodium
kW
9.
40.
222.
160.
9.
441.
884.
1325.
TA
Percent
of total
3
7
0
4
3
7
0
7
2
9
51
36
2
Liquid pumping requirements are set by the required S02 removal; thus
it is unlikely that they will be reduced for a given process design. However,
changes in a process design such as using a different sorbent (for example
lime vs. limestone), a different gas-liquid contactor, or using additives
to enhance the liquid phase alkalinity may all result in reduced liquid
pumping rates. Fan energy requirements are a function of the system pressure
5-25
-------
drop and will vary with the type of gas-liquid contactor and the volume of
flue gas entering the system. Use of an open type contactor (for example
a spray tower) coupled with a decrease in the amount of excess air used by
the boiler will result in decreases in the fan energy requirements.
Energy requirements for the spray drying process are already quite low,
and compare well with the other nonregenerable processes when no reheat is
included. Table 5.2-8 illustrates this for a low sulfur western coal case.
If reheat is required, the spray drying process has a significant advantage
since it exhausts its flue gas at 175°F and requires no reheat.
TABLE 5.2-8. COMPARISON OF SPRAY DRYING PROCESS ENERGY REQUIREMENTS
(44 MW , 0.6% S coal, 75% Removal)
Source of
energy
consumption
Raw Material Handling
Liquid Pumping
Fans
Disposal Pumps
Atomization
Utilities & Services
Total
Reheat steam
Limestone
kWfc
15
133
292
-
-
8
448
775
1223
Percent
of total
3
30
65
-
-
2
Sodium TA
Percent
kW of total
1
29 11
195 78
21 8
-
9 3
254
775
1029
Lime
kW
1.
0.
140
-
120
8
270
—
270
Spray Drying
Percent
of total
5 1
5
52
-
44
3
100
100
Excluding reheat, energy consumption in the Wellman-Lord process is
attributable to four main sources: electricity for pumps and fans, process
steam, and methane for reducing S02 to sulfur. Table 5.2-9 shows the
relative energy requirements of each of the Wellman-Lord process areas for
both eastern and western coal applications.
5-26
-------
For this process, the relative amounts of energy consumed by the various
process areas varies depending upon the sulfur content of the coal being burned
However, for both the eastern 3.5 percent sulfur coal case and the western 0.6
percent sulfur coal cases, the regeneration and sulfur production areas are
the major energy users. It is doubtful that the energy requirements of the re-
generation processing area can be reduced since double effect evaporators were
assumed for these calculations which are some 45 percent more energy efficient
than single effect evaporators.13 If reheat energy requirements are included,
they become the major energy consumer for the low-sulfur western coal case and
serve to increase the energy requirements of both cases by about 900 kW .
TABLE 5.2-9. PERCENTAGE ENERGY CONSUMPTION FOR WELLMAN-LORD PROCESS
(58.6 MW 90% S02 Removal)
Source of
energy
consumption
Raw materials
handling and preparation
Pumps
Fans
Process steam
Methane
Utilities and services
Total
Reheat steam
kWt
1.9
42.2
208.6
469.0
219.0
10.2
950.9
900.0
1850.9
Western coal
(0.6%S)
Percent of total
<1
4
22
49
23
1
48.6
kWfc
9.3
82.9
205.5
2220.0
1048.0
9.6
3575.3
884.0
4459.3
Eastern coal
(3.5%S)
Percent of total
<1
2
6
62
29
<1
20
Energy requirements in the S02 reduction area could be decreased if
sulfuric acid were produced rather than sulfur. Acid production does not
require a reducing gas. Calculations to compare the relative energy con-
sumption of sulfuric acid versus sulfur production were not performed as
part of this study.
5-27
-------
5.3 IMPACT OF CONTROLS FOR OIL-FIRED BOILERS
The energy impact of FGD systems applied to oil-fired boilers will be
similar to that for coal-fired boilers. The major difference will be that
the oil-fired boilers will generally operate with a lower percent excess air
which will result in a lower energy requirement for the fans. The standard
residual oil-fired boiler operates with only 15 percent excess air whereas
the standard coal-fired boiler of the same size, 44 MW (ISOxlO6 Btu/hr)
operates with 50 percent excess air.
Table 5.3-1 presents the results of calculations to estimate the energy
requirements for a limestone system applied to a residual oil-fired boiler.
This table also presents the total energy requirements for the limestone
system applied to a 44 MW (150 x 106 Btu/hr) coal-fired boiler. The system
applied to the residual oil-fired boiler has energy requirements on the order
of 70 percent of those for the systems treating flue gases from coal-fired
boilers. By analogy, energy requirements for other nonregenerable or throw-
away processes should be on the order of 70 percent of those for coal-fired
applications.
For the Wellman-Lord process, a significant amount of the overall pro-
cess energy consumption occurs in the regeneration and S02 conversion sections
of the process. Because of this, energy consumption for a residual oil
application will vary depending upon the amount of S02 processed. As shown
in Figure 5.2-1, the energy consumption of a Wellman-Lord system is strongly
related to the sulfur content of the fuel being burned. From Figures 5.2-2
and 5.2-3, for a 44 MW (150x106 Btu/hr) application, a Wellman-Lord system
applied to an eastern 3.5%S coal-fired boiler would consume about 2700 kW
whereas a Wellman-Lord system applied to a western 0.6% coal-fired boiler
would consume about 700 kW. The eastern coal application would process
5-28
-------
I
ho
TABLE 5.3-1. LIMESTONE PROCESS ENERGY REQUIREMENTS FOR RESIDUAL OIL APPLICATION
[44 MWt(150 x 106 Btu/hr)]
Source of energy consumption
Raw materials handling and
preparation
Pumps
Fans
Disposal
Utilities and Services
Total
Percent of boiler heat input
Total kW for a limestone system
kW
47.7
172.2
263.5
1.0
4.8
489.2
1.1
776.8
90
% of total
9.8
35.2
53.8
.2
1.0
Removal
kW
45.7
135.5
217.0
.9
4.8
403.9
0.92
—
levels, percent
85
% of total
11.3
33.5
53.8
.2
1.2
kW
39.8
99.1
186.0
.8
4.8
330.5
0.75
507.3
75
% of total
12.0
30.0
56.3
.2
1.5
applied to a 44 MW boiler burning
3.5% eastern coal
Total kW for a limestone system
applied to a 44 MW boiler burning
0.6% S western coal
620.7
447.7
-------
15.8 moles/hr SOz whereas the western coal application would process 3.7
moles/hr S02.
Energy requirements for a Wellman-Lord system applied to the standard
residual oil-fired boiler are shown in Table 5.3-2. Energy requirements for
this system for the 90 percent SOa removal case (removal of 6.6 moles/hr SOa)
were calculated to be 1687 kW. This value compares well with the energy
requirements for the coal-fired Wellman-Lord cases as discussed above.
Previous comments for coal-fired FGD system applications concerning major
energy consumption sources and potential methods of energy reduction will
apply to residual oil-fired FGD systems. Also, all of the energy impacts
discussed in this section will apply to both new and retrofit applications.
5-30
-------
TABLE 5.3-2. WELLMAN-LORD PROCESS ENERGY REQUIREMENTS FOR RESIDUAL OIL APPLICATION
[44 MWt(150 x 106 Btu/hr)]
Removal
Source of energy consumption
Raw materials handling and
preparation
Pumps
Fans
Process Steam
Methane
Utilities and services
Total
Percent of boiler heat input
Ui
i
3
54
114
1064
445
_ 4
1686
3
kW
.9
.4
.5
.0
.0
JL§-
.6
.8
90
% of
0
3
6
63
26
0
total
.2
.2
.8
.1
.4
.3
3
53
114
1016
430
4
1622
3
kW
.6
.9
.5
.0
.0
.8
.8
.7
levels, percent
85
% of
0
3
7
62
26
0
total
.2
.3
.1
.6
.5
.3
kW
3.2
53.3
114.5
886.0
416.0
4.8
1477.8
3.4
75
% of
0.
3.
7.
60.
28.
0.
total
2
6
7
0
2
3
-------
REFERENCES
1. Rubin, Edward S. and Due G. Nguyen. "Energy Requirements of a Limestone
FGD System." J. APCA 28 (12), 1207 (1978).
2. Dickerman, J.C. Trip Report. Meeting Notes - Lehigh University
Symposium on Flue Gas Desulfurization Energy Requirements. Austin, TX.
Radian Corporation. October 1978.
3. Thomas, W.C. Energy Requirements for Controlling S02 Emissions from
Coal-Fired Steam/Electric Generators. DCN 78-200-187-08-12. EPA 450/
3_77_050a. EPA Contract No. 68-02-26-8, Task 12. Austin, TX. Radian
Corporation. December 1977.
4. Torstrick, Bob (Tennessee Valley Authority, Chemical Engineering Design
Branch, Muscle Shoals, AL). Private Communication with J. C. Dickerman.
8 January 1979.
5. Radian Corporation. Stack Gas Reheat Evaluation. Draft Final Report.
DCN 78-200-206-25. EPA Contract No. 68-02-2642. Austin, TX. September
1978.
6. Thomas, op.cit. , pp. D1-D13.
7. Ibid.
8. Estcourt, V.F., et al. Tests of a Two-Stage Combined Dry Scrubber/S02
Absorber Using Sodium or Calcium. Presented at the 40th Annual Meeting
American Power Conference, Illinois Institute of Technology. Chicago,
Illinois, p. 42.
9. Ibid., p. 40.
10. Broz, Larry, Charles Sedman, and David Mobley. "Recommendation for
Control Levels to be Considered in Preparation of ITARS and Executive
Summary Outline for ITARS." Raleigh, NC. Acurex Corp. EPA/OAQPS and
EPA/IERL. August 1978.
11. Thomas, op.oit. , p. D13.
12. Ibid..
13. Ottmers, D.M., Jr., et al. Evaluation of Regenerable Flue Gas Desul-
furization Processes, Revised Report, 2 Vols. EPRI RP 535-1, EPRI
FP-272. Austin, TX. Radian Corporation. July 1976. pp. 79-80.
5-32
-------
SECTION 6
ENVIRONMENTAL IMPACT OF CANDIDATES
FOR BEST EMISSION CONTROL SYSTEMS
6.0 INTRODUCTION
Each of the flue gas desulfurization (FGD) processes selected as candi-
dates for industrial boiler application create beneficial air emission impacts
by reducing both the S02 and particulate emissions from industrial boilers.
However, adverse water and solid waste emission impacts can result from FGD
systems if proper design and operating practices are not followed since the
S02 removed from the boiler flue gas is converted to liquid or solid by-
products. The limestone and double alkali processes convert the absorbed S02
into sludges; the spray drying process converts the absorbed S02 into a dry
solid; and the sodium throwaway process converts the absorbed S02 into soluble
sodium salts. Although the Wellman-Lord process produces a salable by-product,
either sulfur or sulfuric acid, a purge sodium sulfate solids stream and a
high chloride aqueous stream are produced which must be properly disposed of.
This chapter examines the air water, and solid waste impacts of each of the
candidate FGD processes.
6.2 ENVIRONMENTAL IMPACT OF CONTROLS FOR COAL-FIRED BOILERS
6.2.1 Air Pollution
Four of the five FGD processes under consideration are wet scrubbing
systems and as such have the capability of removing particulates as well as
3-1
-------
SO2 from industrial boiler flue gases. In fact, several industrial boiler FGD
systems in use today are designed and operated in a dual particulate/S02
removal service. The fifth candidate process, spray drying, is a dry process
with down stream particulate collection (baghouse or ESP) to collect the solid
waste product, generated in the spray dryer along with any accompanying
particulate material.
Although the FGD systems under consideration have the potential for
combined SOj/particulate removal, only the SO2 removal aspects of these systems
will be considered in detail in this Individual Technology Assessment Report
(ITAR)- Particulate removal is considered in detail in the particulate collec-
tion technology assessment report.
Particulate removal in FGD systems has been evaluated by EPA as part of
their efforts to prepare the proposed New Source Performance Standards reported
in the 19 September Federal Register for Electric Utility Steam Generating
Units. As reported in the _19_ September 1978 Federal Register, the EPA has
investigated the following mechanisms to attempt to assess the impacts of FGD
systems upon particulate emissions:
1) FGD system sulfate carryover from the scrubber slurry,
2) particulate matter removal by the FGD systems, and
3) particulate matter generation by the FGD system through
condensation of sulfuric acid (H2SO^) mist.
With regard to the first mechanism, data were taken from steam generators
that had operating FGD systems with low particulate levels at the FGD inlet.
The data indicated that with a properly maintained mist eliminator, particulate
emissions did not increase through the FGD system.2 With regard to the second
mechanism, the FGD system data indicated that scrubbers did indeed reduce the
level of particulate emissions.3
6-2
-------
With regard to the third mechanism, the EPA obtained data that indicated
an FGD system applied to a low sulfur coal-fired boiler will not increase
particulate emissions through sulfuric acid formation. Sufficient data,
however, have not become available for high sulfur coal-fired applications to
fully assess the effect of sulfuric acid emissions on measured particulate
emissions and this matter is currently undergoing further investigation by EPA.4
With regard to SQ? , the air pollution impact of the FGD processes will be
the same for a given SO^ removal level. Tables 6.2-1 through 6.2-5 show the
emission levels achieved by applying varying SG> control levels to the standard
coal-fired boilers. Figures 6.2-1 through 6.2-3 illustrate these impacts
graphically. As shown in Tables 6.2-1 through 6.2-5, emissions from the 3.5
percent sulfur eastern coal fired boilers will be about 4.7 times those from
the 0.6 percent sulfur western coal-fired boilers and 1.7 times those from the
2.3 percent sulfur coal-fired boilers for an equivalent level of control.
This is illustrated in the three figures mentioned previously and, in Figure
6.2-4.
The amount of SOz removal necessary to achieve compliances with an
average State Implementation Plan (SIP) control level was also considered.
The average SIP control level was defined by EPA for this study to be 1075
ng/J (2.5 lb/106 Btu). It is interesting to note that no S02 control would be
required for a boiler firing 0.6 percent sulfur western coal to comply with
this level whereas a boiler burning 3.5 percent sulfur eastern coal would
require an SO 2 control level of 56 percent to meet this same level and a
boile*- burning 2.3 percent sulfur coal would require 24 percent S02 control.
A secondary type of environmental impact resulting from the application
of FGD systems to industrial boilers are problems associated with the forma-
tion of a saturated flue gas plume in the gas-liquid contactor. Stack gases
from industrial boilers without FGD are normally exhausted at temperatures of
between 150-200°C (300-400°F). In this temperature range, the flue gas is
relatively dry and noncorrosive. The wet FGD processes cool the boiler flue
gas to its adiabatic saturation temperature, normally 50-60°C (125-140°F).
This saturated flue gas may cause the followed problems:
6-3
-------
TABLE 6.2-1. AIR POLLUTION IMPACTS OF SO, CONTROL TECHNIQUES FOR COAL-FIRED BOILERS
I
-P-
S02 Emissions
Boiler size Control
and type level (%)
8.8 MWfc(30 X 106 Uncontrolled
Btu/hr)
Underfeed stoker
SIP*
Moderate
75
Intermediate
85
Stringent
90
11,800
Btu/lb
3.5% S eastern coal
g/s
(Ib/hr)
21.3
(169)
9.5
(75)
5.3
(42.2)
3.2
(25.6)
2.1
(16.6)
ng/J
(lb/106 Btu)
2408
(5.6)
1075
(2.5)
606
(1.4)
365
(.85)
236.5
(.6)
9,600
0.6% S
g/s
(Ib/hr)
4.5
(35.8)
1.1
(9.0)
.7
(5.. 4)
.5
(3.6)
Btu/lb
western coal
ng/J
(lb/106 Btu)
516
(1.2)
128.9
(.3)
77.0
(.2)
51.6
(.1)
*56 percent removal required to meet average SIP for 3.5% S eastern coal, no FGD required for
0.6% S western coal.
-------
TABLE 6.2-2. AIR POLLUTION IMPACTS OF S02 CONTROL TECHNIQUES FOR COAL-FIRED BOILERS
I
Ul
11,800 Btu/lb
Boiler
and type
22 MWt
(75 X^IO6 Btu/hr)
Chaingrate Stoker
Control
level (%)
Uncontrolled
SIP*
Moderate
75
Intermediate
85
Stringent
90
3.5% S
g/s
(lb/hr)
53.3
(423)
23.6
(187.5)
13.3
(105.6)
8.0
(63.4)
5.3
(42.2)
eastern coal
ng/J
(lb/10 Btu)
2408
(5.6)
1075
(2.5)
605.3
(1.4)
361.1
(.8)
242.1
(.6)
S02
13,
2.
g/s
(lb/hr)
31.3
(248.3)
23.6
(187.5)
7.8
(62.1)
4.7
(37.2)
3.1
(24.8)
emissions
200 Btu/16
3% S coal
ng/J
(lb/10 Btu)
1421
(3.3)
1075
(2.5)
355.3
(.83)
213.2
(.50)
142.1
(.33)
9,000 Btu/lb
0.6% S
g/s
(lb/hr)
11.2
(88.9)
2.8
(22.4)
1.8
(13.3)
1.1
(9.0)
western coal
ng/J
(lb/106 Btu)
516
(1.2)
128.9
(.3)
77.4
(.2)
51.6
(.1)
*56 percent removal required to meet average SIP for 3.5% S eastern coal, 24 percent removal
for 2.3% S coal, and no FGD required for 0.6% S western coal.
-------
TABLE 6.2-3. AIR POLLUTION IMPACTS FROM "BEST" SO2 CONTROL TECHNIQUES FOR COAL-FIRED BOILERS
SOa Emissions
Boiler size Control
and type level (%)
44 MWt (150 X 10sBtu/Hr) Uncontrolled
Spreader Stoker
SIP*
Moderate
75
Intermediate
85
Stringent
90
11,800
Btu/lb
3.5% S eastern coal
g/s
(lb/hr)
106.4
(844.8)
47.3
(375)
26.6
(211.2)
16.0
(176.7)
10.6
(84.5)
ng/J
(lb/10 Btu)
2408
(5.6)
1075
(2.5)
605.3
(1.4)
361.1
(.8)
242.1
(.6)
9,600
0.6% S
g/s
(lb/hr)
22,4
(177.9)
5.6
(44.2)
3.4
(26.7)
2.3
(17.9)
Btu/lb
western coal
ng/J
(lb/105 Btu)
516
(1.2)
126.6
(.3)
77.4
(0.2)
51.4
(-1)
*56 percent removal required to meet average SIP for 3.5% S eastern coal, no FGD required for
0.6% S western coal.
-------
TABLE 6.2-4. AIR POLLUTION IMPACTS FROM "BEST" S02 CONTROL TECHNIQUES FOR COAL-FIRED BOILERS
SOz Emissions
Boiler size Control
and type level (%)
58.6 MWt(200 X 106Btu/hr) Uncontrolled
Pulverized Coal
SIP*
Moderate
75
Intermediate
85
Stringent
90
11,800
Btu/lb
3.5% S eastern coal
g/s
(Ib/hr)
141.9
(1126.4)
63.0
(500)
35.5
(281.6)
21.3
(169)
14.2
(112.6)
ng/J
(lb/10b Btu)
2408
(5.6)
1075
(2.5)
605.3
(1.4)
363.2
(.8)
242.1
(.6)
9,600
0.6% S
g/s
(Ib/hr)
29.9
(237.4)
7.5
(59.5)
4.5
(35.8)
3.0
(23.7)
Btu/lb
western coal
ng/J
(lb/10 Btu)
516
(1.2)
128.9
(.3)
77.0
(.2)
51.4
(.1)
*56 percent removal required to meet average SIP for 3.5% S eastern coal, no FGD required for
0.6% S western coal.
-------
TABLE 6.2-5. AIR POLLUTION IMPACTS FROM "BEST" S02 CONTROL TECHNIQUES FOR COAL-FIRED BOILERS
i
CO
Boiler
and type
118 MW
(400 X 106Btu/hr)
Pulverized Coal
Control
level (%)
Uncontrolled
SIP*
Moderate
Intermediate
85
Stringent
90
11,800
Btu/lb
3.5% S eastern coal
8/s
(Ib/hr)
284.3
(2254.4)
126.1
(1000.0)
71.1
(563.6)
42.7
(338.2)
28.4
(225.4)
ng/J
(lb/106 Btu)
2423
(5.6)
1075
(2.5)
605.9
(1.4)
363.6
(0.8)
242.3
(0,6)
SO,
' 13,
2.
g/s
(Ib/hr)
166.9
(1323.2)
126.1
(1000.0)
41.7
(330.8)
25.0
(198.5)
16.7
(132.3)
emissions
200 Btu/16
3% S coal
ng/J
(lb/10 Btu)
1422.
(3.3)
1075
(2.5)
335.6
(0.8)
213.4
(0.5)
142.2
(0.3)
9,000
Btu/lb
0.6% S western coal
g/s
(Ib/hr)
59.9
(474.9)
15.0
(118.7)
9.0
(71.2)
6. 0
(47.5)
ne/J
(lb/10 Btu)
510.5
(1.2)
127.6
(0.3)
76.5
(0.2)
51.1
(0.1)
*56 percent removal required to meet average SIP for 3.5% S eastern coal, 24 percent
removal for 2.3% S coal, and no FGD required for 0.6% S western coal.
-------
300(2379)
250(1982) —
200(1586)
E
a
130(1190)—
100(793)
50(396L_
SO- Control Level (,i)
Figure 6.2-1. S02 EMISSIONS VERSUS CONTROL LEVEL FOR STANDARD
BOILERS FIRING 3.5% SULFUR EASTERN COAL
6-9
-------
M
OC
W
O
cn
300(2379)—
250(1982)—i
200(1586)
150(1190)—
100(793)—
50(396) —
50
S02 Control Level (%)
75
85 90
100
Figure 6.2-2. S02 EMISSIONS VERSUS CONTROL LEVEL
FOR STANDARD BOILERS FIRING 2.3% SULFUR COAL.
6-10
-------
300(2379)—
250(1982)—
200(1586)—
e 150(1190)—
100(793)
50(396)—
MW Boiler
118
100
Figure 6.2-3, S02 EMISSIONS VERSUS CONTROL LEVEL FOR STANDARD
BOILERS FIRING 0.6% SULFUR WESTERN COAL.
6-11
-------
2400(5. 6)-i
2100(4.9)-
1800(4.2)_
1500(3.5)_
1200(2.8)-
900(2.1)-
600(1.4) -
300(0.7) -
All Boiler Sizes Firing
High Sulfur Eastern Coal
All Boiler Sizes F\ring
Medium Sulfur Coal
All Boiler Sizes Firing
Low Sulfur Western Coal
I
25
50 56
1
75
I I I
85 90 100
S02 Control Level (%)
Figure 6.2-4. S02 emissions versus control level.
6-12
-------
The occurrence of acid rain in the vicinity of the
plant's stack,
High ground level pollutant concentrations downwind
from the stack due to poor plume buoyancy.
The formation of a visible plume which may be hazardous
to any ground and air traffic in the vicinity of the
power plant, and
The corrosion of equipment downstream of the scrubber
due to the presence of moisture, acid, and chlorides.
Reheating the saturated flue gas to a temperature above its saturation tempera-
ture will lessen the impacts of each of these four potential problems. EPA
is currently investigating the amount of reheat required and the best method
of achieving the desired amount of reheat to prevent these potential problems.5
6.2.2 Water Pollution
Four of the five candidate processes produce waste products that have
the possibility of resulting in aqueous emissions. Table 6.2-6 gives the
major physical properties of these waste product streams.
TABLE 6.2-6. PHYSICAL PROPERTIES OF WASTE PRODUCTS
Waste product
System physical properties
Sodium Throwaway Aqueous stream - 5% dissolved solids
Limestone Sludge - 35-60% solids
Double-Alkali Sludge - 35-60% solids
Wellman-Lord Purge solids
Aqueous stream - low pH
from prescrubber
6-13
-------
Good design and operating practices for the limestone and double-alkali
processes include dewatering the sludge and recycling the supernatant liquid.
Consequently, there should not be water emissions from these systems except
for times of severe rainfall or process upsets. Waste streams from these
processes will be discussed further in the section on solid waste.
The aqueous waste stream from the prescrubber of the Wellman-Lord process
will be characterized by a low pH which results from the chlorides that are
removed from the gas stream. In addition, residual fly ash not removed by
the upstream particulate removal device may be present in the stream. Table
6.2-7, which gives results of the material balance calculation presented in
Appendix A, shows estimated prescrubber discharge rates for the standard size
boilers.
TABLE 6.2-7. WELLMAN-LORD PRESCRUBBER DISCHARGES
Discharge rate
Boiler size
8.8 MW
22 MW
44 MWt
58.6 MWt
£/min
9.1
22.7
45.4
60.6
gal/min
2.4
6.0
12.0
16.0
gal/106Btu
heat input
4.8
4.8
4.8
4.8
This prescrubber discharge stream is for the purpose of maintaining
suspended and dissolved solids at desired levels. Suspended solids (fly ash)
are generally held at less than five percent whereas the dissolved solid
level may be greater than 20,000 mg/£ depending upon the chlorine concentra-
tion in the coal.
Except for the high chloride concentrations and low pH, the quality of
the prescrubber discharge will be very similar to that of the boiler ash
sluice water. Since this stream has been estimated to be approximately one
percent of the ash sluicing requirements for a power plant, it can be used
for ash sluicing where it will become diluted and neutralized with the other
6-14
-------
ash sluice water. Consequently, water emissions from the Wellman-Lord pre-
scrubber stream should thus be limited to intermittent discharges from the
ash pond.
The aqueous stream from a sodium throwaway system will contain about
five percent dissolved solids'. Discharge rates and average stream compositions
for the cases considered in this study are given in Table 6.2-8. The major
differences between the dissolved solid compositions of the eastern and
western coal applications is due to the fact that 20 percent oxidation was
assumed to occur in the western coal cases whereas 10 percent was assumed to
occur for the eastern cases and 15 percent for the 2.3% S coal cases.
In sodium throwaway systems, the absorbed S02 reacts to form Na2S03 and
Na2SO^ which are removed from the system as dissolved solids in an aqueous
waste. Consequently, the amount of aqueous emissions is directly related to
both the S02 control level and the coal's sulfur concentration. The amount
of aqueous wastes from high sulfur eastern coal applications will be approxi-
mately 4.7 times those from western coal-fired applications.
Some of the sodium throwaway systems in use today for controlling SOz
emissions from industrial boilers use a waste process stream that contains
sodium as a feed to the scrubber. Sodium scrubbers applied to boilers in
some paper mills and textile plants are examples. The relatively small
aqueous process waste stream from the FGD system is then recombined with the
industrial process waste streams and discharged to a centralized water treating
plant or a municipal sewer.
Common water treating practice for sodium throwaway systems in use today
is to discharge their wastes to an evaporation pond or to an existing centra-
lized water treating plant. Of the 102 sodium scrubbing systems in use
today, about 80 use evaporation ponds and 10 use centralized water treating
for disposal of their FGD wastes. The remaining systems use varied approaches
8
ranging from discharge to city sewers to deep mine injection.
3-15
-------
TABLE 6.2-8. WATER POLLUTION IMPACTS FOR THE SODIUM THROWAWAY SYSTEM*
Control
Boiler size and type level
8.8 MWt(30 X 10 Btu/hr)
Underfeed stoker
22 MWt(75 X 10 Btu/hr)
Chaingrate Stoker
44 MWt(150 X 10 Btu/hr)
Spreader Stoker
58.6 Mwt(200 X 10 Btu/hr)
Pulverized coal
118 MWf(400 X 10 Btu/hr)
90
85
75
56
90
75
56
90
75
56
90
85
75
56
90
3.5% S eastern coal
H/sec
0.85
0.79
0.71
0.52
2.18
1.77
1.26
4.38
3.53
2.45
5.80
5.26
4.71
3.34
11.7
(gpm)
13.4
12.5
11.2
8.4
34.6
28.1
19.9
69.4
55.9
38.8
92.0
83.4
74.7
52.9
185
2.3% S coal 0.6% S western coal
£/sec (gpm) 9,/sec
0.17
0.16
6.15
_ _
1.28 20.3 0.46
0.37
- - —
0.92
0.75
- - -
1.23
1.11
0.99
_
6.82 108 2.46
(gpm)
2.7
2.6
2.4
-
7.3
5.9
—
14.6
11.8
—
19.5
17.6
15.7
-
38.9
Pulverized coal
Avg.
Avg.
Dissolved
TDS
Solid Compositions
Concentration
(wt. %)
Na2S03=
Na2S01+ =
NA2C03=
5
77
9
14
percent
percent
percent
73
13
14
percent
percent
percent
5
69
17
14
percent
percent
percent
5
*Based on material balance calculations provided in Appendix A.
