EPA-650/2-74-098
September 1974
Environmental Protection Technology Series
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EPA-650/2-74-098
EVALUATION OF R&D INVESTMENT
ALTERNATIVES FOR SOX AIR POLLUTION
CONTROL PROCESSES
by
D. Shore, J. J. O'Donnell, and F. K. Chan
Research and Engineering Development
The M.W. Kellogg Company
Houston, Texas 77046
Contract No. 68-02-1308 (Task 7)
ROAPNo. 21ADE-029
Program Element No. 1AB013
EPA Task Officer: G.J. Foley
Control Systems Laboratory
National Environmental Research Center
Research Triangle Park, North Carolina 27711
Prepared for
OFFICE OF RESEARCH AND DEVELOPMENT
U.S. ENVIRONMENTAL PROTECTION AGENCY
WASHINGTON, D.C. 20460
September 1974
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This report has been reviewed by the Environmental Protection Agency
and approved for publication. Approval does not signify that the
contents necessarily reflect the views and policies of the Agency,
nor does mention of trade names or commercial products constitute
endorsement or recommendation for use.
11
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EVALUATION OF R&D INVESTMENT
ALTERNATIVES FOR SO AIR
X
POLLUTION CONTROL PROCESSES
PART 1
TASK NO. 7 FINAL REPORT
CONTRACT NO. '68-02-1308 & CPA 70-68
by
THE M.W. KELLOGG COMPANY
RESEARCH & ENGINEERING DEVELOPMENT
HOUSTON, TEXAS
PROJECT OFFICER: GARY J. FOLEY
CONTROL SYSTEMS LABORATORY
NATIONAL ENVIRONMENTAL RESEARCH CENTER
RESEARCH TRIANGLE PARK, N.C. 27711
Prepared for
OFFICE OF RESEARCH AND DEVELOPMENT
U.S. ENVIRONMENTAL PROTECTION AGENCY
WASHINGTON, D.C. 20460
SEPTEMBER, 1974
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RESEARCH AND ENGINEERING DEVELOPMENT
EVALUATION OF R&D INVESTMENT
ALTERNATIVES FOR SOV AIR
A,
POLLUTION CONTROL PROCESSES
PART 1
TASK NO. 7 FINAL REPORT
Submitted to
ENVIRONMENTAL PROTECTION AGENCY
OFFICE OF RESEARCH & DEVELOPMENT
CONTROL SYSTEMS LABORATORY
CONTRACT NO. 68-02-1308 & CPA 70-68
Approved
A.G. Sliger
Project Director
W.C. Schreiner
Manager, Chemical Engineering Dept,
M.J. Walr (
Vice-President, R&ED
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THE M. W. KELLOGG COMPANY
A DIVISION OF PULLMAN INCORPORATED
Research & Engineering Development
IKIUOOO]
PAGE NO.
REPORT NO.
EVALUATION OF R&D INVESTMENT
ALTERNATIVES FOR SOX AIR
POLLUTION CONTROL PROCESSES
PART 1
TASK NO. 7 FINAL REPORT
EPA-ORM-CSD CONTRACT NO. 68-02-1308 & CPA 70-68
SEPTEMBER, 1974
D. Shore, J.J. O'Donnell, F.K. Chan, H.A. Khan and
MWK Estimating Department
Staff:
Period Covered:
RDO No.:
Distribution:
October, 1972 to September, 1974
4092-22, 23, 24, 25, 27 & 4118-7
Office of Research & Development (EPA)
L.C. Axelrod
J.S. Burr
F.K. Chan
C.F. Chatfield
A.E
W.C
C.J
J.B
J.A
L.D
S.E
J.J
J.J
W.C
D.
G.T
A.G
M. J
RID
. Cover
. Crady
. Donovan
. Dwyer
. Finneran
. Fraley
. Handman
. McKenna
. O'Donnell
. Schreiner
Shore
. Skaperdas
. Sliger
. Wall
(4)
Authors:
Copy No.
1-100
101
102
103
104 .
105
106
107 .
108
109
110
111
112
113
114
115
116
117
118
-123
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TABLE OF CONTENTS
PAGE NO.
1. Introduction 1
2. Conclusions and Recommendations 5
2.1 Control of Sulfur Dioxide Emissions From 5
Existing Sources
2.2 Stack Gas Scrubbing Costs •*
2.2.1 The Utility Industry
2.2.2. The Industrial Boilers
2.3 Substitute Natural Gas Production 6
2.4 Solvent Refined Coal Production 7
2.5 The Lurgi Gasifier with Combined Cycle 8
2.6 The Fluidized Pressurized Combustor with
a Combined Power Cycle 8
3. Major Sources of Sulfur Dioxide in U. S. 9
3.1 Introduction 9
3.2 Data - 9
3.2.1 Sources
3.2.2 Data Quality and Comparison with
other Published Information
3.3 U. S. Sulfur Dioxide Emis'sions 12
3.3.1 Utility Plant
3.3.2 Smelters (Copper, Zinc, Lead)
3.3.3 Industrial Boilers
3.3.4 Acid Plants
3.3.5 Sulfur Plants
3.4 Summary 19
4. The General Model 54
4.1 The General Process Model 54
4.2 The General Cost Model 55
4.2.1 Basis For Costs
4.2.2 Capital Cost Model
4.2.3 Operating Cost Model
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TABLE OF CONTENTS(CONT'D)
PAGE NO,
4.3 Effect of Location on Plant Cost 66
4.4 Nomenclature 68
The Wet Limestone Process 77
5.1 Process Appraisal 77
5.2 Evaluation of Catalytic Inc. Estimate 80
5.3 Variation of Equipment Costs with Plant
Size 80
5.4 Cost Model 83
5.4.1 Equipment Costs
5.4.2 Other Material Costs and Labor Cost
5.4.3 Raw Materials and Utilities Costs
5.4.4 Total Plant Investment and Total
Capital Required
5.4.5 Operating Costs
5.5 Effect of Various Parameters On Costs 92
5.6 Nomenclature 94
The Wellman/Allied Process 106
6.1 Process Appraisal 106
6.2 Evaluation of the NIPSCO Project Cost Estimate 108
6.3 Variation of Equipment Costs with Plant Size 109
6.4 Cost Model 110
6.4.1 Equipment Costs
6.4.2 Other Material Costs and Labor Costs
6.4.3 Raw Materials and Utilities Costs
6.4.4 Total Plant Investment and Total
Capital Required
6.4.5 Operating Costs
6.5 Effect of Various Parameters on Cost 123
6.6 Wellman/Allied Process Variations and Impact
on Cost Model 123
6.7 Nomenclature 125
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TABLE OF CONTENTS (CONT'D)
PAGE NO,
Application of Stack Gas Scrubbing Models 136
7.1 Stack Gas Scrubbing Applied to Existing
Utilities 136
7.2 Stack Gas Scrubbing Applied to Industrial
Boilers 139
7.2.1 Wet Limestone Process
7.2.2 Wellman/Allied Process
7.2.3 Applicability to Industrial Boilers
Substitute Natural Gas Production Using a Lurgi
Oxygen Gasifier 171
8.1 Process Model 171
8.1.1 Coal Types
8.1.2 Coal, Oxygen and Steam Requirements
for the SNG Plant
8.1.3 Electrical Power and High Pressure
Steam Requirements for the SNG Plant
8.1.4 Sample Calculation of Plant Total
Coal Requirement
8.2 Cost Model 175
8.2.1 Major Equipment Costs, E
8.2.2 Total Net Annual Operating Cost
8.2.3 Total Plant Investment, Total Capital
Required and Total Production Cost
8.2.4 Calculation of Costs for Three Types
of Coal in Three Different Locations
8.2.5 The Influence of Coal Type, Coal Cost,
Percentage Sulfur and Plant Location
on Gas Cost
Solvent Refined Coal Production 198
9.1 Process Appraisal 198
9.2 Process Description 200
9.3 Cost Model 202
9.3.1 Total Plant Investment
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TABLE OF CONTENTS (CONT'D)
PAGE NO,
9.3.2 Total Net Annual Operating Cost,
Total Capital Requirement and
Total Annual Production Cost
9.3.3 Calculation of Costs of Solvent
Refined Coal
9.4 Conclusions 206
10. The Combined Gas Turbine - Steam Turbine Power
Plant Using a Low Btu Lurgi Gasifier 21°
10.1 Introduction 21°
10.2 The Lurgi Gasification Plant 212
10.3 Description of Cycles Studied 214
10.4 Discussion of Results 216
10.5 Cost Model 217
11. Pressurized Fluidized Bed Steam Generator with
Dry Dolomite Injection for SO2 Removal
240
11.1 Process Appraisal
11.2 Process Description 240
11.3 Conclusions 243
11.4 Addendum 243
12. References 253
13. Appendices 257
Appendix A. General Cost Model Derivations 258
Appendix B. Wet Limestone Process-Catalytic Inc.
Estimate, 500 Megawatt, Equipment Costs
(Material & Subcontracts) 264
Appendix C. Wet Limestone Process-Catalytic Inc.
Estimate, Labor and Material Factors 266
Appendix D. A Summary of Comments Made During and
After the Presentation by MWK to EPA,
December 13, 1973 267
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LIST OF TABLES
TABLE NO.
3.1
3.2
3.3
3.4
3.5
3.6
3.7
3.8
3.9
3. .10
3.11
3.12
3.13
4.1
4.2
5.1
5.2
5.3
DESCRIPTION PAGE NO.
Comparison of Sulfur Dioxide Emissions
Between Three Different Sources.
Utilities Fuel Consumption and Sulfur
Emission for 1971.
Statewise Distribution of Fuel Burned
By Utilities in 1971.
U. S. Utility Industry (1971) -
Capacity Distribution.
U. S. Utility Industry Statewise S02
Emissions (1971) .
U. S. Utility Statistics By Plant Size.
U. S. Smelters SO2 Emission.
Industrial Boilers, Coal and Oil
Fired - Statistics By Capacity
Industrial Boilers, Coal. and Oil
Fired - Statistics By States
U. S. Acid Plant Statistics By Plant Size
and Plant Type
U. S. Acid Plants - Statistics By State,
U. S. Sulfur Plants - Statistics By Plant
Size
U. S. Sulfur Plants - Statistics By State
Location Factors for Major U. S. Cities
Average Location Factors for Each State
Boiler Retrofit factors
Unit Costs Used in Illustrative Examples
in Wet-Limestone Stack Gas Scrubbing
Model
Wet-Limestone Process and Cost Model.
Summary of Equations
20
21
22
23
24
25
26
27
28
29
30
31
32
70
71
87
97
98
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LIST OF TABLES (CONT'D)
TABLE NO. DESCRIPTION PAGE NO.
6.1 Unit Costs Used in Illustrative Examples
in Wellman/Allied Stack Gas Scrubbing
Model 127
6.2 Wellman/Allied Process and Cost Model -
Summary of Equations 128
7.1 Boiler Size Distribution For Standard
Size Utility Plant 145
7.2 Summary of Equipment Cost Equations
for Industrial Boilers - Wet Limestone
Process 146
7.3 Summary of Operating Cost Equations For
Industrial Boilers - Wet Limestone
Process 147
7.4 Summary of Equipment Cost Equations for
Industrial Boilers - Wellman/Allied
Process 148
8.1 Summary of Major Equipment Cost Equations -
Substitute Natural Gas Production 190
10.1 Cost of a 1000 Net Megawatt Conventional
Power Station 220
10.2 Cost of a 1000 Net Megawatt Combined Cycle
Power Plant 221
11.1 Gas Stream Compositions - Fluidized Bed
Combustion 245
11.2 Solid Streams Composition - Fluidized Bed
Combustion 246
11.3 Power Generation of an FBC Combined Cycle
Plant 247
11.4 Heat to Steam Cycle 248
11.5 Cost of a 1000 Net Megawatt FBC Combined
Cycle Plant 249
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LIST OF FIGURES
FIGURE NO. TITLE PAGE NO.
3.1
3.2
3.3
3.4
3.5
3.6
3.7
3.8
3.9
3.10
3.11
3.12
3.13
3.14
3.15
3.16
Distribution of U. S. Utility Plants With
Plant Size
Size Distribution of U. S. Utility Boilers
Distribution of Plant Load Factors for the
U. S.. Utility Industry
Variation of Plant Load Factor With Utility
Plant Size
Geographical Distribution of SO- Emissions
From The Utility Industry
S02 Emissions From U. S. Utility Plants
Distribution of Boilers and Average Plant
Age With Utility Plant Size
Distribution of Boilers and Boiler Age With
Utility Boiler Size
Distribution of U. S. Smelters With Plant
Size
S02 Emissions From U. S. Smelters
Geographical Distribution of SO- Emissions
From Smelters
Distribution of U. S. Industrial Boilers
With Boiler Size
Size Distribution of U. S. Industrial
Boilers
Geographical Distribution of SO- Emissions
From Industrial Boilers
SO- Emissions From U. S. Industrial Boilers
Distribution of U. S. Sulfuric Acid Plants
33
34
35
36
37
38
39
40
41
42
43
44
45
46
47
With Plant Size 48
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LIST OF FIGURES (CONT'D)
FIGURE NO. TITLE PAGE NO.
3.17 Geographical Distribution of SO?
Emissions From Sulfuric Acid Plants 49
3.18 SO2 Emissions From U. S. Sulfuric Acid
Plants 50
3.19 Distribution of U. S. Sulfur Plants With
Plant Size 51
3.20 Geographical Distribution of SO»
Emissions From Sulfur Plants 52
3.21 SO_ Emissions From U. S. Sulfur Plants 53
4.1 Relationship Between Capital Cost Factors
in the General Cost Model 72
4.2 Relationship Between Production Cost
Factors in the General Cost Model 73
4.3 Location Factors For Selected Cities 74
4.4 Average Location Factors By State 75
4.5 Effect of Location Factor on Total Plant
Investment 76
5.1 Wet Limestone Process Flowsheet 99
5.2 Method of Varying Equipment Cost With Size 100
5.3 Effect of Boiler Capacity on Total Capital
Requirement - Wet Limestone Process 101
5.4 Effect of Boiler Capacity on Production
Cost - Wet Limestone Process 102
5.5 Effect of Boiler Retrofit Difficulty on Total
Capital Requirement - Wet Limestone Process 103
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LIST OF FIGURES (CONT'D)
FIGURE NO. TITLE PAGE NO,
5.6 Effect of Boiler Retrofit Difficulty on
Production Cost - Wet Limestone Process 104
5.7 Effect of Location Factor on Total
Capital Requirement - Wet Limestone Process 105
6.1 WeiIman/Allied Process Flowsheet 130
6.2 Effect of Boiler Capacity on Total Capital
'Requirement - Wellman/Allied Process 131
6.3 Effect of Boiler Capacity on Production
Cost - Wellman/Allied Process 132
6.4 Effect of Boiler Retrofit Difficulty on
Total Capital Requirement - Wellman/Allied
Process 133
6.5 Effect of Boiler Retrofit Difficulty on
Production Cost - Wellman/Allied Process 134
6.6 Effect of Location Factor on Total Capital
Requirement - Wellman/Allied Process 135
7.1 Average Distribution of Load Factors For
Boilers in a Utility Plant 149
7.2 Average Heat Rates For Utility Boilers 150
7.3 Average Total Capital Requirement For
Installing Wet Limestone System in Existing
Power Plants 151
7.4 Average Unit Cost For Installing Wet
Limestone System in Existing Power Plants 152
7.5 Incremental Operating Cost For Wet Limestone
System in Existing Power Plants 153
7.6 Demand For Clean Fuel as Alternative to
Stack Gas Scrubbing - Wet Limestone System
Applied to Existing Power Plants 154
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LIST OF FIGURES (CONT'D)
FIGURE NO. TITLE PAGE NO.
7.7 Cumulative Average Capital Cost - Wet
Limestone System Applied to Power Plants 155
7.8 Cumulative Incremental Operating Cost -
Wet Limestone System Applied to Power
Plants 156
7.9 Average Total Capital Requirement For
Installing Wellman/Allied System in
Existing Power Plants 157
7.10 Average Unit Cost For Installing Wellman/
Allied System in Existing Power Plants 158
7.11 Incremental Operating Cost For Wellman/
Allied System in Existing Power Plants 159
7.12 Demand For Clean Fuel as Alternative to
Stack Gas Scrubbing - Wellman/Allied
System Applied to Existing Power Plants 160
7.13 Cumulative Average Capital Cost - Wellman/
Allied System Applied to Power Plants
7.14 Cumulative Incremental Operating Cost -
Wellman/Allied System Applied to Power
Plants 162
7.15 Effect of Boiler Capacity on Total Capital
Requirement - Wet Limestone Process Applied
to Large Industrial Boilers I63
7.16 Effect of Boiler Capacity on Total Capital
Requirement - Wet Limestone Process Applied
to Small Industrial Boilers
7.17 Effect of Boiler Capacity on Operating Cost -
Wet Limestone Process Applied to Large
Industrial Boilers 165
7.18 Effect of Boiler Capacity on Operating Cost
Wet Limestone Process Applied to Small
Industrial Boilers I66
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LIST OF FIGURES (CONT'D)
FIGURE NO. TITLE PAGE NO,
7.19 Effect of Boiler Capacity on Total Capital
Requirement - WeiIman/Allied Process
Applied to Large Industrial Boilers 167
7.20 Effect of Boiler Capacity on Total Capital
Requirement - WeiIman/Allied Process
Applied to Small Industrial Boilers 168
7.21 Effect of Boiler Capacity on Operating
Cost - Wellman/Allied Process Applied to
Large Industrial Boilers 169
7.22 Effect of Boiler Capacity on Operating
Cost - WeiIman/Allied Process Applied to
Small Industrial Boilers 170
8.1 High Heating Value of Various Ranks of Coal 191
8.2 Dry Ash Free Coal> Oxygen and Steam Require-
ments For a 250 x 109 BTU/day Lurgi SNG Plant 192
8.3 Lurgi SNG Process Flow Diagram 193
8.4 Effect of Location Factor on Gas Cost -
Subbituminous Coal 194
8.5 Effect of Location Factor on Gas Cost -
Bituminous Coal 195
8.6 Effect of Carbon Content of Coal on Gas Cost 196
8.7 Effect of Carbon Content of Coal on SNG
Capital Costs 197
9.1 Comparison of Solvent Refined Coal and
Substitute Natural Gas Costs 208
9.2 Solvent Refined Coal Process Flow Diagram 209
10.1 Steam Balance For 1000 Megawatt Power Plant 222
10.2 Heat Balance Around Lurgi Low BTU Gasifier 223
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LIST OF FIGURES (CONT'D)
FIGURE NO. TITLE PAGE NO.
10.3
10.4
10.5
10.6
10.7
Advanced Power Cycle -
Section (Cycle 1)
Advanced Power Cycle -
Section (Cycle 2)
Advanced Power Cycle -
Section (Cycle 3)
Advanced Power Cycle -
Section (Cycle 4)
Effect of Gas Turbine
Power Generation
Power Generation
Power Generation
Power Generation
Inlet Temperature
224
225
226
227
on Overall Power Generation of Advanced
Power Cycles 228
10.8 Effect of Gas Turbine Inlet Temperature
on Overall Cycle Efficiency of Advanced
Power Cycles 229
10.9 Compressed Air Flow to Gas Burner vs.
Combustion Temperature 230
10.10 Turbine Exhaust Gas Flow vs. Gas Turbine
Inlet Temperature 231
10.11 Turbine Power Generation and Air Compressor
Power Requirements For Cycle 3 232
10.12 Typical Steam Generation Curve For Cycle 1 233
10.13 Typical Steam Generation Curve For Cycle 2 234
10.14 Typical Steam Generation Curve For Cycle 3
or 4 235
10.15 Approximate Power Generation For Gas
Turbine A 236
10.16 Approximate Exit Temperature For Gas
Turbine A 237
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LIST OF FIGURES (CQNT'D)
FIGURE NO. TITLE PAGE NO,
10.17
10.18
11.1
11.2
11.3
Approximate Enthalpies of Fuel Gas,
Stoichiometric Flue Gas, and Air
Cost of Conventional Coal-Fired Steam
Power Plant
Flow Diagram For Fluidized Pressurized
Coal Combustor Power Generator
Steam Balance For Pressurized Combustor
Gas Turbine Exhaust Cooling Curve
238
239
250
251
252
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1. INTRODUCTION
The work reported herein is a technical and economic evaluation of
the R&D investment alternatives for sulfur oxides pollution con-
trol methods and was performed for the Office of Research &
Development, Environmental Protection Agency under Tasks 22-25, 27,
Contract No. CPA 70-68 and Task 7, Contract No. 68-02-1308.
The primary objective of this work was to provide EPA with cost
information for the control of sulfur oxides, which could be used
to help determine regulations that can be effectively applied to
the existing sulfur dioxide emissions from stationary sources.
This work also attempts to provide EPA with information useful as
a guide for allocating its annual development budget to produce
the optimum short term and long term reduction in emissions of
sulfur oxides.
The work included in this report represents Part I of a two-part
study. Part I was divided into three phases:
Phase 1
To tabulate and assess information on existing sources of sulfur
oxides emissions. Details of all coal-, oil- and gas-fired steam
generating power plants, nonferrous smelters, coal- and oil-fired
industrial boilers, acid plants and Claus plants were to be character-
ized and tabulated according to plant capacity, type of fuel (feed)
used, age of boilers (utilities), stream factor, and geographical
distribution. The results were to be stored in an accessible
computer format so that reference could be made whenever necessary.
Phase 2
To study and evaluate several possible methods of sulfur dioxide
emission control. A significant part of the total effort was to
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be allocated to the selection of processes that would be of signi-
ficance in existing or potential technology.
All processes, some containing many alternate designs, were to be
represented by a process and a cost model. These models relate
the important process variables to the capital and operating costs
of the plant. The models were to be written in such a way as to
facilitate future revisions in the models as dictated by improve-
ments in the processes. The processes were to be classified accord-
ing to the following categories:
1. Stack Gas Scrubbing
Processes representing once through ("throwaway") and
regenerative types were to be evaluated. Wet limestone
(throwaway) and Wellman/Allied (regenerative) were selected
as the candidate processes.
2. Production of Clean and Low Sulfur Fuel
Two different types of processes were selected for evaluation:
• High Btu gas from coal, using a Lurgi gasification unit.
• Highly refined coal by solvent extraction.
3. New Power Plant Designs
Two different concepts were to be evaluated:
• A combined-cycle power plant using low Btu gas from
Lurgi gasifiers.
• A new type of power plant design using a pressurized,
fluidized-bed combustor.
After establishing these models, the cost of installing stack gas
scrubbing for the existing utilities was to be investigated using
the utilities emissions inventory generated from Phase 1 on a
plant basis. In addition, the costs of manufacturing substitute
natural gas as well as solvent refined coal were to be investigated
for different parts of the country. The potential of the new power
plant designs mentioned above were to be assessed and improvements,
if needed, illustrated.
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Phase 3
Assessment of development work in control technology and prediction
of future demands of energy and chemicals. EPA requested that
this be done by a modified Delphi Technique, which involved sending1
questionnaires to a panel of experts and reforming and expanding
the questions on the basis of their answers. The results of this
phase will be reported separately by EPA.
At the completion of Part I, the Environmental Protection Agency
(EPA) requested that the following work be included in a second
part of this study:
• Upgrade the utility boiler data base with the Federal Power
Commission (FPC) Form 67 magnetic tape provided by EPA,
adding to the data base:
1. Boiler load factor
2. Boiler fuel consumption by fuel type
3. Boiler fuel sulfur
• Determine the costs for installing Wet-Limestone and Wellman/
Allied stack gas scrubbing units on existing industrial boilers
on a plant basis for all plants greater than 5 megawatts
(equivalent size) and summarize the results.
• Modify the Wellman/Allied scrubbing process so that it will
be applicable to acid plants, and to determine the costs for
installing the regenerable scrubbing unit on existing acid
plants.
• Upgrade the Claus plant data base by including the number of
reaction stages per plant and investigate the feasibility
of applying the regenerable scrubbing cost model to these
emission sources.
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• Determine the mine-mouth costs of substitute natural gas
(SNG) , and solvent refined coal (SRC) for different parts
of the country where coal exists.
• Incorporate a cost model for the production of low and
intermediate Btu gas production into the SNG cost model.
• Estimate the costs and develop cost models for shop
fabrication and packaging of scrubber units for the throw-
away and regenerable scrubbing processes for non-utility
boilers.
The above mentioned work will be reported as Part II of the study
and will be issued at a later date.
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2. CONCLUSIONS AND RECOMMENDATIONS
2.1 Control of Sulfur Dioxide Emissions From Existing Sources
In every one of the 5 major SO2 source groups studied, 75%
of the emissions come from a relatively small number of the largest
plants. A significant national reduction in S0_ emissions could
be achieved by directing control efforts towards these larger
plants. In the case of the utility industry and industrial
boilers most of these plants are concentrated in the 5 or 6
coal producing states south of the Great Lakes.
The non-ferrous smelting industry is an easy target for
significant SO- reductions. This industry is the second largest
group emitter nationally. There are only about 40 plants all
told and the largest 20 emit 75% of the smelter S02. The cost
of stack gas scrubbing controls for each of these is probably less
than $30 million.
The small industrial boilers do not appear to be an optimal target for
significant S0_ reduction. About 72% of the industrial boiler
population are smaller boilers of 100 MMBtu/HR or below emitting
24% of industrial boiler S02 emissions, and the costs for retrofitting
stack gas scrubbing units for such small size boilers would be very
expensive.
2.2 Stack Gas Scrubbing Costs
2.2.1 The Utility Industry
The largest 200 utility plants, those greater than 400
megawatts, are responsible for 75% of the utilities S02
emissions. These can be controlled by stack gas scrubbing
to an overall plant emission of less than 1.2 Ib SO2/MMBtu
fuel fired for total capital investments ranging from
$40/plant kilowatt to $75/kilowatt. The increases in
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electricity costs by these controls range from 1 to 3
mils/kwh.
2.2.2 The Industrial Boilers
Typical costs were determined for a single boiler, assuming
a load factor of 50%. The equivalent inc rement a 1 fuel costs
for clean fuel which could be absorbed as an alternative to
stack gas scrubbing range from $0.90 to $3.30/MMBtu with
decreasing boiler size. These costs could be lowered in
multi-boiler plants by ducting several boilers to a common
scrubbing unit. Plants with higher load factors would
also have lower costs.
Assuming an average cost for high sulfur fuel of $0.40/MMBtu,
SO- regulations imposed on boilers of less than 100 MMBtu/hr
would create demands for< clean fuel at costs ranging from
$1.30 to $3.70/MMBtu. However the 3600 boilers in this size
range (almost three-fourths of all coal and oil-fired in-
dustrial boilers) account for only 24% of the U.S. industrial
boiler emissions and sensible regulations would not force
controls on these small boilers, unless there were particularly
good local reasons.
2.3 Substitute Natural Gas Production
SNG can be produced from coal at costs ranging from $1.20 to
$1.40/MMBtu provided coal costs are around $3/ton and the location
of the plants is not in a high construction cost area.
The HHV of the product SNG would be about 58% of the HHV of the
input coal to the plant. Although this compares favorably to the
efficiency of a power plant, it could be misleading. Other factors
to be considered are the flexibility of the product, the relative
transportation losses, the relative cost per unit of energy and
the efficiency of the final product consumption for the various
alternatives. The relative final efficiency of utilization is
-------
therefore not calculable without including these additional factors.
Generation of SNG for supplying clean industrial boiler fuel from
low sulfur coal is probably not the best utilization of the low
sulfur coal. It could probably be transported and burned directly
in existing coal-fired industrial boilers at a lower overall cost.
A study of availble coals, sulfur contents, mine-mouth costs and
transportation methods and costs to various areas in U.S. appears
useful. Such a study should provide insights into the optimal
control methods: i.e. stack gas scrubbing versus the production
of clean fuels.
One point which ought to be stressed is the need to ensure that
the costs of mining the low sulfur surface coal include the costs
of returning the landscape to a respectable level and also include
adequate compensation to the inconvenienced residents of the area.
2.4 Solvent Refined Coal Production
Solvent refined coal could undoubtably be produced for much less
than SNG and the SRC plant recovers about 79% of the heat content
of the coal in the products. It therefore has the advantages of
cost and efficiency over SNG. It is, however, a solid and less
flexible fuel, and normally contains about 1% sulfur when produced
from a 4% sulfur coal (DAFB). If raw coal costs were around $3
or $4/ton, it could be produced for $0.7 to $0.9/MMBtu. It appears
that as low as 0.4% sulfur and liquid fuel can be produced by
slight process modification and increase in costs.
The real area for investigation appears to be the market. SRC is
basically an expensive, low sulfur, ash free solid fuel not suitable
for direct use with gas turbines. Since the process can also be
geared towards specialized refinery type products, the question
arises as to whether this would be a more worthwhile direction
than the production of a solid fuel.
-------
2.5 The Lurgi Gasifier with Combined Cycle
It does appear that the Lurgi gasifier with a combined power cycle
compares favorably in both efficiency and operating costs to the
conventional steam cycle power station fitted with Wellman/Allied
stack gas scrubbing. It should be emphasized that the Lurgi com-
bined cycle plant is a base loaded power station and relatively
more complicated to operate than the conventional plant. There
is a need for more detailed technical and cost studies of the
present design. There is a very great incentive for development
of better gas turbine designs with higher inlet temperatures
than presently allowable.
2.6 The Fluidized Pressurized Combustor with A Combined Power
Cycle
This is basically a conventional steam cycle with a small gas
turbine added. Its efficiency is about the same as a conventional
steam plant fitted with Wellman/Allied stack gas scrubbing. The
costs presented here show it to be less expensive, but there are
several areas which have not been proven even at pilot plant
level. The final costs could be several million more than the
figures presented here. In particular the dolomite regeneration
is not proven and there is evidence that regeneration efficiency
falls off with the number of regenerations. The plant is less
flexible and more complicated to operate than the conventional
power station. Its cost .savings are uncertain at this stage and
this is its biggest claim to superiority. It appears to be a
less promising design than the Lurgi combined cycle.
-------
3. MAJOR SOURCES OF SULFUR DIOXIDE EMISSION IN U.S.
3.1 Introduction
One of the main objectives of this study was to tabulate and assess
information on major sources of sulfur dioxide emissions in the U.S. so
that optimal ways of controlling the emissions could be evaluated.
Information on five major sources of sulfur dioxide emissions have
been gathered from steam generating utility plants, coal- and oil-
fired industrial boilers, non-ferrous smelting industry, acid plants
and sulfur (Glaus) plants. The sulfur dioxide emissions studied
in this report are on a national level with emphasis on optimal
reduction.
The total U.S. sulfur dioxide emissions from these major sources
are (in terms of sulfur emitted per year):
Million Tons
Sulfur/Year As % of the Total
Utilities 8.742 64.7
Smelters (Lead, Zinc, Copper) 1.923 14.2
Industrial Boilers 1.761 13.0
Acid Plants 0.654 4.9
Sulfur Plants 0.437 3.2
13.517 100.0
3.2 Data
3.2.1 Sources
The sources of data for these industries are summarized as
followed:
-------
Utilities
The utilities data were generated from a tape supplied by EPA
and containing data from the National Emissions Data System,
Office of Air Quality Planning and Standards (NEDS tape), and
also from a tape prepared by MWK containing data from the FPC
Form 67's and the National Coal Association's 1972 edition
of "Steam-Electric Plant Factors" d).
Smelters
The source of data is the report prepared by Arthur G. McKee
and Company for the National Air Pollution Control Administration
(2 ). Emissions are for the calendar year 1968.
Industrial Boilers
The NEDS tape is the sole source of data for industrial boilers.
Emissions are for the calendar year 1971.
Acid Plants
The sole source of data is Chemico's report for the National
Air Pollution Control Administration (3 ) . The report contains
data for acid plants in 1969.
Sulfur Plants
The NEDS tape as well as the report by Process Research Incor-
porated (4) were used for the sulfur plants emissions. Data
are for 1971.
3.2.2 Data Quality and Comparison with Other Published
Information
Sulfur emissions inventories for this report were either
10
-------
taken from published data or proprietary information from EPA
as itemized previously. The data should not be considered as
absolutely accurate or complete since they are not taken from
a direct census of the emission sources. Additionally, not
all of the data sources for the five industrial categories
use the same base year. Comparison with the official EPA
position on sulfur emissions can only be made when growth rates
for the smelters and acid plants are known. Table 3.1 presents
such a comparison assuming a yearly growth rate of 4% for
these two industries for the purpose of updating the data to
1971 (5), thus providing a consistent time base for all five
industrial categories.
It should be pointed out that there are other limitations for
direct comparison of the two. First, the MWK data came from
actual summations of the NEDS tapes or from other reference
sources while the OAQPS data assumed a uniform emission factor,
which is a simplification of the actual situation. Second,
the categories of industries studied are grouped differently
in the two reports; e.g., the OAQPS report grouped smelters
and acid plants under industrial and chemical processes where-
as this report groups them as separate categories. Third,
there is no direct figure given in OAQPS report for industrial
boilers and the closest comparable figure for industrial boil-
ers in the OAQPS report is that given for other manufacturing
type processes, oil and gas companies as well as steel and
rolling mills.
Despite such limitations, sulfur emission statistics in this
report compare favorably with the official OAQPS figures
(Table 3.1) with the exception of acid plants. The OAQPS report
lists the sulfuric acid production of 29 million tons in 1971
whereas the Chemico report (3) lists annual capacity of 38
million tons of 100% acid equivalent in 1968. The annual ca-
pacity is based upon 330 operating days per year equivalent,
11
-------
which is not an unrealistic figure for acid plants.
Table 3.1 also presents the sulfur oxides emissions from extra-
polation of the InterTechnology Corporation (IT.C) report for
EPA (7) . Note the extrapolated figures are higher than the
corresponding MWK figures for both the utilities and indus-
trial boilers. The MWK figures are probably a little
low considering they are obtained from data files which were
incomplete.
There are situations in the NEDS tape where information such
as yearly net generation, percent of sulfur in coal, yearly
fuel consumption, etc. are missing for a particular plant.
Whenever possible, the MWK data for the utility industry have
been updated by reference to the National Coal Association's
1972 edition of "Steam-Electric Plant Factors" and the Federal
Power Commission (FPC) Form 67. However, for the case of in-
dustrial boilers, there is no comprehensive publication on
sulfur emissions. A task is being undertaken by EPA to obtain
a complete emissions profile for industrial and commercial
boilers. Owing to the limited data available in the NEDS tape,
the sulfur emissions inventory for the industrial boilers is
approximately 86% complete. The emissions inventory can be
updated when more information is obtained.
3.3 U. S. Sulfur Dioxide Emissions
3.3.1 Utility Plants
The utility industry burning coal, oil and gas is the largest
source of the major U.S. sulfur emissions. This industry
alone emitted 8.742 million tons of sulfur in 1971, represent-
ing 65% of the total U.S. sulfur emissions. Only 7.5% of the
total fuel was burned in exclusively gas fired boilers which do
no need stack gas cleaning.
Table 3.2 gives a national breakdown of the utilities' fuel
12
-------
consumption. Tables 3.3-3.6 presents statewise breakdown of
the utility industry. Figures 3.1-3.8 show graphs, histograms
and conclusions which can be drawn from the analysis of these
statistics.
Some 880 utility plants are included in the statistical survey
and their size distribution is given in Fig. 3.1. The average
number of boilers per plant remains fairly constant, at about
4, for all size plants. About 95% of the utility boilers in
the U.S. are less than 400 megawatts. These account for 71%
of total utility boiler capacity (Fig. 3.2). A little more
than 62% of the utility plant population has a plant load
factor of 0.4 or above (Fig. 3.3). Only 2% of the plant po-
pulation has a load factor of 0.8 or above whereas 12% of the
plant population has a load factor of 0.2 or below. The aver-
age load factor for all plants considered is approximately
0.5 (Fig. 3.4).
Six states, representing 29% of the total U.S. utility capacity,
contribute 54.5% of the utility sulfur emissions. These states
are Ohio, Pennsylvania, Illinois, Indiana, Kentucky, and Mic-
higan (Fig. 3.5). It is not surprising that these states are
centered around the major coal fields of U.S. since 92% of the
utilities sulfur emissions come from coal fired boilers.
The 22 states that individually contribute at least 1% of the
total utility sulfur emissions are all in the eastern half of
the country. These states, shown in Figure 3.5, represent
95.4% of the utility industry sulfur emissions. Thus, viewed
on a national level, the utility industry in central and wes-
tern states does not present a problem under existing conditions,
Figure 3.6 shows that significant reductions can be made in
the national utility sulfur dioxide emissions by cleaning up
relatively few of the largest plants. For example:
13
-------
1. The coal and oil fired plants over 750 megawatts
(aooroximately 100 plants) account for 50% of the sul-
fur dioxide emissions in U.S. utility industry.
2. The next 100 plants (400-750 megawatts) emit an ad-
ditional 20% of the sulfur dioxide emissions.
3. The next 15% comes from approximately 180 plants in
the range of 150-400 megawatts.
It is fairly clear then, viewed simply on a national level,
there should be a cut-off point beyond which further sulfur
dioxide reductions become increasingly more difficult to
achieve. Figure 3.7 shows that the older the plant, the smaller
it is and, not surprisingly, the smaller the individual boil-
ers. It is to be expected that the cost per kilowatt of in-
stalling stack gas scrubbing units on a new boiler increases
as the size of the boiler decreases. However, if the boiler
is in existence already, it also becomes increasingly more
difficult to retrofit a stack gas unit with increase in age
and decrease in size.
The average boiler size in Figure 3.7 can be misleading , es-
pecially if the larger size utility plants have one or more
very small boilers. These smaller boilers are usually operated
infrequently. Therefore, the cost of retrofitting them would
further increase the cost of electricity delivered.
Clearly then, a great deal of thought must be given to the
regulations introduced for controlling SC>2 emissions from
existing utility plants. Otherwise some very expensive and
unnecessary modifications may be imposed on the utility industry.
3.3.2 Smelters (Copper, Zinc, Lead)
14
-------
The sulfur dioxide emissions from the nonferrous smelting
industry were analyzed based on a report for EPA by Arthur
G. McKee & Company. The nonferrous smelting industry alone
emits 1.923 million tons of sulfur per year or 14.2% of the
total U.S. sulfur dioxide emissions and is the second largest
group source of sulfur dioxide emissions. The sulfur dioxide
emissions from the respective smelters are (in terms of sulfur
emitted) :
Sulfur Emissions % of Total Smelter
Smelters Mton/year Emissions
Copper 1,471
Zinc 321
Lead 131
1,923
Three states contribute 65% of the total U.S. nonferrous smelting
industry sulfur dioxide emissions. These states are Arizona,
Texas and Montana. Arizona has eight copper smelters which
emit 34% of U.S. smelter sulfur dioxide emissions (Fig. 3.11).
Almost all of the sulfur dioxide emissions from the industry
come from the western and southwestern parts of the country.
Figure 3.9 shows the number of plants and the plant size dis-
tribution. Figure 3.10 shows the relationship between the
number of plants, the range of plant capacity and the cumulative
percentage of the sulfur dioxide emission for a certain range
of plant capacity, beginning with the largest. It can be
seen that 50% of the emissions come from plants over 125 tons/
year capacity (approximately 10 plants). The next 10 plants
(75-125 tons capacity) account for a further 25% of the emissions.
3.3.3 Industrial Boilers
The emissions from U.S. industrial boilers constitute the third
major source of sulfur dioxide emissions based on the data
15
-------
available. The industrial boilers emitted 1.761 million tons
of sulfur or approximately 13% of the total U.S. sulfur di-
oxide emissions in 1971. Industrial boilers are defined as
boilers in manufacturing plants which create or change raw or
unfinished materials into another form or product, including
the generation of electricity (with the exception of the boilers
in utility industry). The gas-fired industrial boilers which
burn clean fuel are excluded from this report.
Figure 3.12 shows the number of boilers in the various ranges
of boiler capacity. Tables 3.8 and 3.9 present data for coal
and oil fired industrial boilers, broken down by state and size.
Based on the available data, six states, representing 50% of the
total industrial boiler capacity, contribute 68% of the U.S.
industrial boiler sulfur dioxide emissions. These states are
Pennsylvania, Ohio, Indiana, Michigan, Illinois and Minnesota.
The first five are the same states which contribute the most
sulfur dioxide emissions in the utility industry. The twenty-
one states which individually contribute at least 1% of the
industrial boilers sulfur dioxide emissions are in the eastern
half of the country (Figure 3.14).
Generally speaking, the percent load factor for the industrial
boilers follows the same trend as for the utility boilers.
The average percent load factor fluctuates between 0.4 and 0.6.
The number of boilers that operate at a load factor below 0.2
or above 0.8 is insignificantly small.
Figure 3.15 shows that significant reductions can be made in
the U.S. industrial boilers sulfur emissions by cleaning up
relatively few boilers. For example:
1. The coal- and oil-fired boilers over 300 MMBtu/hr
capacity (approximately 500 boilers) account for 50%
16
-------
of the U.S. industrial boiler sulfur dioxide emissions.
2. The next 900 boilers (100-300 MMBtu/hr capacity)
account for another 25%.
3. 1600 boilers (50-100 MMBtu/hr capacity) comprise the
next 15%.
(Ideally, the data should be examined on a plant basis rather
than on the individual boiler basis. However, information on
number of boilers per plant, plant size, etc. was not available
from the NEDS tape.)
About 72% of the U.S. industrial boiler population are small
boilers of 100 MMBtu/hr or below. The load factor for these
small size boilers is about 44% (Table 3.9). The cost for
installing stack gas scrubbing units for such small size boil-
ers would be very expensive. Clearly other alternatives such
as burning clean fuel should be considered if regulations are
to be imposed on industrial boilers.
It should be stressed that the emissions inventory for coal-
and oil-fired industrial boilers in this report is incomplete
owing to deficiencies in the NEDS tapes, as previously discussed.
The statistics presented in Tables 3.8 and 3.9 represent all
the data that were available. Approximately 93% of the boilers
had information on sulfur, 91% had information on capacity,
while 86% had data on the amount of fuel burned. Furthermore,
the NEDS tapes have partial or no boiler emissions data for
Iowa, Mississippi, New York, North Carolina, Texas, and West
Virginia, and these have not been included in the statistical
analysis.
3.3.4 Acid Plants
The sulfur dioxide emissions from acid plants constitute the
17
-------
fourth major source of sulfur emissions in. the U.S. The industry
emits a total of 0.654 million tons of sulfur per year or 4.9% of
the total U.S. sulfur emissions. The emissions depend on the
type of plant, the raw feed material and the type of product.
The emissions can be in two forms, acid mist and sulfur dioxide,
but both are expressed in tons of sulfur emitted.
Tables 3.10 and 3.11 show the U.S. acid plant statistics by
plant size, plant type and by state. Three southern states
contribute 41% of the total U.S. acid plant sulfur dioxide
emissions. These states are Florida, Texas and Louisiana.
These are the only states which contribute, individually, 10%
or more of the U.S. acid plant sulfur dioxide emissions (Fig.
3.17). California, Illinois, and New Jersey each emit about
6%.
Figure 3.18 shows that significant reductions in the acid plant
sulfur emissions can be made by cleaning up relatively few of
the larger plants. For example, 50% of the emissions come
from plants over 800 tons per day capacity (approximately 50
plants). The next 50 plants (450-600 tons per day capacity)
account for a further 25% of the emissions.
3.3.5 Sulfur Plants
The sulfur dioxide emissions from the sulfur plants constitute
the fifth major source of U.S. sulfur emissions. The industry
emits a total of 0.437 million tons per year or 3.2% of the
total U.S. sulfur emissions. The sulfur plants statistics are
presented in Tables 3.12 and 3.13, on plant size and geographical
basis, repectively.
Four states emit a total of 62% of the U.S. Sulfur Plant emis-
sions. These states are Texas, California, Mississippi and
Wyoming. Texas alone emits 31% (Fig. 3.20). The major zone
18
-------
centers in the south with six states emitting 50% of the total
sulfur plant S0_ emissions.
Figure 3.19 shows the number of plants and the plant size dis-
tribution. Figure 3.21 shows a significant reduction in the
U.S. sulfur plant sulfur emissions would be achieved by cleaning
up relatively few plants. Sulfur plants over the size of 300
tons per day capacity (18 plants), account for approximately
50% of the sulfur plant emissions. The next 26 plants (100-
300 tons per day of capacity) comprise the next 25%. However,
the next 67 plants emit only 15% of the sulfur plant emissions.
3.4 Summary
In every one of the five major sulfur dioxide source groups,
the majority (about 75%) of the emissions come from a relatively
small number of the largest plants. Significant national re-
duction in sulfur dioxide emissions could be achieved by direct-
ing control efforts towards these larger plants. The costs of
controlling sulfur dioxide emissions for these major sources
will be analyzed and assessed in subsequent sections of the
report.
19
-------
TABLE 3.1
COMPARISON OF SULFUR DIOXIDE EMISSIONS
BETWEEN THREE DIFFERENT SOURCES
Utilities
Industrial Boilers
Smelters
Acid Plant
Sulfur (Glaus) Plant
Total
OAQPSJ
SO2 in 10 TONS
ITC2
20.1
4.2
4.0
0.6
U
28.9
20.5
6.8
U
U
U
27.3
MWK-
17.5
4.3-
1.45
1.0
27.7
NOTES:
U = Unavailable
1. "Data File of National Emissions 1971", Office of Air
Quality Planning and Standards (OAQPS), U.S. Environ-
mental Protection Agency.
2. Extrapolation from InterTechnology Corporation report on
"Energy Scenario Consumption and Consideration" as reported
by G.T. Rochelle in "S02 Control Technology For Combustion
Sources", Task 6 Final Report, EPA contract 68-02-1308.
3. Based on M.W. Kellogg summation of NED tapes and other
reference materials specified previously.
4. Adjusted from MWK figure of 3.5 x 10 tons which represents
86% of the SO2 emissions from Industrial Boilers.
5. Prorated to 1971 assuming a yearly growth rate of 4.0%.
The base year for acid plant is 1969. The base year for
smelters is 1968.
20
-------
TABLE 3.2
PLANT SIZE
(MW)
0-100
101-200
201-400
401-600
601-800
801-1000
1001-1200
1201-1400
1401-1600
1601-3000
UTILITIES FUEL CONSUMPTION
AND SULFUR EMISSION FOR
FUEL BURNED
COAL
2.11
3.02
7.67
9.29
6.37
4.58
7.03
4.02
1.91
8.59
54.60
1971
BY UTILITIES
OIL
0.65
1.57
3.07
3.52
2.33
1.69
0.68
0.60
0.95
1.05
16.11
IN 1971
GAS
2.89
3.06
4.62
4.66
3.03
5.15
1.21
2.03
1.09
1.54
29.2'9
Total sulfur emissions for year were 8742 M tons.
21
-------
TABLE 3.3
STATEWISE DISTRIBUTION OF FUEL
BURNED BY UTILITIES IN 1971
STATE
COAL
OIL
GAS
ALABAMA
ALASKA
ARIZONA
ARKANSAS
CALIFORNIA
COLORADO
CONNECTICUT
DELAWARE
D. C.
FLORIDA
GEROGIA
HAWAII
IDAHO
ILLINOIS
INDIANA
IOWA
KANSAS
KENTUCKY
LOUISIANA
MAINE
MARYLAND
MASSACHUSETTS
MICHIGAN
MINNESOTA
MISSISSIPPI
MISSOURI
MONTANA
NEBRASKA
NEVADA
NEW HAMPSHIRE
NEW JERSEY
NEW MEXICO
NEW YORK
N. CAROLINA
NORTH DAKOTA
OHIO
OKLAHOMA
OREGON
PENNSYLVANIA
RHODE ISLAND
S. CAROLINA
SOUTH DAKOTA
TENNESSEE
TEXAS
UTAH
VERMONT
VIRGINIA
WASHINGTON
W. VIRGINIA
WISCONSIN
WYOMING
2.84
0.00
0.06
0.00
0.00
0.59
0.25
0.27
0.05
0.75
1.54
0.00
0.00
4.31
3.72
0.66
0.07
3.28
0.00
0.00
0.98
0.05
3.66
0.84
0.09
1.90
0.08
0.16
0.22
0.18
0.63
0.90
1.47
3.21
0.45
6.31
0.10
0.00
5.72
0.00
0.82
0.04
.40
.00
0.08
0.01
1.06
0.00
2.81
1.64
0.41
2.
0.
0.20
0.00
0.02
0.12
1.53
0.02
0.97
0.07
0.14
2.17
0.07
0.00
0.00
0.31
0.02
0.00
0.01
0.01
0.02
0.23
0.60
1.96
0.34
0.03
0.04
0.01
0.00
0.00
0.01
0.12
1.54
0.01
3.13
0.01
0.00
0.02
0.00
0.00
1.08
0.12
0.06
0.01
0.00
0.01
0.09
0.00
0.97
0.00
0.02
0.03
0.00
0.14
0.00
0.54
0.65
4.43
0.34
0.00
0.03
0.00
1.90
0.45
0.00
0.00
0.72
0.21
0.48
1.21
0.07
2.73
0.00
0.00
0.06
0.39
0.39
0.71
0.47
0.01
0.34
0.28
0.00
.21
.38
0.61
0.08
0.00
0.09
1.84
0.01
0.06
0.02
0.26
0.02
0.14
8.80
0.01
0.00
0.01
0.00
0.00
0.18
0.02
0,
0.
54.60
16.11
29.29
22
-------
STATE
••" • .' '• ALA BAM A"" ''"'"•••••'•
ALASKA
ARIZONA
ARKANSAS
CALIFORNIA
COLORADO
CONNECTICUT
DELAWARE
D. C.
FLORIDA
: - GEORGIA
: • . HAWAII
IDAHO
3 ILLINOIS
4 INTIANA
IOUA
KANSAS
;.:'.. s KENTUCKY
! LOUISIANA
MAINE
MARYLAND
MASSACHUSFTT
6 MICHIGAN
•~ : ; MINNESOTA
MISSISSIPPI
MISSOURI
MONTANA
NEBRASKA
NEVADA
; \ ;-.(•/ .;NEW HAHPSHIR.
NEW MEXICO
NEW YORK
N. CAROLINA
I ':.'. .-,;::; . NORTH DAKOTA
- . lr,nm
OKLAHOMA
CREGON
2PFNNSYLVANIA
RHPOF ISLAND
: .:.:. S. CAROLINA
i >'-...' SOUTH OAKOTA
' '- •:•'• TFNNPSSFF
TEXAS
UTAH
VERMONT
:. VIRGINIA .