-------
For purposes of this evaluation, onsite treatment of sodium system
aqueous wastes using a basic water treating scheme of sulfite oxidation
and pH neutralization was selected as the treatment method for evaluation.
Although evaporation ponds are currently used in the majority of sodium
system applications, their use is limited to certain geographic areas
of the county where the annual evaporation rate exceeds the annual
rainfall. The water treatment systems selected for this evaluation
will result in a sodium sulfate stream which must be disposed.
6.2.3 Solid Waste
The major solid waste impacts from the five candidate processes result
from the sludges produced in the limestone and double-alkali processes and
the dry solid produced in the spray drying process. A solid purge stream of
Na2SO,, is produced in the Wellman-Lord process, but the stream is relatively
small and should not constitute a major solid waste impact, especially for
the size applications under consideration in this study.
Both the limestone and double alkali sludges are composed primarily of
calcium sulfite and sulfate salts. Significant amounts of fly ash may also
be present, depending on the method of upstream particulate control in use.
The sludges are relatively inert and with proper site selection and proper
disposal procedures, can be disposed of in an environmentally acceptable
manner. The disposal methods currentl}' in use are lined and unlined ponding
and landfilling of treated and untreated materials. Potential adverse
impacts of sludge disposal lie in the areas of disposal acreage requirements,
water contamination through leaching and percolation of soluble components of
the solid waste into the groundwater system, and land use impacts due to poor
structural properties. Treatment techniques to minimize adverse impacts may
involve dewatering, addition of alkaline ash, and/or application of commercial
stabilization technology. These techniques are used to decrease the sludge
volume, decrease its permeability, and improve its structural properties as
6-17
-------
discussed in the following sections. In some instances, forced oxidation can
also be considered a sludge beneficiation technique as it results in a more
easily dewatered material.
The dry solid waste product from the spray drying process will consist
primarily of calcium or sodium salts, depending upon the type of alkali used
as the S02 sorbent. Significant amounts of fly ash will also be present
since the solids collection device associated with the spray drier, probably
a baghouse, will remove the particulates generated from the coal combustion
process along with the spray drying solid wastes. Upstream particulate
removal is not practical for this process since the spray dryer's performance
10
is not adversely affected by the presence of fly ash and dual particulate
removal units would be unattractive from both an energy and economic view.
The alkali, coal, and ash compositions have impacts on waste disposal alterna-
tives which include disposal in both lined and unlined landfills and the
11
potential for producing a low quality cement.
Volume of Sludge Production
As with the sodium throwaway system, all of the S02 absorbed from the
flue gas by a limestone system must leave the process in a waste stream, in
this case as a waste sludge. Consequently, the amount of sludge produced by
a limestone system is proportional to the sulfur content of the coal and the
S02 removal level. Table 6.2-9 presents the results of the limestone process
material balance calculations and shows the variation in sludge production
with coal sulfur content and SO removal. Sludge production rates presented
in this section are on an ash-free basis, since upstream particulate removal
was assumed for this study.
The volume of sludge produced is also important as the sludge volume
will determine the size of the holding pond or landfill area. Table 6.2-10
presents the results of calculations to estimate the sludge volumes produced
by a limestone process for the standard sized boilers. The difference in
sludge densities for the two coals is due to the higher oxidation for the
western coal cases. Results are presented in units of cc/sec, Ib/hr, and
acre-feet/30 years. The last category, acre-feet/30 years gives an indication
of the total volume of sludge to be handled over the life of the plant assuming
a 30 year life and an onstream factor of 60 percent. Figure 6.2-5 illustrates
6-18
-------
TABLE 6.2-9.
SOLID WASTE IMPACT FOR THE LIMESTONE FGD PROCESS
(Ash-Free Basis)
Percent
Boiler size and type removal
8.8 MWt(30xl06 Btu/hr)
Underfeed Stoker
22 MWt(75xl05 Btu/hr)
^ Chaingrate Stoker
I — i
UD
44 MWt(150xl06 Btu/hr)
Spreader Stoker
58.6 MWt(200xl05 Btu/hr)
Pulverized Coal
90
85
75
90
75
90
75
90
85
75
3.5% S eastern coal
g/s
89.8
84.7
74.9
226.5
190.2
457.0
380.6
606.8
576.1
507.0
(Ib/hr)
(712)
(672)
(594)
(1796)
(1508)
(3624)
(3018)
(4812)
(4568)
(4020)
£/min
2.7
2.7
2.3
6.8
5.7
13.6
11.4
18.2
17.4
15.2
(gal/min)
(0
(0
(0
(1
(1
(3
(3
(4
(4
(4
.7)
.7)
.6)
.8)
.5)
.6)
.0)
-8)
.6)
.0)
g/s
20.9
19.4
17.4
52.2
43.6
104.7
87.3
139.5
131.9
116.0
0.6% S western coal
(Ib/hr)
(166)
(154)
(138)
(414)
(346)
(830)
(692)
(1106)
(1046)
(920)
£/min (gal/min)
0.61
0.57
0.53
1.5
1.1
3.0
2.7
4.2
3.8
3.4
(0
(0
(0
(0
(0
(0
(0
(1
(1
(0
.16)
-15)
-14)
.4)
.3)
-8)
.7)
.1)
-0)
-9)
-------
TABLE 6.2-10.
SOLID WASTE VOLUMES FOR THE LIMESTONE FGD PROCESS
(Ash-Free Basis)
i
ho
O
Percent
Boiler size and type removal
8.8 MWt(30xl06 Btu/hr)
Underfeed Stoker
22 MWt(75xl06 Btu/hr)
Chaingrate Stoker
44 MWt(150xl06 Btu/hr)
Spreader Stoker
58.6 MWt(200xl05 Btu/hr)
Pulverized coal
90
85
75
90
75
90
75
90
85
75
3.5% S eastern coal*
Sludge volume
cc/sec
69.1
65.2
57.6
174.2
146.3
351.5
292.8
466.8
443.2
390.0
(ftVhr)
8.8
8.3
7.3
22.1
18.6
44.6
37.2
59.3
56.3
49.5
(acre-ft/
30 yrs)
31.8
30.0
26.4
80.0
67.4
161.4
134.6
214.6
203.8
179.2
0.6% S western coalt
Sludge volume
cc/sec
13.8
12.8
11.5
34.6
28.9
69.3
57.8
92.4
87.4
76.8
(ftVhr)
1.8
1.6
1.5
4.4
3.7
8.8
7.4
11.8
11.1
9.8
(acre-ft/
30 yrs)
6.6
5.8
5.4
16.0
13.4
32:0
26.8
42.8
40.2
35.4
^Eastern sludge bulk density = 1.30 g/cc (81.2 lb/ft3)
^Western sludge bulk density =1.51 g/cc (94.0 lb/ft3;
These are average values from Table 6.2-12.
which follows,
-------
cu
t-t
o c
03 r
O
0)
CO
a
a
C
O
•H
4-1
O
3
13
O
FX<
CU
00
3
rH
CO
500(230)—
400(184)—
300(138)—
200(92)—
100(46),
Eastern Coal
90% Removal
Western Coal
9.0% Removal
Eastern Coal
75% Removal
Western Coal
75% Removal
22 44
Boiler Size, MW
58.6
Figure 6.2-5.
Sludge Production Rates for the Limestone
FGD Process
6-21
-------
the results of these calculations graphically and shows the variation in
sludge production with coal sulfur content, boiler size, and level of removal.
As one would expect, sludge production increases with all of these factors.
The quantity of sludge produced from the double alkali process will also vary
with sulfur removal and fuel sulfur content. Table 6.2-11 presents the solid
waste impacts for the double-alkali process. The major difference between
the amount of sludge produced from the double-alkali and limestone systems is
that the limestone system stoichiometry is based upon 1.2 moles sorbent per
mole SO^ removed whereas the double-alkali system stoichiometry is based upon
1.0 moles sorbent per mole SOj removed.
Volume reduction can be accomplished by one or more sludge dewatering
practices. Methods that have been tested are: thickening, discharge to
settling pond (with or without underdrainage), vacuum filtration, centrifuga-
tion, and addition of dry materials such as fly ash. Although the degree of
volume reduction which can be attained with a given sludge depends on several
individual sludge properties, the relative effectiveness of various dewatering
methods is seen by comparing the bulk densities (wet basis) of treated sludges.
Several sets of data representing various sludge types and dewatering methods
are presented in Table 6.2-12. The commercial identities of the various
methods were not listed in the referenced report.
TABLE 6.2-12. EFFECT OF CHEMICAL FIXATION ON BULK DENSITY12
Bulk density (Ibs/ft )
Untreated Treated*
Average of untreated
and beneficial treatments
for limestone
Eastern lime 52 100 77 101 - -
Eastern limestone 63 108 89 83-63 81.2
Eastern dual alkali 52 96 91 99-53
Western limestone 89 109 80 111 81 57 94.0
Western dual alkali 47 97 82 83-68
*Each column represents one of five different commercial fixation treatments
tested.
6-22
-------
TABLE 6.2-11. SOLID WASTE IMPACT FOR THE DOUBLE ALKALI PROCESS
(Ash free Basis)
Percen"
Boiler Size and type reraova -
8.8 MWt(30 X 106 Btu/hr) 90
Underfeed Stoker
3.5% S eastern coal
0.6% S western coal
g/s (Ib/hr) d/min (gal/mln) g/s " (Ib/hr) t/min (gal/rain)
78.7 (624) 2.3 (0.6) 15.4 (122) 0.4 (0.1)
2.3% S coal
g/s (Ib/hr) ?,/tnln (gal/min)
I
hO
00
22 MWt(75 X 10s Btu/hr) 90 200.0 <1584) 5.7 (1.5) 38.1 (302) 1.1
Chaingrate Stoker
58.6 MWt(200 X 10 Btu/hr) 90
Pulverized Coal
506.0 (4012) 14.4
118 MWt(400 X 106 Btu/hr) 90 1065.6 (84501 30.7
Pulverized Coal
(3.8) 101.6 (806) 2.7
(0.3)
(0.7)
112.0 (8S8) 3.0 (0.8)
(8.1) 203.5 (1614) 5.3 (1.4) 596.8 (4732; 17.1 (4.5)
-------
One of the sludge parameters which determine ease of dewaterability is
the relative sulfate/sulfite contents in the solids. High sulfate sludges
(•I.e.3 with sulfate present as gypsum) tend to dewater much better than
sludges with high sulfite levels. This in turn reduces the volume of material
for disposal and makes a material which may be suitable for landfill without
need of additives. The EPA currently has studies underway at Shawnee to
investigate forced oxidation as a means of achieving a waste with better
settling and dewatering characteristics. Additional activities planned in
this area include a full scale demonstration of the technology at TVA' s
Widows Creek plant.
All of the SOz absorbed in a spray dryer must also exit the process as a
waste stream, in this case as a solid salt. Table 6.2-13 presents the quantity
of solid wastes produced by the spray drying system as a function of S02
removal, coal type, and process size. Solid waste quantities from this
process are, however, a combination of spray dryer solids and fly ash generated
from coal combinations. It is interesting to note that for the cases considered
in this evaluation, the majority of solid wastes from this process resulted
from fly ash and not from the removal of SOa- Table 6.2-14 presents a break-
down of the origin of solid waste material for each of the spray drier cases.
Results of pilot plant testing reported by Basin Electric were that the
spray drying product produced from the coals they tested handled as well as
fly ash and would not require special handling equipment other than the
conventional dry handling equipment used for fly ash.11* Disposal methods
planned for the systems they have under construction are disposal in depleted
mines and landfill after mixing with conventional scrubber sludge.15 The two
spray drying systems under construction for industrial boiler applications
both plan to truck the waste solids to an off site landfill area.16'17
6-24
-------
TABLE 6.2-13. SOLID WASTES FROM SPRAY DRYING
(Total Fly Ash + Alkali Salts)
hO
Ol
Boiler size
MWt (10s Btu/hr)
44
44
44
58.6
17.6
44
44
44
118
118
22
(150)
(150)
(150)
(200)
(60)
(150)
(150)
(150)
(400)
(400)
(75)
Z S
Coal
0.6
0.6
0.6
0.6
0.6
0.6
0.6
0.6
0.6
2.3
2.3
S02
Removal
90
75
50
75
75
90
75
50
70
70
70
Sorbent
Sodium
Sodium
Sodium
Sodium
Sodium
Lime
Lime
Lime
Lime
Lime
Lime
Sorbent
Stoichlometry
1.1
0.8
0.5
0.8
0.8
2.0
1.2
0.65
1.2
1.2
1.2
Solids
g/sec
132
114
99
174
29
134
114
96
344
729
85
Ub/hr7
(1047)
(904)
(785)
(1378)
(227)
(1066)
C905)
(764)
(2725)
(5782)
(675)
cc/sec
111
96
83
146
24
113
96
81
289
613
71
Volume
(ft°/hr)
(14.1)
(12.2)
(10.6)
(18.5)
(3.1)
(14.3)
(12.2)
(10.3)
(36.7)
(77.9)
(9.1)
a
acre-ft/15 years
25.5
22.1
19.2
33.5
5.6
25.9
22-.0
18.6
66.5
141.0
16.5
-------
TABLE 6.2-14. SOLID WASTE BY ORIGIN
Amount of solid waste
MW (106
44
44
44
ON 58.6
N> 17.6
ON
44
44
44
118
118
22
Btu/hr)
(150)
(150)
(150)
(200)
(60)
(150)
(150)
(150)
(400)
(400)
(75)
% S
Coal
0.6
0.6
0.6
0.6
0.6
0.6
0.6
0.6
0.6
2.3
2.3
S02
Removal
90
75
50
75
75
90
75
50
70
70
70
Type
sorbent
Sodium
Sodium
Sodium
Sodium
Sodium
Lime
Lime
Lime
Lime
Lime
Lime
Sorbent
Stoichiometry
1.1
0.8
,0.5
0.8
0.8
2.0
1.2
0.65
1.2
1.2
1.2
g/s
69
69
69
113
11
69
69
69
226
401
24
Fly_ ash
(Ib/hr)
(546)
(546)
(546)
(896)
(84)
(546)
(546)
(546)
(1791)
(3182)
(187)
Percent
of total
52
60
70
65
37
51
60
71 "
66
55
27
Desulfurization
g/s
63
43
30
61
18
66
45
27
118
328
61
products
(Ib/hr)
(501)
(341)
(239)
(482)
(143)
(520)
(359)
(218)
(934)
(2600)
(488)
Percent
of total
48
40
30
35
49
40
29
34
45
-------
Secondary Water Pollution
Secondary water pollution impacts from the disposal of lime/limestone,
double alkali, and spray drying process solid wastes can occur primarily by
landfill runoff, and/or leaching. Currently, there is very little data
existing on the secondary water effects from spray drying wastes due to the
early development stage of the technology. However, because of the similarity
of the chemical composition of the products from a lime spray drying process
and the sludge produced from a lime/limestone process, it has been suggested
that their disposal problems may also be similar. Consequently, the following
discussion on water impacts from calcium based sludge producing processes
should generally hold for calcium based spray drying operations. However,
the waste product from a sodium based spray drying system may require special
handling because of its increased solubility. Acceptable disposal practices
for spray drying products is an area currently undergoing further investigation.
Effects on surface waters by either runoff or overflow from FGD waste
disposal operations are minimized because the runoff/overflow is mixed with
the natural surface water and diluted by incident rainfall. The magnitude of
this impact will depend on site-specific hydrological conditions. This
impact, however, is not receiving as much emphasis as groundwater contamination
because of decreasing probability of incidences occurring. This is due to
zero discharge measures such as recycle of liquor from dewatering and ponding
operations or treatment of concentrated bleed streams. Only in the event of
unusually heavy rainfall or flooding conditions would this become a significant
problem. Present regulations would prevent contaminated runoff from any
storm event less than the 10 year-24 hour storm to be discharged without
receiving treatment at least for solids and pH control.
In rare cases where upset conditions or flooding occurs and water needs
to be purged from the holding pond, the purge streams will be saturated with
dissolved calcium salts and may contain high concentrations of chloride,
6-27
-------
sodium, and other ions. Currently available water treating technology as
discussed in Section 6.2.2 would be applicable to treating any purge liquors
from sludge holding ponds.
Another potential impact of limestone and double alkali sludges is that
the surrounding water quality can be impacted through leaching and percolation
of soluble components of the solid waste into the groundwater system. Some
of the liquor disposed of in unlined sludge ponds and landfills will either
evaporate, percolate down through the sludge and subsoil and potentially to
an underlying aquifer, or overflow the pond. Additional rainfall inflow will
add to the total liquor volume of the contents unless covered. In unlined
pond/landfills, as the liquor moves through the sludge solids, the dissolved
solids components of the liquor and species leached from the solids are
carried into the underlying subsoil. The leachate rate of any given pollutant
is dependent on many factors such as: its solubility in the liquor percolating
through the material, the permeability of the waste, the areas of the landfill
or pond, the permeability of the soil, rainfall, and the hydraulic head.
The permeability of ponded or landfilled material is a measure of the
rate at which leachate passes through a disposal site. Lower permeabilities
result in lower leaching rates. This parameter is generally a function of
the particle or grain size, void ratio, shape and arrangement of pores, the
degree of saturation, fluid viscosity and temperature. Permeability directly
affects the volume of leachate and, in case of high permeability, can result
in physical abrasion of the fill itself.
—If —5
Permeability coefficients for various FGD sludges range from 10 to 10
cm/sec, as shown in Table 6.2-15. According to these investigators, perme-
ability is generally lower for sulfite sludges although close control of
gypsum formation in a dual alkali system can yield a low permeability as
shown by the ADL pilot plant sulfate dual alkali sludge.18 Table 6.2-16 gives
a comparison of permeabilities of settled or drained sludges with compacted
sludges. In general, permeability was shown to decrease with compacted
6-28
-------
TABLE 6.2-15. COEFFICIENTS OF PERMEABILITY FOR FGD SLUDGES
22
Sludge
ADL Pilot Plant
(sulfate dual alkali)
Shawnee TVA Plant
(sulfite-rich lime)
Shawnee TVA Plant
(sulfite-rich limestone)
Duquesne Phillips Plant
(sulfite-rich)
SCE Mohave Plant
(100% sulfate)
% Solids
72.0
45.0
50.0
55.0
67.0
64.0
73.0
73.0
85.0
Coefficient of permeability
(cm/sec)
2.1 x 10~5
2.0 x 10"1*
5.0 x 10 5
2.5 x 10"1*
6.0 x 10~5
1.5 x 10~4
6.0 x 10 5
6.0 x 10"1*
1.5 x 10 "
TABLE 6.2-16. PERMEABILITIES OF FGD SLUDGES
23
Settled
Location
Eastern
Eastern
Western
Eastern
Process
Void
ratio
or drained
Coef. of perm.
(cm/sec)
Compacted
Void
ratio
Coef. of perm.
(cm/sec)
Limestone
Sample
Sample
Lime
Sample
Sample
Sample
1
2
1
2
3
1
2
1
1
1
.53
.07
.83
.65
.25
1
3
2
6
1
x 10 |!
x 10
x 10" g
x 10
x 10
1
1
1
1
0
.24
.56
.68
.42
.97
8
1
5
1
7
x
x
X
X
X
10~5
10 5
10~5
10 5
10~5
Limestone
Sample
Sample
Sample
1
2
3
0
1
0
.96
.20
.75
3
2
8
X 10 5
x 10 ,
x 10
0
1
0
.63
.20
.50
1
1
9
X
X
X
10~5
10'5
10~5
Dual Alkali
Sample
Sample
1
2
5
2
.11
.19
8
2
x 10 I
x 10
4
1
.17
.95
3
8
X
X
10~5
10~5
Western
Dual Alkali 2.77
1 x 10
- 3
2.61
1 x 10
6-29
-------
sludge.19'20 '21 Permeability can also be reduced by addition of fly ash, lime,
and/or commercial fixative as shown in Tables 6.2-17 and 6.2-18.
The composition of the leachate is a function of several factors including
chemical composition of the sludge, pH, solubility of the individual species
present, and age of the disposal site. The nature of the leachate evolving
from untreated sludge can be judged by analysis of liquors associated with
scrubber samples (especially scrubber sludge discharge liquors), laboratory-
generated leachate results, or actual field leachate samples. Since fly ash
will be collected and/or disposed of and mixed with the scrubber sludge in
many cases, it is likely that many of the potentially leachable elements
originate with the ash.
Although the compositions of sludge and sludge-ash leachates can vary
significantly, an overview of the nature of sludge liquors can be seen in
Table 6.2-19 and 6.2-20. These data represent liquors from many types of FGD
systems, coal sources and geographic regions. Chloride and sulfate values
are high and may represent the most significant problem in disposing of FGD
sludges. Since these liquors have to be diluted by rainwater before release
to the environment, the value of detailed comparisons is limited. TDS levels
ranged as high as 95,000 mg/£ . The pH's reported were mostly all neutral or
alkaline. Many elements varied by as many as three orders of magnitude from
sludge to sludge, including As, Be, Cr, Cu, Pb, Mn, Mo, Ni, Se, Na, and Zn.
From the scattering of data presented, it was not possible to identify any
correlatable parameters from adherent sludge liquor analyses. Some indication
was evident that trace element levels may be higher in eastern sludge.
Little documentation is available to determine whether or not leachate
from operating ash disposal sites is a potential source of pollution. However,
this is not to be interpreted that a pollution problem does not exist. The
composition of the fly ash, the concentration of soluble constituents, the
liquor pH, the trace metal solubility, soil attenuation, and/or the lateral
water movement in underlying aquifers may be such that the problem area has
6-30
-------
TABLE 6.2-17. COEFFICIENTS OF PERMEABILITY FOR FGD SLUDGES'
TREATED WITH FLY ASH AND/OR CEMENT
as
Location
Eastern
Eastern
Eastern
Western
Western
Sludge
Sludge
Sludge
Sludge
Process
Lime
Limestone
Dual
alkali
Limestone
Dual
alkali
Sludge
+ 5% cement
+ 65% fly ash
+ 35% fly ash + 5% cement
+ 39% fly ash + 5% lime
TABLE 6.2-18. EFFECT
Compacted
untreated Process "A"
1 x 10~5 2 x 10~6
8 x 1(T5
3 x 10~5 1 x 10~7
1 x 10~5 4 x 10~7
7 x 10~5 9 x 10~7
Coef. of perm.
% Solids (cm/sec)
66 7.6 x 10~6
79 7.2 x 10~6
75 4.0 x 10~6
75 5.6 x 10~6
26-
OF SLUDGE TREATMENT ON PERMEABILITY
Coefficient of permeability
(cm/sec)
Process "B" Process "E" Process "F" Process "G"
2 x lO"4 8 x 10"^
1 x ICf 5 3 x 10~6 - 5 x 10~5
5 x 10~5 5 x 1CT11 - 1 x ICf"
4 x 10~5 4 x 10~8 5 x 10~6 1 x ICT4
7 x 10~5 - x 10~7 - 4 x 10~5
-------
TABLE 6.2-19.
EQUILIBRIUM CONCENTRATIONS OF TRACE ELEMENTS IN FGD SLUDGE
27
LEACHATE (in ppm)
u>
so
Coal source: sulfur
content, % scrubber
absorbent: ash collection
PH
Antimony
Arsenic
Barium
Beryllium
Boron
Cadmium
Chromium
Copper
Fluorine
Germanium
Lead
Manganese
Mercury
Molybdenum
Nickel
Selenium
Vanadium
Zinc
Station 1 sludge
Western low lime
ESP upstream
8.5
0.014
<0.002
2
0.002
2.6
0.0005
0.001
0.031
31.5
<0.01
0.0056
<0.002
0.0005
0.063
<0.05
0.045
<0.1
0.005
Station 4 sludge
Eastern 4 limestone
ESP upstream
9.7
0.013
<0.002
<0.3
0.001
6.3
<0.001
0.011
0.045
8.7
<0.01
0.0033
<0.002
0.001
0.061
<0.05
0.0096
<0.1
0.052
Station 5 sludge
Eastern 5 limestone
with sludge
8.4
0.035
0.03
<0.3
0.002
0.96
0.002
<0.001
<0.005
7.6
0.02
0.0061
<0.002
0.0008
0.075
<0.05
0.016
<0.1
<0.005
Station 1 lime
-
12.6
0.016
<0.002
<0.3
0.001
0.22
<0.001
0.004
0.013
1.2
<0.01
0.0027
<0.002
0.002
0.011
<0.05
0.0005
<0.1
0.11
Source: Holand et al., 1975
Note: Underscored values indicate in excess of National Interim Primary Drinking Vater Standards (NIPDWS).
-------
TABLE 6.2-20.
LEVELS OF CHEMICAL SPECIES IN FGD SLUDGE LIQUORS
AND ELUTRIATES28
Eastern coals
Species
Antimony
Arsenic
Beryllium
Boron
Cadmium
Calcium
Chromium
Cobalt
Copper
Iron
Lead
crs Manganese
1
OJ Mercury
Molybdenum
Nickel
Selenium
Sodium
Zinc
Chloride
Fluoride
Sulfate
TDS
PH
Range in liquor
(ppm)
0.46-1.6
-0.004-1.8
-0.0005-0.05
41
0.004-0.1
470-2,600
0.001-0.5
-0.002-0.1
0.002-0.4
0.02-0.1
0.002-0.55
-0.01-9.0
0.0009-0.07
5.3
0.03-0.91
-0.005-2.7
36-20,000"'"
0.01-27
470-5,000
1.4-70
7 20-30,000"!"
2, 500-70, OOO1"
7.1-12.8
Median
1.2
0.020
0.014
41
0.023
700
0.020
0.35
0.015
0.026
0.12
0.17
0.001
5.3
0.13
0.11
118
0.046
2 , 300
3.2
2,100
7 , 000
-
Total no. of
observations
4
15
16
1
11
15
15
3
15
5
15
8
10
1
11
14
6
15
9
9
13
-
Western coals
Range in liquor
(ppm)
0.09-0.22
-0.004-0.2
0.0006-0.14
8.0
0.011-0.044
240-(%45,000)*
0.024-0.4
0.1-0.17
0.002-0.6
0.42-8.1
0.0014-0.37
0.007-2.5
<0. 01 -0.07
0.91
0.005-1.5
<0. 001-2. 2
0.028-0.88
1,700-4 3, OOO1"
0.7-3.0
2, 100-18, 500+
5,000-95,000*
2.8-10.2
Median Total no. of
(ppm) observations
0.16
0.009
0.013
8.0
0.032
720
0.08
0.14
0.20
4.3
0.016
0.74
0.01
0.91
0.09
0.14
0.18
1.5
3,700
12,000
-
2
7
7
1
7
6
7
2
7
2
7
6
7
1
6
7
2
7
2
3
7
3
*Levels of soluble chloride components in sludges are dependent upon the chloride-to-sulfur ratio
in the coal. The highest levels shown are single measurements for a western limestone scrubbing
system operating in a closed-loop using cooling tower blowdown for process makeup water.
^Levels of soluble sodium salts in dual alkali sludge (filter cake) depend strongly on the degree
of cake wash. The highest levels shown reflect single measurements on an unwashed dual alkali
filter cake.
Source: Lunt, et al. 1977
-------
not been identified. Information on the chemical properties of sulfur oxide
sludge indicate.s a potential pollution problem and a need for careful site
selection, and in some cases installation of a liner or drainage system.
Results of field studies now in progress have not indicated that serious
problems exist with respect to FGD sludge leachate, particularly so with
regard to trace elements. It can be assumed that FGD sludge landfills, at
least at this time, involve placement of chemically treated wastes. In the
future, however, landfilling of gypsum-type sludges generated under forced
oxidation conditions might be feasible. Chemically fixed sludges have been
found to generate leachates lower in IDS over short terms, but long term
behavior is still a matter of concern.
Wellman-Lord Process
Table 6.2-21 presents the results of the calculations to estimate the
quantity of solid waste materials produced in a Wellman-Lord process applied
to the standard size boilers. As can be seen, the quantity of waste material
is quite small varying from about 1 to 20 g/sec (9 to 180 Ib/hr) depending
upon the boiler size and the coal sulfur content. This material is essentially
pure Na2SO^ and Na2S03 for which a market may exist in the wood pulping and
glass making industries. Consequently, due to the small waste stream size
and its potential marketability, no significant solid waste impacts are
foreseen for the boiler sizes in question. If this material can not be sold
it can be disposed of in a lined and covered landfill area.
TABLE 6.2-21. SOLID WASTE PRODUCTION FOR THE WELLMAN-LORD PROCESS
Waste production
Percent
Boiler size and type removal
8.8 MWt(30xl06 Btu/hr)
Underfeed Stoker 90
22 MWt(75xl06 Btu/hr)
Chaingrate Stoker 90
58.6 MWt(200xl06 Btu/hr)
Pulverized Coal 90
3.5% S easterm coal 0.6% S western coal
g/s (Ib/hr)
4.4 (34.2)
20.4 (162)
22.6 (179)
g/s (Ib/hr)
1.2 (9.5)
3.0 (23.7)
7.9 (62.4)
6-34
-------
6.2.4 Environmental Impact on Modified Facilities
The environmental impact of applying FGD controls to modified or recon-
structed boilers will essentially be the same as for newly constructed boilers.
Consequently, previous discussions of air, liquid, and solid waste impacts
are applicable to modified and reconstructed facilities.
6.3 IMPACT OF CONTROLS FOR OIL-FIRED BOILERS
The environmental impact of FGD systems on oil-fired boilers will be
similar to that for coal-fired boilers. The major difference will be that
the oil-fired boilers will generally burn a lower sulfur fuel than the 3.5
percent sulfur eastern coal which will result in lower air, liquid, and solid
emissions. The standard residual oil-fired boiler burns a 3.0 percent sulfur
oil that produces uncontrolled SOZ emissions of 213.4 kg/hr (471.0 Ib/hr).
The standard coal-fired boilers of the same size produced uncontrolled emis-
sions of 283.2 kg/hr (845,9 Ib/hr) for the 3.5 percent sulfur eastern coal
and 8.67 kg/hr (178.1 Ib/hr) for the 0.6 percent sulfur western coal. Conse-
quently, all air, liquid, and solid emissions for the standard residual oil-
fired boiler will be approximately half those of the high sulfur eastern coal
and 2 1/2 those of the western coal. Most of the treatment methods discussed
in the previous sections are equally applicable to oil-fired FGD applications,
however, sludge stabilization techniques using fly ash from coal will not be
practical for oil-fired installations.
6-35
-------
REFERENCES
1. Tuttle, J., et al. EPA Industrial Boiler FGD Survey: Fourth Quarter
1978. Final Report. EPA Contract No. 68-02-2603, Task 45, EPA 600/7-
78-052c. Cincinnati, OH. PEDCo Environmental, Inc. November 1978.
2. Federal Register 43(182), Book 2, (1978), p. 42170.
3. Ibid.
4. Ibid.
5. Radian Corporation. Stack Gas Reheat Evaluation. Draft Final Report.
DCN 78-200-206-25, EPA Contract No. 68-02-2642. Austin, TX. September,
1978. p. 3.
6. Sugarek, R.L. and T.G. Sipes. Controlling S02 Emissions from Coal-Fired
Steam-Electric Generators: Water Pollution Impact. 2 Vols. EPA
Contract No. 68-02-2608, EPA 600/7-78-045a,b. Austin, TX. Radian
Corporation. March 1978. p. 196.
7. Ibid., 197.
8. Tuttle, et al., op.cit.., pp. 19-21.
9. Noe, David M. Memorandum to J.C. Dickerman. Cincinnati, Ohio. PEDCo
Environmental Specialists, Inc. June 12, 1979.
10. Kaplan, Steven M. and Karsten Felsvang. Spray Dryer Absorption of S02
From Industrial Boiler Flue Gas. Presented at the AIChE 86th National
Meeting. Houston, TX. April 1979.
11. Janssen, Kent E., and R.L. Eriksen. Basin Electric's Involvement With
Dry Flue Gas Desulfurization. Presented at the EPA Symposium on Flue
Gas Desulfurization. Las Vegas, Nev. March 1979.
6-36
-------
12. Mahloch, Jerome L. "Chemical Properties and Leachate Characteristics
of FGC Sludges." Paper No. 64a. Presented at the AIChE Symposium,
Atlantic City, NJ. 29 August - 1 September 1976.
13. Crowe, J.L., G.A. Holliden, and Thomas Morasky. "Status Report of
Shawnee Cocurrent and Dowa Scrubber Projects and Widows Creek Forced
Oxidation." Presented at the Industry Briefing Conference, Results
of EPA Lime/Limestone Wet Scrubbing Test Programs. Raleigh, NC.
August 1978.
14. Christman, R.C., et at. Evaluation of Dry Sorbents and Fabric Filtration
for FGD. Draft Report. EPA Contract No. 68-02-2165, Task 10. Vienna,
VA. TRW, Environmental Engineering Division. Undated. p. 52.
15. Janssen, op.c'Lt.
16. Davis, R.A., J.A. Meyler, and K.E. Gude. Dry S02 Scrubbing at Antelope
Valley Station. Presented at the American Power Conference. April 1979.
17. Janssen, op.c'Lt.
18. Hagerty, D. Joseph, C. Robert Ullrich, and Barry K. Thacker. "Engineer-
ing Properties of FGD Sludges." Geotechnical Practice for the Disposal
of Solid Waste Materials. Ann Arbor, MI. June 1977- Conference
Proceedings. New York. ASCE. 1977.
19. Leo, P.P. and J. Rossoff. Control of Waste and Water Pollution from
Power Plant Flue Gas Cleaning Systems: First Annual R and D Report.
EPA 600/7-76-018, EPA Contract No. 68-02-1010. El Segundo, CA.
Aerospace Corporation. October 1976. p. 30.
20. Mahloch, op.c'Lt.
21.. Johnson, Sandra L., and Richard R. Lunt. "Mine Disposal of FGD Waste."
Presented at the Fourth Symposium on Flue Gas Desulfurization. Holly-
wood, FL. November 1977. p. 12.
22. Hagerty, Ullrich, and Thacker, op.c'Lt.
23. Leo and Rossoff, op.dt.-, p. 65.
24. Mahloch, op.oit.
25. Hagerty, Ullrich, and Thacker, op.cit.
26. Mahloch, op.c'Lt.
6-37
-------
27. Holland, W.F., et al. Environmental Effects of Trace Elements in the
Pond Disposal of Ash and Flue Gas Desulfurization Sludge. PB252 090,
EPRI-202. Austin, TX. Radian Corporation. September 1975.
28. Johnson and Lunt, op.ci-t.
6-38
-------
SECTION 7
EMISSION SOURCE TEST DATA
7.1 Introduction
The objective of this chapter is to present continuous monitoring data
for FGD systems applied to industrial boilers that illustrate system per-
formances during long term operations. These data would also show variations
in S02 removal for 8 hour, 24 hour, and 30 day emission averaging times.
There are several FGD systems currently being used in the United States
to treat flue gases from industrial boiler; however, little continuous monitor-
ing data were found to meet the objectives of this chapter. Discrete data
sets from three EPA sponsored test programs are available which illustrate
FGD system performance capabilities for industrial boiler applications.1'2'3
Data from these FGD systems, listed in Table 7.1-1, are presented and dis-
cussed in this section.
7-1
-------
TABLE 7.1-1. SOURCE MONITORING DATA FROM INDUSTRIAL BOILER FGD SYSTEMS
Boiler location FGD type Data status
Rickenbacker A.F.B. Lime/Limestone Test data collected from
Columbus, Ohio March 1976 to May 1977.
Firestone Tire and Rubber Co.* Double Alkali Test data collected from
Pottstown, Pennsylvania 9/27/77 to 10/8/77-
General Motors* Double Alkali Test data collected from
Parma, Ohio 8/19/74 to 9/13/74, 2/17/75
to 3/14/75, and 4/19/76 to
5/14/76.
*These systems had the ability to burn either coal or oil. Data for the
Firestone installation were collected while burning both fuel types.
The best examples of 30-day continuous monitoring data for industrial
boiler FGD systems presently available are from a sodium throwaway system and
a Wellman-Lord system presently operating on industrial boilers in Japan. Dr.
Jumpei Ando graciously provided Radian with these data from 71 MW and 135 MW
G c
installations, respectively, each firing 3 percent sulfur heavy oil. Plots
of these data are shown in Section 7.4.
Continuous monitoring data from seven coal-fired utility FGD systems have
recently been reported by EPA. 5>6'7 These systems include both regenerable and
nonregenerable FGD processes and illustrated that S02 removals in excess of 85
percent are achievable over a 30 day operating period using technology that is
available today. In addition to these data, EPA is currently obtaining contin-
uous monitoring data from one additional lime/limestone systems. Also, data
from performance tests conducted by an outside contractor on one sodium throw-
away system are available. Table 7.1-2 presents a summary of the status of
data gathering efforts for these nine utility FGD systems. Because of the
7-2
-------
TABLE 7.1-2. CONTINUOUS MONITORING DATA FROM
UTILITY BOILER FGD SYSTEMS
Utility
Louisville Gas
& Electric
Pennsylvania
Power Co .
Philadelphia
Electric Co.
Northern Indiana
Public Service
Co.
Columbus and
Southern Ohio
Kansas City
Power & Light
Kansas Power
& Light
Plant site
Cane Run
Unit No. 4
Bruce Mansfield
Unit No. 1
Eddystone
Unit No. 1
D.H. Mitchell
Unit No. 11
Conesville
Unit No. 5
LaCygne
Unit No. 1
Lawrence
Unit No. 4
FGD type
Lime
Lime
Magnesium
Oxide
Wellman-Lord
Lime
Limestone
Lime
Data status
89 days of contin-
uous data
31 days of contin-
uous data
5 days of contin-
uous data
56 days data reported
including 41 days con
tinuous data
34 intermittent
days of data
30 day monitoring
test in progress
22 days continuous
data reported
Tennessee Valley
Authority
Nevada Power
Company
Shawnee Test
Facility
Reid Gardner
Units No. 1 and 2
Limestone
Sodium
Throwaway
42 days continuous
data reported
6 days data reported
including 3 days
continuous data
7-3
-------
similiarity in flue gases from utility and industrial boilers, the utility
data should be useful to indicate the performance levels that can be expected
of industrial sized FGD systems.
7.2 Emission Source Data for Coal-Fired Boilers
In this section, brief descriptions of the test facilities will be pre-
sented, the test results will be discussed, and test procedures will be described,
More detailed descriptions and discussions of the FGD processes being considered
were presented in Section 2.
Data for a lime scrubbing process were collected at the R-C/Bahco scrubbing
system installed at the Central Heat Plant at Rickenbacker Air Force Base near
Columbus, Ohio. The heat plant houses eight coal-fired hot water generators
with a total fuel burning capacity of approximately 330 x 10 Btu/hr. These
stoker-fired generators burn 2.5 to 3.5% sulfur Ohio coal with an average
heating value of 11,300 Btu/lb. 8
The R-C/Bahco system was designed to treat up to 108,000 acfm of flue gas
generated at the peak winter load of approximately 200 x 106 Btu/hr. The
system, which must operate over the relatively narrow range of gas flow of
35,000 to 50,000 scfm, has an essentially unlimited turndown capability for
handling flue gas by mixing air with the flue gas at low boiler loads. This
allows the system to handle seasonal load variations from 20 to 200 x 106
Btu/hr, SOa concentrations from 200 to 2000 ppm and particulate loadings of up
to 2 gr/scfd.9
Data for double-alkali systems were collected from an F.M.C. system
located at Firestone Tire and Rubber Company's Pottstown, Pennsylvania facility.
The test boiler for the FMC system is one of four comprising a steam plant
which supplies process steam and heating steam for the facility. The boiler
is one of three which operates at a fairly constant rate of 45,400 kg/hr
(100,000 Ib/hr) of steam. Process steam demand is relatively steady, since
the plant operates 24 hours per day, seven days per week. Fluctuations in
7-4
-------
heating load are satisfied by either boosting steam generation rates on these
boilers or by operating the fourth boiler (No. 1) . The steam generation rate
of Boiler No. 1 varies from zero to approximately 22,700 kg/hr (50,000 Ib/hr)
of steam.10
The boiler was installed in 1958 and was originally designed as a coal-
fired unit. It was converted to fire either coal or fuel oil in 1967. The
two fuels are usually not burned simultaneously except when converting from
oil to coal firing. The coal is ignited by continuing oil firing until a
stable coal flame is obtained. Oil and coal can be fired simultaneously to
maintain acceptable steam generation rates if coal with a low heat content is
burned.: 1
The flue gases are treated by an air pollution control system consisting
of multiclone units for particulate control and a pilot double alkali FGD unit
designed by FMC. All of the flue gas passes through the multiclones. The
stream then is split and two-thirds of the flue gas ducted to the stack. The
other one-third is ducted to the pilot FGD system. The boiler has no N0v
X.
controls.
Data for double-alkali systems were also collected from the General
Motors system located at their Chevrolet Parma, Ohio plant. Scrubbers were
started up in March 1974 on each of the four boilers in the steam plant. Two
boilers have steam generation rates of 27,300 kg/hr (60,000 Ib/hr) and two
have rates of 45,500 kg/hr (100,000 Ib/hr). All boilers are spreader stoker-
fired, normally burning high sulfur (2-3% S) eastern coal and, on occasion,
lower sulfur waste oil.
The flue gases are first treated by existing mechanical dust collectors
on each boiler for primary particulate control. Then the gases enter the
scrubbers which are three-tray columns since only modest reduction of particu-
lates is required.
7-5
-------
7.3 Data Presentation
Performance data for the R-C/Bahco FGD system is presented in three
tables which represent three separate test periods in the overall program to
evaluate the performance of the R-C/Bahco system. Table 7.3-1 presents the
initial SO2 removal data which were taken during startup and early operation
of the system (Spring 1976). These initial removal efficiencies ranged from
87 to 99 percent and were reported to be affected primarily by the lime/S02
stoichiometry.
A series of lime screening tests were then carried out from December 1976
until February 1977. The objective of these tests was to identify the effects
of major process variables upon S02 removal. Consequently, various independent
process variables such as gas flow rate, pressure drop, liquid pumping rate,
stoichiometry, and slurry concentration were varied in a series of 21 discrete
tests to examine their effect on SO2 removal. Table 7.3-2 presents a summary
of this test data. These tests indicated that the only variable of significance
affecting S02 removal at the 95 percent confidence level was lime/S02 stoichio-
metry. Conclusions from the lime testing were:
1) Virtually any desired S02 removal efficiency can be achieved in the
R-C/Bahco scrubber, when using lime, simply by adjusting the lime/S02
stoichiometry.
2) Lime utilization approaching 100 percent is achieved at stoichiometric
ratios's of up to about 0.9. At stoichiometric ratios up to 1.1,
producing up to 99 percent S02 removal, lime utilization is above 90
percent. l ^
In the Spring of 1977, a second series of variable tests were run using
limestone as the sorbent. This test series was modeled after the lime screening
test and examined the effects of the same process variables. Table 7.3-3
presents a summary of the limestone screening test data. Conclusions from the
7-6
-------
TABLE 7.3-1. S02 REMOVAL EFFICIENCY DATA
12
1976
Date
3/30
4/8
5/19
5/26
5/26
5/26
5/27
5/27
Coal
sulfur
content
3.24
3.24
3.25
2.64
2.64
2.64
2.01
2.01
Coal
firing rate
(MM Btu/hr)
132.
115.
47.
54.
52.
43.
44.
44.
2
2
9
0
8
3
8
3
Inlet SO 2
concentration
(ppm)
1,392
1,200
454
555
489
401
327
323
Outlet S02
concentration
(ppm)
156
45
24
5
8
8
5
5
SO^ removal
efficiency
(%)
87
95
94
99
98
97
98
98
.6
.7
.4
.0
.2
.9
.3
.2
SO 2 emission
rate
(Ibs/MM Btu)
0.
0.
0.
0.
0.
0.
0.
0.
,621
,21
.29
.045
.084
.095
.061
.061
Lime
utilization
(%)
100.0
94.
98.
90,
90,
91
94,
95,
,0
.8
.3
.5
.2
.2
.4
Note: Removal efficiencies were corrected for increased outlet gas volume due to water evaporation in the scrubber.
-------
TABLE 7.3-2.
LIME TEST DATA SUMMARY - RICKENBACKER AIR
FORCE BASE - R-C/BAHCO LIME SCRUBBING SYSTEM13
Test No. and date
1
2R*
3
4
5R*
6
7R*
8
9
10
V 11
12
13
14R*
15
16
17
18
19
20
21
12/16/76
2/17/77
12/17/76
12/15/76
2/17/77
2/16/77
2/18/77
2/17/77
12/15/76
12/18/76
12/18/76
12/19/76
12/19/76
2/13/77
12/13/76
2/13/77
2/15/77
12/20/76
2/15/77
2/10/77
2/14/77
Coal
firing rate
(Ib/hr)
14,098
13,213
14,994
13,991
13,361
13,214
13,368
13,229
12,712
13,229
14,627
15,522
10,789
14,983
16,639
10,448
15,878
20,313
12,539
8,954
10,430
Average
lime/S02
stoichiometry
0.94
0.60
1.00
0.98
0.65
0.93
0.60
0.36
0.97
0.99
0.98
0.99
0.81
0.54
1.06
0.66
0.92
1.08
0.67
0.60
0.85
Flue gas
flow rate
(103 acfm @ 120°
56.2
46.2
47.6
65.4
68.1
46.3
47.7
66.1
68.6
5'5.7
57.9
70.2
67.6
45.9
49.3
70.2
64.5
50.4
45.2
53.4
54.9
Average S02 (ppm)
F) Inlet
1,110
1,500
1,000
992
1,140
1,100
1,350
1,275
910
1,075
1,150
842
940
1,200
1,500
1,095
1,000
1,155
1,045
950
1,075
Outlet
104
550
28
40
365
81
494
742
58
41
70
40
174
546
25
340
115
14
348
370
172
Overall
SOz control
efficiency Average S02 emissions
(percent)
89.4
59.6
96.8
95.4
64.5
92.5
59.8
35.9
92.7
95.6
93.3
94.6
78.8
54.5
98.1
65.8
87.5
98.7
63.4
57.8
82.3
ng/J (lb/105 Btu)
146.2
675.0
30.1
60.2
889.6
103.2
640.6
1298.4
103.2
60.2
94.6
64.5
386.9
610.5
25.8
834.0
180.6
12.9
451.4
451.4'
318.1
0.34
1.57
0.07
0.14
2.07
0.24
1.49
3.02
0.24
0.14
0.22
0.15
0.90
1.42
0.06
1.94
0.42
0.03
1.05
1.80
0.74
*The "R" designation indicates a repeat test
-------
TABLE 7.3-3.
LIMESTONE TEST DATA SUMMARY - RICKENBACKER
AIR FORCE BASE - R-C/BAHCO SCRUBBING SYSTEM 15
Test No. and date
Coal
firing rate
(Ib/hr)
Average
stoichiometry
Flue gas
flow rate
(103 acfm @ 120°
Average
F) Inlet
S02 (ppm)
Outlet
Overall
SO 2 control
efficiency
(percent)
Average S02 emissions
ng/J (lb/106 Btu)
(1977 dates)
38
39
40
41
42
43
44
45
46
V 47
VD
48
49
50
51
52
53
54
55
56
57
58
5/19
5/23
5/24
5/20
5/23
5/25
5/25
5/25
5/25
5/19
5/1
5/1
5/1
5/1
5/1
5/17
5/17
5/1
5/1
5/19
5/19
4312
4928
4531
3955
3338
4158
4992
5252
3984
5006
3949
4765
3958
3960
3631
5408
4165
4158
4092
3393
3966
0.80
1.53
1.55
1.42
1.08
1.41
1.88
1.30
1.14
0.86
0.95
1.19
0.63
0.96
0.72
1.19
0.94
1.38
0.59
0.94
1.01
48.8
36.0
42.8
52.2
54.2
39.5
44.9
60.2
64.0
51.2
45.4
54.2
54.2
37.0
43.3
57.5
58.9
32.6
39.3
465
45.9
438
425
350
290
300
360
410
320
295
390
410
375
254
483
425
350
375
600
525
375
335
109
30
33
31
92
23
57
35
70
90
71
90
125
102
194
71
131
107
284
69
55
75.7
88.8
81.8
92.2
72.9
94.6
80.9
89.0
69.7
86.5
76.7
84.7
79.4
90.2
65.3
84.4
64.4
84.8
50.3
81.7
81.1
438.6
77.4
107.5
146.2
528.9
81.7
189.2
146.2
412.8
326.8
288.1
335.4
636.4
322.5
786.9
266.6
649.3
296.7
958.9
335.4
223.6
1.02
0.18
0.25
0. 34
1.23
0.19
0.44
0.34
0.96
0.76
0.67
0.78
1.48
0.75
1.83
0.62
1.51
0.69
2.23
0.78
0.52
-------
1) Limestone/S02 stoichiometry and slurry pumping rate are the signifi-
cant variables contolling S02 removal efficiency.
2) A considerable excess of limestone is needed to absorbe SO2 , especi-
ally at high SOz removal rates.
3) Limestone can be used to meet the requriements for SO 2 removal at
RAFB (1.0 lb S02/105 Btu).16
Results of one week of testing at the Firestone test facility are pre-
sented in Table 7.3-4. Sulfur dioxide removal data from the Firestone facility
indicated an average scrubber efficiency of about 97 percent. Controlled SO2
emissions averaged 36.3 mg/J (0.08 pounds/10 Btu) which is less than either
existing or proposed standards for utility boilers.
The results of the three one-month intensive testing programs at the GM
Parma facility showed that the stringent level of SO2 control (90% S02 removal)
is obtainable by double alkali systems.18 Some of the process variables studied
include lime and soda ash stoichiometries, recycle pH, scrubber feed location,
filter cake washing, and solids recycle.
Results of all three test programs indicate that high SO2 removal efficien-
cies are achievable in coal-fired industrial boiler installations. Additional
continuous monitoring data will be rquired, however, to evaluate industrial
boiler FGD performance during long-term operations.
7.3.1 Test Methods
SO2 data for the Rickenbacker tests were collected using the test method
described in this section. S02 data for the Firestone tests were collected
using a pulsed fluorescence analyzer.19 Firestone test data reported in Table
7.3-4 are averages of the data collected during each test period. Continuous
monitoring for the GM tests was reported to be an UV absorption analyzer with
semi-continuous determination.
7-10
-------
TABLE 7.3-4. COAL-FIRED EMISSION SOURCE DATA
FIRESTONE TIRE AND RUBBER - FMC DOUBLE-ALKALI FGD SYSTEM
1 7
Coal
feed rate
Test No. kg/hr
200
201-1
201-2
201-3
201-4
Avg . Value
3,629
3,629
3,629
3,175
3,629
3,538
Coal
heating value
kg /kg
29,263
28,872
29,997
29,419
29,878
29,485
S02 inlet
ng/J
1,009
1,284
1,295
1,028
942
1,112
Ib/KT Btu
2.35
2.99
3.01
2.39
2.19
2.59
S02 outlet
ng/g
25.4
39.0
35.5
31.8
49.7
36.3
(lb/106 Btu)
0.06
0.09
0.08
0.07
0.12
0.08
SOz control
efficiency
(percent)
97.5
96.9
97.2
96.9
94.7
96.6
-------
The following discussion describes the test method used at Rickenbacker
A.F.B. to determine the flue gas SO2 content. This method is only approximate
and should be used only as a semiquantitative check on S02 concentrations. No
temperature or pressure corrections have been incorporated, and the method
should not be used below 100 ppm.20
Apparatus Reagents
1) 250 ml impinger with an open 1) 3% Hydrogen Peroxide
glass dip tube 2) 0>1N NaQH or 0>Q1N NaQH
2) A dry test meter 3) Methyl/0range-Xylene Cyanol
3) A source of vacuum indicator
4) 25 ml pipette
5) Vacuum tubing
6) Hose clamp
Procedure:
Inlet Samples - (i.e., 500 + ppm SO 2 ) pipette 25 ml of 0.1N NaOH into
the 250 ml impinger, add 50 ml of 3% hydrogen peroxide. Add approximately
25 ml of deionized water. Add several drops of Methyl/ Orange-Xylene
Cyanol indicator.
Draw the gas sample through the impinger at 0.1 to 0.2 ft/min. Record
the gas meter reading when the indicator turns from green to purple.
Outlet Samples - (100 to 600 ppm S02) substitute 0.01 normal NaOH for 0.1
normal NaOH in the above procedure. Follow the same procedure as above.
The following equation can be used to calculate the SO concentration:
= 10,000 x (NaOH Normality)
2 ppm Meter Volume ft 3
Note: Add the indicator within 15 minutes of running the test. If the
indicator is added at an earlier time, it may be destroyed by the hydro-
gen peroxide in the impinger.
7-1?.
-------
7.4 Emission Source Test Data for Oil-Fired Boilers
Three sets of data (including two sets of one month continuous testing)
were found to illustrate 862 emissions from an oil-fired boiler. The Firestone
facility previously discussed had provisions for firing both coal and oil.
The FMC double alkali test system was operated under both oil- and coal-fired
operations to evaluate its performances. Results shown in Table 7.4-1 indicate
the FMC system is equally effective in controlling S02 emissions from an oil-
fired boiler as it was from a coal-fired boiler. As with the coal-fired tests
at the Firestone facility, S02 emission measurements were taken with a pulsed
22
fluorescent analyzer.
The two examples of 30-day continuous monitoring data (supplied by Dr.
Ando) for industrial boiler FGD systems are shown in Figures 7.4-1 and 7.4-2.
Figure 7.4-1 is a representation of the data (during October 1978) from Kureha
Chemical's sodium throwaway process applied to a 71 MW boiler burning 3
percent sulfur heavy oil at the Nishiki plant. This process uses a packed
tower to remove SOa from the flue gas of an existing boiler (a retrofit FGD
installation). One can see from Figure 7.4-1 that the process operated at 90+
percent S02 removal for all 31 days. It should be pointed out here that the
daily average S02 removal is comprised of three readings, one taken at 6:00
a.m., 1:00 p.m., and 9:00 p.m each day. In a recent one-year time frame, from
April 1, 1977 to March 31, 1978, both the boiler and the FGD system operated
for 8516 hours. 23
Figure 7.4-2 is a representation of the data (during May 1977) from the
Wellman-Lord process applied to a new 135 MW installation burning 3 percent
sulfur heavy oil at Shindaikyowa Petrochemical's Yokkaichi plant. One can see
from Figure 7.4-2 that the process operated at 90+ percent SOz removal for all
31 days. While only one reading per day, using an infrared S02 analyzer, was
taken for the majority of this time, readings were taken every other hour for
three consecutive days to check on fluctuations. These bihourly readings
showed that there was a maximum deviation in SO2 removal during any one day of
7-13
-------
TABLE 7.4-1. OIL-FIRED EMISSION SOURCE DATA
FIRESTONE TIRE AND RUBBER - FMC DOUBLE-ALKALI FGD SYSTEM
21
Oil Oil
feed rate heat value
Test No.
202-1
202-2
202-3
202-4
Average
£ *The heat
(gas/hr) kg/kg ng/J
900
900
880
805
871
content of
* 938
* 1075
* 1085
* 874
40,741* 993
the oil burned is nearly
S02 Inlet
S02 outlet
S02 control
efficiency
lb/106 Btu ng/J lb/10b Btu
2.18
2.50
2.52
2.03
2.31
constant at
32.1
29.2
26.7
19.2
26.8
0.07
0.07
0.06
0.04
0.06
this value; individual values
96.6
97.3
97.5
97.8
97.3
were not
available.
-------
--J
I
100
90
O
01
Csl
w 80
60
70
I i i i i i i
i i i i i i i i i i i i i i
1 2 3 4 5 6 7 8 9 101.112 131415 161718 19202,122232^2526 2728 2930 31
Day
*Note: The average daily removal is comprised of 3 readings, taken
at 8-hour intervals.
Figure 7.4-1. Performance of Kureha Chemical's sodium throwaway process
applied to a 71 MW boiler burning 3 percent sulfur heavy oil at the Nishiki plant during
e October 1978.23
-------
o
e
0)
o
>%
iH
•H
ctf
n
100 r-
90
80
70
I i I I
I I I I
I I I I I I
I I I I I I I I I I I
1234
6 7
10 1112 131415 1617 18 192021 2223242526 2728293031
Day
*Note: The daily removal was taken each day at 1:00 pm. On three consecutive
days (Kay 7,8,9) readings were taken every other hour to gauge amplitude
of fluctuations.
Figure 7.4-2. Performance of Wellman-Lord process applied to a 135 MW boiler
burning 3 percent sulfur heavy oil at Shindaikyowa Petrochemical's Yokkaichi plant during
May 1977.23
-------
only about three percent. Therefore, the once a day readings for the 31
days of testing should adequately represent the operation of the system.
These good examples of continuous monitoring data are for heavy
oil-fired boilers and also for larger boiler sizes than those for the
standard industrial boilers of this study; however, the performance of
these FGD systems certainly demonstrates their ability to achieve the
stringent level of SOo control.