; ... •- WASHINGTON .
• '• s. '&•••: U. VIORINIA
WISCONSIN
WYOMING
u.s
. UTILITY INDU!
TABLE 3.4
STRY (1971)
- CAPACITY
DISTRIBUTION
(All Plants)
TOTAL CAPACITY
NO. OF AVGF SIZE AVGF. PCT AVGE BOILER AVGE NO. (MKW) PCT OF
PLANTS (KKW) LOAD FACTOR SIZE »MKW) ELPS/PLANT TOT U.S.
14.
0.
11.
8.
33.
'.. 19.
12.
4.
2.
36.
12.
.'., 0.
0.
39.
30.
33.
3?.
16.
19.
11.
26.
35.
37.,
27.
17.
... 6.
, '• 'e
Ifc.
15.
33.
14.
.:.. 14.
47.
16.
1.
40.
4.
13.
, 7.
' 7.'
60.
8.
2.
12.
, : 0.
13.
25.
6.
649.3
0.0
183.5
283.8
577.9
iov.6 ;,'
258.1
224.3
412.0
314.2
481.1
0.0
0.0
370.4
354.6
81.6
115.3
571.1
422.7
92.2
427.5
199.3
298.4
SG.O
240.0
254.4
74.8
89.5
415.2
,139.6 ....;'•
418.4 •'•>"•
233.3
444.3
532.6
i .63.6 '•••-.
374.3
284.6
36.0
454.1
80.8
, 247.7
:'; 30.7 ,'';
• .1063.3
346.5
63.4
17.0
432.1 ,
..'' 0.0
• .; 634.0;
209.0
152.3
55.31
0.0
38.65
39.16
44.05
45.62
38.18
53.57
25.35
49.08
53.94
0.0
0.0
46.64
46.02
38.36
37.33
47.64
40.01
49.89
47.45
45.36
46.21
33.62
44.05
39.95
32.39
32.12
51.51 :•'•.
• 46.8i ; :V'
52.59
52.70
45.18
60.41
39.69
41.07
48.70
1.59
53.19
41.79
51.68
33.29
49.18
43.48
37.65
19.88
51.93 i
0.0
57.76 '
46.40
50.10
178.6
0.0
82.6
158.6
122.5
69.3
38.5
78.1
28.4
130.3
155.3
0.0
0.0
89.3
98.9
35.0
43.7
129.1
134.1
32.2
141.3
74.3
81.7
44.1
15J.7
113.7
97.3
0.0
207.6
:;,;"- 106.3 .
99.6
110.3
81. 4
144.4
• 50.0
111.0
30.7
5.1
98.5
26.0
78.6
16.4
0.0
119.9
80.6
O.J
141.4
. . 0.0
311.0
60.2
91.5
3.2
0.0
2.4
2.6
5.3
4.4
8.6
3.7
14.5
3.2
3.4
0.0
0.0
5.8
4.7
4.0
6.3
4.0
4.3
4.3
3.3
4.0
5.6
4.3
2.6
3.6
1.0
0.0
2.0
: .3.0
4.5
3.1
4.3
4.3
2.0
5.1
4.7
7.0
5.3
4.0
3.3
2.8
0.0
3.5
1.0
0.0
3.3
0.0
2.1
4.2
2.2
9090.
0.
2019.
2270.
19070.
2082.
3097.
897.
824.
11310.
5773.
0.
0.
14444.
10637.
2694.
3691.
9138.
8032.
461.
4702.
5181.
10445.
33JO.
2880.
6869.
299.
1521.
2491.
698.
6695.
3499.
14662.
7456.
891.
17592.
4554.
36.
18166.
323.
3220.
215.
7443.
27720.
507.
34.
5185.
0.
8242.
5224.
1218.
3.28
0.0
0.73
0.32
6.89
0.75
1.12
0.32
0.30
4.09
2.09
0.0
\' ^ U.S. utility industry, represent
2.90 j-nnn«i +-lr
0.17 capacity.
1.70
1.87
3.77
1.20
1.04
2.48
0.11
0.55
0.90
0.25
2.42
1.26
5.30
2.69
0.32
6.35
1.65
0.01
6.56
0.12
1.16
0.08
2.69
10.01
0.18
0.01
1.87
0.0
2.98
1.89
0.44
880.
276327. 100.00
-------
TABLE 3.5
U.
STATE NO. CF
PLANTS
r '/•: . >, ''.' '
"""< • ALABAMA
ALASKA
ARIZONA
ARK4NS4S
:.-.- "-:•• .?;.'•" CALIFORNIA
DELAWARE
D. C.
Fl PR TD &
':•••'•• -.'• CF03GIA ./
1 :' "., . ,-' HAKMI .-.,.'
in* HO
3 ILLINOIS
4 INGIANA
T Pbu A
"•'••',' •-; KANSiS
: • 5 KENTUCKY
LOUISIANA
MAJNF
MARYLAND
IkV 6 MICHIGAN
;-*.'.: . M»>H?SOTA
MISSOURI
MONTANA
[• • . NEVADA
!l :;: MFW HA'IPSHIR -
!" " WF'J jPI»$FV
NF.W UF.XICO
NFw YORK
',:'-•' NORTH DAKOTA
i .. • ' 10HIO
' CXI AHilMA
OREGON
2 f'FNNSYLVANIA
KHOIIC TSI «ND
S. CAROLINA
i • SOUTH DAKOTA .
TFXAS
UTAH
VFRMDNT
f VIRGINIA
; WASHINGTON
: U. U tRRTNT A
WISCONSIN
WYOMING
14.
0.
E.
P.
32.
17.
2.
12.
0. ,
37.
30.
'•i "-1 .
IB.
16.
c, t
5.
11.
35.
q.
2.
in.
t.
, 5-
• ' - 1 f .
7.
31.
I1!*.
4t!
S.
1.
40.
i,.
13.
7.
7,
20.
8.
12.
0.
25.
8.
S. UTILITY INDUSTRY STATEWISE S02 EMISSIONS (1971)
(All Plants Except Gas.-Fired Only)
SULFUR EMITTED PER YEAR
/VGF SIZF. AVGF PCT AVGf BOILCR AVGF NO. (PCT OF TOTAL U.S.)
(MKW) LPAD FACTOR SIZF (MKW) »LCS/PL*NT COAL PIL ROTH
649.3
0.0
216.3
PR3.8
•,,.592.8 .
121.5
PRP. i
224.3
412.0
\>\ .n
481.1
0.0
o.n
388.8
354.6
Rl .h
163.9
571.1
~\f-. "\-t &
92.2
427.5
ice;. •?
298.4
94.2
31 n.?
264.3
111.5
I2P.4
415.2
139.6 -..;-
407.4
469.2
63.6
377.7
4SI.B
3t.O
454. I
PC;, fl
247.7
30.7
443. 1
63.4
1 7.O
432.1
0.0
A^4. n
209.0
1E2.3
55.31
0.0
41.57
39.16
44.73
46.62
lfl.1 A
53.57
25.35
49. 7f)
53.94 , .
0.0
48.73
46.02
"^ ft . "•! fe
43.18 ;
47.84
4/>.no
49.89
47.45
45.36
48.21
35.46
41.01
59.10
34. ?0
51.51
46.81
57. =9
64.93
47.59
60*41 . -
39.69
41.38
50.77
1.59
53.19
41.7<3
51.68
33.29
A Q 1 ft
42.77
37.65
19.P.R
51.93
0.0
46.40
50.10
178.6
0.0
92.4
158.6
123.7
69.3
38.5
78.1
28.4
HO. 3
155.8
0.0
95.0
98.9
35.0
44.2
129.1
9 S - h
32.2
141.3
74.1
81.7
44.1
153.7
113.7
111.5
n.o
207.6
106.3
153.2
81.4
5oto""'
114.0
7^ 1 1
5.1
98.5
Ph.O
73.6
lu.4
0.0
127.4
80.8
0.0
141.4
0.0
3 1 1 .0
60.2
91.5
3.2
0.0
2.6
2.6
5.4
4.4
n.6
3.7
14.5
3.2
3.4
0.0
0,0
5.6
4.7
A _ I)
6.6
4.0
^ . H
4.3
3.3
4.0
5.6
4.8
3.6
1.0
0.0
2.0
3.0
4.5
3.4
4.3
"~"j!o
5.1
ft . ?
7.0
5.3
4.0
2.8
o.o
4.3
1.0
0.0
3.3
0.0
4.2
2.2
4.55
0.0
0.02
0.0
0.0
0.23
0.39
0.39
0.04
1.62
1.64
0.0
0.0
9.11
8. 34
1 .40
0. 15
7.16
0.0
0.0
1.22
0.04
6.09
1.28
' ' 0. 14
4.79
0.05
0.07
' 0.24
0.73
0.53
1.81
„._?.. JO,-
0.41
13.52
O.U7
0.0
9.60
0.0
0.64
0.03
0.00
0.04
0.02
0.95
0.0
4.?7
3.00
0.26
0.08
0.0
0.01
_Q,0_S
0.26
0.01
0.65
0.02
0.06
1.33
0.07
0.0
0.0
0.12
0.00
O.oo
0.01
0.00
0.01
0.20
0.32
L^l]
0.22
0.02
J1.JCL3
0.00
0.0
Q^.00.
0.00
0. 10
0-3S
0.01
1.51
0.00
0.00
0.00
O.UO
0.00
0.37
O.ll)
0.04
0.01
0.0
0.00
0.03
0.00
0.93
0.0
0.01
0.00
0.00
4.63
0.0
0.02
0.09
0.26
0.23
1.04
0.41
0.09
3.01
1.71
0.0
0.0
9.23
8.34
1.4O
0. 16
7.16
0.01
0.20
1.54
1.16
6.31
1.30
0. 16
4.79
0.05
0.26
0.07
0. 34
1. 11
0.54
3.32
0.41
13.53
o.oa
0.00
9.96
0. 10
0.68
0.04
4.4Q
0.00
0.06
0.02
1.88
0.0
4.28
3.01
0.26
Six states contribute 54.5%
of the total SO 2 emissions
from the U.S. utility in*
dustry:
1 Ohio 13.5%
2 Pennsylvania 10.0%
3 Illinois 9.2%
4 Indiana o.j*
5 Kentucky 7.2%
6 Michigan 6.3%
54.5%
92% of the U.S.. utilities'
.-_ SOj •tniuioni OORM from
burning coal.
744.
91.99 8.17 100.00
-------
TABLE 3.6
U.S. UTILITY STATISTICS BY PLANT SIZE
A) ALL PLANTS
_.NTU A^GL..
-------
TABLE 3.7
RANGE
TON.S/YR NO. OF PLANTS
0-50 7
51-100 15
101-250 9
151-200 2
201-250 1
251-300 1
301-350 1_
36
U.S. SMELTERS SO,, EMISSIONS
(Lead, Zinc, Copper)
TOTAL CAPACITY
TONS/YR
250
1168
1041
355
215
252
332
AVG. CAP.
TONS/YR
35.7
77.9
115.7
177.5
215
252
332
% OF SMELTER
CAPACITY
6.91
32.32
28.88
9.82
5.95
6.97
9.15
SULFUR
MTON/YR
141.2
542.2
835.1
182.2
7.2
80.8
134.3
1923.0
% OF TOTAL
SULFUR
7.34
28.20
43.42
9.47
0.37
4.20
6.98
100.0%
-------
TABLE 3.8
INDUSTRIAL BOILERS ,- COAL. AND OIL FIRED
RANGE
; o- so
i '..-. 51-^ 100
101- 150
151- 200
201- 250
' • • 251- 300
"..:....- 301- 350
*','.-.« 35i_ 400
401- 450
451- 500
501- 550
: - V 551- 600
i'"--\. •.?!;"••.'. . 601- 650
-••••• •••• '•'•'•• 651- 700
701- 750
751- 800
801- 850
:• B51- 900
! , , :,:,':' L :! .901- 950
• - ••.-•••••• 951-1000
1001-1050
1051-1100
1101-1150
i.. 1151-1200
£j •*.••;•• -' 1201-1250
" , v 1251-1300
1301-1350
1351-1400
1401-1450
?•.;;• •-,,.. v^-:1^ 1451-1500
':':'• -: ":'• ••::,' ""'.. 1501-1550
'. ;:•"• .:.••'.•• ••• 1551-1600
1601-1650
1651-1700
1701-1750
i ;•••/ ••r^'v^. 17 51-1800
;• ••-..-• . /t*>-v .'iiBoi-1850
f. }-.:''.',• "r,'j,V'£ 1851-1900
1901-1950
1951-2000
2001-2050
i, ;:.:.:;. U^ii-: 2051-2100
I '•:•.":. .'$$£X 2101-2150
2201-2250
2251-2300
2301-2350
i -.'.:V.-V 2351-2400
• •••:• •: •.:, 2401-2450
! ?/K.r^.:.24, 5 1-6000
« OF
BOILFRS
2612.
962.
488.
296.
145.
'124.
•'.>:V,-.53i
•'•*"'>"<••> 62.
3e!
20.
.•••''•••:. 26.
'"•..;>. 15.
'-'•'•• •-'-11.
7.
14.
6.
• --L:.; 8.
.',., ;•,•;.;• . 6.
!•:•••-.• • i.
3.
0.
1.
2.
' . •' .' 1.
2.
2.
0.
1.
' v"*-'; 2.
i.
10.
2.
SKI:
1.
0.
0.
:. i , ;,:, , 1 .
Hv'ii-/ 0..
"•i't:'\'&:. 0.;
0.
0.
0.
?•'••-,.::£:.•• a'.:
4960.
AVERAGE *"
CAPACITY
22.8
•••.' 74.9
126.4
175.4
227.4
;,, •;:..;.. ',. ; 276.1
- 381.8
438.2
476.6
520.3
. . 583.4
1 A , 631.8
680.6
742.4
779.6
824.7
:.'. •-•••T-V.:- 881.5
' :••..••'•> I;':.,';.-. 929.0
^ 988.0
1026.7
0.0
1114.0
. 1177.5
': 1240.0
1286.5
1327.5
0.0
1420.0
'^sf-^'lSAsIo"
-1586.0
1640.0
1662.0
1702.5
,-i^fe«v ,1790.0,
1931.0
0.0
0.0
• ^>fei>;;i.::b2ioo.o
•:,'.^';-:::r,-'-V 0.0
^.!;f.T-r.^T.--.;;r. • o.o
0.0
0.0
0.0
,.,--:**:;vV--:. -:0.o
'• -: ?<. V" ',-' . o.o
:;;'-'.v.S^},p4397. 9
119.4
BOIlJiR STATISTICS BY CAPACITY
TOTAL
CAPACITY
50023.
72054.
61661.
51912.
32977.
34232.
17406.
23670.
7887.
18111.
10406.
15168.
9477.
7487.
5197.
10914.
4948.
7052.
.. :' '".'."- 5574.
988.
3080.
0.
1114.
2355.
1240.
2573.
2655.
0.
1420.
•.••••-/•••.: 10292.
A.,;^; 1545.
3172.
1640.
16620.
3409.
'.'^/Xl •: o'.
.•"•.?.•>?•».-. 3740.
1931.
0.
0.
...-;-.. :, 2100.
'.•>•••.••::.•' ' o.
•'.-:. "..''•"•'. 0.
0.
0.
0.
... '-•>••.- '. o.
'.. v .';.„.'•. o.
: r 35183.
542999.
* OF U.S .
CAPACITY
9.212
13.270
11.356
9.560
6.073
6.304
3.206
4.359
1.452
3.335
1.916
2.793
1.745
1.379
0.957
2.010
0.911
1 .299
1.027
0.132
0.567
0.0
0.205
0.434
0.228
0.474
0.489
0.0
0.262
j :: .•.;•• . 1.895
"'."; •'•' '":'"'"•;' 0.285
. U.584
0.302
3.061
0.627
-' : 0.330
r.^f.:'::^-', . o.o
0.356
0.0
0.0
0.387
0.0
0.0
0.0
0.0
C.O
.: .. . 0.0
0.0
6.479
100.000
% OF TOTAL
FUEL BURNED
17.221
11.640
11.410
9.194
6.152
5.400
2.586
. 2.742
0.517
4.271
1.615
2.657
1.369
2.926
0.931
1.635
0.093
1.557
0.235
0.013
0.366
0.0
0.221
0.568
0.479
0.255
0.054
0.0
0.123
1.911
0.010
0.042
0.008
0.589
0.235
0.603
.. .. .0.0
0.490
0.002
0.0
0.0
0.046
. 0.0
0.0
0.0
0.0
0.0
0.0
0.0
9.634
100.000
2268.16
MM-MH9TU/Y
Z OF TOTAL
SULFUR
14.372
10.511
10.501
7.173
4.385
5.296
2.403
2.772
0.395
4.699
1.803
3.431
1.197
3.034
1.211
1.572
0.077
0.762
0.172
0.012
0.069
0.0
0.465
0.895
0.771
0.355
0.035
0.0
0.045
3.330
0.003
0.022
0.003
0.398
0.140
0.0
0.400
0.001
0.0
0.0
0.098
0.0
0.0
0.0
0.0
0.0
0.0
0.0
16.998
100.000
1761325.98
TONS/YEAR
PERCENT
LOAD FACTOR
44.977
42.497
40.866
40.246
41.800
47.622
41.956
36.3d4
19.238
54.561
46.219
53.063
54.798
56.654
20.151
49.246
19.221
53.224
19.494
34.496
70.000
0.0
0.0
0.0
95.741
5.763
Io887
0.0
70.000
82.645
1.691
2.7<>6
O.O
1.986
70.000
69.911
0.0
33.960 •
0.0
0.0
0.0
56.513
0.0
0.0
0.0
0.0
0.0
0.0
0.0
60.845
43.466
-------
TABLE 3.9
INDUSTRIAL BOILERS, COAL AND .OIL FIRED
- BOILER STATISTICS BY STATE
STATE
ALABAMA
AI ASKA
ARIZONA
ARKANSAS
CALIFORNIA
COLORADO
CONNECTICUT
DELAwiRt
D. C.
FLORIDA
r.Fnpr.iA
HAHAII
IDAHO
ILLINOIS
INDIANA
I DMA
KANSAS
KFNTUCKY . :
; LOUIS IANA
MA INF
MARYLAND
KASSACHUSETT
MICHIGAN
I M1NNFSOTA
'M MISSISSIPPI
00 MT^niifci
MONTANA
NF.BPASKA
NFVADA
NFfc HAMPSHIR -.-
NFU JFPSF.Y
NFU M?xirn
NFM YORK
N. CAROLINA
MHPTH DAKOTA
OHIO
OKL4HOMA
QPflJQN
PENNSYLVANIA
RHODE ISLAND
<;. r.ARni TNA
SOUTH DAKOTA .
' TFNNESSFF
TFXAS
UTAH
VERMONT
VIRGINIA
' WASHINGTON
. W. VIRGINIA
hiisrriNSTN
WYOMING
* OF
BOILERS
55.
0.
3.
4.
10.
17.
109.
Sit.
7.
16.
111.
75.
10.
•»Q5.
318.
0.
?3.
100.
3.
1 34-
214.
393.
342.
114.
0.
4|*
10.
11.
3.
72.
326.
i-
0.
0.
fl.
46S.
.••' 5. .
'7 i
499.
61.
f,7.
. • 4.
126.
0- r
20.
36.
l<5fl.
13*.
117.
i.
0»
78807.
3627.
11243.
173.
16817.
0.
930.
799.
17800.
' 9975.
0.
1A170.
995.
_5 42.9.9.9. .
* OF U.S.
CAPACITY
4.882
D.O
0.757
0.193
0.015
0.0
1.513
1.047
0.865
0.030
2.575
0.375
0.419
t.474
9.057
0.0
0.948
2.250
0.116
2.0O7
3.077
4.115
7.070
3.422
0.0
1 . 186
0.134
0.297
0.169
0.867
4.835
n.no*
0.0
0.0
0.184
12.030
0.113
n.n
14.513
0.668
P.071
0.032
3.097
0.0
0.171
0.147
3.278
1.837
0.0
?.97B
0.183
. ._. JLQ.Q..OOO
* OF TOTAL
FU?L BURNFP
1.470
0.0
0.021
0.020
0.360
0.016
0.949
1.025
0.416
0.040
2.056
0.027
0.189
5.619
8.178
0.0
0.100
1.880
0.016
2.S56
0.539.
2.892
10.924
5.037
0.0
O.593
0.161
0.153
0.258
0.872
4.699
0.002
0.0
0.0
0.141
11.936
0.004
n.s?^
21.764
0.504
2.281
0.025
3.282
0.0
0.0
0.115
4.686
0.791
0.0
2.303
0.178
100.000
* OF TOTAL
SULFUR
1.343
0.0
0.009
0.003
0.065
0.062
0.262
0.597
0.154
0.028
1.425
0.505
0.088
6.030
12.155
0.0
0.064
2.508
0.007
2.449
1.089
2.158
10.331
4.788
0.0
0.880
0.110
0.081
0.070
0.619
1.447
0.002
0.0
0.0
0.068
17.335
0.002
nTn
22.964
0.365
1.685
0.015
1.677
0.0
0.0
0.078
2.964
0.646
0.0
2.774
0.096
._ 1.00.000 _
PERCENT
LOAD FACTOR,
70.000
0.0
0.0
85.446
11.945
0.0
40.785
0.0
64.774
38.715
27.618
^5.633
0.0
40.569
37.0^0
0.0
0.0
37.647
67.977
70.000
46.748
13.734
46.^63
i4.251
0.0
36.249
51.589
52.891
84.993 ^
U.O
44.509
O.U
O.O
0.0
39.509
48.U64
12.921
0. fi
53.220
0.0
0.0
43.225
3.494
0.0
U.O
0.0
45. 106
34.1b9
0.0
36.097
35.934
._ 43.466
2268.16 1761325.98
HM-MMBTU/Y TONS/YFAR
-------
TABLE 3.10
U.S. ACID PLANT STATISTICS
?;:;. .:::;;: STATISTICS BY PLANT
SIZE / ..,..;.
•• . .:• • • BY
PLANT SIZE AND PLANT
TYPE
i -•• •'••• .'.. : SULFUR EMITTED / YEAR
SIZE (100 PCT
ACID EOUIV. / DAY)
0- 100 V .':•.
:', - : 101- 200
•' '•"-••' '-•-• 201- 300
301- 400
401- 500
501- 600
r :..-;.' 601- 700 ,...•-,
; .'•.:'. I. 801- 900 :'•
:.- • - 901-1000
1001-2000
2001-3000
3001-4000
;" '• . ".'.-"•. :.,;'.'';:;.4001-5000 .'...,;•.-•
STATISTICS BY PLANT
: ' j'.' ' Vl •- :.'
' V0 . - ! . .•-•.• . •.• '
' ' PLANT TYPE .
-1-
'',- •• f'.f^j''"'^:^V&- —3— - '""ff'^''-'
• ••• •'. '•'.-• .r.-v/^'- '—4— ". *'• '-.•"••
' -5-
251.
NO.
.•"•'- 49.
40.
31.
26.
27.
12.
. 17. .
16.
10.
17.
4.
1.
•••••:. l-
251.
TYPE
NO.
37.
--'• •. 9??
49.
20.
AVG. PLANT SIZE
(TONS / DAY)
54.8
124.6
217.7
321.7
418.5
508.3
. 613.2
762.5
905.0
1371.8
2000.0
3050.0
;••''.; .; 4800.0
442.3
AVG. PLANT SIZE
(TONS / DAY)
69.6
348.1
! :•:,•• •-. 860.5 '
457.7
598.8
442.3
AVG. PLANT AGE
(YEARS)
26.9
20.8
20.0
19.3
18.6
21.8
17.3
15.7
15.5
10.2
11.3
7.0
5.0
19.8
'" AVG. PLANT AGE
(YEARS)
30.9
22.7
••'•;-'• •.'• • 7.7 '.
:.' . • '..' • 22.2 • • ' * -
6.4
19.8
AVG. LOAD FACTOR
(PCT)
94.50
93.34
95.28
95.18
95.23
96.56
95.22
95.03
94.01
94.05
95.90
95.90
95.90
94.73
AVG. LOAD FACTOR
(PCT)
94.26
94.98
94.71
94.31
95.38
94.73 19.88
(PCT. OF
ACID MIST
0.31
0.80
1.31
1.57
2.04
1.59
1.93
2.50
1.25
4.47
1.28
0.33
0.51
19.88
TOTAL U.
S02
2.42
3.6o
5.01
6.07
8.46
4.59
7.83
9.62
6.0J
15.64
5.78
1.96
3.08
80.12
S.)
TOTAL
2.73
4.46
6.32
7.64
10.50
6.18
9.75
12.12
7.25
20.11
7.07
2.29
3.60
100.00
SULFUR EMITTED / YEAR
(PCT. OF TOTAL U.S.)
ACID MIST
0.27
5.65
4.47
6.03
3.46
80.12
SO 2
2.71
21.92
25.15
22.01
8. *i
100.00
TOTAL
2.98
27.57
29.62
28.04
11.78
:'TVPEl=CHAMBERr PLANT TYPE2=SULFUR BURNING WITH 3 CONVERTERS TVPE3=SULFUR BURNING WITH 4 CONVERTERS
-TVPE4~MET GAS'CQNTACT PLANT MITH 3 CONVERTERS TYPE5»MET GAS CONTACT PLANT WITH 4 CONVERTERS
*** THE TOTAL U.S. SULFUR EMISSION FROM ACID PLANTS IS 0.653805E 06 TONS PER YEAR ***
-------
TABLE 3.11
U.S. ACID PLANTS
STATISTICS BY STATE
STATE
• f • •" " . •: •. • •
I ;. ! • ' '•'•." ' • 11 ARAM A ! • • ''!•*' '"-
ALASKA
ARIZONA
ARKANSAS
•• ' ' 6 CALIFORNIA
COLORADO
1 -; . rnNNFCTICUT -
DELAWARE
1 FLORIDA
CFDRGIA
( HAWAII
'.. ' .' IDAHO :.'.-..
5 JLL'NOIS
INDIANA
IOWA
KANSAS
; ' • KENTUCKY ' ; , '
•- . : 3 LOUISIANA '' :
MA INF
MARYLAND
MASS/VCHUSETT
MICHIGAN
;:w -: . MINNESOTA
' 0 MISSISSIPPI
', MISSOURI
MONTANA
NEBRASKA
NFV/ADA
t NEW HAMPSHIR '''.'•
I-'.'-' 4 NEW JERSEY
/ NFy MFxir.a
NEW YORK
N. CAROLINA
NORTH DAKOTA
: OHIO
;. .;.• -, : OKLAHOMA . '.. ;
" ••:"'•' ciRFr.nN
PENNSYLVANIA
RHODE ISLAND
S- CAROLINA
SOUTH DAKOTA
; ; TENNESSEE :; :
2TFXAS
UTAH
VERMONT
VIRGINIA
! "•":'.. WASHINGTON "•'."-
. W. VIRGINIA
WISCONSIN
WYOMING
C. C.
NO.
"' ' " .aa
0.0
2.06
i.oa
5.82
0.31
0.0
1.39
19.79
1.75
0.09
1.95
6.46
2.67
1.40
0.73
0.66
10.02
b'.o'
2.16
0.27
1.04
0.25
0.62
1.45
0.50
0.0
0.42
0.0
6.57
0.42
0.53
2.78
0.0
1.67
0.58
0.0
3.03
0.04
0.49
0.0
3.61
11.20
2.55
0.0
1.61
0.37
0.29
0.27
0.24
0.0
251.
442.3
19.8
94.73
19.88
80.12 100.00
-------
TABLE 3.12
U.S. SULFUR PLANTS
STATISTICS BY PLANT SIZE
•
:. •" .'.. >• •'
.
U) ;
M. -
.
.
.
• . -:•:•;: :.
. .
DAILY SHORT I
TON CAPACITIES
0-100
101- 200
201- 300
301-400
401- 500
501- 600
601- 700
701- BOO
801- 900
901-1000 ,".
1001-1100
1101-1200
1201-1300
•
1301-1400
'
TOTALS
NUMBER OF
PLANTS
____________ __
120.
19.
7.
, ••-•'' ,-.-• •
- ":?-.'! 8. - .'• '•
.,, ,.;.;,, . • - •••-.-.
5.
2.
.•.-••••.
•-••-•- '.i. ' •'-•' ..•
.... •.. •. -
0.
0.
• '.••-•• --.-' •'-.• - - • '•
--• -.:....:: o. , ' ;
• •.. .-. ;-".,..vy •
0.
1.
1 •
0.
1.
.
164.
AVFRAGE
SIZE
36.54
138.00
235.68
329.42
436.58
504.00
645.12
0.0
0.0
0.0
• -.-.:. •:'.:-
0.0
1120.00
' '
0.0
1400.00
107.61
AVERAGE
AG =
8.40
9.95
5.57
3.38
9.00
6.00
9.00
0.0
0.0
0.0
. • •
0.0
4.00
.
0.0
1.00
8.13
PERCENT OF US.
DAILY CAPACITY
24.846
14.857
9.348
14.933
12.369
5.712
3.656
0.0
0.0
0.0
0.0
6.346
0.0
7.933
100.000
TOTAL CAPACITY
17647.504
PERCENT OF US.
ANNUAL SULFUR
25.075
14.609
9.327
14.833
12.329
5.695
3.646
0.0
0.0
0.0
0.0
6.322
0.0
7.914
100.000
TOTAL SULFUR
437399.000
.
-------
TABLE 3.13
U.S. SULFUR PLANTS - STATISTICS BY STATE
1
1
STATES
ALABAMA
ALASKA
(ARIZONA
(ARKANSAS
1 CALIFORNIA
1 COLORADO
ICCNNfCTICUT
(DELAWARE
ID. C.
(FLORIDA
(GEORGIA
(HAWAII
(IDAHO
(ILLINOIS
1 INDIANA
(IOWA
(KANSAS
(KENTUCKY
(LOUISIANA
(MAINE
(MARYLAND
IMASSACHUSSTT
w MICHIGAN
N) (MINNESOTA
(MISSISSIPPI
(MISSOURI
IMCNTANA
INFBPASKA
(NEVADA
IN EH HAMPSHIF.
(NEW JERSEY
INEW MEXICO
INFW YORK
IN. CAROLINA
1 INGRTH DAKOTA
(OHIO
(OKLAHOMA
1 OREGON
IPFNNSYLVANIA
(RHODE ISLAND
IS. CAROLINA
(SOUTH DAKOTA
(TENNESSEE
(TEXAS
IUTAH
(VERMONT
(VIRGINIA
(WASHINGTON
(W. VIRGINIA
IWISfONSIN
(WYOMING
1 TOTALS —
FOUIVALP4TS-
NUMBER OF .
PLANTS
2.
1.
0.
4.
IB.
1.
0.
2.
0.
4.
0.
0.
0.
4.
3.
0.
2.
0.
6.
0.
0.
0.
3.
2.
4.
1.
3.
0.
0.
0.
7.
7.
1.
0.
2.
•'- ' 3. '"
2.
0.
6.
0.
0.
0.
0.
58.
2.
0.
1.
1.
1.
1.
12.
164.
AVERAGE
SIZE
216.2
1C.1
0.0
51.8
167.8
20.2
0.0
434.0
0.0
185.9
0.0
0.0
0.0
159.3
154.6
0.0
24.6
0.0
106.4
0.0
0.0
0.0
33.2
95.2
371.6
89.6
87.0
0.0
0.0
0.0
103.5
23.5
56.0
0.0
136.1
35.5
12.9
0.0
74.7
0.0
0.0
0.0
0.0
94.0
12.3
0.0
56.0
22.4
30.2
16.8
85.8
107.6
AVERAGE
AGE
1.0
1.0
0.0
13.8
8.5
5.0
0.0
9.0
0.0
1.0
0.0
0.0
0.0
4.8
1.3
0.0
5.0
0.0
6.8
0.0
0.0 .
0.0
11.0
7.5
5.8
2.0
4.7
0.0
0.0
0.0
7.0
9.1
4.0
0.0
9.0
5.3
9.0
0.0
8.3
0.0
0.0
0.0
0.0
8.8
3.5
0.0
16.0
11.0
13.0
1.0
13.2
8.1
REFINER>
PERCENT CF
US CAPACITY
2.450
0.0
0.0
1.015
0.0
0.0
0.0
0.0
0.0
4.214
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
. 0.0
0.0
0.0
0.0
0.0
0.0
8.231
0.0
0.0
0.0
0.0
0.0
0.0
0.743
0.0
0.0
0.127
0.038
0.146
0.0
0.698
0.0
0.0
0.0
0.0
23.622
0.063
0.0
0.0
0.127
0.171
0.0
5.832
47.477
8378.578
FEFD
PFRCENT OF
US SULFUR"
2.446
0.0
0.0
1.008
0.0
0.0
0.0
0.0
0.0
4.195
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
8.212
0.0
0.0
0.0
0.0
0.0
0.0
0.737
0.0
0.0
0.127
0.038
0.149
0.0
0.695
0.0
0.0
0.0
0.0
23.537
0.062
0.0
. 0.0
0.126
0.171
0.0
5.818
m
47.323
206990.500
NATURAL (
PERCENT OF
US CAPACITY
0.0
0.057
0.0
0.159
17.110
0.114
0.0
4.919
0.0
0.0
0.0
0.0
0.0
3.611
2.627
0.0
0.279
0.0
3.617
0.0
0.0
0.0
0.565
1.079
0.190
0.508
1.479
0.0
0.0
0.0
4.106
0.190
0.317
0.0
1.415
0.565
0.0
0.0
1.840
0.0
0.0
0.0
0.0
7.286
0.076
0.0
0.317
0.0
0.0
0.095
0.0
52.523
9^68.973
3AS FEED
PERCENT OF
US SULFUR
0.0
0.057
0.0
0.158
17.055
0.114
0.0
4.904
0.0
0.0
0.0
0.0
0.0
3.601
2.618
0.0
0.274
0.0
3.601
0.0
0.0
0.0
0.560
1.075
0.190
0.503
1.475
0.0
0.0
0.0
4.092
0.189
0.320
0.0
1.416
0.568
0.0
0.0
2.162
0.0
0.0
0.0
0.0
7.258
0.075
0.0
0.320
0.0
0.0
0.091
0.0
—
52.677
230408. 500
BOTH
PERCENT OF
US CAPACITY
2.450
0.057
0.0
1.174
17.110
0.114
0.0
4.919
0.0
4.214
0.0
0.0
0.0
3.611
2.627
0.0
0.^79
0.0
3.617
0.0
0.0
0.0
0.565
1.079
8.422
0.508
1.479
0.0
0.0
0.0
4.106
0.933
0.317
0.0
1.542
0.603
0.146
0.0
2.539
0.0
0.0
0.0
0.0
30.907
0.140
0.0
0.317
0.127
0.171
0.095
5.832
100.000
17647.651
FEEDS
PERCFNT OF
US SULFUR
2.446
0.057
0.0
1.166
17.055
0.114
0.0
4.904
0.0
4.195
0.0
0.0
0.0
3.601
2.618
0.0
0.274
0.0
3.601
0.0
0.0
0.0
0.560
1.075
8.402
0.503
1.475
0.0
0.0
0.0
4.092
0.926
0.320
0.0
1.543
0.606
0.149
0.0
2.658
0.0
0.0
0.0
0.0
iO.795
0.137
0.0
0.320
0.126
0.171
0.091
5.818
i 00. 000
4J7399.00U
-------
FIGURE 3.1
DISTRIBUTION OF U.S. UTILITY PLANTS WITH PLANT SIZE
500 _
400
300 _
200 _
NO. OF
PLANTS
100
90
80
70
60
50 _
40
30 _
20 _
10
TOTAL NO. OF PLANTS = 880
(NOTE: IN ADDITION TO THOSE SHOWN,
THERE ARE 26 PLANTS OF 1400 MEGAWATTS
AND ABOVE. SEE TABLE 3.6)
200 400 600 800 1000 1200 1400
PLANT SIZE, MEGAWATTS
33
-------
FIGURE 3.2
SIZE DISTRIBUTION OF U.S. UTILITY BOILERS
20_
u>
PERCENT OF
TOTAL U.S.
CAPACITY
2 -
200
GAS FIRED ONLY
OIL FIRED ONLY
COAL FIRED ONLY
MORE THAN ONE FUEL
400
600 800
BOILER SIZE, MEGAWATTS
1000
1200
14OO
-------
FIGURE 3.3
DISTRIBUTION OF PLANT LOAD FACTORS
FOR THE U.S. UTILITY INDUSTRY
200
100 _
90 _I
80 —
70 —
60 _
P
2 50 —
3
a.
u.
O 40 _
6
z
30 —
20 _
D 1
0 2
O 3
0 4
o &
0 61
0 7
} 8
.-*
1
1 1
0 90 100
NUMBER DENOTES
ACTUAL FIGURE
PLANT LOAD FACTOR, %
35
-------
FIGURE 3.4
VARIATION OF PLANT LOAD FACTOR
WITH UTILITY PLANT SIZE
70 —,
60 —
50 —
40
3O.
AVERAGE PLANT
LOAD FACTOR, %
20 —
10.
200 400 600 800 1000 1200 1400 1600 &• ABOVE
PLANT SIZE', MEGAWATTS
36
-------
FIGURE 3.5
GEOGRAPHICAL DISTRIBUTION OF SO2
EMISSIONS FROM THE UTILITY INDUSTRY
1.2%
NUMBERS ON MAP DENOTE
PERCENT OF TOTAL U.S. SO2
EMISSIONS FROM UTILITY PLANTS
1 -5%
5- 10%
> 10%
-------
UJ
I
D
D
u
1000':
900
800
700
600
500
400
300
200
100
90
80
70
60
50
40
30
20
10
fiiGURE 3.6
SO2 EMISSIONS FROM U.S. UTILITY PLANTS
% OF TOTAL UTILITY PLANT S02 EMISSIONS
a: i
100
200
400
600
I
o
UJ
N
1/9
800
1000
10 20 30 40 ,_ 60 60 70 80 90 100
o o
-------
FIGURE 3.7
DISTRIBUTION OF BOILERS AND AVERAGE PLANT
AGE WITH UTILITY PLANT SIZE
200 400 600 800 1000 1200 1400 1600 & ABOVE
PLANT SIZE, MEGAWATTS
39
-------
FIGURE 3.8
DISTRIBUTION OF BOILERS AND BOILER
AGE WITH UTILITY BOILER SIZE
1600
V)
DC
O
m
1000 .
900 .
800 .
700 .
600 .
500 .
400 —
300 —
200
100 -
90 .
80 •
70 •
60 •
50 —
40 —
30 —
20 —
10
O
Z
<
S
o
10
Z
X
UJ
_l
in
88
LL 0
O in
8
\
— 40
_30
-20
. 10
. 9
8
7
6
_ 5
_ 4
_ 3
_ 2
HI
o
cc
UJ
_J
5
o
NUMBERS DENOTES
ACTUAL FIGURE
100 200 300 400 500 600 700
BOILER SIZE, MEGAWATTS
800
900 1000
40
-------
FIGURE 3.9
DISTRIBUTION OF U.S. SMELTERS WITH PLANT SIZE
20—,
NO. OF
SMELTERS
10:
9
8
7
6
5 —
4 _
3 _
2 —
FIGURE DENOTES
ACTUAL NUMBER
50 100 150 200 250 300
PLANT SIZE. TONS OF PRODUCT/YEAR
350
41
-------
FIGURE 3.10
SO2 EMISSIONS FROM U.S. SMELTERS
40.
30_
20.
uj O
5
u.
O
u
10-
9.
8.
7-
6.
50
100
UJ
D
O
O
IT
Q. •
Ui
I
UJ
N
CO
I-
150 <
200
10 20 30 40 50 60 70 80
% OF TOTAL SMELTER SO2 EMISSIONS
90
100
42
-------
FIGURE 3.11
GEOGRAPHICAL DISTRIBUTION OF SO2 EMISSIONS FROM SMELTERS
(jj
FIGURES ON MAP DENOTE PERCENT OF
TOTAL U.S. SOX EMISSIONS FROM SMELTERS
5-10%
> 10%
-------
3574
1000 -3
900 -I
800 —
700 -I
600 _I
5OO _
400 _
300 _
200 _
100 —
90 _I
5 80 -
£ 70 ^
u. —
60 _
0 50 _
_i 40 _
8
30 _
DC
LU _
5
20 _
U-
o
o -
10 _
9 —
8 _I
7 ~
6 .1
5 _I
4 _
3
2
•^
i^
FIGURE 3.12
DISTRIBUTION OF U.S. INDUSTRIAL
BOILERS WITH BOILER SIZE
(NOTE: IN ADDITION TO THOSE SHOWN,
THERE ARE APPROXIMATELY 40 BOILERS
OF 1,300 MM BTU/HR AND ABOVE. SEE
TABLE 3.8)
0
100
200
300
400
500
600
700
800
900
1000
1100 1200
1300
BOILER CAPACITY, MM BTU/HR
44
-------
FIGURE 3.13
SIZE DISTRIBUTION OF U.S. INDUSTRIAL BOILERS
30 _
a
Ul
en
UL
J
o
eB
O
O
a.
<
O
CO
a.
O
20 _
10
9
8
7
6
5 _
4 —
2 _
(NOTE: IN ADDITION TO THOSE SHOWN,
THERE ARE 40 BOILERS IN THE RANGE
OF 1000 - 6000 MM BTU/HR WHICH MAKE
UP 17.6% OF TOTAL U.S. CAPACITY.
SEE TABLE 3.8.)
OIL FIRED
COAL FIRED
100 200 300 400 500 600 700
BOILER CAPACITY, MM BTU/HR
800
900
1000
45
-------
ft
FIGURE 3.14
GEOGRAPHICAL DISTRIBUTION OF SO2 EMISSIONS FROM INDUSTRIAL BOILERS
1.0%
FIGURES ON MAP DENOTE
PERCENT OF TOTAL U.S. SOX
EMISSIONS FROM INDUSTRIAL BOILERS
1 -5%
5- 10%
2.1%
-------
FIGURE 3.15
SO2 EMISSIONS FROM U.S. INDUSTRIAL BOILERS
5000
4000
3000
2000
o
oc
UJ
I
I-
I
t-
o
z
cc
cc
UJ
o
m
Q
Z
O
o
1000
900
800
700
600
500
400
300
200
100
90
80
70
60
50
40
30
20
10
100
200
400
IT
I
3
m
5
u
o
oc
o
CD
600
800
1000
O
I
10 20 30 40 50 60 70 80 90
% OF TOTAL INDUSTRIAL BOILER SO2 EMISSIONS
47
100
-------
FIGURE 3.16
DISTRIBUTION OF U.S. SULFURIC ACID PLANTS
WITH PLANT SIZE
(NOTE: IN ADDITION TO THOSE SHOWN,
THERE ARE 23 PLANTS IN THE RANGE
OF 1000 -5000 TONS/DAY OF 100% ACID
EQUIVALENT. SEE TABLE 3.10.)
50
40 _
30 _
20 —
NO. OF
PLANTS
10 _
9 _I
8 _
7 _
6 _
100 200 300 400 500 600 700 800
PLANT SIZE, TONS/DAY OF 100% ACID EQUIVALENT
900
1000
48
-------
i VO
FIGURE 3.17
GEOGRAPHICAL DISTRIBUTION OF SO2 EMISSIONS FROM SULFURIC ACID PLANTS
FIGURES ON MAP DENOTE PERCENT
OF TOTAL U.S. SOX EMISSIONS FROM
ACID PLANTS
1-5%
> 5%
.4%
2.2%
-------
FIGURE 3.18
S02 EMISSIONS FROM U.S. SULFURIC ACID PLANTS
300
O
DC
UJ
I
K
I
H
3
o
1C
fc
LL.
O
u
2
O
200
100
90
80
70
60
50
40
30
-
<
Q
N
800
1000
_L
_L
J_
10
20 30 40 50 60 70 80
% OF TOTAL ACID PLANT EMISSIONS
90
100
50
-------
FIGURE 3.19
DISTRIBUTION OF U.S. SULFUR PLANTS WITH PLANT SIZE
NO. OF
PLANTS
200 _
100 .
90
80
70
60 .
50
40 _
30 _
20 _
10
9
8
7
6
5
3 —
2 _
(NOTE: IN ADDITION TO THOSE SHOWN,
THERE ARE 2 PLANTS WITH SIZE GREATER
THAN 100 TONS/DAY. SEE TABLE 3.12.)
100 200 300 400 600
PLANT SIZE, TON/DAY
800
700
51
-------
1m
FIGURE 3.20
GEOGRAPHICAL DISTRIBUTION OF SO2 EMISSIONS FROM SULFUR PLANTS
. •.(T^T7 •'..'.. -.FLORIDA . -.
4.9%
FIGURES ON MAP DENOTE PERCENT
OF TOTAL U.S. SOX EMISSIONS FROM
SULFUR PLANTS
1 -5%
> 5%
4.1%
-------
FIGURE 3.21
SO2 EMISSIONS FROM U.S. SULFUR PLANTS
200
O
tn
O
0.
-------
4. THE GENERAL MODEL
4.1 The General Process Model
The plants in the models have, as far as possible, been made
self-contained apart from the intake of basic raw feed materials;
i.e., the plant should not be buying natural gas or electricity.
If possible, it should not even be buying desulfurized fuel oil
since supply cannot be assumed. There are obviously exceptions
if the plant is an addition to a larger conventional plant; e.g.,
with stack gas scrubbing for a power plant it would be illogical
not to assume a supply of power. In general, a large plant having
a coal feed will generate its own power, steam and heat requirements
by burning coal and scrubbing the stack gases.
It was not a primary concern to provide special chemical by-products
from any process, but to avoid additional treatment facilities
for impure materials by routing these side streams back to the
plant fuel supply where possible. This approach simplifies the
models and minimizes the effect of credits for special chemical
by-products on the plant costs.
The cost of equipment and raw material, utility and waste product
quantities have all been related to one or more basic process
parameters; e.g., in the stack gas scrubbing models, the basic
process parameters are flue gas flow rate and sulfur content of
the fuel. For a plant producing high quality fuel, the basic
process parameters are product flow rate and properties of the
raw feed materials.
Where possible, equipment costs were related directly to the basic
process parameters. However, the format of some of the estimates
used to develop the models prevented this. In these cases, the
available cost information was carefully examined relative to the
General Cost Model to determine exactly what the costs included.
54
-------
The equipment costs were extracted from these estimates by using
the relationships between construction labor costs, other material
costs and equipment costs given in the General Cost Model.
Each plant design was examined to fix maximum train sizes for
each group of equipment. It has been assumed that N trains cost
N times the cost of one train. Where a plant is largely made up
of several trains, size variations were only taken in increments
of their size.
For the smaller plants, it was possible to examine the cost of
every item of equipment and assign an exponent of size to give
cost variations. However, for the larger plants, whole sections
have been grouped together. The following is given as a general
guide to the exponents for equipment cost vs. size ( 9,14,21) :
cost
n,
- _ /Size,Yl
[ - \Sl^l) \
Increasing number of trains of equipment 1.0
Blowers 0.9
Solids grinding equipment 0.8
Steam generation equipment 0 . 8
Process furnaces and reformers 0.7
Compressors 0.7
Power generation equipment 0.7
Solids handling equipment 0.6
Offsites 0.6
Other process units 0.6
4.2 The General Cost Model
4.2.1 Bases For Costs
All costs in the models are those in existence at the end
of 1973. To update prior cost information used in the con-
struction of the models, an annual inflation multiplication
55
-------
factor of 1.05 has been used. All costs other than unit
costs for labor, raw materials, etc., are shown in thousands
of dollars (M$).
The direct field construction labor cost, L, and the direct
cost of operating labor, CO, both refer-to a Gulf Coast
(Houston) location. For any other location, they are adjusted
through the use of a location factor, F, which is explained
in section 4.3.
Whenever possible in the development of the cost models dis-
cussed in this report, major equipment costs, E, have been
related to plant size variations. The reference values of E
have been taken from actual plant cost estimates when these
were available. Sometimes, however, the cost estimates were
not available in such a detailed breakdown. In such cases,
the relationships developed in the General Cost Model were
used to analyze the cost data. The relationships in the
General Cost Model were developed based on procedures reported
and recommended in the literature ( 9,13) and on Kellogg's
general experience.
4.2.2 Capital Cost Model
Major equipment costs, E, represent the cost of major
equipment delivered to the site, but not located, tied-in
to piping, instruments, etc., or commissioned. It includes
material costs only. Major equipment is defined to include
furnaces, heat exchangers, converters, reactors, towers,
drums and tanks, pumps, compressors, transportation and
conveying equipment, special equipment (filters, centrifuges,
dryers, agitators, grinding equipment, cyclones, etc.), and
major gas ductwork.
Other material costs, M, represent the cost of piping,
electrical, process instrumentation, paint, insulation,
foundations, concrete structures, and structural steel
56
-------
for equipment support. It does not include such items as
site preparation, steel frame structures, process buildings,
cafeterias, control rooms, shops, offices, etc.
M has been taken as a fixed fraction of E. Whenever possible,
this fraction has been determined from an estimate covering
the particular plant under consideration. This fraction is
often different for each section of the plant. if particular
details were not available, the following relationships have
been assumed ( 9) :
Solids handling plant: M = 0.40E
Chemical process plant: M = 0.80E
Direct field construction labor costs, L, are based on Gulf
Coast rates and productivities. Again, L has been taken
as a fixed fraction of E. Wheneve.r possible, it has been
derived from an estimate covering the particular plant under
consideration. This fraction is often different for each
section of the plant. If particular details we're not available,
the following relationships have been assumed (9 ):
Solids handling plant: L = 0.40E
Chemical process plant: L = 0.60E
Indirect costs associated with field labor have been assumed
as follows:
Fringe benefits and payroll burden = 0.12 L
Field administration, supervision
temporary facilities
Construction equipment and tools
Total field labor indirect costs
57
-------
Home office engineering includes home office construction,
engineering and design, procurement, client services,
accounting, cost engineering, travel and living expenses,
reproduction and communication. This could range from under
10% to almost 20% of the major equipment and other material
costs. In the model, this has been assumed to be 15% of the
total direct material cost (E + M).
The bare cost of the plant, BARC, is defined as the sum of
equipment costs, other material costs, construction labor
and labor indirects, and home office engineering. For a
Gulf Coast location, it is given by:
BARC = E + M+L + 0.43L + 0.15 (E + M)
= 1.15 (E + M) + 1.43 L
For any other location, it is given by:
BARC = 1.15 (E + M) + 1.43 L-F
where F is the location factor (see section 4.3).