7-17
-------
REFERENCES
1. Biedell, E.L., et al. EPA Evaluation of Bahco Industrial Boiler Scrubber
Systerm at Rickenbacker AFR. Report No. EPA-600/7-78-115. June 1978.
2. Leavitt, C., et at. Environmental Assessment of Coal- and Oil-Firing in
a Controlled Industrial Boiler. Report No. EPA-600/7-78-164b. August
1978.
3. Interess, E. Evaluation of the General Motors' Double Alkali SO2 Control
System. Report No. EPA-600/7-77-005. January 1977.
4. Ando, Dr. Jumpei: (Chuo University, Tokyo, Japan). Private communication
with Gary D. Jones (Radian Croporation, Austin, Texas) , February 1979.
5. Kelly, W.E., et at. Air Pollution Emission Test, First Interim Report.
Vol. 1, Continuous Sulfur Dioxide Monitoring at Steam Generators. EPA
Contract No. 78-02-2818, Work Assignment No. 2, EPA/EMB Report No. 77SPP23A.
Research Triangle Park, NC. EPA, Office of Air Quality Planning and
Standards, Emission Measurement Branch. August 1978.
6. Kelly, W.E., et at. Air Pollution Emission Test, Second Interim Report.
Vol. 1, Continuous Sulfur Dioxide Monitoring at Steam Generators.
EPA/EMB Report No. 77SPP23B. March 1979.
7. Kelly, W.E., et al. Air Pollution Emission Test, Third Interim Report.
Vol. 1, Continuous Sulfur Dioxide Monitoring at Steam Generators.
EPA/EMB Report No. 77SPP23C. March 1979.
8. Biedell, E.L., op.eit., p. 22
9. Ibid., P- 22
10. Leavitt, C., op.ait., Vol. II, p. 3-1.
11. Ibid. , p. 3-1.
12. Biedell, E.L. , op.oit. , p. 57.
7-18
-------
13. Ibid., p. 166-172.
14. Ibid., p. 56-63
15. Ibid., p. 188-195.
16. Ibid., p. 64-68.
17. Leavitt, C., op.ait., VIII, p. 5-21.
18. Interess, E. , op.oit., p., 2.
19. Ibid.t VII, p. 4-24.
20. Biedell, E.L., op.sit., p. 152.
21. Leavitt, C., op.ait., VII, p. 4-24.
22. JMcZ., VIII, p. 6-9, 6-10.
23. Ando, op.oit.
7-19
-------
SECTION 8
ADDITIONAL FGD TOPICS
This section provides supplemental evaluations of the candidate control
systems previously considered for application to small industrial boilers.
Included in this section are evaluations of partial scrubbing for limestone,
double alkali, and sodium processes; and evaluations of high sulfur coal lime-
stone applications. The technologies are assessed in light of their economic,
energy, and environmental considerations in the same manner as the candidate
systems evaluated in Sections 4, 5, and 6 of this report.
Three additional scrubbing cases which treat only part of the flue gas
were considered for these processes in order to develop cost and energy
correlations for partial scrubbing. The results of material and energy ba-
lances are presented in this section for cases treating 75, 50, and 25 percent
of the flue gas from coal-fired boilers (3.5% S coal, 90% removal). To study
the effect of coal variability on scrubber operations, the costs and energy
requirements for a limestone scrubbing system applied to a 5 percent sulful
coal are included.
8.1. PARTIAL SCRUBBING
8.1.1. Control Costs
The cost information for the partial scrubbing cases was developed in
the same manner as in Section 4 of this report. Material and energy balances
were calculated for each additional case and the equipment items discussed
-------
in Section 4 were sized according to the material balance results. Equipment
cost data obtained from vendors was then used to estimate costs for each of
the FGD processes. The base cases for the partial scrubbing comparisons are
the 90 percent SOa removal cases treating 100 percent of the flue gas from
boilers burning 3.5 percent sulfur coal. The calculations in this section
were extended to evaluate the impacts of treating 75, 50, and 25 percent of
the flue gas from these boilers.
Tables 8.1-1 through 8.1-3 summarize the capital and annualized costs
for the partial scrubbing cases. These data are also plotted in Figures 8.1-1
and 8.1-2 for the sodium throwaway process to illustrate the cost sensitivity
of partial scrubbing. Throughout this discussion, the sodium throwaway pro-
cess will be used as the example to illustrate the impacts of partial scrubbing.
Tabular results of the impacts of partial scrubbing on the limestone and
double alkali processes are provided so that one can verify that the trends
shown by the sodium throwaway process are also true for the other FGD pro-
cesses being evaluated.
For the 58.6 MW (200 x 106 Btu/hr) applications, the sodium throwaway
capital costs ranged from $1,214,000 (100 percent treated) to $580,000 (25
percent treated). Annualized costs ranged accordingly from $1,207,000 to
$520,000. Capital costs were reduced roughly by one-half in going from 100
percent scrubbing to 25 percent scrubbing, a reduction by 4 in the amount of
S02 removal. The same factor is also evident for the annualized costs and
simply illustrates the economy of scale for the process. The other processes
also show similar economies of scale.
Data are provided to compare the impact of'removing 75 percent of the
S02 from all the gas versus removing 90 percent of the S02 from 75 percent of
the gas (67.5 percent removal). Results presented in Table 8.1-1 indicate
that removing 90 percent of the SO from a portion of the gas (partial scrub-
bing) is more economical than low efficiency (75%) full scrubbing. This is
because a significant portion of the process capital costs (scrubber, fans)
-------
TABLE 8.1-1. SODIUM THROWAWAY PARTIAL SCRUBBING COST SUMMARY
(Eastern 3.5% S Coal)
Boiler Size
& Type
22 MW
(75xl06Btu/hr)
Chaingrate Stoker
44 MW
oo (150xl06Btu/hr)
w Spreader Stoker
58,6 MW
(200xl06Btu/hr)
Pulverized Coal
118 MW
t6
(400x10 Btu/hr)
Pulverized Coal
Percent of
Gas Scrubbed
100
100
75
50
100
100
75
50
25
100
100
75
50
75
100
75
50
25
Percent SO 2
Remova 1
90
75
67.5
45.0
90
75
67.5
45
22.5
90
75
67.5
45
22.5
90
67.5
45
22.5
Total Capital
Investment
(io3 $)
730
705
590
500
1077
1033
890
7-20
510
.1214
1160
1080
840
580
1911
1600
1214
840
Annualized Costs
(io3 $)
646
596
540
460
1010
917
822
640
440
1207
1080
1010
740
520
2063
1625
1207
740
Percent increase
over uncontrolled
33
31
28
24
33
30
27
21
14
27
25
23
17
12
27
21
16
10
-------
TABLE 8.1-2. DOUBLE ALKALI PARTIAL SCRUBBING COST SUMMARY
(Eastern 3.5% S coal)
I
-C-
Boiler Size Percent of Gas
and Type Scrubbed
22 MW
(75 x 106Btu/hr)
Chaingrate Stoker
58.6 MW
(200 x 106Btu/hr)
Pulverized Coal
118 MW
(400 x 106Btu/hr)
Pulverized Coal
100
75
50
100
75
50
25
100
75
50
25
Percent S02
Removal
90
67.5
45
90
6 7'. 5
45
22.5
90
67.5
45
22.5
Total Capital
Investment
(103$)
960
840
750
1422
1220
1050
812
2064
1740
1420
1050
Annualized Costs
Percent Increase
(103$) Over Uncontrolled
625
530
460
1053
870
695
505
1778
1420
1053
695
32
27
24
25
20
16
12
23
18
14
9
-------
TABLE 8.1-3. LIMESTONE PARTIAL SCRUBBING COST SUMMARY
(Eastern 3.5% S Coal)
Annualized Costs
Boiler Size
& Type
22 MW
(75xl06Btu/hr)
Chaingrate Stoker
oo
1
Ln
44 MW
(ISOxlO6 Btu/hr
Spreader Stoker
58.6 MW
(200xl06 Btu/hr)
Pulverized Coal
Percent of
Gas Scrubbed
100
100
75
50
100
100
75
50
25
100
100
75
50
25
Percent S02
Removal
90
75
67.5
45
90
75
67.5
45
22.5
90
75
67.5
45
22.5
Total Capital
Investment
(103 $)
987
913
860
730
1385
1270
1180
970
730
1530
1392
1385
1120
820
(103$)
650
593
560
468
974
865
805
630
470
1155
1014
960
735
520
Percent increase
over uncontrolled
boilers
34
31
29
24
32
28
26
20
15
26
23
22
17
12
-------
2000 r
1500 -
n.
CD
o
1000 _
500 _
Full
Scrubbing
75% of
Gas Scrubbed
50% of
Gas Scrubbed
25% of
Gas Scrubbed
FGD Size, MW (106 Btu/hr)
Figure 8.1-1. Sodium Partial Scrubbing Capital Investment Costs
(3.5% S Coal, 90% Removal)
-------
2000
1500
m
o
u
-d
0)
N
a
3
I
1000 —
500 -,
Full
Scrubbing
75% of
Gas Scrubbed
50% of
Gas Scrubbed
25% of
Gas Scrubbed
FGD Size, MW (106 Btu/hr)
Figure 8.1-2. Sodium Partial Scrubbing Annual Costs (3.5% S Coal, 90% Removal)
3-7
-------
are based on gas flow rate and remain constant for a given sized application
whereas costs for partial scrubbing processes take advantage of a reduction
in the volume of gases to be treated.
8.1.2 Energy Requirements for Partial Scrubbing
Tables 8.1-4 through 8.1-6 present a summary of the energy requirements
for the partial scrubbing cases. Figure 8.1-3 presents the partial scrubbing
energy requirements for the sodium throwaway process. These cases are based
on 90 percent SOa removal from boilers burning 3.5 percent S coal. Energy
requirements were calculated from the material balances using the same cal-
culation bases presented in Table 5.2-1 of this report.
Results of these calculations show essentially a linear use of energy
with size for the various partial scrubbing cases which indicates that there
is no economy of scale or energy savings for the sodium throwaway process
for the partial scrubbing cases. However, if stack gas reheat were required
for the full scrubbing cases, the partial scrubbing cases would show a sig-
nificant energy advantage. This is because the reheat energy requirements
could be saved for the partial scrubbing cases since bypassing 19 percent
of the flue gas provides sufficient heat to maintain the stack gas exit
temperature above 353°K (175°F).
8.1.3. Environmental Impacts
Tables 8.1-7 through 8.1-9 present a summary of the wastewater and sludge
production rates for the partial scrubbing cases. Figure 8.1-4 illustrates
the partial scrubbing wastewater production rates graphically for the sodium
throwaway process. As with the economic and energy impacts, the environmen-
tal impacts are based on 90 percent SO removal from boilers burning 3.5% S
coal. Wastewater and sludge production rates were calculated using the
material balances found in Appendix A of this report.
8-8
-------
TABLE 8.1-4. SODIUM THROWAWAY PARTIAL SCRUBBING ENERGY REQUIREMENTS
(Eastern 3.5% S coal)
Boiler Size Percent of
and Type Gas Scrubbed
22 MW
(75 x 10s Btu/hr)
Chaingrate Stoker
44 MW
(150 x 106Btu/hr)
CO
1 Spreader Stoker
58.6 MW
(200 x 10s Btu/hr)
Pulverized Coal
118 MW
(400 x 10s Btu/hr)
Pulverized Coal
100
100
75
50
100
100
75
50
25
100
100
75
50
25
100
75
50
O T-
Percent S0:
Removal
90
75
67.5
45
90
75
67.5
45
22.5
90
75
67.5
45
22.5
90
67.5
45
22.5
; Energy (
KW 1
181
168
125
85
363
337
265
168
85
442
409
360
222
112
873
660
440
220
Consumption
U06Btu/hr) Ov
0.6
0.6
0.4
0.3
1.2
1.1
0.9
0.6
0.3
1.5
1.4
1.2
0.8
0.4
3.0
2.2
1.5
0.8
Percent Increase
er Uncontrolled Boiler
0.8
0.8
0.5
0.4
0.8
0.7
0.6
0.4
0.2
0.8
0.7
0.6
0.4
0.2
0.8
0.6
0.4
0.2
-------
TABLE 8.1-5. DOUBLE ALKALI PARTIAL SCRUBBING ENERGY REQUIREMENTS
(Eastern 3.5% S Coal)
CO
I
o
Boiler Size Percent of
& Type Gas Scrubbed
22 MW
(75xl06Btu/hr)
Chaingrate Stoker
58.6 MW
(200xl06Btu/hr)
Pulverized Coal
118 MW
(400xl06 Btu/hr)
Pulverized Coal
100
75
50
100
75
50
25
100
75
50
25
Percent SOz
Removal
90
67.5
45
90
67.5
45
22.5
90
67.5
45
22.5
Energy Consumption
kW (106Btu/hr)
130
90
63
307
234
157
82
606
460
307
157
0.4
0.3
0.2
1.0
0.8
0.5
0.3
2.1
1.6
1.0
0.5
Percent increase
over uncontrolled
boiler
0.5
0.4
0.3
0.5
0.4
0.3
0.2
0.5
0.4
0.3
0.1
-------
TABLE 8.1-6. LIMESTONE PARTIAL SCRUBBING ENERGY REQUIREMENTS
(Eastern 3.5% S coal)
00
i
Boiler Size
& type
22 MW
(75 x 106 Btu/hr)
Chaingate Staker
44 MW
(150xl06 Btu/hr)
Spreader Stoker
58.6 MW
(200x10 Btu/hr)
Pulverized Coal
Percent of
Gas Scrubbed
100
100
75
50
100
100
75
50
25
100
100
75
50
25
Percent S02
Removal
90
75
67.5
45
90
75
67.5
45
22.5
90
75
67.5
45
22.5
Energy
KW
387
253
290
195
111
507
570
380
195
912
599
770
510
245
Consumption
(106 Btu/hr)
1.3
0.9
1.0
0.7
2.7
1.7
1.9
1.3
0.7
3.1
2.1
2.6
1.7
0.8
Percent Increase
over uncontrolled
boiler
1.7
1.2
1.3
0.9
1.8
1.1
1.3
0.9
0.5
1.6
1.1
1.3
0.9
0.4
-------
looo r
750 I
I
to
s
o
o
0)
e
w
500 \-
250
Full
Scrubbing
75% of
Gas Scrubbed
50% of
Gas Scrubbed
25% of
Gas Scrubbed
79
(100)
FGD Size, MW (10G Btu/hr)
Figure 8.1-3. Sodium Partial Scrubbing Energy Consumption
(3.5% S Coal, 90% removal)
5-12
-------
TABLE 8.1-7.
SODIUM THROWAWAY PARTIAL SCRUBBING WASTEWATER PRODUCTION RATES
(Eastern 3.5% S coal)
CO
I
Boiler Size
& Type
22 MW
(75 x 10EBtu/hr)
Chaingate Stoker
44 MW
(150 x 106 Btu/hr)
Spreader Stoker
58.6 MW
(200 x 10s Btu/hr)
Pulverized Coal
118 MW
(400 x 105 Btu/hr)
Pulverized Coal
Percent of gas
Scrubbed
100
100
75
50
100
100
75
50
25
100
100
75
50
25
100
75
50
25
Percent of S02
Removal
90
75
67.5
45
90
75
67.5
45
22.5
90
75
67.5
45
22.5
90
67.5
45
22.5
Wastewater Production Rate
H/sec (gpm)
2.2
1.8
1.5
1.0
4.4
3.5
3.2
2.2
1.0
5.8
5.3
4.4
2.9
1.4
11.7
8.8
5.8
2.9
34.6
28.1
23.8
16.5
69.4
55.9
50.5
34.6
16.5
92.0
83.4
69.4
45.3
22.2
185.0
139.0
92.0
45.3
-------
TABLE 8.1-8,
DOUBLE ALKALI PARTIAL SCRUBBING SLUDGE PRODUCTION RATES
(Eastern 3.5% S coal)
00
i
Boiler Size Percent of
and Type Gas Scrubbed
22 MW
(75 x 106Btu/hr)
Chaingrate Stoker
58.6 MW
(200 x 106Btu/hr)
Pulverized Coal
118 MW
(400 x 10sBtu/hr)
Pulverized Coal
100
75
50
100
75
50
25
100
75
50
75
Percent SOz
Removed
90
67.5
45
90
67.5
45
77.5
90
67.5
45
77.5
Sludge Production Rates
g/sec (Ib/hr)
200
145
95
506
391
259
126
1066
788
506
259
1584
1150
750
4012
3100
2050
1000
8450
6250
4012
2050
-------
TABLE 8.1-9.
CO
I
LIMESTONE PARTIAL SCRUBBING SLUDGE PRODUCTION RATES
(Eastern 3.5% S Coal)
Boiler Size
& Type
22 MW
(75xl06Btu/hr)
Chaingrate Stoker
44 MW
(150xl06Btu/hr)
Spreader Stoker
58.6 MW
(200xl06Btu/hr)
Pulverized Coal
Percent of
Gas Scrubbed
100
100
75
50
100
100
75
50
25
100
100
75
50
25
Percent S02
Remova 1
90
75
67.5
45
90
75
67.5
45
22.5
90
75
67.5
45
22.5
Sludge Production Rates
g/sec
227
190
169
108
457
381
343
220
108
607
576
458
300
148
(Ib/hr)
1796
1508
1340
860
3624
3020
2720
1750
860
4812
4568
3630
2380
1170
-------
200 i—
150
S
p.
bO
I
4J
O
3
T3
O
M
PM
cu
4-1
tfl
100
50
Full
Scrubbing
75% of
Gas Treated
50% of
Gas Treated
25% of
Gas Treated
I
29 58.6
(100) (200)
FGD Size, MW (106 Btu/hr)
88
(300)
118
(400)
Figure 8.1-4.
Sodium Partial Scrubbing Wastewater Production
(3.5% S Coal, 90% removal)
8-16
-------
Results of these calculations show a linear production of waste materials
with size which indicates that there is no economy of scale or environmental
benefit resulting from partial scrubbing. This is because the waste produc-
tion rates vary linearly with the amount of S02 removal; that is, a reduction
by 50 percent in the SOa being removed will result in a 50 percent reduction
in the waste production rate for all the throwaway processes.
8.2 LIMESTONE SCRUBBING OF FLUE GAS FROM FIVE PERCENT SULFUR COAL
The purpose of this section is to further examine the effects of coal
sulfur content on the capital and annualized costs of a limestone scrubber.
Material and energy balances for this case were performed on a limestone
scrubber treating the flue gas from a 58.6 MW (200xl05 Btu/hr) boiler burning
5.0% sulfur coal. Table 8.2-1 presents the bases for the material and energy
balance for this case.
TABLE 8.2-1 ASSUMPTIONS AND BASES FOR MATERIAL AND ENERGY BALANCES
Process parameters
Limestone process
L/G, £/m3 (gal/103 acf)'
Particulate Removal
Stoicbiometry
(moles sorbent/mole sorbed S02)
Q
Gas Pressure Drop Pa (in HaO)
Pump Discharge Pressure Pa(Psi)
Pumping Height M (ft)
Stack Gas Reheat °C (°F)b
13.3 (100)
99 Percent Upstream
of Scrubber
1.2
214 (17)
5227 (15)
6 (20)
28 (50)
Based on process data presented in Section 2
Radian assumption
Q
Based on TVA empirical relationship (Reference 4)
3-17
-------
8.2.1 Control Costs
Table 8.2-2 tabulates the cost data for three cases relating capital and
annualized costs to variations in sulfur content of coal. Figure 8.2-1
illustrates this relationship. As shown in the plots there is a slight
flattening of the capital cost curve as higher sulfur coals are burned.
This is due in part to economy of scale coming into effect for the raw mat-
erial handling facilities. Capital costs increase about 27 percent in going
from a 0.6 to 3.5 percent sulfur coal. From 3.5 to 5.0 percent sulfur coal,
the capital costs increase by an additional 10 percent. Increased sulfur
content to 5 percent should affect the capital costs of other FGD systems in
the same manner as for limestone systems. Therefore, the relative ranking of
the processes with regard to capital cost are not expected to change at the
5 percent coal level. The increased slope in the annualized cost curve is
due to the increased solids disposal cost for the high sulfur coal systems.
The relative ranking of FGD processes shown in Section 4 should also not
change with regard to annualized costs for the high sulfur coals. Conse-
quently, the sodium throwaway process would still be the most costly for
the high sulfur applications.
8.2.2 Control Energy Requirements
An approximate 19 percent increase in energy use is required to go from
3.5 to 5 percent sulfur coal in a 58.6 MW (200xl06 Btu/hr) boiler assuming
90 percent S02 removal. The energy requirements are calculated using the
same bases as given in Section 5 of the report. The energy increase is
primarily due to increased feed preparation and liquid pumping rates (in-
creased sorbent). Table 8.2-3 shows the energy requirements for a limestone
control system treating the flue gas from a 58.6 MW (200xl06 Btu/hr) boiler
burning coals with three different sulfur levels. The other processes show
less of an energy impact with changes in coal sulfur content as discussed
in Section 5.
8-18
-------
TABLE 8.2-2. COST SUMMARY FOR LIMESTONE SCRUBBING
Boiler Size
Total Capital
SO2 Removal Investment
Coal Type level
(10 $)
Annualized Costs
Percent Increase
(103$) Over Uncontrolled Use
58.6 MWt
(200xlOS Btu/hr) 0.6% S
90
1201
634
14
58.6 MW
(200xlQ6 Btu/hr) 3.5% S
90
1530
1155
26
CD
58.6 MW
(200xl06 Btu/hr) 5.0% S
90
1689
1426
32
-------
2000
1500
1000
CO
S-l
n)
O
P
500
Total Capital
Investment
Costs
Annualized Costs
Coal Sulfur Content
(% Sulfur)
Figure 8.2-1.
Capital and Annualized Costs for Limestone Scrubbing
versus Coal Sulfur Content
8-20
-------
TABLE 8.2-3. ENERGY SUMMARY FOR LIMESTONE SYSTEMS
Energy Consumption Percent Increase
Boiler Size Coal Type SOa Removal KW 106 Btu/hr Over Uncontrolled
58.6 MW
(200xl06 Btu/hr) 0.6% S 90 711 2.4 1.2
58.6 MW
(200xlOG Btu/hr) 3.5% S 90 912 3.1 1.6
58.6 MW
? (200xlOs Btu/hr) 5.0% S 90 1076 3.7 1.9
fO
-------
8.2.3. Environmental Impact of Burning Five Percent Sulfur Coal
The solid waste production rates for burning high sulfur coal in a
limestone process are shown in Table 8.2-4. This table shows a proportional
increase in solid waste with coal sulfur content which is expected since
waste production rates vary linearly with the amount of 862 removed. This
same trend is also true with regard to waste production rates for the other
FGD processes.
TABLE 8.2-4. SOLID WASTE SUMMARY FOR LIMESTONE SYSTEMS
Waste Study Production
Boiler Size Coal Type S02 Removal s/sec(Ib/hr)
58.6 MW 0.6% S 90 139 1106
(200xl06 Btu/hr)
58.6 MW 3.5% S 90 607 4812
(200xl06 Btu/hr)
58.6 MW 5.0% S 90 867 6879
(200xl06 Btu/hr)
8-22
-------
APPENDIX A
MATERIAL BALANCE CALCULATIONS
A-l
-------
>
N3
^CL&UV fllf
eAW fLUC GAS >
TCfJM WMB£A j DCSCAIPTKH
TOTAL /Z.OK/
35,910
Voo
(/t,
L/OUiO fMAJf fLOtV MTS
SOL/O PMASf rtO*S KATf
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38,33^
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OO/LCA s/ze: 3O * /o'
COAL TYPE: £/ieT-£TtV
-------
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w
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KM GAS
WATtX
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(GOM)
QAi nOk/ RATE"
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33,334.
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ttf.,771
771.3
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BOTTOMS
4-30, 1
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7
«.\
f££D
s
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373 (,
33 S3
371-
0.1
33k
2.77^
S.S
2.738
L/MESTONCFGD PROCESS
BO/LCO. s/zei 30* l°b
COAL TYPE; £/tSTHKA/
SO2
-------
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PAW fLUC GAS
HUMBBO. j DexfuPTKu
TOTAL flOH KA7€
(GOM)
GAl no V GATE-
(ACFM*)
LX3U/D PHASE FLOW RATG
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too
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fETUKN
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L/MESTONE 'FGD WOC£SS
oon.cn
COOL TYPE:
SOj RBMOVAt:
-------
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Ul
^CL£AA/ fiUf GAS
PAU fLUf G4J
c«.co
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STBfJM NUMBER j D€SCAIPTKH
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(ACFM)
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L/MESTONFFGD PROCESS
COAL TYPE:
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OOILCH.
COOL TYPE:
SOf REMOVAL:
-------
t*
I
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fAW fLUf GAS
STVfAM NUMBED, f DeSCRIPT/OV
TOTAL FLOVJ ZATE
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OO/LCO. s/af.- tfO x/
-------
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-------
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f-A\fJ fLUf GAS
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7V700
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SO/LCA S/&: f-00 Y/O
COOL TYPE:
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-------
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TOTAL flow
(GOM)
TfMfBtATUK '(°f)
GAS
C/t Mcx.es/kr)
(scrri)
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400
700
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BOILCR.
COAL TYPS-.
-------
>
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TOTAL flOW KATE"
f/b.//,*)
(GCM)
TfMfEX A TUff
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(it. MOLES/hr)
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(SCFM)
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COOL TYPE:
-------
>
I
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NUMBER f nescwnou
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V
OOUBLf ALKALI
FG£>
B0/L£7f
COAL
-------
I
01
K!
-*• CLEAN rive a AS
XAO
Ftufsas
NUMBER t Dcxe/rmiN
TOTAL f~LOH
ZJS?(,0
TEMfCKATVtf (¥)
GAJ FLOW KATE
(/6. MOifS/tir)
(ACFM)
L/QU/O PHASE FLOW
SOLID PHASC ttOW KATE
(?*/**)
/HLET
GAS
EXIT
GAS
/Z.S
SCKUMEK.
sorrows
0-7
scweaex
73*-
V/77AV
4W3
75%
ATION
fffO
7
LtME-
MAKE-UP
#10
& 'CYCLE
L/QVOK.
3169X1
374932
/3oo
35-
DOUBLE ALKALI
BO/LEK
Co*>i.
-------
I
LO
/gs >
afx/v /Y«/f s/«
KifGAS
STAf AM HUMdtK ( OfXe/rmn
TOTAL FLOW f/ITf
(f)
GAJ /^OK/ >P/(7r
(7X, /hr)
{/(>. MOi.CS /Ar.)
(ACFM)
L /a U/D
FLOW £AT£
P/SXM.VCD StX/OJ V7T r<,
SOl/D PHASC H.OM KAT£
INLET
GJS
$00
rs.sto
GAS
/2?r
^
eorroMs
s-~~
67/QfJ
fffO
LIME:
I77F
MAKE-UP
K£ 'CYCLE
L/QUOK.
/£>
/73
DOUBLtT ALKALI
eo/tee
-------
Ln
HVMB&!
TOTAL FLOW f/ITT
(/6,/Ar.)
(°F)
GAJ fLOW RATE
<"> /**)
(/b.
(ACFM)
*•CLEAN FLUE &AS
MAXfUP
now KATE
sat/as
SOLID PHASC flOW KATE
IHLET
G4S
#b
£X/T
GAS
SCXUBOfR
7
?
ATIQN
1
LIME:
MAKE-UP
i/QUOK.
7
DOUBLE ALfTAL/
FGD
TYPE,
-------
I
Ln
-. MO itS //ir)
(ACfM)
L/QUID PHASE FLOW KATE
sa/aj
50L/D PHASC fZOH KATE
/HLET
GAS
K/J20
erxir
GAS
sorrows
isto
AT/QN
J37
Kf CYCLE
/J2/
to
CLAf/FIEe.
/3
7/
7/
DOUBLE ALKALI
BO/LEU s/fci t/oo&t
C**(~ TYPIT:
-------
Ln
ALKALI
MAKE-UP WATER
vi
CL6A04 FLUE GAS
RAW
FLUE.
\
SPRAY
s
/
DRYER
BAGH
SPRAY
Cool Ty ft
•* SOLIDS
FLovJ
Ib/Vtr
f'F")
FLOW RATE
SOL to riovJ RMt
J'S/ooo
Slurry 4x> D'yir
-------
I
Un
ALKALI
MAKE-UP WATER
n
CLE/XN fLUE GAS
FLUE.
GAS
V
SPRAY
DRYER
r^
BA6H
SPRAY DRYER FGD
Co o 1 Ty pe :
SOLIDS
r t
TOTAL FLOW PAtg
C»F)
now RATS
UIQOtD
SOLID
ExiV
SS/7P3
13*1
Wafer
Slurry ^« D'Vtr
Solids
-------
i
Ul
oo
ALKALI
MAKE-UP WATER
CLEAN FLUE GAS
RAW
FLUE.