Taxes and insurance can be 1-4% of the bare cost. In the
model, they have been assumed to be 2%. Contractor's
overheads and profit could depend on several factors, but
are generally in the range of 6-13% of the bare cost. A
value of 10% was chosen for the model.
A contingency has been included in the model and is expressed
as a fraction of the bare cost. It represents the degree
of uncertainty in the process design and the cost estimate.
The contingency, CONTIN, could range from zero for a well-
established process to 0.20 or more for a process still under
development.
58
-------
The total plant investment, TPI, is defined as the sum of
the bare cost (including contingency), taxes and'insurance,
and contractor's overheads and profit. It is therefore
given by:
TPI = (1.0 + CONTIN) BARC + 0.02 (1.0 + CONTIN) BARC
+ 0.10 (1.0 + CONTIN) BARC
=1.12 (1.0 + CONTIN) BARC
In order to obtain the total capital required for construction
of a particular plant, some additional costs should be added
to the total plant investment. These costs are:
1. Start-up costs
2. Working capital
3. Interest during construction
Start-up costs, STC, have been assumed to be 20% of the total
net annual operating cost, AOC (see section 4.2.3 for
explanation of AOC). Thus:
STC = 0.20 AOC
Working capital, WKC, is required for raw materials inventory,
plant materials and supplies, etc. For simplification, it
has also been assumed to be 20% of the total net annual
operating cost, AOC.
Thus:
WKC =0.20 AOC
Interest during construction, IDC, obviously increases with
the length of the construction period which, to some extent,
is a function of the size of the plant. The construction
of plants the size of the stack gas scrubbing units is now
taking about 2-3 years and projects of the magnitude and
59
-------
complexity of a substitute natural gas plant or a power
station are taking 4-5 years. Two different values for the
interest during construction have therefore been assumed.
The first is intended to be used for stack gas scrubbing
units fitted to existing power plants or for constructions
well under $100 million:
IDC =0.12 TPI*
The second is for the larger, more complex plants such as
substitute natural gas, solvent refined coal, and power plants:
IDC = 0.18 TPI*
The total capital required, TCR, is equal to the sum of the
total plant investment, start-up costs, working capital, and
interest during construction.
Thus:
TCR = TPI + STC + WKC + IDC
For stack gas scrubbing units, this can be reduced to:
TCR = TPI +0.20 AOC +0.20 AOC +0.12 TPI
= 1.12 TPI +0.40 AOC
For the larger plants, this can be reduced to:
TCR = TPI +0.20 AOC + 0.20 AOC +0.18 TPI
= 1.18 TPI +0.40 AOC
From section 4.2.3, AOC is calculated from:
AOC = 0.078 TPI + 2.0 TO'CO (1.6 + F) + ANR
*See Appendix A for derivation of equation
60
-------
where TO = total number of shift operators
ANR = Annual cost of raw materials, utilities, and
waste disposal, less by-product credits.
Therefore, for stack gas scrubbing units, the equation for the
total capital required becomes: :
/
TCR = 1.12'i'TPI + 0.40 [0.078 TPI + 2.0 TO'CO (1.0 + F) + ANR]
* 1.12 TPI + 0.03 TPI + 0..8 TO'CO (1.0 + F) + 0.40 ANR
= 1.15 TPI + 0.8 TO. CO (1.0 + F) + 0.40 ANR
For the larger plants, the equation for the total capital
required becomes:
TCR = 1.18 TPI + 0.40 [0.078TPI + 2.0 TO-CO (1.0 + F) + ANR]
= 1.18 TPI + 0.03 TPI + 0.8 TO-CO (1.0 + F) + 0.4 ANR
= 1.21 TPI + 0.8 TO-CO (1.0 + F) + 0.4 ANR
The buildup of costs to determine the total capital required is
illustrated in Figure 4.1.
4.2.3 Operating Cost Model
The total net annual operating cost, AOC, is the total cost of
operating the plant less the credits from the sale of by-products,
It does not include return of capital, payment of interest on
capital, income tax on equity returns or depreciation. The total
net annual operating cost is made up of the following items:
1. Annual cost of raw materials, utilities, and waste
disposal, less by-product credits
2. Annual cost of operating labor and supervision
3. Annual cost of maintenance labor and supervision
4. Annual cost of plant supplies and replacements
5. Annual cost of administration and overheads
6. Annual cost of local taxes and insurance
61
-------
The annual cost of raw materials, utilities, and waste disposal,
less by-product credits, ANR, is clearly a function of the
particular process under consideration. It is given by
different relationships for each model.
The total number of operators employed on all shifts, TO,
is different for each process and is either given as an
equation or number for each particular model . It has been
assumed that each operator works 40 hours per week for 50
weeks per year (2000 hours per year) . If CO is the hourly
rate for an operator (Gulf Coast basis) , then the annual
cost of operating labor is given by:
Operating labor (Gulf Coast) =
= 2 TO. CO M$/yr
The annual cost of operating labor for any other location
has been assumed to be:
Operating labor = 2 TO- CO (0.5 + 0.5 F)
Supervision was assumed to be 15% of operating labor. Thus,
the total cost of operating labor and supervision, AOL, is
given by:
AOL = 1.15 [2 TO-CO (0.5 + 0.5 F) ]
= 2.3 TO-CO (0.5 + 0.5 F)
The annual cost of maintenance labor has been assumed to be
1.5% of the total plant investment. Maintenance supervision
is 15% of maintenance labor. Therefore, the total annual
cost of maintenance labor and supervision, AML, is:
62
-------
AML = 1.15 (0.015 TPI)
= 0.018 TPI (rounded up)
Plant supplies and replacements include charts, cleaning
supplies, miscellaneous chemicals, lubricants, paint, and
replacement parts such as gaskets, seals, valves, insulation,
welding materials, packing, balls (grinding), vessel lining
materials, etc. The annual cost of plant supplies and re-
placements, APS, has been assumed to be 2% of the total plant
investment. Thus:
APS =0.02 TPI
Administration and overheads include salaries and wages
for administrators, secretaries, typists, etc., office
supplies and equipment, medical and safety services, trans-
portation and communications, lighting, janitorial services,
plant protection, payroll overheads, employee benefits, etc.
The annual cost of administration and overheads, AOH, has
been assumed to be 70% of the annual operator, maintenance
labor, and total supervision costs. Thus:
AOH = 0.70 [2.3 TO-CO (0.5 + 0.5F) + 0.018 TPI]
=1.7 TO-CO (0.5 + 0.5F) + 0.013 TPI (rounded up)
Local taxes and insurance include property taxes, fire and
liability insurance, special hazards insurance, business
interruption insurance, etc. The annual local taxes and
insurance, ATI, have been assumed to be 2.7% of the total
plant investment. Thus:
ATI = 0.027 TPI
The total net annual operating cost, AOC, is therefore given
by:
63
-------
AOC = ANR + AOL + AML + APS + AOH + ATI
= ANR + 2.3 TO.CO (0.5 + 0.5F) + 0.018 TPI
+ 0.02 TPI + 1.7 TO.CO (0.5 + 0.5F) + 0.013 TPI
-I- 0.027 TPI
= 0.078 TPI + 4.0 TO-CO (0.5 + 0.5F) + ANR
= 0.078 TPI + 2.0 TO.CO (1.0 + F) + ANR
In order to obtain the total annual production cost, the
following items must be added to the total net annual
operating cost:
1. depreciation
2. average yearly interest on borrowed capital
3. average yearly net return on equity
4. average yearly income tax
The straight-line method was used to determine depreciation,
based on the total capital required less the working capital,
For stack gas scrubbing units (15 year life), the annual
depreciation, ACR, is:
ACR = 1/15 (TCR-WKC)
= 0.067 (TCR-0.20 AOC)
For substitute natural gas and solvent refined coal plants
(20 year life), it is given by:
ACR = 0.050 (TCR - 0.20 AOC)
For power plants, both conventional and combined cycle (28
year life), it is:
ACR = 0.036 (TCR - 0.20 AOC)
64
-------
Interest on debt and return on equity are calculated following
a procedure recommended in the literature (13) and illustrated
in Appendix A. The procedure assumes a fixed debt-to-equity
ratio, an interest rate on debt, and the required net (after
tax) rate of return on equity. Interest on debt and return
on equity are calculated over the plant life, and the yearly
average is expressed as a percentage of the total capital
required (TCR). Assuming a 75%/25% debt-to-equity ratio,
a 9% per year interest rate, and a 15% per year net rate of
return on equity, the annual interest and return, AIC, is
given by:
AIC = 0.054 TCR
Federal income tax is the average yearly income tax over the
plant life, expressed as a percentage of the total capital
required. The calculation of income tax is illustrated in
Appendix A. Based on the assumptions listed in the preceding
paragraph and an assumed tax rate of 48%, the annual federal
income tax, AFT, is given by :
AFT = 0.018 TCR
The total annual production cost, TAG, is given by:
TAG = AOC + ACR + AIC + AFT
For stack gas scrubbing plants, this can be reduced as
follows:
TAG = AOC + 0.067 (TCR - 0.20 AOC) + 0.054 TCR + 0.018 TCR
= AOC + 0.067 TCR - .013 AOC + 0.054 TCR +0.018 TCR
= 0.139 TCR +0.99 AOC
65
-------
Substituting for TCR and AOC from preceeding equations:
TAG = 0.139 [1.15 TPI + 0.8 TO-CO (1.0 + F) + 0.40 ANR]
+ 0.99 [0.078 TPI + 2.0 TO-CO (1.0 + F) + ANR]
= 0.237 TPI + 2.1 TO-CO (1.0 + F) + 1.04 ANR
Making the appropriate substitutions, the total annual
production cost for substitute natural gas and solvent
refined coal plants is:
TAG = 0.225 TPI + 2.1 TO-CO (1.0 + F) + 1.04 ANR
For power plants, this equation becomes:
TAG = 0.208 TPI + 2.1 TO-CO (1.0 + F) + 1.04 ANR
The buildup of costs to determine the total annual production
cost is illustrated in Figure 4.2.
4. 3 Effect of Location on Plant Cost
The cost models have been developed using U.S. Gulf Coast 1973
costs as a basis. In order to predict plant costs for other
locations, factors have been developed which relate construction
labor costs at various locations to Gulf Coast labor costs. By
multiplying the field labor construction portion of plant cost
by this location factor, the total plant cost is adjusted to
the desired location.
Labor rates for different crafts were obtained from the literature
(10) and escalated to the end of 1973. Using an average craft
mix obtained from in-house information (12), an average construction
labor rate was obtained for each location. Productivity factors
for the various locations, also obtained from in-house data, were
used to create the rate for equal work output. These rates were
66
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then normalized, using Houston (Gulf Coast) as a basis, to yield
relative field labor construction costs.
Table 4.1 lists the relative labor costs determined for twenty
cities. They range from 1.0 for Houston to 2.08 for New York.
Costs are generally highest in the Northeastern quarter of the
country and lowest in the South. These factors are shown on a
map of the U.S. in Figure 4.3.
Table 4.2 lists average location factors for each state. Allowance
has been made in the factor for the importation of temporary labor
to the more remote states. The factors are shown on a map of the
U.S. in Figure 4.4.
Figure 4.5 gives the relationship between major equipment
cost, E, total plant investment, TPI, and location factor, F,
when the contingency, CONTIN, is zero.
67
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4.4 Nomenclature
E
M
BARC
Major equipment costs
Other material costs
Direct field labor costs (Gulf Coast)
Bare cost
Location factor
M$
M$
M$
M$
CONTIN
TPI
STC
WKC
IDC
TCR
ANR
AOL
AML
APS
Contingency
Total plant investment
Start-up costs
Working capital
Interest during construction
Total capital required
Annual cost of operating labor and
supervision
Annual cost of maintenance labor and
supervision
Annual cost of plant supplies and re-
placements
M$
M$
M$
M$
M$
Annual cost of raw materials, utilities,
and waste disposal, less by-product
credits M$/year
M$/year
M$/year
M$/year
68
-------
AOH
ATI
AOC
TO
CO
ACR
AIC
AFT
TAG
COHP
TAXI
FLIC
ENGR
Annual cost of administration and
overheads
M$/year
Annual cost of local taxes and insurance M$/year
Total net annual operating cost
Total number of shift operators
Hourly rate for shift operators (Gulf
Coast)
Annual depreciation
Annual interest on debt and return on
capital
Annual federal income taxes
Total annual production cost
Contractor overhead & profits
Taxes and insurance
Field Labor Indirect Cost
Engineering Fees
M$/year
$/hour
M$/year
M$/year
M$/year
M$/year
M$/year
M$/year
M$/year
M$/year
69
-------
TABLE 4 .1
LOCATION FACTORS FOR MAJOR U. S. CITIES
Location
Atlanta
Baltimore
Birmingham
Boston
Chicago
Cincinnati
Cleveland
Dallas
Denver
Detroit
Kansas City
Los Angeles
Minneapolis
New Orleans
New York
Philadelphia
Pittsburgh
St. Louis
San Francisco
Seattle
Houston
Location Factor F
1.10
1.41
1.16
1.23
1.52
1.53
1.86
1.07
1.03
1.73
1.37
1.44
1.54
1.16
2.08
1.82
1.52
2.01
1.45
1.21
1.00
70
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TABLE 4^2
AVERAGE LOCATION FACTORS FOR .EACH STATE
State Location Factor
Alabama
Alaska
Arizona
Arkansas
California
Colorado
Connecticut
Delaware
D.C.
Florida
Georgia
Hawaii
Idaho
Illinois
Indiana
Iowa
Kansas
Kentucky
Louisiana
Maine
Maryland
Massachusetts
Michigan
Minnesota
Mississippi
Missouri
Montana
Nebraska
Nevada
New Hampshire
New Jersey
New Mexico
New York
N. Carolina
North Dakota
Ohio
Oklahoma
Oregon
Pennsylvania
Rhode Island
S. Carolina
South Dakota
Tennessee
Texas
Utah
Vermont
Virginia
Washington
W. Virginia
Wisconsin
Wyoming -,-,
1.2
2.1
1.3.
1.2
1.5
1.2
1.7
1.4
. 1.4
1.2
1.1
2.0
1.3
1.7
1.6
1.5
1.4
1.5
1.1
1.2
1.4
1.3
1.7
1.5
1.1
1.6
1.3
1.4
1.4
1.2
2.1
1.3
2.1
1.2
1.3
1.6
1.4
1.2
1.6
1.3
1.1
1.3
1.2
1.1
1.2
1.2
1.4
1.2
1.5
1.5
1.3
-------
FIGURE 4.1
RELATIONSHIP BETWEEN CAPITAL COST FACTORS IN THE GENERAL COST MODEL
NJ
MAJOR
EQUIPMENT
COSTS (E)
OTHER
MATERIAL COSTS (M)
DIRECT FIELD CONSTRUCTION
LABOR COSTS (L)
DIRECT PLANT
CONSTRUCTION COSTS
FIELD LABOR INDIRECT COSTS
[FLIC = 0.43 L]
ENGINEERING FEES
[ENGR = 0.15 (E + M)]
FRINGE BENEFITS &
PAYROLL BURDEN
FIELD ADMINISTRATION,
SUPERVISION & TEMPORARY
FACILITIES
CONSTRUCTION EQUIPMENT
& TOOLS
INDIRECT COSTS
OF CONSTRUCTION
TAX & INSURANCE
[TAXI = 0.02 BARC]
BARC PLANT COST
[BARC = 1.15
(E + M) + 1.43 L)
2
COST OF SITE
WORKING CAPITA^
[WKC = 0.20 AOC]
CONTRACTOR
OVERHEADS & PROFITS
[COHP = 0.10 BARC]
CONTINGENCY 1
(CONTIN)
TOTAL PLANT
INVESTMENT (TPI)
STARTUP COSTS
[STAR = 0.20 AOC]
INTEREST ON 4
CONSTRUCTION
CAPITAL
TOTAL CAPITAL REQUIREMENT
(TCR)
1. SEE DEFINITION ON PAGE 58.
2. COST WOULD NORMALLY BE INCLUDED ONLY IF PURCHASE IS REQUIRED. COST IS USUALLY SMALL AND HAS NOT BEEN INCLUDED IN MODEL.
3. SEE NOTE 3 OF FIGURE 4.2.
4. SEE FIGURE 4.2.
-------
FIGURE 4.2
RELATIONSHIP BETWEEN PRODUCTION COST FACTORS IN THE GENERAL COST MODEL
RAW'MATERIALS
UTILITIES
CATALYSTS & CHEMICALS
-J
CO
WASTE DISPOSAL
BY-PRODUCT CREDIT
COST OF MATERIALS LESS
BY-PRODUCT CREDITS (ANR)
OPERATING LABOR &
SUPERVISION (AOL)
MAINTENANCE LABOR &
MATERIALS [AML = 0.018 TPI]
PLANT SUPPLIES &
REPLACEMENTS [APS = 0.02 TPI]
ADMINISTRATIVE & PLANT
OVERHEADS
[AOH = 0.70 (AOL + AMD]
DIRECT & INDIRECT COST
DEPRECIATION
[ACR = (TCR-WKO/YEARS]
COST OF MONEY
[AIC = 0.054 TCR]
FEDERAL INCOME TAX
[AFT = 0.018 TCR]
LOCAL TAX & INSURANCE
[ATI = 0.027 TPI]
FIXED COST
TOTAL ANNUAL PRODUCTION COST
[TAC]
1. AVERAGE OVER THE PLANT LIFE, ASSUMING 75% DEBT AT 9% INTEREST RATE PER YEAR, AND 25% EQUITY GIVING A NET RETURN OF 15%.
2. AVERAGE OVER THE PLANT LIFE, ASSUMING 48% FEDERAL INCOME TAX RATE.
3. ANNUAL OPERATING COST IS: AOC = ANR + AOL + AML + APS + AOH + ATI.
-------
FIGURE 4.3
LOCATION FACTORS FOR SELECTED CITIES
1.82
MISSOURI
KANSAS
CITY
1.37
-------
I-J
01
FIGURE 4.4
AVERAGE LOCATION FACTORS BY STATE
1.3
Y / /i / / QKLAHOHf A
> 1.75
-------
FIGURE 4.5
EFFECT OF LOCATION FACTOR ON TOTAL PLANT INVESTMENT
(CONTINGENCY = 0)
TPI = C • E
SCALE UP
FACTOR C
4.4 . -
4.2
4.0
3.8 - -
3.6 - -
3.4 - -
3.2 - -
3.0
2.8 - -
2.6 - -
2.4 . -
2.2 . _
2.0
CHEMICAL
PROCESSING
PLANT
SOLID
HANDLING
PLANT
-I 1
1.0 1.1 1.2 1.3 1.4 1.5 1.6 1.7 1.8 1.9 2.0
LOCATION FACTOR F
76
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5. THE WET LIMESTONE PROCESS
5.1 Process Appraisal
The wet limestone process has been re-examined in the light of
the experiences of the Will County Plant, the Shawnee Plant and
the estimate study carried out for EPA by Catalytic Inc. (16)
Several process alternatives were considered for various sections
of the plant and the conclusion was that the process flowsheet
presented by Catalytic Inc. was the best available solution for
the wet limestone process.
Heat recovery from the flue gas can be dismissed as a practical
possibility. The flue gas would be cooled below its acid dew
point and consequently expensive alloy exchangers would be required.
In order to achieve the required exchanger area, finned tubes
would be required. These would soon be blocked by wet fly ash.
The most sensible method of cooling is obviously by direct scrubbing
in a venturi. The slurry temperature is allowed to rise to 130°F
and all the heat is removed in saturating the flue gas at this
temperature.
A turbulent contact absorber has been selected to handle the
limestone slurry which both silts up and scales equipment. This
is also the design used by Catalytic Inc. In view of the operating
difficulties experienced by Will County with blockage of the
Chevron demister, it appears sensible to make this separate from
the TCA and put considerable thought into its design.
The clean flue gas must be reheated from 130°F to 200°F to restore
its buoyancy and reduce its relative humidity, so there is no
alternative but to provide this heat by burning extra fuel. Al-
though this could amount to 1 or 2% of the boiler fuel, it is
still the most sensible and economic design. The in-line burner
appears a better idea than indirect heating with steam or air,
since the direct heat exchange is more efficient and the additional
77
-------
equipment required is less. A small quantity of low sulfur ash
free fuel is required for this purpose.
Because of the limited nature of the results at this point in
time it is desirable to opt for a safe design. This appears to
be a slurry flowrate to the venturi of 11 gpm/MACFM inlet gas
and a slurry flowrate to the TCA of 55 gpm for the same quantity
of gas. The pressure drop in the venturi should be about 10 inch
w.g. and the gas velocity in the TCA 10 ft/sec. With a venturi
and a 2 stage TCA, the absorption efficiency at 2500 ppm SO2
inlet is about 87%, and with a venturi and 3 stage TCA the efficiency
is about 90%. Although 90% absorption would only be required
when burning 6 or 7% sulfur coals there is little point in design-
ing for anything less than 90%; the slurry flowrate could not be
significantly reduced and the cost of an additional TCA stage is
relatively small. So a venturi and a 3 stage TCA with the above
flowrate will be a constant unit in the wet limestone model. The
process then has the flexibility in the scrubbing section to clean
up flue gas to within the federal limit of 1.2 lb SO_/MMBtu even
when burning 6-7% sulfur coal. A change to a coal containing
more sulfur could be easily accomodated by installing additional
limestone slurry preparation and waste removal units. Short term
changes could even be handled by using the spare limestone grinding
and slurrying equipment.
Two changes have been made to the Catalytic design:
1. The slurry flowrate to the TCA was doubled and the
number of stages increased from 2 to 3.
2. The differential produced by the induction fan was in-
creased from 18 to 30 inch w.g. which is made up of the
following:
78
-------
inches w.g.
Venturi 10
3 stage TCA 8
Entrainment separator 2
New ductwork 7
Plugging allowance __3
30
The maximum sized venturi and TCA unit will handle a gas flow
to the venturi of 550,000 ACFM. This flowrate corresponds to
approximately one sixth of the total gas flow rate at 300°F
from a 1000 megawatt coal fired power station with a heat rate
of 9650 Btu/kwh, assuming a 10% increase due to leakage in the
air preheater.
The process flowsheet is shown in Figure 5.1. The process has
been divided into 3 basic sections:
1. The scrubbing system, including flue gas reheat and dis-
charge
2. The limestone handling and slurry preparation section
3. The waste disposal system including the settling pond.
The first section consists mainly of parallel trains of equipment.
The reference cost for each item of equipment in a train will be
for the size which handles 550,000 ACFM flue gas to the venturi.
The reference cost for equipment in the scrubbing section which
is not required in parallel trains will be the size handling
3,300,000 ACFM flue gas to the Venturis. This corresponds approx-
imately to the size of a 1000 megawatt facility.
The second and third sections will be in one train whatever the
size of the wet limestone facility. Costs in these sections
will be referenced to a sulfur flow rate in the fuel burned of
28,000 Ib/hr, or approximately a 1000 megawatt power station burn-
ing a 4% sulfur coal.
79
-------
5 . 2 Evaluation of Catalytic Inc. Estimate
A general examination of the Catalytic Inc. estimate by the MWK
Estimating Department showed it to be sound. A breakdown of the
total plant cost agrees closely with the equivalent MWK figures.
A close examination of all equipment material and subcontract
costs and comparison with quotations made to MWK shows them to
be reasonable, with the exception of the induction fans and motors.
The cost of a 1500 BHP, 380,000 ACFM fan and motor for the TVA
Gallatin 1050 megawatt plant built in 1955 was $81,000 and so the
Catalytic figure of $90,000 (end of 73) for a 3000 BHP 360,000
ACFM I.D. fan and motor appears to be much too low. Quotations
made recently to MWK in fact show this to be so, the cost of a
4,500 BHP 475,000 ACFM I.D. fan and motor was $210,000 (end of
73). For the maximum sized unit 550,000 ACFM (almost the same
ACFM as to the venturi inlet) , the BHP will be about 4,000. So
a safe assumption for the cost of the incremental I.D. fan and
motor appears to be $200,000 (end of 73).
The basis for all other equipment costs is the Catalytic Inc.
figure listed in Appendix B.
5.3 Variation of Equipment Costs with Plant Size
An article by K.M. Guthrie in the March 1969 issue of Chemical
Engineering, "Data and Techniques for Preliminary Capital Cost
Estimating" (9) , has been used to establish how equipment costs vary
with size.
For most types of equipment, the cost does not vary with size in
the same way over the whole size range. For example the cost of
centrifugal pumps and motors varies with the 0.4 power of the BHP
over the lower size range and the power increases to 0.6 for
larger machines. This can make cost prediction from one quotation
using one exponent very inaccurate especially where large variations
in size are involved.
80
-------
Referring to Figure 5.2, the curve shown represents the variation
of equipment cost with size. This' curve is not an easy thing to
establish and various sources of information have different ideas
on the relationship between equipment cost and size. However
assuming it to be available and accurate, it is still difficult
to use it safely in simple form, when the quotations could be for
equipment on any part of the curve.
If the equipment cost which is to be used as a basis for the cost
equation is in the middle of the curve, then using the average
exponent results in seriously underestimating the cost of larger
and smaller equipment. A safer method is illustrated by considering
the example of the centrifugal pump and motor in Figure 5.2.
The known equipment cost ($5,000 for 100 BHP) is scaled up to the
cost for the maximum required size (200 BHP) using the exponent
for the higher end of the curve 0.6. This maximum cost ($7,600)
is used in the equipment cost equation which is then based on
the average exponent of 0.5. Thus:
£>.5
Pump/Motor Cost = *" ™«'-"^\
In this way the estimated value is always greater than or equal
to the cost curve, which is better in view of the accuracy of
the exponents. However, this method should be used with care,
keeping the overall equipment size range to about 2.0:1.
If the known equipment cost is for a size much smaller than the
required maximum size, then it would be preferable to obtain
another cost estimate closer to the maximum size. In lieu of
this, however, the average exponent could be used for both scaling
up and scaling down.
The equipment required in this process is listed below with the
exponent relating cost to size.
81
-------
Cost Proportional to
Distillation and absorber tower shells
(this applies to the venturi, TCA, and
ductwork)
(ACFM)
0.4-0.6
Distillation column trays
(this applies to the entrainment
separator)
Centrifugal pumps and agitators
(ACFM)
0.8-1.0
(BHP)
0.4-0.6
Horizontal pressure vessels - constant
pressure
Storage tanks - up to 200,000 gal
Storage silos
(volume)
(volume)
(volume)
0.4-0.6
0.4-0.6
0.8-1.0
Conveyors and feeders
(length and height fixed, quantity
handled varies)
Tube mill wet grinders
Separating ponds
(quantity)
(quantity)
(quantity)
0.8-1.0
0.8-1.0
0.8-1.0
The total flue gas flow rate, to the Venturis in the Catalytic
Inc. design is 1,520,000 ACFM and there are 4 scrubbing trains
which handle 380,000 ACFM each. The total sulfur flow into the
control process is 13,000 Ib/hr. Therefore the factors used to
scale up the Catalytic"s costs to the maximum size unit are:
550
= 1.45
3,300 =
1,520
28
IT
= 2.16
(1.45)0-6 = 1.25
(2.17)0'6 = 1.59
(2.16)0*6 = 1.58
82
-------
5.4 Cost Model
5.4.1 Equipment Costs
Using Catalytic's estimate and the scale-up factors given
in the preceding section, equipment costs have been calculated
for the maximum or reference size units and are shown below.
1. The Scrubbing System
Cost of Maximum Cost
Size Train Relationship
M$ (end of 73) with GT
A. Venturi, 3 stage TCA
(extra stage, 15%) and
sumps 1.25 [273 + (908 x
1.15) + 274]/4
= 499
GT
,0.5
B. Entrainment Separator
1.45 (574)/4
= 208
GT
0.9
C. Venturi recirculation
tank, agitator and pumps
1.25 [92 + 28.5 + 51.5]/4
= 54
GT
0.5
D. TCA recirculation tank,
agitator and pumps 1.25
[131 + 33.5 + (98.3 x
= 95
GT
0.5
E. Ductwork (including dampers)
and reheater 1.25 (1,085 +
169)/4 = 393
GT
,0.5
F. I.D. fans and motors (in-
cremental for control
facility) = 200
GT
0.9
83
-------
The cost of equipment in each scrubbing
& reheat train
= 1041 (GT/550)0'5 + 408 (GT/550) °<9 in M$ where GT is
the gas flow rate to each train in MACFM.
In addition to the equipment in the parallel trains there is
the emergency ammonia injection system (G) , the entrainment
separator recirculation system (H) , the reheat fuel storage
and delivery system (J) .
The total cost of these three units for the reference size
plant is
1.59 (10.8 + 64.2 + 75) = 238 M$
The cost of these three units for a. plant handling a total
gas flowrate of GP MACFM
0.5
- 238 - M$
2. The Limestone Handling and Slurry Preparation System
Cost of Reference Cost
Size Unit Relationship
M$ (end of 73) with SF
K. Limestone silo conveyor
and stockpile feeder
2.16 (61.7) = 133 SF°'9
L. Limestone silo and feeders
2.16 (82 + 23.1) = 227 SF°'9
M. 3 tube mill wet grinders and air
compressor 2.16 (595 + 13.5) =1320 SF°'9
84
-------
Cost of Reference Cost
Size Unit Relationship
M$ (end of 73) with SF
P. Slurry hold up tanks,
agitators & feed pumps
1.58 (1.6 +50 + 29.1 +
2.4 + 3.2) = 136 SF°'5
The cost of all the equipment in the limestone handling system
= 1680 (SF/28)0'9 + 136 (SF/28) °'5 M$
where SF is the total sulfur flow into the control unit
in M Ib/hr.
3. The Waste Disposal System Including the Settling Pond
Cost of Reference Cost
Size Unit Relationship
M$ (end of 73) with SF
Q. Surge tanks and pumps
serving pond 1.58 (6.7 +
34.5) = 65 SF0*5
R. Separating Pond (80%
load factor) 0.7 x
2.16 (4,000) = 6,000 SF0'9
NB The Catalytic design included approx.
30% space for fly ash, which is not an
SO control cost. The cost of the sep-
X
arating pond represents the cost in
Cincinnati.
The cost of the surge tanks and pumps serving the pond
= 65 (SF/28)0'5 M$
85
-------
The cost of the separating pond
'>'9
Where LF is the load factor.
The total equipment cost (material and subcontract) for
chemical processing plant (EC) in the Wet Limestone process
NA
RB [1041 (GT/550)0'5 + 408 (GT/550) ° '.9]
n=l
+ 238 RP (GP/3,300)0'5 + 201 (SF/28)0'5 M$
where NA is the number of scrubbing trains and RB and RP are
retrofit difficulty factors as explained below.
The total equipment cost for solid handling plant (ES)
= 1680 (SF/28)0'9 M$
In addition to this, the material and construction costs of
the reference size separating pond, adjusted to 100% load
factor and Gulf Coast location is $5,000 M. Thus:
P = 5,00o M$
RB is the retrofit difficulty factor of the individual boiler.
The increased difficulty is not so much reflected in the actual
major equipment costs as in the increase in other material
and labor costs associated with them. However this is a
convenient place to introduce the factor.
RP is the retrofit difficulty factor of the rest of the
scrubbing section which is not in parallel trains. This
has been assumed equal to the highest RB in the plant.
86
-------
An examination of the MWK reports for EPA, "Applicability
of SO2 Control Processes to Power Plants" and "Evaluation
of the Controllability of Power Plants Having a Significant
Impact on Air Quality Standards" (17, 18), produces a simpli-
fied table of boiler retrofit factors:
TABLE 5.1 Boiler Retrofit Factors
Boiler Size Boiler Age
(Megawatts) (Years) RB
< 50
50-100
101-200
201-500
>500
All new boilers
Other Material Costs
>10
> 1 0
<10
All
All
All
-
and Labor
2.0
1.8
1 R
-L . O
1.6
1.6
1.5
1.4
1.0
Costs
The Guthrie paper ( 9) indicates that different relationships
exist between major equipment costs, other material costs
and labor costs for chemical process plant and for solid
handling plant. This was found to be true for the Catalytic
Inc. estimate although the relationships did not agree with
the Guthrie paper. This is not really surprising as it
depends on how the job is contracted out and estimated. The
obvious solution is to use the relationships generated from
the Catalytic figures since they will be used with Catalytic's
major equipment costs. These costs are listed in Appendix C.
Major Equipment Costs, E:
E = EC + ES
87
-------
Field Labor Costs, L (U.S. Gulf Coast):
LC = 0.39 EC
LS = 0 . 18 ES
L = LC + LS
Other Material Costs, M:
MC = 0.82 EC
MS = 0.09 ES
M = MC + MS
The letter C after the letters E, L and M denotes chemical
process type plant. The letter S denotes solid handling
plant.
5.4.3 Raw Material and Utilities Costs
1. Limestone
The quantity of limestone used by the process during the
year is directly proportional to the sulfur flow into the
control unit, SF, and the boiler load factor, LF. The
Catalytic plant uses 32 tons/hr of limestone for a sulfur
flowrate into the control unit of 13,000 Ib/hr. The reference
flowrate is 28,000 Ib/hr.
The limestone used for the reference flow at 100% load factor
= 32 x ^8 x 8760 = 600 M tons/year
13
the cost of limestone, Al = 600 CL-LF (SF/28) M$
where CL is the purchase price of limestone, $/ton.
88
-------
2 . Ammonia
Ammonia is used intermittently and the yearly consumption
is estimated to be 200 tons. The number of upsets requiring
ammonia injection probably will not reduce with reduction
of the load factor, since startups and shutdowns represent
unsteady conditions. Since the cost is small, for simplifi-
cation the use of ammonia will be assumed directly proportional
to the sulfur flow into the control unit.
Ammonia used for reference flow
= 200 x 28 = 0.43 M tons/year
IT
The cost of ammonia, AA = 0.43 CA (SF/28) M$/year
where CA is the purchase price of ammonia, $/ton.
3. Process Water
The consumption of process water is 400 gpm, which is lost
almost equally between the settling pond and the exhausting
flue gas.
The scale-up factors for the reference flows are:
3300 _ 9 ._ - 28 _ j ,,
TF2Q- - 2.17 and - 2.16
The water consumption per year at the reference flowrate and
100% load factor
= 2.17 x 400 x 60 x 8760
= 460,000 M Gal/year
89
-------
The cost of process water,
AW = 230 CW.LF [(GP/3,300) + (SF/28)] M$/year
where CW is the purchase price of water, $/M Gal.
4. Fuel Oil
The consumption of fuel oil in the catalytic design is 95
MMBtu/hr.
The consumption of fuel oil at the reference flowrate and
100% load factor
= 2.17 x 95 x 8760 MMBtu/year
= 1,800,000 MMBtu/year
The cost of fuel oil,
AF = 1,800 CF-LF (GP/3,300) M$/year
where CF is the purchase price of fuel oil, $/MMBtu.
5. Electricity
The electricity used in the scrubbing section has been
increased by 1100 kw to cover doubling the slurry flow to
the TCA. The rating of the flue gas fans has been increased
by 4300 kw to cover the increased pressure differential.
The total electricity consumed is now 13,050 kw. Of this
11,210 kw are proportional to GP and 1,840 kw are proportional
to the sulfur flow.
The electricity consumption per year at the reference flowrates
and 100% load factor
90
-------
= 2.17 x 11.21 x 8760 (proportional to GP)
= 213,000 M. kwh
and 2.17 x 1.840 x 8760 (proportional to SF)
= 35,000 M. kwh
The cost of electricity,
AE = CE-LF [213 (GP/3,300) + 35 (SF/28)] M$/year
where CE is the purchase price of electricity, mils/kwh.
The total incremental energy consumption of the Wet Limestone
scrubbing unit amounts to about 5% of the HHV of the coal feed
to the power plant.
The total annual cost of raw materials and utilities, ANR,
is given by:
ANR = AL + AA + AW + AF + AE
5.4.4 Total Plant Investment and Total Capital Required
The main costs of the separating pond (P) are the construction
labor costs and land cost and have been assumed to be dependent
on the location at which the Wet Limestone unit is to be built.
The bare cost of the unit can be derived from the General
Cost Model.
BARC = 1.15 (E + M) + (P + 1.43 L) F
The Total Plant Investment is given by:
TPI = 1.12 (1.0 + CONTIN) BARC
91
-------
The contingency CONTIN, represents the degree of uncertainty
in the process design and the cost estimate.
The Total Capital Required is given by the appropriate equation
in the General Cost Model.
TCR =1.15 TPI + 0.8 TO-CO (1 + F) + 0.4 ANR
5.4.5 Operating Costs
The total net annual operating cost, AOC, is the total cost
of operating the plant less the credits from the sale of by-
product. It does not include return of capital, payment
of interest or income tax on equity return. The total net
annual operating cost for the Wet Limestone process is given
by:
AOC = 0.078 TPI + 2TO-CO (1 + F) + ANR
The total number of shift operators, TO, for the Wet Limestone
process is 8 (2 men per shift) for plant capacities of 200
megawatts or above. For plants below 200 megawatts, the
cost for operating labor is assumed to decrease linearly
with size. The hourly wage of the operators, CO, is expressed
in $/hr.
The Total Annual Production Cost, TAG, including the return
of capital, payment on interest and income tax on equity
return is given by:
TAG = 0.237 TPI + 2.1 TO-CO (1 + F) + 1.04 ANR
5.5 Effect of Various Parameters on Costs
In Figures 5.3-5.7 typical costs which were calculated from the
model have been plotted to illustrate the effects of different
variables on plant costs. Unit values for raw materials and utilities,
92
-------
which were used in determining operating costs, are as shown in
Table 5.2
These plots are not for actual, existing plants, but have been
included merely to illustrate typical cost variations predicted
by the model. Although the figures are self-explanatory, some
of the more significant conclusions should be noted.
Figure 5.3 shows the large effect of plant capacity (i.e., gas
flow) on capital required. Small plants are far more expensive
to control then large ones. While a new 1000 MW plant (4% S,
80% load factor) could be controlled for about $46/KW, for a 10
MW plant it would cost almost three times as much. The sulfur
content of the coal has a noticeable but minor effect on cost,
particularly at small plant capacities.
Figure 5.4 illustrates the pronounced effect of load factor on
operating cost. In fact, decreasing the load factor from 80% to
40% is more significant than quadrupling the sulfur content of the
coal. Plant capacity has an effect on operating cost similar
to that on capital required.
Figures 5.5-5.6 show the influence of the retrofit factor on costs.
As it becomes more difficult to install a wet limestone unit at
an existing plant, capital required increases substantially.
For a 10 MW plant, it could be more than $250/KW. Even for a
large 1000 MW plant, capital required could be as much as 70%
more than for a new plant. The increase in operating cost is due
to the fixed charges on the additional capital.
In figures 5.7, the effect of location factor on capital required
is shown. Basically, this shows the influence of higher labor
rates on the construction cost of the plant. Relative to a
Gulf Coast location, costs could be as much as 25-35% higher at
other locations.
93
-------
5.6 Nomenclature
GP
GT
Total gas flow into all Venturis MACFM
Total flow of gas into each venturi MACFM
(Maximum value of GT = 550)
NA
SF
LF
RB
RP
Number of venturi/TCA trains
(GT = GP/NA 'for a' new plant)
Maximum flow of sulfur into the control M Ib/hr
unit
Load factor of the power station
The retrofit difficulty factor of a
boiler
The retrofit difficulty factor of all
scrubbing equipmen't which is not in parallel
trains. Assumed.to be equal to the highest
RB
CL
CA
CW
CF
CE
The purchase price of limestone
The purchase price of ammonia
The purchase price of process water
The purchase price of fuel oil
The purchase price of electricity
$/ton
$/ton
$/M Gal
$/MM Btu
Mils/kwh
94
-------
CO
M
c, s
AL
AA
AW
AF
AE
TPI
TAG
BARC
The direct cost of operating labor
Major equipment cost
(Material and subcontract)
Other material costs
(Piping, instruments, electrical
civil etc.)
Direct field labor costs
Letters follows E, M and L
C refers to chemical process type
equipment
S refers to solid handling equipment
The total cost of the settling pond
(Material and total labor)
Total annual cost of limestone
Total annual cost of ammonia
Total annual cost of process water
Total annual cost of fuel oil
Total annual cost of electricity
The total plant investment
Total annual production cost of wet
limestone SO_ control unit
The bare cost of the control unit
$/hour
M$
M$
MS
M$
M$/year
M$/year
M$/year
M$/year
M$/year
M$
M$/year
M$
95
-------
AOC Annual net operating cost M$
TCR Total capital required M$
CONTIN Contingency
F Location Factor
96
-------
TABLE 5.2
UNIT COSTS USED IN ILLUSTRATIVE
EXAMPLES - WET"LIMESTONE STACK
GAS SCRUBBING MODEL
Purchased Price of Limestone ($/Ton) 4.00
Purchased Price of Ammonia ($/Ton) 50.00
Purchased Price of Water ($/MGal) 0.20
Purchased Price of Fuel Oil ($/MMBtu) 0.80
Purchased Price of Electricity (mils/Kwhr) 8.00
Average Hourly Wages Per Gulf Coast ($/Hr) 7.00
Interest on Capital During Construction (%) 12.00
97
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TABLE 5.3
WET LIMESTONE PROCESS AND COST MODEL
SUMMARY OF EQUATIONS
AL = 600 CL-LF (SF/28) M$/year
AA = 0.43 CA (SF/28) M$/year
AW = 230 CW-LF [ (GP/3,300) + (SF/28)] M$/year
AF = 1,800 CF-LF (GP/3,300) M$/year
AE = CE-LF [213 (GP/3,300) + 35 (SF/28)] M$/year
ANR = AL + AA + AW + AF + AE M$/year
NA
EC = ^7 RB [1041 (GT/550)U'b + 408 (GT/550)0'9] M$
*^ i n
n=l
+238 RP (GP/3,300)0-5 + 201 (SF/28)0'5
ES = 1680 (SF/28)0'9 M$
P = 5,000 [2|^J M$
BARC = 1.15 (E + M) + (P + 1.43 L)F M$
TPI = 1.12 (1.0 + CONTIN) BARC M$
TAG = 0.237 TPI + 2.1 TO-CO (1+F) +1.04 ANR M$/year
TCR = 1.15 TPI +0.8 TO-CO (1+F) +0.4 ANR M$
98
-------
FIGURE 5.1 WET LIMESTONE PROCESS FLOWSHEET
VO
DIRECT FIRED
RECIRCULATION
TANKS AND
PUMPS
ENTRAPMENT
AND DAMPERS
FLUE GAS
TOTAL FLOW TO ALL TRAINS Gp ACFM
SULFUR FLOW Sc M LB/HR
AGITATOF
AND PUMP
WET GRINDER.
SLURRY TANKS
SETTLING POND
SECTION II
SECTION
-------
FIGURE 5.2
METHOD OF VARYING EQUIPMENT COST WITH SIZE
PUMP/MOTOR COST = K(BHP)n
LOG [COST]
$7,600 MAXCOST
PUMP/MOTOR
COST EQUATION = 7,600
•$7,000
ACTUAL
CURVE
tOO 200
/SIZE WITH \ /MAXIMUM\
I KNOWN COST/ \SIZE )
LOG [BHP]
MAXCOST = 5000
COST EQUATION
/200 \
VfooJ
PUMP/MOTOR COST = 7-
$7,600
0.5
100
-------
FIGURE 5.3
EFFECT OF BOILER CAPACITY ON TOTAL CAPITAL REQUIREMENT
WET LIMESTONE PROCESS
200.
UJ
IT
a
Ul
cc
5
_l
<
100.
90.
80.
70.
60
50 .
40.
30.
20
10
80% BOILER LOAD FACTOR
40% BOILER LOAD FACTOR
I I I I
_L
I I I I I I III
20
40
60 80 100
BOILER CAPACITY, MW
200
400
600 800 1000
BASIS OF CALCULATION = NO RETROFIT, NO CONTINGENCY, U.S.
GULF COAST LOCATION, END OF 1973 FIGURE, BOILER HEAT
RATE 9,500 BTU/KWH, HEATING VALUE OF COAL (HHV) OF
11.000 BTU/LB.
-------
FIGURE 5.4
EFFECT OF BOILER CAPACITY ON PRODUCTION COST
WET LIMESTONE PROCESS
200,
o
M
8
o
o
o
tr.
100.
90
80
70.
jO 70
30
20
10
10
.^COAL SULFUR CONTENT 8%
— —. 4%
• .T" —* • 2%
80% LOAD FACTOR
40% LOAD FACTOR
20
40 60 80 100 200
BOILER CAPACITY. MW
400
600
800 1000
(SEE FIGURE 5.3 FOR BASIS OF CALCULATION.)
-------
FIGURE 5.5
EFFECT OF BOILER RETROFIT DIFFICULTY ON TOTAL CAPITAL REQUIREMENT
WET LIMESTONE PROCESS
300
200
1-'
HI
Ul
x
5
a
tU
-------
FIGURE 5.6
EFFECT OF BOILER RETROFIT DIFFICULTY ON PRODUCTION COST
WEffl LIMESTONE PROCESS
200
z
cc
m
8
u.
O
HI
u
1
§
8
en
O.
20
1.0
10 MW
50 MW
100 MW
200 MW
400 MW
600 MW
1000 MW
1.2
1.4 1.6
RETROFIT FACTOR
(SEE FIGURE 5.5 FOR BASIS OF CALCULATION.)
1.8
2.0
104
-------
FIGURE 5.7
EFFECT OF LOCATION FACTOR ON TOTAL CAPITAL REQUIREMENT
WET LIMESTONE PROCESS
300
NO RETROFIT
200
Ul
ui
a
UJ
IT
a.
o
10 MW BOILER SIZE
50 MW
100 MW
200 MW
400 MW
600 MW
800 MW
1000 MW
40
30
1.0
1.2
1.4 1.6
LOCATION FACTOR
1.8
2.0
BASIS OF CALCULATION: 4% SULFUR COAL. HEATING
VALUE 11,000 8TU/LB, INDIVIDUAL BOILER WITH HEAT
RATE OF 9,500 BTU/KWH AND LOAD FACTOR OF 0.7,
CONTINGENCY 10% IN CAPITAL INVESTMENT, U.S. GULF
COAST LOCATION, END OF 1973 FIGURE.
105
-------
6. THE WELLMAN/ALLIED PROCESS
6.1 Process Appraisal
The basis for the Wellman/Allied process and cost models is the
design proposed for the demonstration plant to be installed at
the D.H. Mitchell plant of the Northern Indiana Public Service
Company. This system is a combination of the Wellman-Lord S02
recovery process and the Allied Chemical SO reduction process,
producing elemental sulfur as an end product. The process and
cost models were developed for this combined system. For simplicity,
it will hereafter be referred to in this report as the Wellman/Allied
system or process.
The NIPSCO design, which is for a 115 MW plant burning 3.2% sulfur
coal, is the only one available which is sufficiently detailed for
use in deriving the models. While a review of the design showed
it to be reasonable, a number of process changes were made and are
discussed below.
The NIPSCO design provides for an absorber capable of handling the
maximum flow of flue gas from the power plant whereas the SO..
recovery system is designed for the average flow rate, corresponding
to an 80% load factor. The difference between these design capacities
is handled by providing large surge capacity for the sodium sulfite
solution. For the model, this has been simplified to allow the en-
tire regeneration plant to run at full absorber capacity, i.e., at
100% load factor. Correspondingly, surge capacity has been reduced.
In addition to being simpler to model (i.e., does not require know-
ledge of short-term variation of boiler load factor) , this type of
design provides the capability of operating the regeneration and
recovery system at peak load conditions if it were necessary, thus
giving the total system a greater flexibility and operating range.
Under normal operation, the regeneration and recovery system would
operate at some reduced steady state level, while allowing the scrub-
bing section to fluctuate in response to varying demands from the
boiler. The reduced level of operation would have to be adjusted
106
-------
periodically, depending on solution inventory and anticipated boiler
operation. Of course, this type of design would be somewhat more
expensive than designing for the average flow (see section 6.6).
The NIPSCO absorber was designed to remove about 91% of the S02 in
the flue gas and included three contacting stages plus space for
a fourth. In lieu of operating data, it was assumed that four trays
would be a safe design for 90% removal. An additional tray was
added for the model and it was assumed that this would give an overall
SO removal efficiency of 95%. This high removal efficiency was chosen
as the "standard" design to permit use of the model not only in
utility applications, but also in other applications (smelters,
Glaus plants, etc.) where high S02 removal efficiencies would be
particularly desirable. However, high removal efficiency is also
a useful device for investigating utility plant applications, since
the computer cost program has been designed to consider sequential
control of plant boilers until the desired emissions limitations
are achieved (see section 7.1) .
The flue gas blower has been changed from upstream of the absorber
to downstream of the reheater. Pressure drop has been increased from
18" HO to 30" HO and is distributed as follows:
£ £
InchesH O
'^ ~~ •
Pre scrubber 6
5 Absorption Stages 15
Demister 2
Ductwork 7
30
Some items which are identical in design and operation to those
which were included in the wet limestone model previously developed
have been based on that model rather than the NIPSCO design. These
are: the flue gas ductwork and dampers, the gas reheater, and the
reheat fuel storage and delivery system.
All pumps, fans, blowers, and compressors have been spared with the
107
-------
exception of the flue gas fan and the SO_ compressor. In cases
^
where multiple units are used for a single service, it was assumed
that one common spare would be adequate.
The NIPSCO design included a separate stack for the clean flue
gas. This was deleted from the model since the clean flue gas would
ordinarily be diverted back to the main stack, which is properly
cos ted to the source plant and not the control unit.
The process flow sheet is shown in Figure 1 and is divided into
four main areas:
1) The absorber area, including gas reheat and compression
2) The SO2 regeneration area
3) The purge/make-up area
4) The SO_ reduction area (Allied Process)
6.2 Evaluation of the NIPSCO Project Cost Estimate
The cost estimate made by Davy Powergas and Allied Chemical for
the NIPSCO project has been used as a basis for the cost model.
Since process equipment costs for this project were more than 85%
quoted, they should form a sound basis for defining equipment costs
in the model. Quotes were received primarily during the latter part
of 1972 and have been assumed to be valid as of the end of 1972.
Before using these costs, they have been increased by 5% to allow
for escalation to the end of 1973, which is the reference time
chosen for the model.