SPRAY DRYER
SPRAY DRYER FGD
Co* I TV pe : U>£S7&e/i/
Sot
Typ*
- SOLIDS
ToTAL FLOW PATE
3P
&AS FLOW RATE
^5-32
«?•*??
Slurry 4» Dtytr
b
Solid*
-------
Ln
ALKALI
CLEAN FLUE GAS
RAW
FLUE
GAS
TOTAL FLOVJ
Ib/Vxr
FLOW
SCFM
SOUD
350
42,8
BA6HOUSE
SOLIDS
Slurry io D'Vtr
426
SPRAY DRYER
Cool
-------
I
c^
o
ALKALI
MAKE-UP WATER
CLEAN FLUE GAS
RAW
FLOE.
GAS
V
SPRAY
^
DRYER
^
BAGH
ToTAu FLovJ
FLOW RATE
185
ns
SPRAY DRYER FGD
t
SOUOS
t>
Solids
-------
ALKALI
MAKE-UP WATER
T]
CLEAN FLUE GAS
RAW
FLUE.
GA.S
SPRAY DRYER
SPRAY
1
Cool Type :
Typ*. Avicitli; SODIUM
- SOLIDS
TOTAL FLOW If AT6
f'F)
GAS FLOW RAT 6
185,0^7.
548
Exi'V
vns
I 8
3
Altai;
7,\M
Slurry -to D'ytr
b
Solids
-------
c^
K)
ALKALI
MAKE-UP WATER
vl
CLEAN FLUE GAS
RAW
FLU£
^
SPRAY
X
DRYER
BA6H
SPRAY
Cool TV{*:
SOLIDS
TOTAL FLovJ f?AT£
ib/Kr
3?^
FLOW RATC
LIQUID
Soc »D FiowJ RMt
TJ.,1.00
^00
ExiV
feO.MOS
3
A i Kali
/me
Slurry
415
b
Solids
FGD
-------
I
U)
ALKALI
RAW
FLOE
GAS
GAS
SPRAY DRYER FGD
• SOLIDS
TOTAL FLovJ SAT6
&A*> FLOW RATK
IWW
UQOiD
T--.TI.1
84
Exit
ns
11.15
0.<4
3
Altaii
V2-3
Waftr
7.9
Slurry +« D»v«r
3^3
t>
Solid*
"
7
-------
ALKALI
MAKE-UP WATER
RAW
FLUE.
GAS
TOTAL FLovJ PATE
f'F")
FLOW RATE
\_IQO\O
SOLID F»-ovO RM6.
(oT.81.
SM8
ExiV
"JU
CLEAN FLUE GAS
^
SPRAY
^
X
DRYER
r-^
BAGH
SPRAY DRYER FC.D
Cool Type:
SOj.
-• SOUDS
Slurry \9 D'V*r
Solids
-------
i
On
ALKALI
CLEAN FLUE GAS
RAW
FLUE.
6A.S
TOTAL FLowJ «?AT£
Ib/Kr
Cfl
GAS FLOW RATE
LIQUID
$01-»D
Whr
3 SO
6,381.
ExiV
ns
(oTl S
HT..MH.
181
BA&HOUSe
» SOLIDS
Slurry +« Dryer
181
SSS
SPRAY DRYER FGD
: ts~OKio
Cool Type :
SotRem««fc\: IS
Typ* AMfttli :
-------
ALKALI
MAKE-UP WATER
Ti
CLE*N FLUE GAS
RAW
FLUE.
1
SPRAY
^
X
DRYER
r^
BA6H
TOTAL FLovsl RATE
Ib/Hr
GAS FLOW RATE
Sou ID
RX16.
fcl.fioo
S"48
ExiV
ISO.T20
ns
fcTIS.S
Si. BOO
4^.400
\o\
\o\
SPRAY DRYER FC.D
B«i\e : ISO
Cool Type :
Typ* Avtcali :
SOLIDS
Slurry
\0\
b
Solid-i
TI4-
-------
Ct&SH FlUC
CC
PfiOCFSS
30/LEK
COAL JYPE:
STREAM fWM8£ff 4 OfSCK/PF/OfJ
/NLCT
CAJ
FXIT
G4J
3
M/Xf-UP
BtOWOOWl 30T7UMS
fffO
7
PG
SfCT/ON
8
PUKGE
SECT/ON
PETURN
DKYK
STEAM
SOUOS
PURGE
P6ODUC.T
-SO,
FEED
/4
PCGt'NCR-
AT/OU
STEAM
/5
TOTAL fLOW if ATE
473
o.fe
392
173
3//S~
173
,AS FLOW RATE
173
(/b.
(ACFM)
(SCFM)
/Z.OS"
L/QUfD PHA3E
PATE
/z/z
D/SSOtV£C> SOL/OS IV 7.
Zl-0
(/t./tlr)
3-1
-------
3SO-
eaw /\ 400.,=
fLUE
UELLMAN - LOPD
FQD P/2OCFSS
COAL TYPE:
502 QtMOVAL • 90 fO
STREAM /vuvsee 4 OESCKIPF/ON
/NLET
CAJ
^XIT
GAS
H2O
4.
fiessctueasi
SOT7VMS
fffa
FEED
3
PUK.GE
seer/on
STEAM
souos
f/eooucT
so.
XffVCBCA
FEED
/4
AT/OH
STEAM
STK/PP/NG
STCAM
TOTAL FLOW BATE
3700
1ZOD
JD2.
O.I
Si'7
&.I
202.
2.&0
KATE
36977
(fb. MOtfS/Ah)
(ACFM)
(SCFM)
'3
FLOW
SOIL/OJ U/T,
3 ?OD
/zoo
9?
8^.7
0.3
.0
2.3.O
-------
+CLG4/V ficV G4S
UELLMAN-LOPD
?-GD PfiOCfSS
-*§H
CHILLER X-'
£TH. GLYCOL
COOLANT
STREAM
4 O£SCRIPr/ON
ex IT
GAS
MAXf'UP
H2O
4.
Pf£SCtl>SB&
BiOWCOUto BOTTOMS
A770H
7
mess
SfCT/ON
e
PUAGE
SECT/ON
PETUKN
STEAM
m
SOL/OS
//
ffODUCT
/3
XffvBBEA
FEED
6T/QH
STEAM
/5
STK/PP/NO
TOTAL fiOW PAT'S
&'0
//PA-
S'7
773%
Af
260
>AS FLOW RATE
(/to.
(ACFM)
1S7Z]
338$
L/QU/D PHASE FLOW ffATF
o/ssoivea SOL/OS wr.
731Z
Z?,0
//7O
Z9.D
432
;r.o
77/D
13,0
SOL SO
ftOW
It-
-------
SEffA'.
35O-
gAW /\ 400.f
FLUe(/
G^5 N-/
Miff UP,
WAT£f
'€>
' ' C.J '
frD. fAM
\
Y
«h
i
UELLMAN -LOPD
/->/?O
£7W. G^VOM.
O»0«AT
STREAM /VW/SS? 4 DESCRIPTION
/HL.CT
GAS
2
FX/T
3
MMZ-UP
pffcsc/eueaci SCRUBBER
A7/QH
fffO
7
G
S&C7/ON
a
PUKGE
eETURN
9
tyf
STEAM
SOL/OS
//
PfODLKT
/3
XGU6BEA
FEED
AT/QN
STEAM
IS
STK/PP/NG
•STCAM
TOTAL FLOW KATE.
3000
(.,0
/700
0.4-
0.3
O, Z,
(f)
3575
z&o
flow
/3S-
(/b. MOt.es /^
(ACFM)
(SCFM)
3/5-2-
33 %o
L/QU/D PHASE FLOW RATE
DISSOLVED SOL/OS W7.
"5,000
?/
-------
MAKEUP.
WATEK
UELLMAN-LOPD
FGD P/2OCFSS cuioftee
SOILED j/«-, zoo vo***./»*.
TYPE:
STREAM /WM8£8 4 DESCRIPTION
FX/T
GAJ
3
MAK£-UP
PfX3C(VSeet SCfUBKR
BiOWDOUk BOTTOMS
AT/ON
ffff)
SfCT/ON
FEED
seer/on
eETUKN
9
Wt
STEAM
SOL/OS
PURGE
PRODUCT
SO,
/3
XKUB&fJt.
FEED
AT/W
STEAM
STR/PPWG
STCAM
TOTAL FLOW HATE
1707
7/5-3
("f)
26,0
&AS FLO*/ KATE
(/b.
(ACFM)
(SCFM)
7^37
/8.7
57.$
L/QU/O PHA-3E FLOW KATE
D/SSOLVEQ SOL/OS IVT.
7???
Z037Z
Zf.O
SOL/O
-------
xr/vw:
V/ELLMAN-LOPD
FGD P/?OCE~SS
BO,LEX
caa. TYPE:
S02
£>> f-SOC/OS fM&X-
\s
9^ y.JTF/l/fl
STREAM
4 DESCRIPT/ON
/
'L£
G/kT
FXIT
GAS
3
MAK£ -UP
4-
Cfl
BiOWDOUtt
30T7VMS
7
6
SK7/ON
secr/ofj
PETVRN
•STEAM
/O
SOL/OS
It
peocucT
FEED
/4
AT/OH
STEAM
15
'/PP/,
S7CAM
TOTAL FLOW RATE
7997
3C7C*
s/.r
0,7
0-7
7.0
35-0
/zS
7AJ FLO*/ RATE
(lh.M
(^3. MOtFS/An
(ACFM)
(SCFM)
77/3
79-7
717/0
L/OUID PHASE FLOW
SOL/OS WT.
Z.3
-------
APPENDIX B
CAPITAL AND ANNUALIZED COST FOR FGD PROCESSES
B-l
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
Limestone
FGD Type:
Boiler Capacity: 30xlQ6Btu/hr (8.R MWt)
Coal Feedstock: Eastern 3.5% S
SO2 Control Level:_
90%
Item
Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling
S02 Scrubbing
Fans
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services
Total Direct Costs (TDC)
Indirect Capital Costs
a
Engineering
Construction and Field Expenses (0.1 TDC)
Construction Fees (0.1 TDC)
Start-up (0.02 TDC)
Performance Test (0.01 TDC)
Total Indirect Cost (TIC)
Contingencies (0.2 (TDC+TIC))
TOTAL TURNKEY COSTS (TTC)
i S
Land (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)C
TOTAL CAPITAL INVESTMENT (TCI)
59
149
80
109
20
83
36
36
358
166
105
629
0.5
62
691
a. Engineering Costs
0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOa removal. This cost remains constant for the smaller cases.
For the 118 MW (400xl06 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-2
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Limestone
30xl06Btu/hr (8.8
Eastern 3.5% S
.90%
.60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.3385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
14
54.7
kW
14
7
GJ/hr
1.4
m3/hr
0.2
GJ/hr
m3/hr
323
kg/hr
75
kg/hr
127
kg/hr
10
kg/hr
JicL
40
246
78
117
441
$/106 Btu
$/kg S0:
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-3
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level: 85%
Limestone
3QxlQ6Btu/hr (8.8
Eastern 3 . S% S
Item
Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling
S02 Scrubbing
Fans
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services
Total Direct Costs (TDC)
Indirect Capital Costs
a
Engineering
Construction and Field Expenses (0.1 TDC)
Construction Fees (0.1 TDC)
Start-up (0.02 TDC)
Performance Test (0.01 TDC)
Total Indirect Cost (TIC)
Contingencies (0.2 (TDC+TIC))
TOTAL TURNKEY COSTS (TTC)
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)°
TOTAL CAPITAL INVESTMENT (TCI)
57
142
20
107
20
83
35
35
346
163
102
611
0.5
60
671
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOa removal. This cost remains constant for the smaller cases.
For the 118 MW (400xl06 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-4
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Limestone
30xlObBtu/hr (8.8 MWt)
Eastern 3.5% S
85%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
45
kW
GJ/hr
1.36m3/hr
GJ/hr
105
21
14
14
0.3
m3/hr
305
kg/hr
70
kg/hr
120
kg/hr
9
kg/hr
38
40
$/106 Btu
$/kg SO;
239
78
114
431
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-5
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Limestone
30xl06Btu/hr (8.8 MWt)
Eastern 3.5% S
75%
Item
Direct Capital Costs
Raw Material Handling
SO 2 Scrubbing
Fans
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services
Total Direct Costs (TDC)
Indirect Capital Costs
a
Engineering
Construction and Field Expenses (0.1 TDC)
Construction Fees (0.1 TDC)
Start-up (0.02 TDC)
Performance Test (0.01 TDC)
Total Indirect Cost (TIC)
Contingencies (0.2 (TDC+TIC))
TOTAL TURNKEY COSTS (TTC)
, ^
Land (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)
TOTAL CAPITAL INVESTMENT (TCI)
Cost (Thousands of dollars)
55
131
20
106
,
19
332
83
33
33
7
3
159
98
589
0.5
57
646
a. Engineering Costs =0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
SO? removal.
b. Reference: 4
c. From Annual Cost Table
5-6
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coa'l Feedstock:
S02 Control Level:
Operating Factor:
Limestone
30xl06Btu/hr (8,8 MWt)
Eastern 3.5% S
75%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
NazCOa ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
270
GJ/hr
1,34 m3/hr
_GJ/hr
_m3/hr
kg/hr
105
21
13
0.3
63
kg/hr
106
kg/hr 8
kg/hr
38
40
$/106 Btu
$/kg SOj
228
78
110
416
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-7
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Limestone
Boiler Capacity: 75xlQSBtu/hr (22 MWt"L
Coal Feedstock: Eastern 3.5% S
S02 Control Level: _ 9Q%
Item Cost (Thousands of dollars)
Direct Capital Costs
99
Raw Material Handling
SOa Scrubbing _ 244
Fans 40
Wastewater Pumps ____
Regeneration
Solids Separation 121
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services 30
534
Total Direct Costs (TDC)
Indirect Capital Costs
Engineering 83
Construction and Field Expenses (0.1 TDC) 53
Construction Fees (0.1 TDC) 53
Start-up (0.02 TDC) 11
Performance Test (0.01 TDC) 5
Total Indirect Cost (TIC) 205
Contingencies (0.2 (TDC+TIC)) 148
Q Q -J
TOTAL TURNKEY COSTS (TTC) _._
Landb (0.00084 TTC) 0-7
Working Capital (0.25 Direct Operating Costs) 1QQ
TOTAL CAPITAL INVESTMENT (TCI) 987_
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOa removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal @ 90%
b. Reference: 4
c. From Annual Cost Table
B-8
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Limestone
75xl06Btu/hr (22 MW )
Eastern 3.5% S
90%
60%
Item
Cost (Thousands of dollars)
105
21
21
19
GJ/hr
3.5 mVhr
0.7
816
!3/hr
_kg/hr
188
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh) 137 kW
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ) GJ/hr
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
NaaCOs ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs $/106 Btu $/kg S02
kg/hr
322
kg/hr
24
kg/hr
38
44
400
82
168
650
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-9
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Limestone
Boiler Capacity: 7 5x1 n6Kt-n /hr (?7 MM )
Coal Feedstock: Eastern 3.5% S
S02 Control Level: _ 75%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 89
S02 Scrubbing 216
40
Fans
Wastewater Pumps
Regeneration
Solids Separation 118
Solids Collection
Purge Treatment __^
Sulfur Production
Utilities and Services 28
Total Direct Costs (TDC) 490
Indirect Capital Gusts
Q
Engineering _ 83 _
Construction and Field Expenses (0.1 TDC) _ 4_2 _
Construction Fees (0.1 TDC) 49
Start-up (0.02 TDC) 10_
Performance Test (0.01 TDC) 5
Total Indirect Cost (TIC) 196
Contingencies (0.2 (TDC+TIC)) 137
TOTAL TURNKEY COSTS (TTC) 823
Landb (0.00084 TTC) Q. 7
Working Capital (0.25 Direct Operating Costs)0 89
TOTAL CAPITAL INVESTMENT (TCI) 913
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOa removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: A
c. From Annual Cost Table
B-10
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level;
Operating Factor:
Limestone
75xl06Btu/hr (22
Eastern 3.5% S
75%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
685 kg/hr
105
21
20
20
86
kW
12
GJ/hr
3.4
m3/hr
]
GJ/hr
m3/hr
158
kg/hr
270 kg/hr 20
kg/hr _____
38
43
357
81
_Q )
155
593
$/106 Btu
$/kg S02
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-ll
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Limestone
Boiler Capacity: 150xlQ6Btu/hr (44 MWt)
Coal Feedstock: Eastern 3.5% S
SO, Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 147
S02 Scrubbing 368
Fans 69
Wastewater Pumps
Regeneration
Solids Separation 133
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services 43
Total Direct Costs (TDC)
Indirect Capital Gusts
Engineering 83
Construction and Field Expenses (0.1 TDC) 76
Construction Fees (0.1 TDC) 76
Start-up (0.02 TDC) 15
Performance Test (0.01 TDC) R
Total Indirect Cost (TIC) 258
Contingencies (0.2 (TDC+TIC)) 204
TOTAL TURNKEY COSTS (TTC) 1221
Landb (0.00084 TTC) 1
Working Capital (0.25 Direct Operating Costs) 163
TOTAL CAPITAL INVESTMENT (TCI) 1385
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal @ 90%
S02 removal.
b. Reference: A
L-. From Annual Cost Table
B-12
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coa'l Feedstock:
S02 Control Level:
Operating Factor:
Item
Limestone
150xl05Btu/hr (44 MW,.)
Eastern 3.5% S
90%
60%
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h) 105
Supervision (15.63/m-h) 21
Maintenance Labor (.04 TDC) 30
Maintenance Materials (.04 TDC) 3Q
Electricity (25.8 mills/kWh) 265 kW 36
Steam ($1.84/GJ) GJ/hr
Proc. Water ($.04/m3) 6.9 m3/nr 1_
Methane ($2.05/GJ) , GJ/hr
Wastewater Treating m3/hr
Solids Disposal ($.044/kg) 1647 kg/hr 381
Chemicals
Lime ($.-0385/kg) kg/hr
Limestone ($.0143/kg) 647 kg/hr 49
Na2COs ($.0991/kg) kg/hr
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above) 38
Plant (.26x(1+2+3+4) above) 48
Total Overhead Costs 86
By-Product Credits ( 0 )
Capital Charges
Capital Recovery (.17 TCI) 235
974
TOTAL ANNUALIZED COSTS
Annual Unit Costs $/106 Btu $/kg
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-13
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
Limestone
FGD Type:
Boiler Capacity:
Coal Feedstock: _ Eastern 3.5% S
S02 Control Level:
150xl06Btu/hr (44
75%
Item
Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling
S02 Scrubbing
Fans
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services
Total Direct Costs (TDC)
Indirect Capital Costs
a
Engineering
Construction and Field Expenses (0.1 TDC)
Construction Fees (0.1 TDC)
Start-up (0.02 TDC)
Performance Test (0.01 TDC)
Total Indirect Cost (TIC)
Contingencies (0.2 (TDC+TIC))
TOTAL TURNKEY COSTS (TTC)
Land5 (0.0008A TTC)
Working Capital (0.25 Direct Operating Costs)°
TOTAL CAPITAL INVESTMENT (TCI)
134
325
69
129
39
83
70
70
14
696
244
188
1128
141
1270
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-14
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Limestone
150xl06Btu/hr (44 MWt
Eastern 3.5% S _ _
75%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
173
540
kW
_kg/hr
_kg/hr
_kg/hr
105
21
28
28
23
GJ/hr
6.9
m3/hr
1
GJ/hr
rnVhr
L372
kg/hr
318
40
38
47
$/106 Btu
$/kg S02
564
85
216
865
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-lf
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Limestone
Boiler Capacity: 200xlQ6Btu/hr (58.6
Coal Feedstock: Eastern 3.5% S
S02 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 171
401
S02 Scrubbing
76
Wastewater Pumps
Regeneration
Solids Separation 137
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services 4_7
832
Total Direct Costs (TDC)
Indirect Capital Costs
Engineering 83
Construction and Field Expenses (0.1 TDC) 83
Construction Fees (0.1 TDC) 83
Start-up (0.02 TDC) 17
Performance Test (0.01 TDC) 8
Total Indirect Cost (TIC) 274
Contingencies (0.2 (TDC+TIC)) 221
TOTAL TURNKEY COSTS (TTC) 1327
Landb (0.00084 TTC)
c
Working Capital (0.25 Direct Operating Costs) 202
TOTAL CAPITAL INVESTMENT (TCI)
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-16
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Item
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor ( . 04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh) 311
Steam ($1.84/GJ)
Proc. Water ($.04/m3) 8.
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg) 2187
Chemicals
Lime ($.£385/kg)
Limestone ($.0143/kg) 862
NaaCOs ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (. 2 6x( 1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Limestone
200xl06Btu/hr (58.6 MW^)
Eastern 3.5% S
90%
60%
Cost (Thousands of dollars)
105
21
33
33
kw 42
GJ/hr
ImVhr 2
GJ/hr
m3/hr
kg/hr 5Q6
kg/hr
kg/hr ftS
kg/hr
807
38
50
88
( 0 )
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
260
1155
$/106 Btu
$/_kg SOj
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-17
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Limestone
Boiler Capacity: 20x10 Btu/hr (58.6 MW )
Coal Feedstock: Eastern 3.5% S
S02 Control Level: 85%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 166
S02 Scrubbing 375
Fans 7J2
Wastewater Pumps
Regeneration
Solids Separation 135
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services 45
Total Direct Costs (TDC) 797
Indirect Capital Costs
Engineering 83
Construction and Field Expenses (0.1 TDC) 80
Construction Fees (0.1 TDC) 80
Start-up (0.02 TDC) -| fi
Performance Test (0.01 TDC) 8
Total Indirect Cost (TIC) 267
Contingencies (0.2 (TDC+TIC)) _ 21
TOTAL TURNKEY COSTS (TTC) -] 277
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)C 192
TOTAL CAPITAL INVESTMENT (TCI) 1470
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOz removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: A
c. From Annual Cost Table
B-18
-------
ANMJALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Item
Limestone
200xl06Btu/hr (58.6
Eastern 3.5% S
85%
60 %
Cost (Thousands
MWt)
of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
^5 . 4
kw
GJ/hr
8.2
3/hr
105
21
32
32
34
GJ/hr
m3/hr
2076
kg/hr
479
kg/hr
818
kg/hr
61
kg/hr
38
49
766
87
0/
250
11Q2
$/l(T' Btu
$/kg SO 2
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-19
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
Limestone
FGD Type:
Boiler Capacity: 200xlQ6Btu/hr (58.6 MW_ )
Coal Feedstock: Eastern 3.5% S
S02 Control Level: 75%
Item
Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling
S02 Scrubbing
Fans
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services
Total Direct Costs (TDC)
jndirect Capital Costs
a
Engineering
Construction and Field Expenses (0.1 TDC)
Construction Fees (0.1 TDC)
Start-up (0.02 TDC)
Performance Test (0.01 TDC)
Total Indirect Cost (TIC)
Contingencies (0.2 (TDC+TIC))
TOTAL TURNKEY COSTS (TTC)
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)0
TOTAL CAPITAL INVESTMENT (TCI)
155
354
76
131
42
_8_3_
76
76
15
757
258
203
1218
173
1392
a. Engineering Costs =0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (400xl06 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-20
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Limestcme
200xKrBtu/hr (58.6 MW )
Eastern 3.5% S
75%
60%
Item
Cost (Thousands of dollars)
105
21
30
30
28
GJ/hr
7.9m3/hr
1827 kg/hr 422
kg/hr
720 kg/hr
kg/hr
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh) 205 kW
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ) GJ/hr
Wastewater Treating m3/hr
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs $/106 Btu $/kg SOg
54
38
48
692
86
236
1014
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-21
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Limestone
Boiler. Capacity: 3f)x1 O6 Rf ii/hr Cf
Coal Feedstock: Western 0.6% S
SO2 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
33
Raw Material Handling
1 4?
S02 Scrubbing
Fans 20
Wastewater Pumps
Regeneration
Solids Separation °4
Solids Collection
Purge Treatment ,
Sulfur Production
Utilities and Services 17
Total Direct Costs (TDC) 305
Indirect Capital Costs
a
Engineering
Construction and Field Expenses (0.1 TDC) 30
Construction Fees (0.1 TDC) 30
Start-up (0.02 TDC) 5_
Performance Test (0.01 TDC) 3
1 52
Total Indirect Cost (TIC)
Contingencies (0.2 (TDC+TIC)) 91
TOTAL TURNKEY COSTS (TTC) 548
Landb (0.00084 TTC) 0.5
Working Capital (0.25 Direct Operating Costs)0 46
TOTAL CAPITAL INVESTMENT (TCI) 594
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
S02 removal.
b. Reference: A
c. From Annual Cost Table
B-22
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coa'l Feedstock:
SOa Control Level:
Operating Factor:
Limestone
30xl06Btu/hr (8.8 MWfc)
jfestern 0.6% S
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/ni3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Line ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
42 kW
GJ/hr
1.2m3/hr
GJ/hr
105
21
12
0.2
m3/hr
75
kg/hr
18
kg/hr
27
kg/hr
2
kg/hr
38
38
186
76
0 )
100
362
$/106 Btu
$/kg SO;
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-23
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Limestone
Boiler Capacity: 30x10 Btu/hr (8.8 MWt)_
Coal Feedstock: Western 0.6% S
S02 Control Level:
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 31
S02 Scrubbing 128
Fans 20
Wastewater Pumps
Regeneration
Solids Separation 94
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services 16
Total Direct Costs (TDC) 289
Indirect Capital Costs
a 83
Engineering
Construction and Field Expenses (0.1 TDC) 29
Construction Fees (0.1 TDC) 29
Start-up (0.02 TDC) £_
Performance Test (0.01 TDC) •}
Total Indirect Cost (TIC) 150
Contingencies (0.2 (TDC+TIC)) 88
TOTAL TURNKEY COSTS (TTC)
Landb (0.00084 TTC) 0.5
Working Capital (0.25 Direct Operating Costs)° 42
TOTAL CAPITAL INVESTMENT (TCI) 569
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-24
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Limestone
30xl06Btu/hr
Western 0.6% S
75%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
kg/hr
23 kg/hr
kg/hr
105
21
12
31
kW
12
4
GJ/hr
1.
Om3/hr
0
.2
GJ/hr
m3/hr
63
kg/hr
14
38
39
170
87
97
$/106 Btu
$_/kg SO;
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-25
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Limestone
Boiler- Capacity: 30xlQ6Btu/hr (8.8 MWfc)
Coal Feedstock: Western 0.6% S
S02 Control Level: 85%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 32
SOa Scrubbing 136
Fans 20
Wastewater Pumps
Regeneration
Solids Separation 94
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services ]J_
Total Direct Costs (TDC) 298
Indirect Capital Costs
Engineering 83
Construction and Field Expenses (0.1 TDC) 3Q
Construction Fees (0.1 TDC) 3Q
Start-up (0.02 TDC) £_
Performance Test (0.01 TDC) 3
Total Indirect Cost (TIC) 152
Contingencies (0.2 (TDC+TIC)) 9Q
TOTAL TURNKEY COSTS (TTC) 540
Landb (0.00084 TTC) 0.5
Working Capital (0.25 Direct Operating Costs)C 43
TOTAL CAPITAL INVESTMENT (TCI) 583
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal @ 90%
b. Reference: A
c. From Annual Cost Table
B-26
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
Limestone
10xlQ6BtU/hr (8,8 MWt)
-n D
S02 Control Level:
Operating Factor:
85%
60%
Item
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh) 35
Steam ($1.84/GJ)
Proc. Water ($.04/m3) ] . 1
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg) 70
Chemicals
Lime ($.O385/kg)
Limestone ($.0143/kg) 25
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (. 2 6x( 1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Cost (Thousands of dollars)
105
21
12
1?
kW 5
GJ/hr
m3/hr 0.2
GJ/hr
m3/hr
kg/hr T fi
kg/hr
kg/hr 2
kg/hr
173
38
39
87
( 0 )
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
99
359
$/106 Btu
$/kg SO;
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-27
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
Type: Limestnnp
Boiler Capacity: 75xlD6Btu/hr (22MWt)
Coal Feedstock: Western 0.6% S
S02 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 48
S02 Scrubbing 233
Fans 40
Wastewater Pumps
Regeneration
Solids Separation 102
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services 25
Total Direct Costs (TDC)
Indirect Capital Costs
Engineering "3
Construction and Field Expenses (0.1 TDC) ^ ^
Construction Fees (0.1 TDC) 45
Start-up (0.02 TDC) 9_
Performance Test (0.01 TDC) A
Total Indirect Cost (TIC) 186
Contingencies (0.2 (TDC+TIC)) 127
TOTAL TURNKEY COSTS (TTC)
Landb (0.00084 TTC) 0.6
Working Capital (0.25 Direct Operating Costs)0 60
TOTAL CAPITAL INVESTMENT (TCI) 821
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOa removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-28
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Limestone
75xl06Btu/hr (22 MWt)
Western 0.6% S
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
18
18
106
kW
14
GJ/hr
2.7
m3/hr
1
GJ/hr
m3/hr
188
kg/hr
57
kg/hr
68
kg/hr
5
kg/hr
38
42
239
80
139
458
$/106 Btu
$/kg S02
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-29
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Limestone
Boiler Capacity: 75xlQ6Btu/hr (22 MW )
Coal Feedstock: Western 0.6% S
S02 Control Level: 75%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 44
SO2 Scrubbing 211
Fans 40
Wastewater Pumps
Regeneration
Solids Separation 101
Solids Collection __
Purge Treatment
Sulfur Production
Utilities and Services 24
Total Direct Costs (TDC) 420
Indirect Capital Costs
Engineering 83
Construction and Field Expenses (0.1 TDC) 42
Construction Fees (0.1 TDC) 42
Start-up (0.02 TDC) 8
Performance Test (0.01 TDC) 4
Total Indirect Cost (TIC) 179
Contingencies (0.2 (TDC+TIC)) 120
TOTAL TURNKEY COSTS (TTC)
Landb (0.00084 TTC) • 0.6
Working Capital (0.25 Direct Operating Costs)° 52
TOTAL CAPITAL INVESTMENT (TCI) 771
a. Engineering Costs =0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOa removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-30
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Limestone
75xl06Btu/hr (22 MW )
. ^_
Western 0.6% S
75%
fiOZ
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
57
_kg/hr
_kg/hr
_kg/hr
105
21
17
17
10
38
42
210
80
0 )
131
421
$/106 Btu
$/kg S02
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-31
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Limestone
Boiler Capacity: 150xl06 Btu/hr (44 MWt)
Coal Feedstock: Western 0.6% S
S02 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 67
SO2 Scrubbing 351
Fans 66
Wastewater Pumps
Regeneration
Solids Separation 112
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services 35
Total Direct Costs (TDC) 632
Indirect Capital Costs
Engineering
a 83
Construction and Field Expenses (0.1 TDC) 63
Construction Fees (0.1 TDC) 63
Start-up (0.02 TDC) 13_
Performance Test (0.01 TDC) 6_
Total Indirect Cost (TIC)
Contingencies (0.2 (TDC+TIC)) 172
TOTAL TURNKEY COSTS (TTC) 1032
Landb (0.00084 TTC) 0.8
Working Capital (0.25 Direct Operating Costs)C 76
TOTAL CAPITAL INVESTMENT (TCI) 1108
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-32
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type: LimesrnnP
Boiler Capacity: 150xlQ6BtU/hr (44
Coa'l Feedstock: Western 0.6% S
S02 Control Level: 90%
Operating Factor: 60%
Item Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h) 105
Supervision (15.63/m-h) 21
Maintenance Labor (.04 TDC) 25
Maintenance Materials (.04 TDC) 25
Electricity (25.8 mills/kWh) 212 kw -29
Steam ($1.84/GJ) GJ/hr
Proc. Water ($.04/m3) 5.4m3/hr 1
Methane ($2.05/GJ) GJ/hr
a 3
Wastewater Treating m /hr
Solids Disposal ($.044/kg) 377 kg/hr 88
Chemicals
Lime ($.-0385/kg) kg/hr
Limestone ($.0143/kg) 136 kg/hr 10
Na2C03 ($.0991/kg) kg/hr
Total Direct Operating Cost 304
Overhead
Payroll (.3x(l+2) above) 38
Plant (.26x(1+2+3+4) above) 4fi
Total Overhead Costs 84
By-Product Credits ( 0 )
Capital Charges
Capital Recovery (.17 TCI) 188
TOTAL ANNUALIZED COSTS 576
Annual Unit Costs $/10 Btu $/kg SO
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-33
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Limestone
Boiler Capacity: 150xl06Btu/hr (44 MWt)_
Coal Feedstock: Western 0.6% S
S02 Control Level: 75%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 62
SO2 Scrubbing 319
Fans 66
Wastewater Pumps
Regeneration
Solids Separation 111
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services 33
Total Direct Costs (TDC) 591
Indirect Capital Costs
Engineering 83
Construction and Field Expenses (0.1 TDC) 59
Construction Fees (0.1 TDC) 59
Start-up (0.02 TDC) 12
Performance Test (0.01 TDC) f.
Total Indirect Cost (TIC) 219
Contingencies (0.2 (TDC+TIC)) 162
TOTAL TURNKEY COSTS (TTC) 972
Landb (0.00084 TTC) 0.8
Working Capital (0.25 Direct Operating Costs)°
TOTAL CAPITAL INVESTMENT (TCI) 1041
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOa removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Limestone
Boiler Capacity: 200x10 Btu/hr (58.6 MWt)
Coal Feedstock: Western 0.6% S
S02 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 7^
SOa Scrubbing 384
Fans 7^
Wastewater Pumps
Regeneration
Solids Separation 116
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services 39
Total Direct Costs (TDC) 686
Indirect Capital Costs
_ . .
Engineering
Construction and Field Expenses (0.1 TDC)
a 83
Construction Fees (0.1 TDC) 69
Start-up (0.02 TDC) 14_
Performance Test (0.01 TDC) 7
Total Indirect Cost (TIC) 242
Contingencies (0.2 (TDC+TIC)) Iftfi
TOTAL TURNKEY COSTS (TTC) 1114
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs) 86
TOTAL CAPITAL INVESTMENT (TCI) 1201
a. Engineering Costs =0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOz removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
S02 removal.
b. Reference: A
c. From Annual Cost Table
B-35
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coa'l Feedstock:
S02 Control Level:
Operating Factor:
Li'mpsfrmp
20Qxl06Btu/hr (58.6
Western 0,6% S
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 raills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2COs ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
GJ/hr
105
21
27
27
247 kw 33
GJ/hr
5 m3/hr
1
m3/hr
503
kg/hr
117
kg/hr
182
kg/hr
14
kg/hr
38
47
345
( 0 )
204
634
$/10b Btu jS/kgSO;
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-36
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Limestone
150xl06Btu/hr (44 MWt)
Western 0.6% S
75%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.JD385/kg)
Limestone ($.0143/kg)
NazCOs ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
-24.
24
153
kW
21
GJ/hr
5
.4m3/hr
1
GJ/hr
m3/hr
314
kg/hr
72
kg/hr
114
kg/hr
9
kg/hr
38
45
277
83
177
537
$/106 Btu
$/kg SO,
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-37
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
Limestone
FGD Type:
Boiler Capacity: 200xlQ6Btu/hr (58.6 MWt)
Coal Feedstock: Western 0.6% S
S02 Control Level: 85%
Item
Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling
SO2 Scrubbing
Fans
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services
Total Direct Costs (TDC)
Indirect Capital Gusts
a
Engineering
Construction and Field Expenses (0.1 TDC)
Construction Fees (0.1 TDC)
Start-up (0.02 TDC)
Performance Test (0.01 TDC)
Total Indirect Cost (TIC)
Contingencies (0.2 (TDC+TIC))
TOTAL TURNKEY COSTS (TTC)
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)0
TOTAL CAPITAL INVESTMENT (TCI)
73
367
73
116
38
83
67
67
13
666
237
181
1084
83
1168
a. Engineering Costs =0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOa removal. This cost remains constant for the smaller cases.
For the 118 MW (400xlOs Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal @ 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-38
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Limestone
200xl06Btu/hr (58.6 MWt)
Western 0.6% S
85%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2COa ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
27
27
208
kW
28
GJ/hr
5
m3/hr
1
GJ/hr
m3/hr
47S
kg/hr
inq
kg/hr
172
kg/hr
13
kg/hr
38
47
331
85
198
614
$/106 Btu
I/kg SOj
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-39
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Limestone
Boiler Capacity: 200xlQ6 Btu/hr (58.6
Coal Feedstock: Western 0.6% S
S02 Control Level: 75%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 69
S02 Scrubbing 348
Fans 73
Wastewater Pumps
Regeneration
Solids Separation 116
Solids Collection
Purge Treatment ____
Sulfur Production
Utilities and Services 36
642
Total Direct Costs (TDC)
Indirect Capital Costs
Engineering 83
Construction and Field Expenses (0.1 TDC) 64
Construction Fees (0.1 TDC) 64
Start-up (0.02 TDC) 13
Performance Test (0.01 TDC) 6
Total Indirect Cost (TIC) 230
Contingencies (0.2 (TDC+TIC)) 174
TOTAL TURNKEY COSTS (TTC) 1046
Landb (0.00084 TTC) 1
Working Capital (0.25 Direct Operating Costs) 78
TOTAL CAPITAL INVESTMENT (TCI) 1125
a. Engineering Costs =0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOa removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-40
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Limestone
200xl06Btu/hr (58.6 MW )
Western 0.6 % S
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
26
26
175
kW
24
GJ/hr
5
m3/hr
1
GJ/hr
m3/hr
418
kg/hr
97
kg/hr
151
kg/hr
n
kg/hr
38
46
311
84
191
586
$/10s Btu
$/kgS02
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-41
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
Sodium Throwawav
FGD Type:
Boiler Capacity:
Coal Feedstock: 3.5% S Eastern
S02 Control Level:
MWfc (30 MBtu/hr
90%
Item
Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling
SO2 Scrubbing
Fans
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services
Total Direct Costs (TDC)
Indirect Capital Costs
a
Engineering
Construction and Field Expenses (0.1 TDC)
Construction Fees (0.1 TDC)
Start-up (0.02 TDC)
Performance Test (0.01 TDC)
Total Indirect Cost (TIC)
Contingencies (0.2 (TDC+TIC))
TOTAL TURNKEY COSTS (TTC)
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)C
TOTAL CAPITAL INVESTMENT (TCI)
63
109
20
18
13
62
22
22
223
112
67
402
0.3
55
457
d. Engineering Costs =0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (400xl06 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
S02 removal.
b. Reference: A
c. From Annual Cost Table
B-42
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium-Throwaway
8.8
(30 MBtu/hr)
3.5% S Eastern
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
.Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(l+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
24.3
kW
9
3
GJ/hr
4,25
m3/hr
1
GJ/hr
3.04
m3/hr
_5
kg/hr
kg/hr
kg/hr
131
kg/hr
68
38
37
221
75
78
374
$/106 Btu
$/kg SO;
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-43
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Throwaway
Boiler Capacity: 8.8 MW (30 MBtu/hr)
Coal Feedstock: 3.5% S Eastern
S02 Control Level: 85%
Item Cost (Thousands of dollars)
Direct Capital Costs
• f\ ^
Raw Material Handling _
SO2 Scrubbing 109
Fans 20
Wastewater Pumps 18
Regeneration ^_
Solids Separation ^_
Solids Collection ~
Purge Treatment ^_
Sulfur Production -
Utilities and Services 13
Total Direct Costs (TDC) 223
Indirect Capital Costs
. a 62
Engineering
Construction and Field Expenses (0.1 TDC) 22
Construction Fees (0.1 TDC) 22
Start-up (0.02 TDC) ^_
Performance Test (0.01 TDC) 2
Total Indirect Cost (TIC) 112
Contingencies (0.2 (TDC+TIC)) 67
TOTAL TURNKEY COSTS (TTC) 402
Landb (0.00084 TTC) 0-3
Working Capital (0.25 Direct Operating Costs) 55
TOTAL CAPITAL INVESTMENT (TCI)
456
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOz removal. This cost remains constant for the smaller cases.
For the 118 MW (400xl06 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: A
c. From Annual Cost Table
B-44
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
8.8 MWt (30 MBtu/hr)
3.5% S Eastern
85%
602
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
23.7
kW
9
3
GJ/hr
4.Q4
m3/hr
1
GJ/hr
2.84
m3/hr
5
kg/hr
kg/hr ^
kg/hr
125 kg/hr 65
38
37
218
75
78
371
$/106 Btu $/kg_Sp2
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-45
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Throwaway
Boiler Capacity: 8.8 MWf (30 MBtu/hr)
Coal Feedstock: 3.5% S Eastern
S02 Control Level: 75%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 60
S02 Scrubbing 109
Fans 20
Wastewater Pumps 17
Regeneration ^_
Solids Separation ^_
Solids Collection ^_
Purge Treatment ~^_
Sulfur Production ~
Utilities and Services 12
Total Direct Costs (TDC) 218
Indirect Capital Costs
Engineering 62
Construction and Field Expenses (0.1 TDC) 22
Construction Fees (0.1 TDC) 22
Start-up (0.02 TDC) ^_
Performance Test (0.01 TDC) 2
Total Indirect Cost (TIC) 112
Contingencies (0.2 (TDC+TIC)) 66
TOTAL TURNKEY COSTS (TTC) 396
Landb (0.00084 TTC) 0-3
Working Capital (0.25 Direct Operating Costs)C 53
TOTAL CAPITAL INVESTMENT (TCI)
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOa removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal @ 90%
b. Reference: 4
<~ • From Annual Cost Table
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
8.8 MWt (30 MBtu/hr)
3.5% S Eastern
" 75% ~
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.O385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
22
.9 kW
9
3
GJ/hr
3
•75m3/hr
1
GJ/hr
2
.54m3/hr
4
kg/hr
kg/hr
kg/hr
116
kg/hr
60
38
37
212
75
76
363
$/106 Btu
$/kg SO;
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-47
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Throwaway
Boiler Capacity: 22 MWfc (75 MBtu/hr)
Coal Feedstock: 3.5% S Eastern
SO? Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 104
S02 Scrubbing 182
Fans
Wastewater Pumps ^-*-
Regeneration ~
Solids Separation ~
Solids Collection ~
Purge Treatment
Sulfur Production -
Utilities and Services 21
Total Direct Costs (TDC) 368
Indirect Capital Costs
Engineering 62
Construction and Field Expenses (0.1 TDC) 37
Construction Fees (0.1 TDC) 37
Start-up (0.02 TDC) 7_
Performance Test (0.01 TDC) 4
Total Indirect Cost (TIC) 147
Contingencies (0.2 (TDC+TIC)) 103
TOTAL TURNKEY COSTS (TTC) 618
Landb (0.00084 TTC) 0.5
Working Capital (0.25 Direct Operating Costs)°
TOTAL CAPITAL INVESTMENT (TCI) 705
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xlOb Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (400xl06 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal @ 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-48
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coa'l Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
22 MWt (75 MBtu/hr)
3.5% S Eastern
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
61.7 kw
GJ/hr
10.9 m3/hr
GJ/hr
7.85m3/hr
_kg/hr
$/106 Btu
105
21
15
15
kg/hr -_
kg/hr ^
337 kg/hr 175
38
41
$/kgSO_2
348
79
120
547
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-49
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Throwaway
Boiler Capacity: 22 MWt (75 MBtu/hr)
Coal Feedstock: 3.5% S Eastern
S02 Control Level: 75%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 96
S02 Scrubbing 182
Fans 40
Wastewater Pumps 21
Regeneration ^_
Solids Separation ~__
Solids Collection ^_
Purge Treatment ~^_
Sulfur Production -
Utilities and Services 20
Total Direct Costs (TDC) 359
Indirect Capital Costs
Engineering 62
Construction and Field Expenses (0.1 TDC) 36
Construction Fees (0.1 TDC) 36
Start-up (0.02 TDC) ~l_
Performance Test (0.01 TDC) 4
Total Indirect Cost (TIC) 145
Contingencies (0.2 (TDC+TIC)) ]_Q1
TOTAL TURNKEY COSTS (TTC) 605
Landb (0.00084 TTC) 0.4
Working Capital (0.25 Direct Operating Costs)° 80
TOTAL CAPITAL INVESTMENT (TCI)
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (400xl06 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: 4
c. From Annual Cost T.ible
B-50
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
22 MWt (75 MBtu/hr)
3.5% S Eastern
75%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(l+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
kg/hr
kg/hr
kg/hr
291 kg/hr
105
21
14
57
.5 kW
14
8
GJ/hr
9
•42ra3/hr
2
GJ/hr
6
•38mVhr
6
151
38
40
321
78
116
515
$/106 Btu $/kg S0_;
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-51
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Throwaway
Boiler Capacity: 44 MWt (150 MBtu/hr)
Coal Feedstock: 3.5% S Eastern
S02 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 154
S02 Scrubbing _ 275
Fans 68
9 c
Wastewater Pumps J
Regeneration
Solids Separation
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services 31
Total Direct Costs (TDC) 553
Indirect Capital Costs
„ . . a 62
Engineering
Construction and Field Expenses (0.1 TDC) 55
Construction Fees (0.1 TDC) 55
Start-up (0.02 TDC) 11
Performance Test (0.01 TDC) 6
Total Indirect Cost (TIC) 189
Contingencies (0.2 (TDC+TIC)) 148
TOTAL TURNKEY COSTS (TTC) 890
Landb (0.00084 TTC) 0.7
Working Capital (0.25 Direct Operating Costs)C 138
TOTAL CAPITAL INVESTMENT (TCI) J.Q28
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200x10" Btu/hr) case burning 3.5% S coal
@ 90% SOa removal. This cost remains constant for the smaller cases.
For the 118 MW (400xl06 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal @ 90%
b. Reference: A
c. From Annual Cost Table
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
44 MWt (150 MBtu/hr)
3.5% S Eastern
90%
60%
Item
Cost (Thousands of dollars)
105
21
22
22
17
GJ/hr
21.8 m3/hr
GJ/hr
10
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh) 124 kW
Steam ($1.84/GJ) _
Proc. Water ($.0
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Throwaway
Boiler Capacity: 44 MWt (150 MBtU/hr)
Coal Feedstock: 3.5% S Eastern
S02 Control Level: 75%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 142
S02 Scrubbing ___Z25
Fans 68
Wastewater Pumps 24
Regeneration -
Solids Separation ~
Solids Collection -
Purge Treatment _—
Sulfur Production
Utilities and Services 3Q
Total Direct Costs (TDC) 539
jndirect Capital Costs
a
Engineering 62
Construction and Field Expenses (0.1 TDC) 54
Construction Fees (0.1 TDC) 54
Start-up (0.02 TDC) 11
Performance Test (0.01 TDC) 5
Total Indirect Cost (TIC) 186
Contingencies (0.2 (TDC+TIC)) 143
TOTAL TURNKEY COSTS (TTC) 868
Land (0.00084 TTC) Q. 7
Working Capital (0.25 Direct Operating Costs)° 125
TOTAL CAPITAL INVESTMENT (TCI) 993
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
Q 90% SOa removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
b. Reference: 4
c. From Annual Cost Table
B-54
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
44 MWt (150 MBtu/hr)
3.5% S Eastern
75%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
22
22
115
kW
16
GJ/hr
18.8
m3/hr
4
GJ/hr
12.7
ma/hr
9
kg/hr
kg/hr ^_
kg/hr -_
579 kg/hr 302
38
44
501
82
169
752
$/106 Btu
$/kgS02
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-5f
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Thrnwaway
Boiler Capacity: 58.6 MWt (200 MBtu/hr)_
Coal Feedstock: 3,5% S Eastern
S02 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 178
S02 Scrubbing 301
Fans 76
Wastewater Pumps 26
Regeneration ^_
Solids Separation -._
Solids Collection -_
Purge Treatment .-_
Sulfur Production —_
Utilities and Services ^5
Total Direct Costs (TDC)
Indirect Capital Costs
Engineering 62
Construction and Field Expenses (0.1 TDC) 62
Construction Fees (0.1 TDC) 62
Start-up (0.02 TDC) 12_
Performance Test (0.01 TDC) 6
Total Indirect Cost (TIC) 204
Contingencies (0.2 (TDC+TIC)) 164
TOTAL TURNKEY COSTS (TTC) 984
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs) 167
TOTAL CAPITAL INVESTMENT (TCI) 115L
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xlOb Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: A
c. From Annual Cost Table
B-56
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
58.6 MWt (200 MBtu/hr)
3.5% S Eastern
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
NaaCOs ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(l+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
151
kW
GJ/hr
27.9 m3/hr
GJ/hr
20.9 mVhr
kg/hr
$/106 Btu
105
21
27
27
20
11
kg/hr ^_
kg/hr ^_
864 kg/hr 450
38
47
$/kg 50^
667
85
196
947
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-57
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Throwaway
Boiler Capacity: 58.6 MW (200 MBtu/hr)
Coal Feedstock: 3.5% S Eastern
SO2 Control Level: 85%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 171
S02 Scrubbing 301
Fans 76
Wastewater Pumps 26
Regeneration ~2_
Solids Separation ^_
Solids Collection ^_
Purge Treatment
Sulfur Production
Utilities and Services ^4.
Total Direct Costs (TDC)
Indirect Capital Costs
Engineering 62
Construction and Field Expenses (0.1 TDC) 61
Construction Fees (0.1 TDC) 61
Start-up (0.02 TDC) 12
Performance Test (0.01 TDC) 6
Total Indirect Cost (TIC) 202
Contingencies (0.2 (TDC+TIC)) 162
TOTAL TURNKEY COSTS (TTC) ^72
Landb (0.00084 TTC) 0.8
Working Capital (0.25 Direct Operating Costs)C 157
TOTAL CAPITAL INVESTMENT (TCI) 1129
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3-5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
S02 removal.
b. Reference: A
c. From Annual Cost Table
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
58.6 MWt (200 MBtu/hr)
J3.55 S Eastern
85%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 raills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
24
24
145
kW
20
GJ/hr
26.0
m3/hr
5
GJ/hr
18.9
m 3 / hr
10
kg/hr
kg/hr ^_
kg/hr ^_
803 kg/hr 418
38
45
627
83
192
902
$/106 Btu
$/kg so.
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-59
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Throwaway
Boiler Capacity: 58.6 MWfc (200 MBtu/hr}_
Coal Feedstock: 3.5% S Eastern
S02 Control Level: 75%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 163
S02 Scrubbing 3Q1
76
Wastewater Pumps 25
Regeneration —_
Solids Separation ^_
Solids Collection ^_
Purge Treatment -_
Sulfur Production
Utilities and Services i/.
Total Direct Costs (TDC) 599
Indirect Capital Costs
Engineering 62
Construction and Field Expenses (0.1 TDC) 60
Construction Fees (0.1 TDC) 60
Start-up (0.02 TDC) 12
Performance Test (0.01 TDC) 6
Total Indirect Cost (TIC) 200
Contingencies (0.2 (TDC+TIC)) 160
TOTAL TURNKEY COSTS (TTC) 959
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs) 148
TOTAL CAPITAL INVESTMENT (TCI) 1108
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal @ 90%
S02 removal.
b. Reference: A
c. From Annual Cost Table
-------
ANMJALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
58.6 MW (200 MBtu/hr)
3.5% S Eastern
75%
60%
Item
Cost (Thousands of dollars)
105
21
24
24
19
24.0 myhr
GJ/hr
10
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh) 139 kW
Steam ($1.84/GJ) GJ/hr
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
Wastewater Treating 17.0 m3/hr
Solids Disposal ($.044/kg) kg/hr
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs $/106 Btu $/kg SOj
kg/hr ^_
kg/hr ~__
741 kg/hr 386
38
45
594
83
188
865
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G5.A, property taxes and insurance. Federal and State income
taxes are not included.
B-61
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD TyPe: JSod-nnn Thrnwaway
Boiler Capacity: 118 MWt (400 MBtu/hr)
Coal Feedstock: 3.5% S Eastern
S02 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 264
S02 Scrubbing 465
Fans 126
Wastewater Pumps 30
Regeneration 2_
Solids Separation ^_
Solids Collection ^_
Purge Treatment —_
Sulfur Production ^_
Utilities and Services S^
Total Direct Costs (TDC) 938
Indirect Capital Costs
Engineering 94
Construction and Field Expenses (0.1 TDC) 94
Construction Fees (0.1 TDC) 94
Start-up (0.02 TDC) 19
Performance Test (0.01 TDC) 9
Total Indirect Cost (TIC) 310
Contingencies (0.2 (TDC+TIC)) 250
TOTAL TURNKEY COSTS (TTC) 1498
Landb (0.00084 TTC) 1.3
c
284
Working Capital (0.25 Direct Operating Costs)
TOTAL CAPITAL INVESTMENT (TCI) 1783
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOz removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: A
c. From Annual Cost Table
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
118 MWf (400 MBtu/hr)
3.5% S Eastern
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
NaaCOs ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(l+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
298
_kW
GJ/hr
54.0 mVhr
GJ/hr
42.0 m3/hr
kg/hr
105
21
37
37
II
15
kg/hr ~_
kg/hr ^_
1680 kg/hr 873
38
52
1139
90
303
1532
$/10b Btu $/kg
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-63
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Throwaway
Boiler Capacity: 22 MWt
Coal Feedstock: 2.3% S
Boiler Capacity: 22 MWt (75 MBtU/hr)
S02 Control Level: 9Q%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 84
S02 Scrubbing _ 183
Fans 40
Wastewater Pumps 19
Regeneration ^_
Solids Separation ^_
Solids Collection ^_
Purge Treatment
Sulfur Production ~
Utilities and Services 20
Total Direct Costs (TDC) 346
Indirect Capital Costs
Engineering
Construction and Field Expenses (0.1 TDC) 35
Construction Fees (0.1 TDC) 35
Start-up (0.02 TDC) 7_
Performance Test (0.01 TDC) 3
Total Indirect Cost (TIC) 142
Contingencies (0.2 (TDC+TIC)) 98
TOTAL TURNKEY COSTS (TTC) 586
Landb (0.00084 TTC) 0.5
Working Capital (0.25 Direct Operating Costs)C '_
TOTAL CAPITAL INVESTMENT (TCI) 657
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-64
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
22 MWt (75 MBtu/hr)
2.3% S
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam (S1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
NaaCOs ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
14
52
.1
kW
14
7
GJ/hr
7
.26
m3/hr
2
GJ/hr
4
.54
m3/hr
6
kg/hr
kg/hr 2.
kg/hr ^
225 kg/hr 117
38
40
286
78
112
476
$/106 Btu
$/kg S02
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-65
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type:
Sodium Throwaway
Boiler Capacity: 118 MWt (400 MBtu/hr)
Coal Feedstock: _2.3% S
S02 Control Level:_
90%
Item
Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling
SO2 Scrubbing
Fans
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services
Total Direct Costs (TDC)
Indirect Capital Costs
a
Engineering
Construction and Field Expenses (0.1 TDC)
Construction Fees (0.1 TDC)
Start-up (0.02 TDC)
Performance Test (0.01 TDC)
Total Indirect Cost (TIC)
Contingencies (0.2 (TDC+TIC))
TOTAL TURNKEY COSTS (TTC)
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)C
TOTAL CAPITAL INVESTMENT (TCI)
209
444
127
27
49
94
86
86
17
856
292
230
1378
1.2
209
1588
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOz removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal @ 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-66
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
118 MWt (400 MBtu/hr)
2.3% S
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/ra3)
Methane ($2.05/GJ)
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
_kg/hr
105
21
34
34
249
kW
34
GJ/hr
36.8
m3/hr
8
GJ/hr
24.5
m3/hr
12
kg/hr ^_
kg/hr ^_
1130 kg/hr 591
38
50
839
270
1197
$/106 Btu
I/kg SO 2
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-67
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Throwaway
Boiler Capacity: 8.8 MWt (30 MBtu/hr)
Coal Feedstock: 0.6% S Western
S02 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 37
S02 Scrubbing _ 107
Fans 20
Wastewater Pumps 13
Regeneration ~__
Solids Separation ^_
Solids Collection ^_
Purge Treatment ^_
Sulfur Production -
Utilities and Services J_J_
Total Direct Costs (TDC) 188
jndirect Capital Costs
Engineering 62
Construction and Field Expenses (0.1 TDC) 19
Construction Fees (0.1 TDC) 19
Start-up (0.02 TDC) ^
Performance Test (0.01 TDC) 2
Total Indirect Cost (TIC) 106
Contingencies (0.2 (TDC+TIC)) 59
TOTAL TURNKEY COSTS (TTC) 353
Landb (0.00084 TTC) 0.3
Working Capital (0.25 Direct Operating Costs)° ^*-
TOTAL CAPITAL INVESTMENT (TCI) 394
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
SO2 removal.
b. Reference: 4
i_. From Annual Cost Table
B-68
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
8.8 MWt (30 MBta/hr)
0.6% S Western
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
17.
6 kW
7
2
GJ/hr
1.