For some pieces of equipment in the absorber area that are common
in design and operation to both the Wellman/Allied system and the
wet limestone process, costs have been derived from the wet limestone
model rather than the NIPSCO estimate. These are:
1. Induction fan
2. Reheater, ductwork, and dampers
3. Fuel oil system
108
-------
6.3 Variation of Equipment Costs with Plant Size
In order to determine exponents relating cost to size for different
types of equipment, several sources were consulted (9, 14, 21).
This resulted in the following variation of equipment cost with size
for the different types of equipment used in the Wellman/Allied
system:
Tower shells (including lining)
Tower internals
Centrifugal pumps
Tanks and drums
Agitators
Pressure filter
Fans, blowers, and compressors
Direct-fired heaters
Ductwork and dampers
Heat exchangers
Forced-circulation evaporators (complete
system)
Storage silos and bins
Entrainment separators
Pressure vessels
Pressure vessel internals
Sulfur pit
Miscellaneous solids handling equipment
Cost Proportional to
0.4-0.6
(ACFM)
(BHP)
0.4-0.6
0.4-0.6
(volume)
0.4-0.6
(BHP)
0.5-0.7
(flow)
0.8-1.0
(BHP)
(duty)
(ACFM)
0.4-0.6
0.4-0.6
0.5-0.7
(surface)
0.5-0.7
(duty)
0.8-1.0
(volume)
0.8-1.0
(ACFM)
0.4-0.6
(volume)
0.8-1.0
(ACFM)
0 .8-1.0
(volume)
..... .0.8-1.0
(flow)
Equipment sizes for the, process were related to either of two basic
variables: the flue gas flow or the flow of sulfur in the flue gas,
The absorber and related equipment, which are all included in the
absorber area, are proportional in size to the flue gas flow. The
remainder of the process equipment in the plant is proportional to
sulfur flow. Each process section was reviewed to determine its
maximum train size, based on the equipment sizes shown for the
NIPSCO design.
An analysis of the absorber area showed that the maximum size
109
-------
absorber can handle a gas flow of about 550,000 ACFM. The other
gas-related equipment has also been limited to this maximum size,
with the exception of the fuel oil system. The latter, which was
taken from the wet limestone model, has been assumed to be single
train, regardless of plant size. For the sulfur-related equipment
in the absorber area, an upper limit on equipment sizes was found
at a train size corresponding to about 7,000 Ibs/hr of sulfur.
This was also the case for the SO regeneration area. Equipment
£
in the purge/make-up area can be single train at the reference
plant flow of 28,000 Ibs/hr of sulfur, but this is about the max-
imum practical size for this section. Within the size ranges of
interest, no upper limit was found for single train operation in
the SO reduction area (Allied plant).
Considering the reference plant size (3,300 MACFM of flue gas and
28,000 Ibs/hr of sulfur), and based on the NIPSCO design, scale-
up factors to maximum or reference size trains were determined to
be:
Number of Trains in
Area scale-up Factor Reference Size Plant
Absorber Area
gas-related equipment 1.18 6
sulfur-related equipment 3.04 4
SO Regeneration Area 3.04 4
Purge/Make-up Area 12.16 1
SO Reduction Area (Allied plant) 12.16 1
6.4 Cost Model
6.4.1 Equipment Costs
Using the NIPSCO estimate as a basis and the sacle-up factors
given in the preceding section, equipment costs have been
calculated for the maximum or reference size units and are
shown below.
110
-------
1. The Absorber Area
Cost of Maximum
Size Train Cost
M$ (end of 73) Relationship
A. Absorber shell, lining
and circulation pumps ;
prescrubber circulation
pumps; reheater, duct-
work and dampers 726 GT *
B. Vessel internals; induc-
tion fan 639 GT0*9
C. Fuel oil system2 119 GP°'9
D. Tanks , pumps, and agita-
tors 113 S7°'5
E. Fly ash filter system 127 S70-6
where GT is the flue gas flow per train in MACFM, GP is
the total plant flue gas flow in MACFM, and S7 is the
sulfur flow rate per train in Mlbs/hr.
For the reference size plant, six maximum size absorber trains,
one fuel oil system, and four maximum size sulfur trains are
required. Thus, the total equipment cost for the absorber
area, EA, is:
1 NIPSCO costs have been adjusted to provide for an additional
absorption stage.
2 Unit cost taken from wet limestone model.
3 NIPSCO costs have been adjusted to reflect reduced surge
capacity.
Ill
-------
EA = (726 + 639) x (6 trains)
+ (119) x (1 train)
+ (133 + 127) x (4 trains)
= $9394M
In general, for a plant with a total gas flow of GP MACFM,
a gas flow per train of GT MACFM (with NA absorber trains) ,
a total sulfur rate of SF M Ibs/hr , and a sulfur rate per
train of S7 M Ibs/hr (with N7 sulfur trains) , the total
equipment cost for the absorber area is:
NA
0*5
=^Z I726 (GT/550)0'5 + 639 (GT/550)0*9 + 119 (GP/3300)
n=l L Jn
+ [133 (S7/7)°>5 + 127 IF (S7/7)°>6j
N7 M$
where
S7 = SF/N7
N7 = SF/7 (rounded to next higher integer)
IF is merely an index used to include or delete the cost of
a fly ash filter system, as necessary.
IF = 1 if particulates are present in the flue gas
IF = 0 if particulates are absent from the flue gas
2. The SO Regeneration Area
112
-------
Cost of Maximum
Size Train, M$ Cost
(end of 73) Relationship
A. Vessels, agitators and
pumps1 209 S7°'5
B. Heat exchangers, evapora-
tor system 618 S7°'6
C. Compressor, vessel inter-
nals 157 S7°*9
For the reference size plant, four trains are required,
each handling the maximum sulfur rate per train of 7 M Ibs/hr.
The total equipment cost for the SO regeneration area is:
ES = (209 + 618 + 157) x (4 trains)
= $3936M
In general, for a plant handling a sulfur rate per train
of S7 M Ibs/hr (with N7 sulfur trains), the total equipment
cost for the SO regeneration area is:
£»
ES = [209 (S7/7)0'5 + 618 (S7/7)°t6 + 157 (S7/7)°'9]N7 M$
3. The Purge/Make-up Area
Cost of Maximum
Size Train, M$ Cost
(end of 73) Relationship
A. Pumps, tanks, agitators,
heat exchangers, and dryer 525 S28 '
B. Separating Equipment 380 S28
1 NIPSCO costs for absorber feed tank and agitator have been
adjusted to reflect reduced surge capacity.
113
-------
Cost of Maximum
Size Train, M$ Cost
(end of 73) Relationship
C. Special equipment
D. Packaged heat exchanger
86
306
S28
S28
0.7
0.8
E. Fan and miscellaneous solids
handling equipment
519
S28
0.9
For the reference size plant, one train is required, handling
the maximum sulfur rate per train of 28 M Ibs/hr. The total
equipment cost for the purge/make-up area is:
EP = (525 + 380 + 86 + 306 + 519) x (1 train)
= $1816M
In general, for a plant handling a sulfur rate per train of
S28 M Ibs/hr (with N28 sulfur trains), the total equipment
cost for the purge/make-up area is:
EP
= [525 (S28/28)0'5 + 380 (S28/28) °' 6_ + 86 (S28/28)0'7
+ 306 (S28/28)
0.8
+ 519 (S28/28)
0.9"!
N28
M$
where N28 = SF/28 (rounded to next higher integer)
S28 = SF/N28
4. The SO» Reduction Area (Allied plant)
114
-------
Cost of Reference
Size Train, M$ Cost
(end of 73) Relationship
A. Pumps, fired heaters,
vessels, and ductwork
998
SF
,0.5
B. Heat exchangers
287
SF
0.6
C. Compressors, mist eliminator,
sulfur pit, and vessel inter-
nals
683
SF
,0.9
For the reference size plant, this area is sized for an equiva-
lent sulfur rate of 28 M Ibs/hr. The total equipment cost for
the SO reduction area is:
ER = 998 + 287 + 683
= $1968M
In general, for a plant handling a sulfur rate of SF M Ibs/hr
the total equipment cost for the SO- reduction area is:
ER = 998 (SF/28)0*5 + 287 (SF/28)0'6 + 683 (SF/28)0*9 M$
The total equipment costs for the Wellman/Allied system can be
summarized as follows:
NA
EA = ^> ' RB J726 (GT/550)0*5 + 369 (ST/550)0*9]
n=l
n
+ 119 RP (GP/3300)0'5 + Il33 (S7/7)°'5 + 127 IF (S7/7)0'6JN7 M$
ES =1209 (S7/7)°*5 + 618 (S7/7)°t6 + 157 (S7/7) "' *\ N7 M$
4
4
0.9"]
EP =1525 (S28/28)0'5 + 380 (S28/28)0'6 + 86 (S28/28)0-7
+ 306 (S28/28)0*8 + 519 (S28/28)°* * |N28
115
M$
-------
ER = 998 (SF/28)0'5 + 287 (SF/28)0'6 + 683 (SF/28)0'9 M$
RB and RP are the retrofit difficulty factors as described in
the wet limestone process model.
6.4.2 Other Material Costs and Labor Costs
Costs for labor and other materials generally can be estimated
as a percentage of major equipment costs. Since the NIPSCO
estimate was broken down by plant area, factors were obtained
for these costs for each area. The factors derived from the
data are shown below, where E is the major equipment cost,
L is the labor cost, and M is the cost of other materials.
The letters A, S, P, and R refer to the absorber area, the
SO- regeneration area, the purge/make-up area, and the SC-
reduction area respectively. Labor costs are based on the
Gulf Coast area. Field materials include only piping, instru-
ments, electrical, insulation, painting, concrete, and structural
steel.
LA = 0.224 EA MA = 0.429 EA
LS = 0.310 ES MS = 0.742 ES
Lp = 0.433 Ep Mp = 0.827 Ep
LR = 0.623 ER MR = 0.772 ER
6.4.3 Raw Materials and Utilities Costs
1. Sodium Carbonate
Sodium carbonate make-up is required to replenish the sodium
values lost by oxidation of the scrubbing solution. The
quantity used is directly proportional to the sulfur rate,
SF. For the NIPSCO design, 0.265 tons/hr were required.
Since for the reference size plant the scale-up factor on
the sulfur rate is 12.16, the sodium carbonate make-up for the
reference plant at 100% load factor is:
116
-------
Consumption = x 12.16 x 8760 = 28.2 M tons/yr
In general, the annual cost of sodium carbonate, AS, for
a power plant having a load factor of LF is:
AS = 28.2 CS-LF (SF/28) M$/yr
where CS is the purchase price of sodium carbonate in $/ton.
2. Natural Gas
Natural gas is used in the S02 reduction area to convert the
SO,, to elemental sulfur. The amount consumed is pro-
portional to the sulfur rate, and for the NIPSCO plant
equals 13.7 MSCFH. The annual consumption for the reference
plant at 100% load factor is:
Consumption = j^jj x 12.16 x 8760 = 1460 MMSCF/yr
The annual cost of natural gas, AN, is:
AN = 1460 CN-LF (SF/28) M$/yr
where CN is the purchase price of natural gas in $/MSCF.
3. Filter Aid
Filter aid, which is used in the fly ash filter system, is
of course needed only if particulates are present in the
flue gas. The quantity required is assumed to be proportional
to the gas flow. A design rate of 40 Ibs/hr was shown for
the NIPSCO design. The scale-up factor to the reference
size plant, on gas flow, is 6 x 1.18, or 7.08. For the
reference plant, therefore, the annual consumption at 100% load
factor is:
117
-------
Consumption = 20QQ x j^QO" x 876° = -1*24 M tons/yr
The annual cost of filter aid, AFA,is:
AFA = 1.24 CFA-LF-IF (GP/3300) M$/yr
where CFA is the purchase price of filter aid in $/ton and
IF is the fly ash index previously defined.
4. Power
The power consumption shown for the NIPSCO design has been
adjusted to reflect some process and equipment changes (an
additional absorption stage, increased gas pressure drop,
etc.) which were incorporated in the model, as discussed
previously. The adjusted power requirement for the NIPSCO
design is 3220 KW of which 2480 KW are proportional to the
gas flow rate and 740 KW are proportional to the sulfur rate.
The annual power consumption of the reference plant at 100%
load factor is:
Consumption = 2.480 x 7.08 x 8760 (proportional to GP)
+ 0.740 x 12.16 x 8760 (proportional to SF)
= 154,000 MKWH/hr. + 79,000 MKWH/yr
The annual power cost, AE, is:
AE =[l54 (GP/3300) + 79 (SF/28)] CE-LF M$/yr
where CE is the purchase (or transfer) price of electricity
in mills/KWH.
5. Steam
The steam consumption shown for the NIPSCO design has been
adjusted because of the deletion of steam turbine drives
118
-------
on the flue gas fan and the SO- compressor. The adjusted
value is 51.0 M Ibs/hr. and is proportional to the sulfur
rate. For the reference plant at 100% load factor, the
steam consumption is:
Consumption = jjj^j- x 12.16 x 8760 = 5430 MM Ibs/yr.
The annual cost of steam, AH, is:
AH = 5430 CH.LF (SF/28) M$/yr
where CH is the purchase (or transfer) price of steam in $/M Ibs
6. Cooling Water
The total cooling water requirement for the NIPSCO plant
is 3.34 MGPM of which 0.23 MGPM is proportional to the gas
flow and 3.11 MGPM is proportional to the sulfur rate. Cooling
water required for the reference plant at 100% load factor is:
Consumption = 0.23 x 7.08 x 60 x 8760 (proportional to GP)
+ 3.11 x 12.16 x 50 x 8760 (proportional to SF)
= 856,000 M gal/yr + 19,900,000 M gal/yr
The annual cost of cooling water, ACW, is:
ACW =[856 (GP/3300) + 19,900 (SF/28)"] CCW-LF M$/yr
where CCW is the cost of cooling water in $/M gal.
7. Process Water
Small amounts of process water are used in the purge and
make-up systems and are proportional to sulfur rate. For
NIPSCO, process water use is about 10 GPM. The quantity
required for the reference plant at 100% load factor is:
119
-------
Consumption = 1QOQ x 12.16 x 60 x 8760 = 64,000 M gal/yr
The annual cost of process water, AW, is:
AW = 64 (SF/28) CW-LF M$/yr
where CW is the cost of process water in $/M gal.
8. Fuel Oil
Since the fuel oil system for reheating the flue gas is
identical to that included in the wet limestone model, the
oil consumption and cost will be the same. For the reference
plant at 100% load factor:
Consumption = 1,800,000 MM Btu/yr.
The cost of fuel oil, AF, is:
AF = 1,800 (GP/3300) CF-LF M$/yr
where CF is the purchase price of fuel oil in $/MM Btu.
9. Credits
The process produces two materials: sulfur, and a dry purge
solids stream consisting of sodium sulfite, sodium sulfate, and
and sodium thiosulfate. The product sulfur would normally
be listed as a credit. However, the purge solids may have
positive or negative value depending upon whether or not they
are salable. Normally, it is expected that a waste disposal
cost would be incurred. The cost treatment of the purge
solids can be handled by insertion of a positive or negative
unit value in the model.
120
-------
a. Sulfur
The sulfur production for the NIPSCO plant is 21.5 long
tons/day and is proportional to the sulfur rate. For
the reference plant at 100% load factor:
Production = j^ffi * 12.16 x 365 = 95.4 M long tons/yr
The sulfur credit, ASC, is:
ASC =95.4 (SF/28) VSC'LF M$/yr
where VSC is the unit value of sulfur in $/long ton.
b. Purge Solids
The NIPSCO design shows a purge solids production rate
of 0.35 tons/hr which is proportional to the sulfur rate.
The purge solids flow for the reference plant at 100%
load factor is:
Production = ^-jj- x 12.16 x 8760 = 37.3 M tons/yr
The purge solids credit (or debit), APS, is:
APS =37.3 (SF/28) VPS-LF M$/yr
where VPS is the unit value of the purge solids in $/ton.
If the purge solids are listed as a credit (debit) ,
VPS would be positive (negative).
The total cost of raw materials and utilities less credits,
ANR, is:
ANR = AS + AN + AFA + AE + AH + ACW + AW + AF - ASC - APS M$/yr
121
-------
6.4.4 Total Plant Investment and Total Capital Required
The bare cost (BARC) , total plant investment (TPI) , and total
capital required (TCR) for the WeiIman/Allied system can be
calculated from the appropriate equations in the General Cost
Model. Thus,
BARC =1.15 (E+M) + 1.43 L«F M$
TPI = 1.12 (1.0 + CONTIN)-BARC M$
TCR =1.15 TPI + 0.8 TO.CO (1.0+F) +0.4 ANR M$
where E = EA + ES + EP + ER M$
M = MA + MS + MP + MR M$
L = LA + LS + LP + LR M$
6.4.5 Operating Costs
The total net annual operating cost, AOC, represents the total
cost of running the plant, excluding depreciation, interest,
and income tax. It is given by the following equation from the
General Cost Model:
AOC = 0.078 TPI + 2.0 TO-CO (1.0 + F) + ANR M$/yr
where TO = total number of shift operators
CO = hourly rate of operators
For plants larger than 200 MW, the Wellman/Allied process requires
16 operators (4 per shift) . It has been assumed that for plants
less than 200 MW, operating labor costs are directly proportional
to plant size.
The total annual production cost, TAG, including the return on
capital, interest, and income tax is given by:
TAG = 0.237 TPI + 2.1 TO-CO (1.0 + F) + 1.04 ANR M$/yr
122
-------
6.5 Effect of Various Parameters on Costs
Figures 6.2-6.6 show typical costs which were calculated from the
model illustrating the effects of different variables on plant
costs. Unit values used fir raw materials and utilities are listed
in Table 6.1. (
In general, the effects of variables are similar to those noted for
the Wet Limestone process. However, there are some important
differences. Costs for Wellman/Allied are much greater than Wet
Limestone for small boilers and high percent sulfur. For all sizes,
percent sulfur has a greater impact on costs in the Wellman/Allied
system.
Load factor has a significant effect on operating costs, as in the
Wet Limestone process, particularly for small sizes. For large
plants, percent sulfur has a greater effect than load factor.
Retrofit factor is less significant for the Wellman/Allied process
compared to Wet Limestone. This is due to the fact that much of
the cost is in sulfur recovery rather than scrubbing. The latter
is usually where the retrofit difficulty occurs.
6.6 Wellman/Allied Process Variations and Impact on Costs
For convenience and simplicity, the model developed for the Wellman/
Allied system uses a single processing scheme. There are, however,
process modifications which could be made and which could effect
costs.
The model assumes that the SO recovery and reduction sections would
^
be designed for full capacity. Designing for less than full
capacity is possible, as was done for the NIPSCO project. This
requires some accurate knowledge of the expected variation of
load factor with time. If this information is available and indicates
123
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that a less than full capacity design is possible, the model
could easily be changed to accomodate this. The sulfur flow
in the gas would be reduced appropriately before using the equip-
ment cost equations. For example, if the recovery sections were
to be designed for 80% of full capacity, the design sulfur rate
would be:
Sp' = 0.8 Sp
S ' would then be used to cost equipment. Calculations with the
model show that sizing recovery sections for 80% of capacity would
reduce capital costs by about 6-8% below costs for a 100% design.
The model uses a single effect evaporator (for SO regeneration)
£
and assumes electric drive for the flue gas fans. Steam economies
could be achieved by double-effect evaporation and steam drive for
the fans. Low pressure exhaust steam from the fans would be used
in the first effect of the evaporator. By this method, overall
energy consumption of the process could be reduced from more than
11% of the power plant heat input to perhaps 8-9%. However, it is
likely that capital costs would increase. In addition, a double-
effect evaporator with its first effect operating at a higher
temperature suggests the likelihood of increased sulfite oxidation
losses. This would increase the make-up cost. Although there were
not enough data available to estimate costs for this type of design,
it appears unlikely that costs could be reduced significantly.
It has been assumed for all models that costs of multiple trains
are direct multiples of single train costs. Since there are several
sections in the Wellman/Allied system where multiple trains may
occur, it was decided to investigate this assumption in some
detail.
Potential savings exist only in engineering costs and possibly field
supervision, if the multiple units are constructed concurrently.
Estimates were made indicating potential cost reductions are quite
small. Multiple units of 2-6 trains show a possible reduction in
investment of 1-3% compared with the basic assumption of multiple
124
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train costs.
6.7 Nomenclature
GP
GT
NA
SF
S7
S28
Total flue gas to control plant MACFM
Total flue gas to each absorber MACFM
train (maximum value of GT = 550)
Number of absorber trains
Total sulfur flow in flue gas to
control plant M Ibs/hr
Total sulfur flow in flue gas to
control unit per train of sulfur-
related equipment in absorber and S0_
regeneration areas (maximum value of
S7 = 7) . M Ibs/hr,
Total sulfur flow in flue gas to
control unit per equipment train
in the purge/make-up area
(maximum value of S28 = 28) M Ibs/hr,
N7
N28
Number of trains of sulfur-related
equipment in the absorber and SO-
regeneration areas.
Number of equipment trains in the
purge/make-up area
M
L
A,S,P,
IF
RB
RP
Major equipment cost
(direct material and subcontracts) $M
Field Materials Costs $M
Field Labor Costs
Letters following E,M,L
A refers to absorber area
S refers to S02 regeneration area
P refers to purge/make-up area
R refers to S02 reduction area
No letter following refers to total for
all areas
Particulate index (IF = 1 if par-
ticulates are present in flue gas.
IF = 0 if particulates are absent)
Retrofit difficulty factor of each
boiler
Retrofit difficulty factor of gas-
related equipment in the absorber area
which is not in parallel trains, i.e.,
the fuel oil system; assumed to be equal
to the highest RB
125
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BARC
TPI
TCR
CONTIN
AS
AN
AFA
AE
AH
ACW
AW
AF
ASC
APS
CS
CN
CFA
CE
CH
CCW
CW
CF
VSC
VPS
TO
CO
LF
AOC
TAG
F
Bare cost of the control unit
Total Plant Investment
Total Capital Required
Contingency
Annual cost of sodium carbonate
Annual cost of natural gas
Annual cost of filter aid
Annual cost of electric power
Annual cost of steam
Annual cost of cooling water
Annual cost of process water
Annual cost of fuel oil
Annual sulfur credit
Annual purge solids credit or debit
Purchase price of sodium carbonate
Purchase price of natural gas
Purchase price of filter aid
Purchase (or transfer) price of
electricity
Purchase (or transfer) price of steam
Cost of cooling water
Cost of process water
Purchase price of fuel oil
Unit value of sulfur (negative if
credit)
Unit value of purge solids (negative
if credit)
Total number of operators
Unit cost of operating labor
Load factor of the power plant
Annual net operating cost
Total annual production cost
Location Factor
$M
$M
$M
$M
$M/Yr
$M/YR
$M/YR
$M/yr
$M/Yr
$M/Yr
$M/Yr
$M/Yr
$M/Yr
$M/Yr
$/ton
$/MSCF
$/ton
mills/KWH
$/M Ibs.
$/M gal.
$/M gal.
$/MM Btu
$/long ton
$/ton
$/hr
$M/Yr
$M/Yr
126
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TABLE 6.1
UNIT COSTS USED IN ILLUSTRATIVE
EXAMPLES - WELLMAN/ALLIED STACK
GAS SCRUBBING MODEL
Purchased Price of Sodium Carbonate ($/Ton) 40.00
Purchased Price of Filter-Aid ($/Ton) 50.00
Purchased Price of Natural Gas ($/MSCF) 0.50
Purchased Price of Electricity (Mils/KWHR) 8.00
Purchased Price of Steam ($/MLB) 0.50
Purchased Price of Cooling Water ($/MGal) 0.02
Purchased Price of Process. Water ($/MGal) 0.20
Purchased Price of Fuel Oil ($/MMBtu) 0.80
Sulfur Credit ($/LT) 5.00
Unit Cost of Solid Disposal ($/Ton) 1.00
Average Hourly Wages Per Gulf Coast ($/Hr) 7.00
Interest on Capital During Construction (%) 12.00
127
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TABLE 6.2
WELLMAN/ALLIED PROCESS AND COST MODEL
SUMMARY OF EQUATIONS
Capital Cost Model
NA
EA =
RB [726 (GT/550)0'5 + 639 (GT/550) °' 9]n + 119 RP (GP/3300)0'
N7
ES
EP
ER =
+ [l33 (S7/7)0'5 + 127 IF (Sl/l)°'6~\l
[209 (S7/7)0'5 + 618 (S7/7)0'6 + 157 (S7/7)°'9jN7
[525 (S28/28)0'5 + 380 (S28/28)0'6 + 86 (S28/28)0'7
+ 306 (S28/28)0-8 + 519 (S28/28)°'9 ] N28
998 (SF/28)0'5 + 287 (SF/28)0'6 + 683 (SF/28)°'!
M = 0.429 EA + 0.742 ES + 0.827 EP + 0.772 ER
L = 0.224 EA + 0.310 ES -I- 0.433 EP + 0.623 ER
BARC =1.15 (E+M) + 1.43 L'F
TPI =1.12 (1.0 + CONTIN) BARC
TCR =1.15 TPI + 0.8 TO-CO (1+F) +0.4 ANR
$M
$M
$M
$M
$M
$M
$M
$M
$M
Operating Cost Model
AS = 28.2 CS-LF (SF/28)
AN = 1460 CN-LF (SF/28)
AFA = 1.24 CFA-LF'IF (GP/3300)
AE = [l54 (GP/3300) + 79 (SF/28)] CE-LF
AH
= 5430 CH-LF (SF/28)
ACW = ^856 (GP/3300) + 19,900 (SF/28)] CCW-LF
AW
64 (SF/28) CW-LF
$M/yr.