70m3/hr
0
GJ/hr
0.
66m3/hr
2
kg/hr
kg/hr
kg/hr
34.
6 kg/hr
18
38
36
162
74
67
303
$/106 Btu $/kg S02
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-69
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Throwaway
Boiler Capacity: 8.8 MWt (30 MBtu/hr)
Coal Feedstock: 0.6% S Western
S02 Control Level: 85%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 36
107
S02 Scrubbing __
Fans 20
1 Q
Wastewater Pumps
Regeneration ^_
Solids Separation ^_
Solids Collection ~^_
Purge Treatment
Sulfur Production ^_
Utilities and Services 11
Total Direct Costs (TDC) 187
Indirect Capital Costs
a r 0
Engineering b/
Construction and Field Expenses (0.1 TDC) 19
Construction Fees (0.1 TDC) 19
Start-up (0.02 TDC) 4_
Performance Test (0.01 TDC) 2
Total Indirect Cost (TIC) 106
Contingencies (0.2 (TDC+TIC)) 59
TOTAL TURNKEY COSTS (TTC) 352
Landb (0.00084 TTC) 0.3
Working Capital (0.25 Direct Operating Costs)C
TOTAL CAPITAL INVESTMENT (TCI) 392
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOa removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
S02 removal.
b. Reference: 4
i- • From Annual Cost Table
B-70
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
8.8 MWr (30 MBtu/hr)
0.6% S Western
85%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
0.59m3/hr
_kg/hr
105
21
17.4 kwe
GJ/hr
1.63m3/hr
2
_
0
GJ/hr
kg/hr
kg/hr
33.2
kg/hr
17
38
36
155
74
67
302
$/106 Btu
$/kg
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-71
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Throwaway
Boiler Capacity: 8.8 MWfc (30 MBtu/hr)
Coal Feedstock: 0-6% S Western
SO? Control Level: '-"A
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 36
S02 Scrubbing _ 107
Fans 20
Wastewater Pumps 13
Regeneration ^_
Solids Separation ^_
Solids Collection ^_
Purge Treatment ~^_
Sulfur Production -
Utilities and Services 11
Total Direct Costs (TDC) 187
Indirect Capital Costs
Engineering 62
Construction and Field Expenses (0.1 TDC) ^
Construction Fees (0.1 TDC) 19
Start-up (0.02 TDC) 4
Performance Test (0.01 TDC) 2
Total Indirect Cost (TIC) 106
Contingencies (0.2 (TDC+TIC)) 59
TOTAL TURNKEY COSTS (TTC) 352
Landb (0.00084 TTC) 0.3
Working Capital (0.25 Direct Operating Costs)C
TOTAL CAPITAL INVESTMENT (TCI)
d. Engineering Costs - 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOz removal. This cost remains constant for the smaller cases,
For the 118 MW (400xl06 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: A
c. From Annual Cost Table
B-72
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
8.8 MWt (30 MBtu/hr)
0.6% S Western
75%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ,($2.05/GJ)
Waste-water Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
17
.2 kW
7
2
GJ/hr
1
.59m3/hr
0
GJ/hr
0
.54m3/hr
2
kg/hr
kg/hr
kg/hr
kg/hr
17
38
36
155
74
67
302
$/106 Btu $/kg S02
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-73
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Throwaway
Boiler Capacity: 22 MWt (75 MBtu/hr)
Coal Feedstock: 0.6% S Western
SO, Control Level:90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling _ -^ _
S02 Scrubbing _ 168
Fans JJ
Wastewater Pumps
Regeneration ^_
Solids Separation ~^_
Solids Collection ^_
Purge Treatment ~^_
Sulfur Production ~^_
Utilities and Services I/
Total Direct Costs (TDC) 294
Indirect Capital Costs
Engineering 62
Construction and Field Expenses (0.1 TDC) 29
29~
Construction Fees (0.1 TDC)
Start-up (0.02 TDC)
Performance Test (0.01 TDC)
Total Indirect Cost (TIC) 129
Contingencies (0.2 (TDC+TIC)) 85
TOTAL TURNKEY COSTS (TTC) 508
Landb (0.00084 TTC) 0-4
Working Capital (0.25 Direct Operating Costs)C 51
TOTAL CAPITAL INVESTMENT (TCI) 559
a. Engineering Costs 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
(? 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (400xl06 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: 4
c . Froir Annual Cost Table
B-74
-------
ANMJALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
22 MWt (75 MBtu/hr)
0.6% S Estern
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor ( . OA TDC)
Maintenance Materials (.OA TDC)
Electricity (25.8 mills/kWh)
Steam ($1.8A/GJ)
Proc. Water ($.QA/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.04A/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.01A3/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(l+2+3+A) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
_kg/hr
_kg/hr
kg/hr
85.4 kg/hr
105
21
J_2_
44
• 2 kw
12
6
GJ/hr
4
.20m3/hr
1
GJ/hr
1
.fifimVhr
4
44
38
39
205
77
95
377
$/106 Btu
$/kg
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-75
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Throwaway
Boiler Capacity: 22 MWf (75 MBtu/hr)
Coal Feedstock: 0.6% S Western
S02 Control Level: 75%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 52
S02 Scrubbing 168
Fans 39
Wastewater Pumps 15
Regeneration ~^_
Solids Separation ^_
Solids Collection ^_
Purge Treatment ^_
Sulfur Production —
Utilities and Services 17
Total Direct Costs (TDC) 291
Indirect Capital Costs
Engineering 62
Construction and Field Expenses (0.1 TDC) 29
Construction Fees (0.1 TDC) 29
Start-up (0.02 TDC) 6_
Performance Test (0.01 TDC) 3
Total Indirect Cost (TIC) 129
Contingencies (0.2 (TDC+TIC)) 84
TOTAL TURNKEY COSTS (TTC)
Landb (0.00084 TTC) 0.4
Working Capital (0.25 Direct Operating Costs)
50
TOTAL CAPITAL INVESTMENT (TCI) 554
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases,
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
b. Reference: 4
c. From Annual Cost Table
B-76
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
22 MWt(75 MBtu/hr)
0.6% S Western
75%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
12
43.
3 kW
12
6
GJ/hr
3.
95m3/hr
1
GJ/hr
1.
34m3/hr
3
kg/hr
kg/hr
kg/hr
80.
1 kg/hr
42
38
39
202
77
95
374
$/106 Btu $/kg SO2
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-77
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Throwaway
Boiler Capacity: 44 MW (150 MBtu/hr)
Coal Feedstock: 0.6% S Western
S02 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 75
S02 Scrubbing 267
Fans 66
Wastewater Pumps 18
Regeneration ^_
Solids Separation ~^_
Solids Collection 2_
Purge Treatment
Sulfur Production -
Utilities and Services 26
Total Direct Costs (TDC) 452
Indirect Capital Costs
Engineering 62
Construction and Field Expenses (0.1 TDC) 45
Construction Fees (0.1 TDC) 45
Start-up (0.02 TDC) 9_
Performance Test (0.01 TDC) 5
Total Indirect Cost (TIC) 166
Contingencies (0.2 (TDC+TIC)) 124
TOTAL TURNKEY COSTS (TTC) 742
Landb (0.00084 TTC) 0.6
Working Capital (0.25 Direct Operating Costs)
c
68
TOTAL CAPITAL INVESTMENT (TCI) 810
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (400xl06 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-78
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
44 MWt (150 MBtu/hr)
0.6% S Western
_ 90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.fl385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(l+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
18
88
.2 kw
18
12
GJ/hr
8
.51m3/hr
2
GJ/hr
3
•31m3/hr
5
kg/hr
kg/hr
kg/hr
173
kg/hr
90
38
42
271
80
138
489
$/106 Btu
$/kg SO,
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-79
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Thrnwaway
Boiler Capacity: 44 MWfc (150 MBtu/hr)
Coal Feedstock: 0.6% S Western
S02 Control Level:_
75%
Item
Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling
S02 Scrubbing
Fans
Waste-water Pumps
Regeneration
Solids Separation
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services
Total Direct Costs (TDC)
Indirect Capital Costs
a
Engineering
Construction and Field Expenses (0.1 TDC)
Construction Fees (0.1 TDC)
Start-up (0.02 TDC)
Performance Test (0.01 TDC)
Total Indirect Cost (TIC)
Contingencies (0.2 (TDC+TIC))
TOTAL TURNKEY COSTS (TTC)
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)
TOTAL CAPITAL INVESTMENT (TCI)
72
267
66
17
25
62
45
45
447
165
122
734
0.6
66
8DIL
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOz removal. This cost remains constant for the smaller cases,
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
SO-, removal.
b. Reference: A
c. From Annual Cost Table
B-80
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
44 MWt (150 MBtu/hr)
0.6% S Western
75%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.0^/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
18
18
86
.5 kW
12
GJ/hr
7
_QQm3/nr
2
GJ/hr
2
,68m3/hr
4
kg/hr
kg/hr ^
kg/hr ^_
160 kg/hr 83
38
42
263
80
479
$/106 Btu
$/kg S0;
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-81
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Throwaway
Boiler Capacity: 58.6 MWt (200 MBtu/hr)
Coal Feedstock: 0.6% S Western
S02 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 84
SO2 Scrubbing 294
Fans 74
Wastewater Pumps 19
Regeneration ~^_
Solids Separation ^_
Solids Collection ^_
Purge Treatment ~_
Sulfur Production ~
Utilities and Services 28
Total Direct Costs (TDC) 499
Indirect Capital Costs
Engineering 62
Construction and Field Expenses (0.1 TDC) 50
Construction Fees (0.1 TDC) 50
Start-up (0.02 TDC) 10
Performance Test (0.01 TDC) 5
Total Indirect Cost (TIC) 177
Contingencies (0.2 (TDC+TIC)) 135
TOTAL TURNKEY COSTS (TTC) 811
Landb (0.00084 TTC) 0.7
Working Capital (0.25 Direct Operating Costs)C 75
TOTAL CAPITAL INVESTMENT (TCI) 886
a. Engineering Costs 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOz removal. This cost remains constant for the smaller cases.
For the 118 MW (400xl06 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-82
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwawav
58.6 MW (200 MBtu/hr)
0.6% S Western
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.Q4/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
10.5 m3/hr
GJ/hr
4.43m3/hr
kg/hr
105
21
20
20
14
kg/hr -_
kg/hr ~_
214 kg/hr 111
38
43
$/106 Btu
$/kg SO;
29g
81
150
529
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-83
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Throwaway
Boiler Capacity: 58.6 MWfc (200 MBtu/hr)
Coal Feedstock: 0.6% S Western
S02 Control Level: 85%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 82
SO2 Scrubbing 294
Fans 74
Wastewater Pumps 19
Regeneration —_
Solids Separation ^_
Solids Collection ^_
Purge Treatment 2_
Sulfur Production ~~
Utilities and Services 28
Total Direct Costs (TDC) 497
Indirect Capital Costs
Engineering 62
Construction and Field Expenses (0.1 TDC) 50
Construction Fees (0.1 TDC) 50
Start-up (0.02 TDC) 10
Performance Test (0.01 TDC) 5
Total Indirect Cost (TIC) 177
Contingencies (0.2 (TDC+TIC)) 335
TOTAL TURNKEY COSTS (TTC) 809
Landb (0.00084 TTC) 0. 7
Working Capital (0.25 Direct Operating Costs)C _
TOTAL CAPITAL INVESTMENT (TCI) 882
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (400xl06 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
b. Reference: A
c. From Annual Cost Table
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
58.6 MWfc (200 MBtu/hr)
0.6% S Western _
85%
_ 60% _.
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(l+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
103
kW
105
21
20
14
GJ/hr
10.1 m3/hr
2
GJ/hr
4.00m3/hr
5
kg/hr
kg/hr ^
kg/hr -_
205 kg/hr 107
38
43
294
81
150
525
$/106 Btu $/kg SO^
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-85
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Throwaway
Boiler Capacity: 58.6 MWf (200 MBtU/hr)
Coal Feedstock: 0.6% S Western
SO? Control Level: 75%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 80
294
S02 Scrubbing _
Fans 74
Wastewater Pumps 18
Regeneration ^_
Solids Separation ~2_
Solids Collection ^_
Purge Treatment ^_
Sulfur Production ~~
Utilities and Services 28
Total Direct Costs (TDC) 494
Indirect Capital Costs
Engineering 62
Construction and Field Expenses (0.1 TDC) 49
Construction Fees (0.1 TDC) 49
Start-up (0.02 TDC) 10
Performance Test (0.01 TDC) 5
Total Indirect Cost (TIC) 175
Contingencies (0.2 (TDC+TIC)) 134
TOTAL TURNKEY COSTS (TTC) 803
Landb (0.00084 TTC) 0.7
Working Capital (0.25 Direct Operating Costs) 72
TOTAL CAPITAL INVESTMENT (TCI) 875
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal @ 90%
S02 removal.
b. Reference: A
c. From Annual Cost Table
B-86
-------
ANMJALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway _
58.6 MWt (200 MBtu/hr)
0.6% S Western
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2COa ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(l+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
102
kW,
e
GJ/hr
9.65 m3/hr
3.56
_GJ/hr
_in3/hr
_kg/hr
105
21
20
20
14
kg/hr ^
kg/hr ^
196 kg/hr 1Q2
38
43
$/106 Btu $/kg SO^
289
81
149
519
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-87
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Sodium Thrnwaway
Boiler Capacity: 118 MWt (kC>C> MRf.ii/hr)
Coal Feedstock: Q.6% S We atern
S02 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 119
SOz Scrubbing 458
Fans 1?4
Wastewater Pumps 22
Regeneration ~
Solids Separation -
Solids Collection -
Purge Treatment ^
Sulfur Production -
Utilities and Services 43
Total Direct Costs (TDC) 766
Indirect Capital Costs
a
Engineering 94
Construction and Field Expenses (0.1 TDC) 77
Construction Fees (0.1 TDC) 77
Start-up (0.02 TDC) 15
Performance Test (0.01 TDC) 8
Total Indirect Cost (TIC) 271
Contingencies (0.2 (TDC+TIC)) 207
TOTAL TURNKEY COSTS (TTC) 1244
Landb (0.00084 TTC) J..Q
Working Capital (0.25 Direct Operating Costs)0
TOTAL CAPITAL INVESTMENT (TCI) 1355
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xlOb Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (400xl06 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-8
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Sodium Throwaway
118 MWt (400 MBtu/hrl
0.6% S Western
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(l+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
206
GJ/hr
20.2 m3/hr
GJ/hr
8.85 m3/hr
105
21
28
_ kg/hr
kg/hr
kg/hr -_
413 kg/hr 215
443
38
49
87
230
760
$/106 Btu $/kg SO;
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-89
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Dual Alkali
Boiler Capacity: 8.8 MWt (30 MBtU/hr)
Coal Feedstock: 3.5% S Eastern
S02 Control Level: 9Q%
Item Cost (Thousands af dollars)
Direct Capital Costs
Raw Material Handling 76
108
S02 Scrubbing __
Fans 20
Wastewater Pumps -_
Regeneration 5_
Solids Separation 139
Solids Collection ^_
Purge Treatment ^_
Sulfur Production -
Utilities and Services 21
Total Direct Costs (TDC) 369
Indirect Capital Costs
a 7 Q
Engineering ' °
Construction and Field Expenses (0.1 TDC) 37
Construction Fees (0.1 TDC) 37
Start-up (0.02 TDC) 7
Performance Test (0.01 TDC) ^
Total Indirect Cost (TIC) 163
Contingencies (0.2 (TDC+TIC)) 106
TOTAL TURNKEY COSTS (TTC) 633
Landb (0.00084 TTC) ®' 5
Working Capital (0.25 Direct Operating Costs)
TOTAL CAPITAL INVESTMENT (TCI) 699
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SO? removal. This cost remains constant for the smaller cases.
For the 118 MW (400xl06 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
-------
ANMJALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Dual Alkali
8.8 MWt (30 MBtu/hr)
3.5% S Eastern _
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/.GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
15
15
17
.8 kW
2
GJ/hr
1
.30m3/hr
0
GJ/hr
m3/hr
?S3
59
kg/hr
.5 kg/hr
66
12
kg/hr
5
. 90kg/hr
3
38
41
239
79
119
437
$/106 Btu $/kgSO_2
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: JDual Alkali
Boiler Capacity: 8.8 MWt (30 MBtu/hr)
Coal Feedstock: 0.6% S. Western
S02 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 55
S02 Scrubbing 106
Fans 20
Wastewater Pumps —
Regeneration 2
Solids Separation 114
Solids Collection ~2_
Purge Treatment ^_
Sulfur Production -
1 o
Utilities and Services -1-0
Total Direct Costs (TDC) 315
Indirect Capital Costs
Engineering 78
Construction and Field Expenses (0.1 TDC) -37
Construction Fees (0.1 TDC) 32
Start-up (0.02 TDC) 6
Performance Test (0.01 TDC) 3
151
Total Indirect Cost (TIC)
Contingencies (0.2 (TDC+TIC)) 93
TOTAL TURNKEY COSTS (TTC) 559
Landb (0.00084 TTC) 0.5
Working Capital (0.25 Direct Operating Costs)° 43
TOTAL CAPITAL INVESTMENT (TCI) 603
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xlOs Btu/hr) case burning 3.5% S coal
@ 90% SOz removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
b. Reference: 4
c. From Annual Cost Table
B-92
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Dual Alkali
8.8 MWt (30 MBtu/hr)
0.6% S Western
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
Na2COa ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
16.0
GJ/hr
1.03 m3/hr
GJ/hr
m3/hr
2.41 kg/hr
$/106 Btu
105
21
13
13
55.4 kg/hr 13
11.4 kg/hr
kg/hr
38
$/kg
170
78
102
350
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-93
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Item
Direct Capital Costs
Raw Material Handling
SO 2 Scrubbing
Fans
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services
Total Direct Costs (TDC)
Indirect Capital Costs
a
Engineering
Construction and Field Expenses (0.1 TDC)
Construction Fees (0.1 TDC)
Start-up (0.02 TDC)
Performance Test (0.01 TDC)
Total Indirect Cost (TIC)
Contingencies (0.2 (TDC+TIC))
TOTAL TURNKEY COSTS (TTC)
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)°
Dual Alkali
22 MWr (75 MBtu/hr)
3.5% S Eastern
90%
Cost (Thousands of dollars)
106
108
40
-
10
157
_
-
-
30
523
78
52
52
10
5
197
144
864
0.7
95
TOTAL CAPITAL INVESTMENT (TCI) 960
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (400xl06 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-94
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Dual Alkali
22 MWt (75 MBtu/hr)
3.5% S Eastern
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (. OA TDC)
Maintenance Materials (.OA TDC)
Electricity (25.8 mills/kWh)
Steam (?1.84/GJ)
Proc. Water ($.OA/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.OAA/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.01A3/kg)
NaaCOs ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
GJ/hr
105
21
21
21
3.
27m3/hr
1
GJ/hr
m3/hr
719
151
kg/hr
kg/hr
166
31
kg/hr
14.
5 kg/hr
8
38
380
82
163
625
$/106 Btu
$/kg SO2
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-95
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Item
Direct Capital Costs
Raw Material Handling
S02 Scrubbing
Fans
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services
Total Direct Costs (TDC)
Indirect Capital Costs
a
Engineering
Construction and Field Expenses (0.1 TDC)
Construction Fees (0.1 TDC)
Start-up (0.02 TDC)
Performance Test (0.01 TDC)
Total Indirect Cost (TIC)
Contingencies (0.2 (TDC+TIC))
TOTAL TURNKEY COSTS (TTC)
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)°
TOTAL CAPITAL INVESTMENT (TCI)
Dual Alkali
22 MW,. (75 MBtu/hiO
2.3% S
90%
Cost (Thousands of dollars)
86
182
41
-,
7
145
_
.-
-
28
489
78
49
49
10
5
191
136
816
0.7
72
889
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases,
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal @ 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Dual Alkali
22 MWt (75 MBtu/hr)
2.3% S
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
.Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
NaaCOa ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
20
20
42.0 kWg
6
GJ/hr
2.87m3/hr
1
GJ/hr
m3/hr
403 kg/hr
84.4 kg/hr
93
17
kg/hr
8. 17 kg/hr
4
38
43
287
81
151
519
$/106 Btu
$/kg SO;
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-97
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Dual Alkali,
Boiler Capacity: 22 MWt (75 MBtn/hr)
Coal Feedstock: 0.6% S Western
S02 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 67
180
SOa Scrubbing _
Fans
Wastewater Pumps ~^_
Regeneration £_
Solids Separation 127
Solids Collection ^_
Purge Treatment ^_
Sulfur Production ~
Utilities and Services 25
Total Direct Costs (TDC) 443
Indirect Capital Costs
a 7ft
Engineering
Construction and Field Expenses (0.1 TDC) 44
Construction Fees (0.1 TDC) 44
Start-up (0.02 TDC) 9_
Performance Test (0.01 TDC) 4
Total Indirect Cost (TIC) 179
Contingencies (0.2 (TDC+TIC)) -| 74
TOTAL TURNKEY COSTS (TTC) 746
Landb (0.00084 TTC) °'6
Working Capital (0.25 Direct Operating Costs)° 52
TOTAL CAPITAL INVESTMENT (TCI) 799
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-98
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Dual Alkali
_22 MWt (75 MBtu/hr)
0.6% S Western
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
NasCOs ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
18
18
40.
3 kWg
5
GJ/hr
2.
55m3/hr
1
GJ/hr
m3/hr
137
28.
kg/hr
6 kg/hr
32
6
kg/hr
5.
90kg/hr
3
38
209
80
( - )
136
425
$/106 Btu $/kg SO^
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-99
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
Dual Alkali
FGD Type:
Boiler Capacity:
Coal Feedstock: 3.5% S Eastern
S02 Control Level:
58.6 MWt (200 MBtu/hrl
90%
Item
Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling
S02 Scrubbing
Fans
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services
Total Direct Costs (TDC)
Indirect Capital Costs
a
Engineering
Construction and Field Expenses (0.1 TDC)
Construction Fees (0.1 TDC)
Start-up (0.02 TDC)
Performance Test (0.01 TDC)
Total Indirect Cost (TIC)
Contingencies (0.2 (TDC+TIC))
TOTAL TURNKEY COSTS (TTC)
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)0
TOTAL CAPITAL INVESTMENT (TCI)
162
297
76
18
178
44
78
78
78
16
775
258
207
1240
1.0
181
1422
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
S02 removal.
b. Reference: A
c. From Annual Cost Table
B-100
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Dual Alkali
58.6 MWt (200 MBtu/hr)
3.5% S Eastern
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
GJ/hr
7.04 m3/hr
1820
_GJ/hr
_m3/hr
_kg/hr
402
$/106 Btu
105
21
31
31
14
421
81
_kg/hr
kg/hr -
37.2 kg/hr 19
38
49
$/kg SO 2
725
87
241
1053
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-101
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Dual Alkali
Boiler Capacity: 58.6 MWfc (200 MBtu/hr)
Coal Feedstock: 0.6% S Western
S02 Control Level: 90%
Item Cost (Thousands Of dollars)
Direct Capital Costs
Raw Material Handling 90
S02 Scrubbing 292
Fans 74
Wastewater Pumps —_
Regeneration 6_
Solids Separation 143
Solids Collection
Purge Treatment ~^_
Sulfur Production -
Utilities and Services 36
Total Direct Costs (TDC) 641
Indirect Capital Costs
Engineering 78
Construction and Field Expenses (0.1 TDC) 64
Construction Fees (0.1 TDC) 64
Start-up (0.02 TDC) 13
Performance Test (0.01 TDC) 6
Total Indirect Cost (TIC) 225
Contingencies (0.2 (TDC+TIC)) 2.73
TOTAL TURNKEY COSTS (TTC) 1039
Landb (0.00084 TTC) 0.9
Working Capital (0.25 Direct Operating Costs)° 75
TOTAL CAPITAL INVESTMENT (TCI) 1115
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal @ 90%
S02 removal.
b. Reference: A
c. From Annual Cost Table
B-102
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Dual Alkali
_58.6 MW
0.6% S Western
90%
60%
t (200 MBtu/hr)
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
93.8
105
21
26
26
13
GJ/hr
5
.98m3/hr
1
GJ/hr
m3/hr
366
76
kg/hr
.3 kg/hr
85
15
kg/hr
15
.9 kg/hr
8
38
300
84
190
574
$/106 Btu
$/kg
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-103
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Dual Alkali
Boiler Capacity: 118 MW (400 MBtu/hr)
Coal Feedstock: 3.5% S Eastern
S02 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 229
S02 Scrubbing
Fans 126
Wastewater Pumps ^_
Regeneration 30
Solids Separation 198
Solids Collection ^_
Purge Treatment ^_
Sulfur Production ~
Utilities and Services 63
Total Direct Costs (TDC) 1110
Indirect Capital Costs
Engineering 111
Construction and Field Expenses (0.1 TDC) 13-1
Construction Fees (0.1 TDC) 111
Start-up (0.02 TDC) 22
Performance Test (0.01 TDC) 11
366
Total Indirect Cost (TIC)
Contingencies (0.2 (TDC+TIC)) 295
TOTAL TURNKEY COSTS (TTC) 1771
Landb (0.00084 TTC) 1.5
Working Capital (0.25 Direct Operating Costs) 334
TOTAL CAPITAL INVESTMENT (TCI) 2105
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases,
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B--104
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Dual Alkali
JU8 MW (400 MBtu/hr)
3.5% S Eastern
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.0i/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
NazCOs ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
43
43
207
k^
28
GJ/hr
14.0
m3/hr
3
GJ/hr
mVhr
3840
806
kg/hr
kg/hr
887
163
kg/hr
78.5
kg/hr
41
38
55
1334
93
351
1778
$/106 Btu
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
3-105
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Dual Alkali
Boiler Capacity: 118 MWf (400 MBtu/hr)
Coal Feedstock: 2.3% S
S02 Control Level: 9Q%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 173
S02 Scrubbing 440
Fans 1?7
Wastewater Pumps -
70
Regeneration ^-u
Solids Separation
Solids Collection
Purge Treatment ^_
Sulfur Production —_
Utilities and Services
Total Direct Costs (TDC)
Indirect Capital Costs
Engineering
a 108
Construction and Field Expenses (0.1 TDC) 100
Construction Fees (0.1 TDC) 100
Start-up (0.02 TDC) 20
Performance Test (0.01 TDC) 10
Total Indirect Cost (TIC) 338
Contingencies (0.2 (TDC+TIC)) 268
TOTAL TURNKEY COSTS (TTC) 1606
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)°
TOTAL CAPITAL INVESTMENT (TCI) 1819
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
b. Reference: A
c. From Annual Cost Table
B-106
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Dual Alkali
118 MWt (400 MBtu/hr)
2.3% S
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
195
GJ/hr
13.3 m3/hr
2150
_GJ/hr
_m3/hr
kg/hr
105
21
40
40
26
497
449
kg/hr
91
kg/hr
44.9
kg/hr
23
38
54
$/106 Btu $/kg S02
846
92
309
1247
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-107
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Dual Alkali
Boiler Capacity: 118 MW^ (400 MBtu/hr
Coal Feedstock: 0.6% S Western
S02 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 118
S02 Scrubbing 453
Fans 1 24
Wastewater Pumps —
Regeneration -L"
Solids Separation
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services 52
Total Direct Costs (TDC) 912
Indirect Capital Costs
Engineering
a 108
Construction and Field Expenses (0.1 TDC) 91
Construction Fees (0.1 TDC) 91
Start-up (0.02 TDC) 18
Performance Test (0.01 TDC) 9
Total Indirect Cost (TIC) 317
Contingencies (0.2 (TDC+TIC)) 246
TOTAL TURNKEY COSTS (TTC) 1475
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)0 111
TOTAL CAPITAL INVESTMENT (TCI)
a.. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (400xl06 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
SO2 removal.
b. Reference: 4
c. From Annual Cost Table
B-108
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Dual Alkali
118 MWt (400 MBtu/hr)
0.6% S Western
90%
60%
Item
Cost (Thousands of dollars)
GJ/hr
11.8 m3/hr
GJ/hr
mVhr
Direct Costs
Operating Labor (12.02/m-h) 105
Supervision (15.63/m-h) 21
Maintenance Labor (.04 TDC) 36
Maintenance Materials (.04 TDC) 36
Electricity (25.8 mills/kWh) 185 kH^. 25
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg) 733 kg/hr 169
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs $/106 Btu $/kg S02
153
kg/hr
31
kg/hr
32.2
kg/hr
17
38
51
442
269
800
a.. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-109
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Spray Drying (Lime)
Boiler Capacity: 150xlQ6 Btu/hr (44
Coal Feedstock: Western 0.6% S
S02 Control Level: 75%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 79
S02 Scrubbing 230
Fans 66
Wastewater Pumps
Regeneration
Solids Separation 426
Solids Collection
Purge Treatment .
Sulfur Production
Utilities and Services 48
Total Direct Costs (TDC) 850
Indirect Capital Costs
Engineering 92
Construction and Field Expenses (0.1 TDC) 85
Construction Fees (0.1 TDC) 85
Start-up (0.02 TDC) 17
Performance Test (0.01 TDC) 9
288
Total Indirect Cost (TIC)
Contingencies (0.2 (TDC+TIC)) 228
TOTAL TURNKEY COSTS (TTC) 1366
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs) 64
TOTAL CAPITAL INVESTMENT (TCI) 1431
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOz removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
S02 removal.
b. Reference: 4
<_. From Annual Cost Table
B-110
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Spray Drying CLime)
150xl06Btu/hr (44
Western 0.6% S
75%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
NazCOs ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
34
34
92 kvt
12
GJ/hr
4.4 m3/hr
1
GJ/hr
m3/hr
163 kg/hr
85 kg/hr
38
17
kg/hr
kg/hr
38
50
261
88
242
592
$/106 Btu
S/kg SO 2
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-lll
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: _Spray Drying (Sodium)
Boiler Capacity: 150xl06 Btu/hr (44 MW£)
Coal Feedstock: Western 0.6% S
S02 Control Level: 75%
Item
Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling
SO2 Scrubbing
Fans
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services
Total Direct Costs (TDC)
Indirect Capital Costs
a
Engineering
Construction and Field Expenses (0.1 TDC)
Construction Fees (0.1 TDC)
Start-up (0.02 TDC)
Performance Test (0.01 TDC)
Total Indirect Cost (TIC)
Contingencies (0.2 (TDC+TIC))
TOTAL TURNKEY COSTS (TTC)
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)
TOTAL CAPITAL INVESTMENT (TCI)
65
230
66
426
47
92
83
83
17
834
283
223
1340
79
1420
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (400xl06 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-112
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Spray Drying (Sodium)
150xl06 Btu/hr (44 MWt)
Western 0.6% S
75%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
NazCOa ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
92
GJ/hr
m3/hr
GJ/hr
105
21
33
33
13
m3/hr
155
kg/hr
36
kg/hr
kg/hr
141
kg/hr
74
38
50
240
644
$/10e Btu $/kg S02
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
3-113
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Spray Drying (Sodium)
Boiler Capacity: 20Qxl06 Btu/hr (58.6 MWf)
Coal Feedstock: Western 0.6% S
S02 Control Level: 75%
Item
Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling
S02 Scrubbing
Fans
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services
Total Direct Costs (TDC)
Indirect Capital Costs
a
Engineering
Construction and Field Expenses (0.1 TDC)
Construction Fees (0.1 TDC)
Start-up (0.02 TDC)
Performance Test (0.01 TDC)
Total Indirect Cost (TIC)
\
Contingencies (0.2 (TDC+TIC))
TOTAL TURNKEY COSTS (TTC)
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)
TOTAL CAPITAL INVESTMENT (TCI)
75
247
73
470
59
92
92
92
18
917
303
244
1464
91
1556
a. Engineering Costs =0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal @ 90%
S02 removal.
b. Reference: A
c. From Annual Cost Table
B-114
-------
ANNUALIZED COSTS FOR.FGD PROCESSES.
FGD Type: Spray Drying (Sodium)
Boiler Capacity: 200xl06 Btu/hr (58.6
Coal Feedstock: Western 0.6% S
S02 Control Level: 75%
Operating Factor: 5Q%
Item Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h) 105
Supervision (15.63/m-h) 21
Maintenance Labor (.04 TDC) 37
Maintenance Materials (.04 TDC) 37
Electricity (25.8 mills/kWh) 105 kV^ 14
Steam ($1.84/GJ) GJ/hr
Proc. Water ($.04/m3) 5.2 m3/nr 1
Methane ($2.05/GJ) GJ/hr
Wastewater Treating ma/hr
Solids Disposal ($.044/kg) 219 kg/hr 51
Chemicals
Lime ($.0385/kg) kg/hr
Limestone ($.0143/kg) kg/hr
Na2C03 ($.0991/kg) 188 kg/hr 98
Total Direct Operating Cost 364
Overhead
Payroll (.3x(l+2) above) 38
Plant (.26x(1+2+3+4) above) 52
Total Overhead Costs 90
By-Product Credits ( )
Capital Charges
Capital Recovery (.17 TCI) 264
718
TOTAL ANNUALIZED COSTS
Annual Unit Costs $/106 Btu $/kg SO?
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
3-115
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: _Spray Drying (Lime)
Boiler Capacity: 4QOxl06 Btu/hr (117 MWt)
Coal Feedstock: Eastern 2.3% S
S02 Control Level: 70%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 179
349
S02 Scrubbing ___
Fans
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection 830
Purge Treatment
Sulfur Production
QQ
Utilities and Services
Total Direct Costs (TDC) 1,573
Indirect Capital Costs
Engineering 157
Construction and Field Expenses (0.1 TDC) 157
Construction Fees (0.1 TDC) 157
Start-up (0.02 TDC) 31_
Performance Test (0.01 TDC) 16
Total Indirect Cost (TIC) 518
Contingencies (0.2 (TDC+TIC)) 418
TOTAL TURNKEY COSTS (TTC) 2509
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)°
TOTAL CAPITAL INVESTMENT (TCI) 2682
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOa removal. This cost remains constant for the smaller cases
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
b. Reference: 4
<_. From Annual Cost Table
B-116
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Spray Drying (Lime)
400x106Btu/hr (118 MW )
Eastern 2.3% S
70%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
63
63
209
kWp
28
GJ/hr
10
m3/hr
2
GJ/hr
m3/hr
1181
631
kg/hr
kg/hr
273
128
kg/hr
kg/hr
38
66
683
104
455
1242
$/106 Btu
$/kg S0:
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD T^ype: Spray Drying (Sodium)
Boiler Capacity: 150xlO? Btu/hr (44 MWt)
Coal Feedstock: Western 0.6% S
S02 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 75
S02 Scrubbing 230
Fans 66
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection 426_
Purge Treatment
Sulfur Production
Utilities and Services 48
Total Direct Costs (TDC) 845
Indirect Capital Costs
Engineering 92
Construction and Field Expenses (0.1 TDC) 85
Construction Fees (0.1 TDC) 85
Start-up (0.02 TDC) 17
Performance Test (0.01 TDC) 9
Total Indirect Cost (TIC) 288
Contingencies (0.2 (TDC+TIC)) 226
TOTAL TURNKEY COSTS (TTC) 1359
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)° 90
TOTAL CAPITAL INVESTMENT (TCI) 1450
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOa removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
b. Reference: 4
c. From Annual Cost Table
B-118
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Spray Drying (Sodium)
150xl06 Btu/hr (44 MW
Western 0,6% S
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.OA TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.OA/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
34
34
93
Mfe
13
GJ/hr
4.4
m3/hr
1
GJ/hr
m3/hr
227
kg/hr
53
kg/hr
kg/hr
194 kg/hr 101
38
50
362
245
695
$/106 Btu
$/kg SO2
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-119
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Spray Drying (Sodium)
Boiler Capacity: 150xl06 Btu/hr (44 MWt)
Coal Feedstock: Western 0.6% S
S02 Control Level: 50%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 56
S02 Scrubbing 230
Fans 66
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection 426
Purge Treatment
Sulfur Production
Utilities and Services 47
826
Total Direct Costs (TDC)
Indirect Capital Costs
Engineering 92
Construction and Field Expenses (0.1 TDC) 83
Construction Fees (0.1 TDC) 83
Start-up (0.02 TDC) 16
Performance Test (0.01 TDC) 8
Total Indirect Cost (TIC) 282
Contingencies (0.2 (TDC+TIC)) 222
TOTAL TURNKEY COSTS (TTC) 1330
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)° 70
TOTAL CAPITAL INVESTMENT (TCI) 1401
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOa removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
b. Reference: 4
c. From Annual Cost Table
B-120
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Spray Drying (Sodium)
150xl06 Btu/hr (44 MWfc)
Western 0.6% S
50%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
N32C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
92
GJ/hr
4.4 m3/hr
GJ/hr
m'/hr
109 kg/hr
$/106 Btu
105
21
33
33
13
25
kg/hr
kg/hr
97 kg/hr 51
38
50
I/kg SOj,
78?.
88
237
607
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-121
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Spray Drying (Lime)
Boiler Capacity: 150xlQ6 Btu/hr (44 MWfc)
Coal Feedstock: Western 0.6% S
S02 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 94
S02 Scrubbing 230
Fans 66
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection 426
Purge Treatment
Sulfur Production
Utilities and Services 49
865
Total Direct Costs (TDC)
Indirect Capital Costs
Engineering 92
Construction and Field Expenses (0.1 TDC) 86
Construction Fees (0.1 TDC) 86
Start-up (0.02 TDC) 17_
Performance Test (0.01 TDC) 9
Total Indirect Cost (TIC) 290
Contingencies (0.2 (TDC+TIC)) 231
TOTAL TURNKEY COSTS (TTC) 1386
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)C 73
TOTAL CAPITAL INVESTMENT (TCI) 1460
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOa removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: A
c. From Annual Cost Table
B-122
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
_Spray Drying (Lime)
_150xl06 Btu/hr (44
_Western 0.6% S
90%
__ 60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.-0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
93
GJ/hr
4.4 m3/hr
236
_GJ/hr
_m3/hr
_kg/hr
105
21
35
35
13
141
kg/hr
29
kg/hr
kg/hr
38
51
$/106 Btu
1/kgSOj
294
247
630
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-123
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Spray Drying (Lime)
Boiler Capacity: 150xl06 Btu/hr (44 MW,J
Coal Feedstock: Western 0.6% S
S02 Control Level: 50%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 69
S02 Scrubbing _ 230
Fans 66
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection 426
Purge Treatment
Sulfur Production
Utilities and Services 47
Total Direct Costs (TDC) 839
Indirect Capital Costs
a 92
Engineering
Construction and Field Expenses (0.1 TDC) 84
Construction Fees (0.1 TDC) 84
Start-up (0.02 TDC) 17_
Performance Test (0.01 TDC) 8
Total Indirect Cost (TIC) 285
Contingencies (0.2 (TDC+TIC)) 225
TOTAL TURNKEY COSTS (TTC) 1349
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)C 58
TOTAL CAPITAL INVESTMENT (TCI) 1408
d. Engineering Costs = 0.1 TDC for the 58.6 MW (200xlOb Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
b. Reference: 4
<_. From Annual Cost Table
B-124
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Spray Drying (Lime)
150x106 Btu/hr (44 MWt)
_Western 0.6% S
50%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h) 105
Supervision (15.63/m-h) 21
Maintenance Labor (.04 TDC) 34
Maintenance Materials (.04 TDC) 34
Electricity (25.8 mills/kWh) _9_2__k% 12
Steam ($1.84/GJ) GJ/hr
Proc. Water ($.04/m3) 4.4 m3/hr 1_
Methane ($2.05/GJ) GJ/hr
a i
Wastewater Treating m /hr
Solids Disposal ($.044/kg) 99 kg/hr 23
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
Na2COs ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above) 38
Plant (.26x(1+2+3+4) above) 50
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs $/106 Btu $/kg_S02
46
kg/hr
9
kg/hr
kg/hr
239
0 )
238
565
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-125
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Spray Drying (Lime)
Boiler Capacity: 400xlQ6 Btu/hr (118 MWt)
Coal Feedstock: Western 0.6% S
S02 Control Level: 70% ^
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 116
S02 Scrubbing 337
Fans 124
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection 840^
Purge Treatment
Sulfur Production
Utilities and Services 85
Total Direct Costs (TDC) 1501
Indirect Capital Costs
Engineering 150
Construction and Field Expenses (0.1 TDC) 150
Construction Fees (0.1 TDC) 15Q
Start-up (0.02 TDC) 30
Performance Test (0.01 TDC) 15
Total Indirect Cost (TIC) 495
Contingencies (0.2 (TDC+TIC)) 399
TOTAL TURNKEY COSTS (TTC) 2395
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs) 106
TOTAL CAPITAL INVESTMENT (TCI) 2503
=1. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOa removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal @ 90%
S02 removal.
b. Reference: 4
i-, FroTTi Annual Cost Table
B-126
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Spray Drying (Lime)
400xl06 Btu/hr (118 MW )
^Western 0.6% S
70%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam (51.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
60
203
k%
60
27
GJ/hr
10
rnVhr
2
GJ/hr
m3/hr
424
227
kg/hr
kg/hr
98
46
kg/hr
kg/hr
38
64
419
102
425
946
$/106 Btu
$/kg SO,
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-127
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Spray Drying (Sodium)
Boiler Capacity: 60xlQ6 Btu/hr (17 MW
Coal Feedstock: Western 0.6% S
S02 Control Level: 75%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 42
S02 Scrubbing
Fans 33
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection 200
Purge Treatment
Sulfur Production
Utilities and Services 25
450
Total Direct Costs (TDC)
Indirect Capital Costs
Engineering 92
Construction and Field Expenses (0.1 TDC) 45
Construction Fees (0.1 TDC) 45
Start-up (0.02 TDC) 9_
Performance Test (0.01 TDC) 5
Total Indirect Cost (TIC) 196
Contingencies (0.2 (TDC+TIC)) 129
TOTAL TURNKEY COSTS (TTC) 775
Landb (0.00084 TTC) 0.6
Working Capital (0.25 Direct Operating Costs)
c 53
TOTAL CAPITAL INVESTMENT (TCI) 828
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOa removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-128
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Spray Drying (Sodium)
60xl06 Btu/hr (17 MWt)
Western 0.6% S
75%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.Q4/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
45
GJ/hr
1.8 m3/hr
GJ/hr
i3/hr
65 kg/hr
$/106 Btu
105
21
18
18
0.4
15
kg/hr
kg/hr
56 kg/hr 29
38
42
$/kg S0_;
212
80
140
432
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-129
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Spray Drying
Boiler Capacity: 75xlQ6 Btu/hr (22 MWtl
Coal Feedstock: Eastern 2.3% S
S02 Control Level: 70%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 87
S02 Scrubbing
Fans
Waste-water Pumps
Regeneration
Solids Separation
Solids Collection 250
Purge Treatment .
Sulfur Production
Utilities and Services 32
Total Direct Costs (TDC) 582
Indirect Capital Costs
Engineering 92
Construction and Field Expenses (0.1 TDC) 58
Construction Fees (0.1 TDC) 58
Start-up (0.02 TDC) 12
Performance Test (0.01 TDC) 6
Total Indirect Cost (TIC) 226
Contingencies (0.2 (TDC+TIC)) 162
TOTAL TURNKEY COSTS (TTC) 970
Landb (0.00084 TTC) 1_
c 65
Working Capital (0.25 Direct Operating Costs)
TOTAL CAPITAL INVESTMENT (TCI) 1036
a. Engineering Costs - 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-130
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Spray Drying (Lime)
75xl06 Btu/hr (22 MI
Eastern 2.3% S
70%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
NazCOa ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
62
GJ/hr
105
21
25
25
2.2
m3/hr 0.5
GJ/hr
m3/hr
222
119
kg/hr S2
kg/hr 24
kg/hr
kg/hr
38
46
260
84
176
520
$/106 Btu
$/kg S02
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-131
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Wellman-Lord
Boiler Capacity: 8.8 MWfc (30 MBtu/hr)
Coal Feedstock: 3.5% S Eastern
S02 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 23
S02 Scrubbing 104
Fans 32
Wastewater Pumps —
Regeneration 176
Solids Separation ^_
Solids Collection ^_
Purge Treatment 113
Sulfur Production 298
Utilities and Services 48
796
Total Direct Costs (TDC)
Indirect Capital Costs
Engineering 257
Construction and Field Expenses (0.1 TDC) 80
Construction Fees (0.1 TDC) 80
Start-up (0.02 TDC) 16
Performance Test (0.01 TDC) 8
Total Indirect Cost (TIC) 441
Contingencies (0.2 (TDC+TIC)) 247
TOTAL TURNKEY COSTS (TTC) 1484
Landb (0.00084 TTC) 1.2
Working Capital (0.25 Direct Operating Costs)° 54
TOTAL CAPITAL INVESTMENT (TCI) 1539
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-132
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Wellman-Lord
8.8 MWt (30 MBtu/hr)
3.5% S Eastern
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
32
32
17
1.2
4.1
0.6
k^g
GJ/hr
m3/hr
GJ/hr
2
12
1
6
m3/hr
kg/hr
kg/hr
kg/hr
12
kg/hr
6
38
49
217
87
262
558
$/106 Btu $/kg SO 2
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-133
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: ' Wellman-Lord
Boiler Capacity: 8.8 MWfc (30 MBtu/hr^
Coal Feedstock: 0.6% S Western
S02 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 9
SOa Scrubbing 104
Fans _ 32
Wastewater Pumps ^_
Regeneration 67
Solids Separation ^_
Solids Collection ~
46
Purge Treatment
Sulfur Production 128
Utilities and Services 16
Total Direct Costs (TDC) 370
Indirect Capital Costs
Engineering 257
Construction and Field Expenses (0.1 TDC) 37
Construction Fees (0.1 TDC) ^7
Start-up (0.02 TDC) ^_
Performance Test (0.01 TDC) 4
Total Indirect Cost (TIC) 342
Contingencies (0.2 (TDC+TIC)) 142
TOTAL TURNKEY COSTS (TTC) 854
Landb (0.00084 TTC) 0.7
Working Capital (0.25 Direct Operating Costs) 4]
TOTAL CAPITAL INVESTMENT (TCI) 896
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOa removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: A
c. From Annual Cost Table
B-134
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Wei] man-Lord
8.8 MWt (30 MBtu/hr)
0.6% S Western
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
NazCOa ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
15
0.25GJ/hr
1.7 mVhr
Q.13GJ/hr
Vhr
kg/hr
105
21
15
15
kg/hr
kg/hr
2
kg/hr
1
38
41
162
79
152
385
$/106 Btu $/kg S02
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-135
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: __Wellman-Lord
Boiler Capacity: 22 MWfc (75 MBtu/hr)
Coal Feedstock: 3.5% S Eastern
S02 Control Level: 9Q%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 38
S02 Scrubbing 217
Fans 59
Wastewater Pumps
Regeneration 317
Solids Separation ~
Solids Collection -
Purge Treatment 199
Sulfur Production 497
Utilities and Services 93
Total Direct Costs (TDC) 1420
Indirect Capital Costs
Engineering 257
Construction and Field Expenses (0.1 TDC) 142
Construction Fees (0.1 TDC) 142
Start-up (0.02 TDC) 28
Performance Test (0.01 TDC) 14
Total Indirect Cost (TIC) 583
Contingencies (0.2 (TDC+TIC))
TOTAL TURNKEY COSTS (TTC) 2404
Landb (0.00084 TTC) 2.0
Working Capital (0.25 Direct Operating Costs)° 77
TOTAL CAPITAL INVESTMENT (TCI) 2483
d. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
B-136
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Wellman-Lord
22 MWt (75 MBtu/hr)
3.5% S Eastern
90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
S7
57
43.5
3.0
4.5
1.4
kH=
GJ/hr
m3/hr
GJ/hr
6
29
1
15
m3/hr
kg/hr
kg/hr -_
kg/hr -_
30 kg/hr 16
38
62
$/106 Btu
$/kg
307
100
( -20
422
809
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-137
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
Wellman-LorcL
FGD Type:
Boiler Capacity: 22 MW (75 MBtU/hr)
Coal Feedstock: 0.6% S Western
S02 Control Level:
90%
Item
Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling
S02 Scrubbing
Fans
Wastewater Pumps
Regeneration
Solids Separation
Solids Collection
Purge Treatment
Sulfur Production
Utilities and Services
Total Direct Costs (TDC)
Indirect Capital Costs
a
Engineering
Construction and Field Expenses (0.1 TDC)
Construction Fees (0.1 TDC)
Start-up (0.02 TDC)
Performance Test (0.01 TDC)
Total Indirect Cost (TIC)
Contingencies (0.2 (TDC+TIC))
TOTAL TURNKEY COSTS (TTC)
Landb (0.00084 TTC)
Working Capital (0.25 Direct Operating Costs)
TOTAL CAPITAL INVESTMENT (TCI)
16
217
59
119
79
211
257
73
73
15
732
425
231
1388
1.2
51
1440
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
S02 removal.
b. Reference: 4
c. From Annual Cost Table
E-138
-------
ANOTJALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Wei Iman-T.nrrl
22 MWj- (75 MBtu/hr)
0.6% S Western
90%
-GUI-
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
NaaCOs ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(1+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
29
29
38
0
3
0
.1 kVfe
.64GJ/hr
.4 m3/hr
.30GJ/hr
5
6
1
3
m3/hr
kg/hr
kg/hr
kg/hr
6
kg/hr
3
38
48
202
86
-4
245
529
$/106 Btu
$/kg S02
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-139
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: Wellman-Lord
Boiler Capacity: 58.6 MWfc (200 MBtu/hr)_
Coal Feedstock: 3.5% S Eastern
S02 Control Level: 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 58
432
S02 Scrubbing
Fans 106
Wastewater Pumps ~^_
Regeneration 587
Solids Separation ^_
Solids Collection -
Purge Treatment 354
Sulfur Production 852
Utilities and Services 184
Total Direct Costs (TDC) 2573
Indirect Capital Costs
a
Engineering 257
Construction and Field Expenses (0.1 TDC) 257
Construction Fees (0.1 TDC) 257
Start-up (0.02 TDC) 51
Performance Test (0.01 TDC) 26
Total Indirect Cost (TIC) 848
Contingencies (0.2 (TDC+TIC)) 534
TOTAL TURNKEY COSTS (TTC)
Landb (0.00084 TTC) 3.4
Working Capital (0.25 Direct Operating Costs)C -105
TOTAL CAPITAL INVESTMENT (TCI) 4233
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% SOz removal. This cost remains constant for the smaller cases.
For the 118 MW (AOOxlO6 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal Q 90%
S02 removal.
b. Reference: A
c. From Annual Cost Table
B-140
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
S02 Control Level:
Operating Factor:
Wellman-Lord
58.6 MWL (200 MBtu/hr)
3.5% S Eastern
_90%
60%
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(l+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Charges
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
105
21
103
103
105
8.
9.
3.
0
6
8
Mfe
GJ/hr
m3/hr
GJ/hr
14
77
2
41
m3/hr
kg/hr
kg/hr ^_
kg/hr ^_
61 Jcg/hr 32
38
86
498
124
-53
720
1289
$/106 Btu
_$/kgS02
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-141
-------
CAPITAL INVESTMENT COSTS FOR FGD PROCESSES
FGD Type: _Wellman-Lord
Boiler Capacity: 58.6 MWfc (200 MBtu/hrJ
Coal Feedstock: 0.6% S Western
S02 Control Level:_ 90%
Item Cost (Thousands of dollars)
Direct Capital Costs
Raw Material Handling 24
S02 Scrubbing 438
Fans 107
Wastewater Pumps —
?9O
Regeneration ^u
Solids Separation ~^_
Solids Collection ^_
Purge Treatment L^L
Sulfur Production 362
Utilities and Services ft?
Total Direct Costs (TDC) 1354
Indirect Capital Costs
Engineering
a 257
Construction and Field Expenses (0.1 TDC) 135
Construction Fees (0.1 TDC) 135
Start-up (0.02 TDC) ^
Performance Test (0.01 TDC) 14
Total Indirect Cost (TIC) 568
Contingencies (0.2 (TDC+TIC)) 334
TOTAL TURNKEY COSTS (TTC) 2306
Landb (0.00084 TTC) 1.9
Working Capital (0.25 Direct Operating Costs)° 7Q
TOTAL CAPITAL INVESTMENT (TCI) 2378
a. Engineering Costs = 0.1 TDC for the 58.6 MW (200xl06 Btu/hr) case burning 3.5% S coal
@ 90% S02 removal. This cost remains constant for the smaller cases.
For the 118 MW (400xl06 Btu/hr) cases, the engineering cost is
constant and equal to that case burning 3.5% S coal 0 90%
SO2 removal.
b. Reference: 4
c. From Annual Cost Table
B-142
-------
ANNUALIZED COSTS FOR FGD PROCESSES
FGD Type:
Boiler Capacity:
Coal Feedstock:
SO; Control Level:
Operating Factor:
Wellman-Lord-
58.6 MWt (200 MBtu/hr^
0.6% S Western
90%
_6XLZ
Item
Cost (Thousands of dollars)
Direct Costs
Operating Labor (12.02/m-h)
Supervision (15.63/m-h)
Maintenance Labor (.04 TDC)
Maintenance Materials (.04 TDC)
Electricity (25.8 mills/kWh)
Steam ($1.84/GJ)
Proc. Water ($.04/m3)
Methane ($2.05/GJ)
a
Wastewater Treating
Solids Disposal ($.044/kg)
Chemicals
Lime ($.0385/kg)
Limestone ($.0143/kg)
Na2C03 ($.0991/kg)
Total Direct Operating Cost
Overhead
Payroll (.3x(l+2) above)
Plant (.26x(l+2+3+4) above)
Total Overhead Costs
By-Product Credits
Capital Chargesb
Capital Recovery (.17 TCI)
TOTAL ANNUALIZED COSTS
Annual Unit Costs
89.7 k%
1.7 GJ/hr
105
21
54
54
12
16
7.2
0.79
m3/hr
GJ/hr
2
8
mVhr
kg/hr
kg/hr
kg/hr
13
kg/hr
7
38
61
$/106 Btu
$/kg S0:
279
99
( -11
404
771
a. Reference: 5
b. Capital charges include investment return of 10%; straight-line depreciation,
(15 yrs. FGD life) G&A, property taxes and insurance. Federal and State income
taxes are not included.
B-143
-------
TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
1 REPORT NO
EPA-600/7-79-1781
3. RECIPIENT'S ACCESSION NO.
4. TITLE AND SUBTITLE
Technology Assessment Report for Industrial Boiler
Applications: Flue Gas Desulfurization
5. REPORT DATE
November 1979
6. PERFORMING ORGANIZATION CODE
7. AUTHOR(S)
J. C. Dickerman and K. L. Johnson
8. PERFORMING ORGANIZATION REPORT NO
9. PERFORMING ORGANIZATION NAME AND ADDRESS
Radian Corporation
P.O. Box 8650
Durham, North Carolina 27707
10. PROGRAM ELEMENT NO.
EHE624
11. CONTRACT/GRANT NO.
68-02-2608, Task 47
12. SPONSORING AGENCY NAME AND ADDRESS
EPA, Office of Research and Development
Industrial Environmental Research Laboratory
Research Triangle Park, NC 27711
13. TYPE OF REPORT AND PERIOD COVERED
Task Final; 6/78 - 10/79
14. SPONSORING AGENCY CODE
EPA/600/13
IB.SUPPLEMENTARY NOTES JERL-RTP project officer is John E. Williams, Mail Drop 61,
919/541-2483.
16. ABSTRACT
The report gives results of an assessment of the applicability of flue gas
desulfurization (FGD) technology to industrial boilers and is one of a series to aid in
determining the technological basis for a New Source Performance Standard for In-
dustrial Boilers. The development status and performance of alternative FGD con-
trol techniques were assessed and the cost, energy, and environmental impacts of
the most promising were identified. The study concluded that there is no best FGD
technology for application to industrial boilers: each alternative has advantages and
disadvantages which could make it best for a specific application. Cost estimates of
applying FGD processes indicated that the cost effectiveness varies significantly
depending on the fuel fired, boiler size, and control level. However, boiler size is
the most significant factor affecting cost effectiveness: the economy of scale causes
control of large sources to be the most effective. The energy requirement of applying
FGD processes varied from about 0. 5% to 6% of boiler capacity, excluding stack gas
reheat. The environmental impacts of each alternative were evaluated: each could be
applied in an environmentally acceptable manner under existing regulations. The
report does not consider combinations of technology to remove all pollutants, and
these findings have not undergone detailed assessments for regulatory action.
17.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.IDENTIFIERS/OPEN ENDED TERMS
c. COSATI Field/Group
Pollution
Flue Gases
Desulfurization
Assessments
Boilers
Pollution Control
Stationary Sources
Industrial Boilers
13B
2 IB
07A,07D
14 B
13A
13. DISTRIBUTION STATEMENT
Release to Public
19. SECURITY CLASS (This Report)
Unclassified
21. NO. OF PAGES
664
20. SECURITY CLASS (This page)
Unclassified
22. PRICE
EPA Form 22^0-1 (9-73)
B-144
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