$M/yr.
$M/yr.
$M/yr.
$M/yr,
$M/yr.
$M/yr,
128
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TABLE 6.2 (Cont'd)
AF = 1,800 (GP/3300) CF-LF $M/yr.
ASC =95.4 (SF/28) VSC-LF $M/yr.
APS =37.3 (SF/28) VPS'LF $M/yr.
ANR = AS + AAO -I- AN+ AFA + AE + AH + ACW
+ AW -I- AF + ASC + APS $M/yr.
AOC = 0.078 TPI + ' 2-TO-CO (1+F) + ANR $M/yr.
TAG = 0.237 TPI + 2.1-TO-CO (1+F) +1.04 ANR $M/yr.
129
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FIGURE 6.1
WELLMAN/ALLIED PROCESS FLOWSHEET
I API
AREA I
FLUE GAS
REHEAT
FLUE GAS
H2O-
FLUE GAS
COMPRESSION
FLUE GAS
TO STACK
AREA III
1
MAKE-UP
SYSTEM
PRESCRUBBING
AND
S02 REMOVAL
SULFITE1 SOL'N
EVAPORATION
AND
CRYSTALLIZATION
FLY ASH
SLURRY
AREA I -ABSORBER
AREA II -S02 REGENERATION
AREA III - PURGE/MAKE-UP
AREA IV - SO2 REDUCTION
| AREA
J L
VENT GAS
TO ABSORBER
S02
SO2 PURIFICATION
(CONDENSATION/
STRIPPING)
CONDENSATE
I ARI
AREA IV
PURGE SYSTEM
(CRYSTALLIZATION
AND DRYING)
T
AREA II
1
S02
NATURAL
GAS
S02
REDUCTION
TAIL GAS I
TO ABSORBER I
PURGE
SOLIDS
SULFUR
-------
FIGURE 6.2
EFFECT OF BOILER CAPACITY ON TOTAL CAPITAL REQUIREMENT
WELLMAN/ALLIED PROCESS
OJ
I
oc
o
o.
O
400.
300.
200.
30.
20.
COAL SULPUR CONTENT
10
80% BOILER LOAD FACTOR
40% BOILER LOAD FACTOR
I I i
I
20
40
60 80 100
BOILER CAPACITY, MW
200
400
600 800 1000
BASIS OF CALCULATION: NO RETROFIT, NO CONTINGENCY,
U.S. GULF COAST LOCATION, END OF 1973 FIGURE, BOILER
HEAT RATE 9,500 BTU/KWH, HEATING VALUE OF COAL
(HHV) 11,000 BTU/LB.
-------
FIGURE 6.3
Ul
111
U
W
O
U
O
O
Q
O
fC
400.
300-
200.
40
30.
20.
10.
10
EFFECT dF BOILER CAPACITY ON PRODUCTION COST
WELLMAN/ALLIED PROCESS
80% BOILER LOAD FACTOR
40% BOILER LOAD FACTOR
I
20
40
60 80 100 200
BOILER CAPACITY, MW
400
600 800 1000
(SEE FIGURE 6.2 FOR BASIS OF CALCULATION)
-------
FIGURE 6.4
EFFECT OF BOILER RETROFIT DIFFICULTY ON TOTAL CAPITAL REQUIREMENT
WE LLM AN/ALL I ED PROCESS
400.
300.
200.
LU
LU
cr
o
LU
cc
a.
o
100.
90.
80.
70.
60.
50.
40-
BOILER CAPAC|TV_
10 MW
50 MW
100 MW
200 MW
400 MW
600 MW
1000 MW
1.0
1.2
1.4 1.6
RETROFIT FACTOR
1.8
2.0
BASIS OF CALCULATION: 4% SULFUR COAL, HEATING
VALUE OF 11,000 BTU/LB, INDIVIDUAL BOILER WITH
HEAT RATE OF 9,500 BTU/KWH AND LOAD FACTOR
OF 0.7, CONTINGENCY 10%, U.S. GULF COAST LOCATION,
END OF 1973 FIGURE.
133
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FIGURE 6.5
EFFECT OF BOILER RETROFIT DIFFICULTY ON PRODUCTION COST
WELLMAN/ALLIED PROCESS
200
Q
HI
z
a:
O
o
o
t-
12
5
o
o
o
£
20
1.0
10 MW
1.2
1.4 1.6
RETROFIT FACTOR
1.8
50 MW
100 MW
200 MW
400 MW
600 MW
800 MW
2.0
(SEE FIGURE 6.4 FOR BASIS OF CALCULATION.!
134
-------
FIGURE 6.6
EFFECT OF LOCATION FACTOR ON TOTAL CAPITAL RBJ1UIREMENT
WELLMAN/ALLIED PROCESS
300
z
LU
a
UJ
tr
t
o
_l
<
200
100
90
80
70
60
50
40
BOILER CAPACITY_
10 MW
50 MW
100 MW
200 MW
400 MW
600 MW
1000 MW
1.0
1.2
1.4 1.6
LOCATION FACTOR
1.8
2.0
BASIS OF CALCULATION: 4% SULFUR COAL, HEATING
VALUE OF 11,000 BTU/LB, INDIVIDUAL BOILER WITH
HEAT RATE OF 9.500 BTU/KWH AND LOAD FACTOR
OF 0.7, CONTINGENCY OF 10%, END OF 1973 FIGURE.
135
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7. APPLICATION OF STACK GAS SCRUBBING MODELS
7.1 Stack Gas Scrubbing Applied to Existing Utilities
,j
One of the main concerns in this study was to investigate the
cost of retrofitting stack gas scrubbing units, which have been
described in the previous sections, to existing utility boilers.
This exercise was carried out for the existing utilities in 1971
using the data from the statistics discussed in section 3. The
cost of retrofitting stack gas scrubbing units was analyzed and
averaged for the different plant sizes on a national basis.
A program was written first to analyze the existing data. These
data include plant size, the types and amounts of fuel burned in
the plant, the heating value of the respective fuels, the
number of boilers, boiler sizes, etc. If part of these data
was missing, the plant data were upgraded by the program using
the average statistics for that size plant. Table 7.1 presents
part of these statistics. The number of boilers, the percent of
the plant capacity attributable to the largest boiler size, the
percent attributable to the second largest boiler size etc., for
various size ranges of utility plants are given. The average number
of boilers per plant is fairly constant and is equal to four for
all cases studied except plants below 50 megawatts size. The
figures presented are the averages for U. S. utilities. There
may be some exceptional cases where the fifth and sixth boilers
are still significant.
Figure 7.1 represents a highly realistic situation in which the
largest boiler is base loaded and the load factor decreases with
size. The smaller boilers are used for peaking only. The load
factor is defined as follows:
LOAD FACTOR = Total Yearly Generation
Maximum Capacity x 8760 hours/yr
If the load factor for the largest boiler is unity, the load factor
136
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for the second largest boiler is about 0.75. The load factor for
the third boiler is 0.55 and the remaining boiler or boilers
are below 0.15. The procedure for calculating individual boiler
load factor has been set up to give the exact load factor for the
plant.
Figure 7.2 shows the average heat rate of a boiler versus
boiler size. The smaller boilers are generally the older ones
and consequently less efficient than the newer, larger boilers.
In this exercise an average heat rate of 13,000 Btu/kwh, representing
an overall cycle efficiency of 26%, has been taken as the upper
limit, and an average heat rate 6f 9,000 Btu/kwh, representing
an overall cycle efficiency of 38%, as the lower limit. Any
value outside this range is viewed as an error in the data collected
and has been adjusted before the actual calculations. Using
the boiler load factor and the boiler heat rate, the fuel demand
for the boilers can be calculated for the utilities. Fuel is
then allocated in the order of coal, oil and gas starting with
the largest boiler and working toward the sjnallest boiler. The next
step is to investigate the overall plant SO- emission. If the overall
plant SO_ emission is above the specified level of 1.2 Ib/MMBtu,
a stack gas scrubbing unit (either the Wet Limestone or the Wellman/
Allied) is fitted first to the largest sulfur emitting boiler, then
the second largest, then the third largest etc. until the overall
plant SO- emission is below the emission level specified. The cost
for installing stack gas scrubbing units described in this manner
should be fairly realistic and represents the minimum cost in terms
of $/KW of plant capacity. Figures 7.3 to 7.14 present graphically
the results from this exercise.
Figures 7.3 and 7.9 show the total capital required for installing
Wet Limestone and Wellman/Allied stack gas scrubbing units in
existing utilities of various sizes. Cost differences between the
two processes are quite small and well within the order of accuracy
of the models.
137
-------
Figures 7.4 and 7.10 present the total capital required, expressed
as $/KW of plant capacity, for the two processes. The cost in
terms of $/KW of plant capacity increases gradually with decreasing
plant size from 2,000 megawatts to 100 megawatts but rises sharply
below 100 megawatts plant size.
Figures 7.5 and 7.11 show the estimated incremental cost of
electricity delivered in mils/kwh for installing stack gas scrub-
bing in the existing utilities. For the two processes studied,
the incremental cost varies from about 5 mils/kwh to 1.5 mils/kwh
for plant sizes of 100 megawatts to 1500 megawatts. Below 100
megawatts size, the incremental cost rises sharply with decrease
in size. It would be practical to require utility plants to be
fitted with Wet Limestone or Wellman/Allied process stack gas
scrubbing units if the incremental cost of electricity delivered
is less than 4 mils/kwh. Above this value, other alternatives,
such as burning clean fuel, should be investigated if the control
on SO,, emissions is to be imposed.
Figures 7.6 and 7.12 show the total cummulative demand for clean
fuel versus the incremental cost which could be paid for the fuel
as an alternate to stack gas scrubbing, and the range of plant
sizes in which the clean fuel would be burned. For both processes,
the conclusions are about the same. If clean fuel is available
at an incremental cost below $0.30/MMBtu, there is a potential
Q
market of 6 x 10 MMBtu/year. However, if the clean fuel is
available at an incremental cost above $2.50/MMBtu, the potential
Q
market decreases to 1.50 x 10 MMBtu/year, corresponding to a
reduction of 97.5%.
Figures 7.7, 7.8, 7.13, 7.14 show the relationship between the
average cost of stack gas scrubbing for the existing utilities
(starting with the largest plants) and the cumulative percent of
total U.S. capacity under control. For the Wet Limestone process, the
average cost for controlling 10% of the U.S. capacity to meet the
emission standard of 1.2 Ib S02 per MMBtu of fuels burned is $40/KW
of plant capacity and 1.5 mil/Kwh of electricity delivered. To
138
-------
control 100% of the U.S. capacity, the average cost increases to
$64/KW of plant capacity and 2.8 mils/Kwh of electricity delivered.
For the Wellman/Allied process, the corresponding figures are $42/KW
and 1.7 mils/Kwh for controlling 10% of the U.S. capacity; $68/KW
and 3.0 mils/Kwh for controlling 100% of capacity to meet the
emission standard.
It must be stressed that the figures presented here are not the
cost for controlling the total plant capacity but rather controlling
enough boilers to meet the specified SO~ emission level. The location
factor, which is described in the General Cost Model (section 4),
is incorporated into the calculations and the figures are the average
costs for the U.S. utilities on a national basis. The overall
conclusion is that it is economically preferable to install stack
gas scrubbing units on larger size utility plants while for small
size utility plants (below 100 megawatts) the better alternative is
to burn low sulfur fuel.
7.2 Stack Gas Scrubbing Applied to Industrial Boilers
An investigation was made to determine the costs of fitting stack
gas scrubbing processes to coal- and oil-fired industrial boilers
in the United States. As an initial phase, the process and cost
models for the Wet Limestone and Wellman/Allied processes were
reviewed to determine their applicability to industrial boilers.
Of particular interest were the smaller boilers, since these
represent a large extrapolation of the models from the type of
application for which they were initially developed, viz., large
utility boilers.
As a result of this review, a number of changes were incorporated
into the Wet Limestone and Wellman/Allied models. These changes
are briefly discussed below.
7.2.1 Wet-Limestone Process
A review of the Wet Limestone model prompted the following
139
-------
changes:
1) Replacement of the sludge pond with a thickener and
temporary sludge disposal pit.
2) Elimination of onsite limestone grinding at small
limestone design rates (low sulfur flows).
3) Reduction of some of the scrubbing equipment costs
for small boiler sizes.
It was felt that, for industrial boiler applications, a thickener
circuit for sludge handling would be more universally applicable
than a large sludge pond. The sludge would be periodically
hauled offsite and disposal treated as an operating cost.
Grinding of limestone becomes increasingly expensive as the
limestone design rate decreases, and it was assumed that at
sulfur flows of less than 2000 Ibs/hr, grinding would be
eliminated in favor of purchasing pulverized stone. Costs
for some of the scrubbing equipment were found to be high for
boiler sizes less than about 400 MM BTU/hr and were reduced
accordingly.
Table 7.2 summarizes the changes made to the equipment cost
portion of the model. The equation for chemical process
equipment costs, EC, now includes the factor FC which reduces
the cost of some of the scrubbing equipment for small boilers.
This factor varies with boiler capacity as shown, ranging in
value from 1.0 to 2.25. The term ISF is an index used to delete
the grinding equipment costs when the sulfur rate falls below
2000 Ibs/hr. P now represents the cost of the small temporary
storage pit for limestone sludge.
The raw materials and utilities cost equations are presented in
Table 7.3. The first equation represents the annual cost of
sludge disposal, ASL, in terms of the unit cost, CSL, in $/ton.
The equation for the annual cost of limestone, AL, remains
\
140
-------
unchanged in form. However, when pulverized limestone is pur-
chased (i.e., when grinding is eliminated), the numerical value
of the unit cost of limestone, CL, would be increased appropriately.
The last equation in the table gives electric power costs and
now includes a term which reduces the process power consumption
when limestone grinding is eliminated.
In Figures 7.15-7.18 typical costs which were calculated from
the cost model have been plotted to illustrate the effects
of different variables on the Wet-Limestone stack gas scrubbing
costs. The figures have been separated into 10-100 MM Btu/hr
and 100-1000 MM Btu/hr size ranges because of difficulty in
scaling.
Figures 7.15 and 7.16 illustrate the effect of boiler capacity
on the total capital required (TCR) with load factor and sulfur
content of coal shown as parameters. For small boilers (10-
100 MM Btu/hr), the load factor and percent sulfur have insignificant
effects on capital required. The effect becomes noticeable
when the boiler size becomes larger. For a boiler size of 1000
MM Btu/hr, doubling the sulfur content in coal (from 2% to
4% or from 4% to 8%) increases the capital required by approximately
10%. The effect of load factor remains minor.
Figure 7.17 and 7.18 illustrate the pronounced effects of
sulfur content and load factor on the operating cost (TAG).
Generally speaking, doubling the sulfur content in coal in-
creases the operating cost by 10% for the small boilers (10-50
MM Btu/hr) and the percent gradually increases as boiler capacity
increases, to as much as 20% for 1000 MM Btu/hr. The operating
cost can be increased by as much as 100% for smaller boilers
(10 MM Btu/hr) when the load factor is reduced to half (0.8
to 0.4) and as much as 60% for larger boilers (1000 MM Btu/hr).
7.2.2 Wellman/Allied
A review of the Wellman/Allied model suggested that two changes
141
-------
could be made for application to industrial boilers. First,
it was found that the predicted costs of some of the scrubbing
equipment were high for small boiler sizes, as in the Wet
Limestone model. It was also found that the same adjustment
factor used in the Wet Limestone model could be used in the
Wellman/Allied model.
The second change would affect the SO- regeneration and sulfur
recovery areas. Most industrial boilers operate at a fairly
low load factor, indicating a significant variation in their
operating rate throughout the year. Since their operation is
tied exclusively to a particular plant or plant site, this
variation in operating rate might be more reliably predicted
than for a utility plant which is tied into a grid system.
Consequently, it might be possible for industrial boiler
applications to size the regeneration and recovery areas for
somewhat less than peak sulfur load by providing adequate surge
capacity between the absorber and regeneration plant. For
purposes of illustrating the effect of this type of design
on costs, a design point 25% above the average sulfur flow
has been assumed adequate for the regeneration plant. Surge
capacity has been increased accordingly.
The resultant equipment cost equations are summarized in Table
7.6. The first equation now includes the term, FC, to reduce
some of the scrubbing equipment costs for small boiler applica-
tions. The primed variables (S71, S281, SF') in the equations
reduce the size and cost of the regeneration and sulfur recovery
plant and are related in the same manner as the unprimed
variables (see Section 6).
In Figures 7.19-7.22 typical costs which were calculated from
the cost model have been plotted to illustrate the effects of
different variables on the Wellman/Allied stack gas scrubbing
process. The figures have been separated into 10-100 MM Btu/hr
and 100-2500 MM Btu/hr size ranges.
142
-------
Figures 7.19 and Figures 7.20 illustrate the effect of boiler
capacity on the total capital required (TCR), with percent
sulfur and load factor as parameters. To some extent, the
load factor and percent sulfur have a larger effect on the
capital required in the Wellman/Allied system than in the
Wet Limestone process. Generally speaking, doubling the
percent sulfur in coal or decreasing the load factor to half
increases the total capital required by 20% to 40%; the effect
is more pronounced for smaller boilers (10-100 MM Btu/hr)
than for large boilers (100-2500 MM Btu/hr) .
Figures 7.21 and 7.22 illustrate the effect of load factor and
percent sulfur in coal on the operating costs (TAG). It
can be seen from these figures that doubling the sulfur percent
in coal increases the operating cost by about 30% for large
boilers (100-2500 MM Btu/hr) and by as much as 50% for small
boilers. The effect of load factor is more pronounced than
percent sulfur with the increase in operating cost, for decreasing
the load factor to half, ranging from 50% to 70%. Again the
effect is more pronounced for small boilers (10-100 MM Btu/hr).
7.2.3 Applicability to Small Industrial Boilers
Coal- and oil-fired industrial boilers in the United States
number more than 5,000. Based on the statistical analysis of
the available boiler population, these range in capacity up
to 4400 MM Btu/hr. The small boilers, i.e., those with a
capacity of 100 MM Btu/hr or less, represent almost three-fourths
of the total population. However, these boilers emit less than
one-fourth of the total sulfur emissions from all coal- and
oil-fired industrial boilers.
Figures 7.19-7.22, which show typical costs of stack gas scrub-
bing, indicate that the costs incurred by small size industrial
boilers (<_ 100 MM Btu/hr) are very high. Depending on boiler
size, clean fuel at incremental prices of roughly $l-3/MM Btu
(or less) would be preferable to scrubbing as a control method.
143
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However, considering the limited emissions from these small
boilers, it would be difficult to justify either type of control
unless there were particularly good local reasons.
The preceding discussion is based on single boiler installations
There were no data available to permit estimation of costs
on a plant basis (i.e., considering the total number of boilers
per plant). This type of analysis should be done to obtain
more meaningful costs.
144
-------
TABLE 7.1 BOILER SIZE DISTRIBUTION FOR
STANDARD SIZE UTILITY PLANT
PLANT SIZE
AVERAGE
NO. OF BOILERS
AVERAGE DISTRIBUTION-PLANT GENERATION CAPACITY, %
Megawatts
0 to 10
11 to 50
51 to 100
101 to 200
201 to 300
:H 301 tO 400
£ 401 to 500
501 to 600
601 to 700
701 to 800
801 to 900
901 to 1000
1001 to 1200
1201 to 1600
1401 to 1600
Over 1600
1
3
4
4
4
4
4
4
4
4
4
4
4
4
4
4
BOILER 1
100.0%
65.0
58.0
57.0
53.0
52.0
53.0
58.0
56.0
49.0
58.0
52 .0
47.0
46.0
45.0
41.0
BOILER 2
26.0
24.0
25.0
27.0
27.0
26.0
23.0
27.0
34.0
25.0
28.0
30.0
35.0
28.0
38.0
BOILER 3
9.0
12.0
11.0
12.0
13.0
14.0
12.0
11.0
12.0
9.0
14.0
13.0
10.0
20.0 .
13.0
BOILER 4
-
6.0
7.0
8.0
8.0
7.0
7.0
6.0
5.0
7.0
6.0
10.0
9.0
7.0
8.0
-------
TABLE 7.2
WET LIMESTONE PROCESS
SUMMARY OF EQUIPMENT COST EQUATIONS FOR INDUSTRIAL BOILERS
NA
EC = ^>~ RB [(629 + p^-) (GT/550)0'5 + (200 + ^-) (GT/550) °'9]n
n=l
+ 238 RB (GP/3300)0'5 + (201-7ISF) (SF/28)0'5 SM
ES = (1680-1180 ISF) (SF/28)0'9 + 120 (SF/28)0'7 $M
P = 40 (SF/28)0'5 $M
FC = 2.25-0.003 CAP (FC> 1.0)
where CAP = boiler capacity, MM BTU/hr input
ISF = 1 when SF < 2
ISF = 0 when SF > 2
146
-------
TABLE 7.3
WET LIMESTONE PROCESS
SUMMARY OF OPERATING COST EQUATIONS FOR INDUSTRIAL BOILERS
RAW MATERIALS AND UTILITIES
ASL = 2210-CSL-LF-(SF/28) $M/year
AL = 600-CL-LF-(SF/28) $M/year
AA = 0.43-CA-(SF/28) $M/year
AF = 1800-CF-LF-(GP/3300) $M/year
AW = 230-CW.LF-[(GP/3300) + (SF/28)] $M/year
AE = CE.LF-[213(GP/3300) + (35-23.5 ISF) (SF/28)] $M/year
147
-------
TABLE 7.4
WELLMAN/ALLIED PROCESS
SUMMARY OF EQUIPMENT COST EQUATIONS FOR INDUSTRIAL BOILERS
NA
EA = \ RB f(407 + |~) (GT/550)0'5 + (200 + ^-) (GT/550) ° '
L FC FC
FC = 2.25 - 0.003 CAP (FC> 1.0)
where CAP = boiler capacity, MM BTU/hr input
SF1 = SF if LF > 0.8
SF' = 1.25-LF-SF if LF < 0.8
n
n=l
+ 119 RB (GP/3300)0'5
+ [l90 (S7/7)0'5 + 50 (S7'/7)°'5 + 127IF (S7 '
$M
ES = [209 (S7'/7)°'5 + 618 (S7'/7)0'6 + 157 ( S7 ' /7) ° ' 9JN7 $M
EP = [525 (S28'/28)°'5 + 380 (S28'/28)°'6 + 86 (S28'/28)°'7
+ 306 (S28'/28)°*8 + 519 (S28 '/28) ° *9jN2 8 $M
ER = 998 (SF'/28)0'5 + 287 (SF'/28)0'6 + 683 (SF'/28)°<9 $M
148
-------
FIGURE 7.1
AVERAGE DISTRIBUTION OF LOAD FACTORS
FOR BOILERS IN A UTILITY PLANT
LOAD FACTOR OF BOILER N = RATIO X LOAD FACTOR OF BOILER 1
O
oc
•j 1 . V "
OC
UJ
§ 0.9 .
in
U.
0
oc 0.8
O
O
u.
0 0.7 .
O
2 0.6 •
OC
UJ
§ 0.5 •
LL
0
oc
0
S 0.4 .
u.
Q
3 0.3
u.
O
g
£ 0.2 .
oc
0.1
n
1st 2nd 3rd 4th & Up
BOILER NUMBER STARTING WITH THE LARGEST
149
-------
o
<
cc
FIGURE 7.2
AVERAGE HEAT RATES FOR UTILITY BOILERS
14000 1-
13000
12000
11000
10000 . _
9000
8000
100
200
300
400
500
BOILER SIZE, MEGAWATTS
150
-------
FIGURE 7.3
AVERAGE TOTAL CAPITAL REQUIREMENT FOR INSTALLING
WET LIMESTONE SYSTEM IN EXISTING POWER PLANTS
EMISSION STANDARD = 1.2 LBS SO2/MM BTU OF FUEL BURNED
V)
5
UJ
UJ
-------
O
20
FIGURE 7.4
AVERAGE UNIT COST FOR INSTALLING WET
LIMESTONE SYSTEM IN EXISTING POWER PLANTS
EMISSION STANDARD = 1.2 LBS SOj/MM BTU OF FUEL BURNED
(/I
<
Q
HI
c/>
X
o.
X
ai
uj
QC
a
a.
<
O
300
200
100
90
80
70
60
50.
40 .
30
-------
FIGURE 7.5
INCREMENTAL OPERATING COST FOR WET
LIMESTONE SYSTEM IN EXISTING POWER PLANTS
EMISSION STANDARD = 1.2 LBS SO2/MM BTU OF FUEL BURNED
I
8
o
<
z
cc
u
10
20
40 60 80 100 200
PLANT SIZE, MW
400 600 8001000
2000
153
-------
FIGURE 7.6
WET
DEMAND FOR CLEAN FUEL AS ALTERNATIVE
TO STACK GAS SCRUBBING
LIMESTONE SYSTEM APPLIED TO EXISTING POWER PLANTS
6000 ,_
5000 . .
4000 . .
3000 • -
2000 . _
1000 . _
-r 1500
500
400
300
200
100
50
50
100
150
200
250
EQUIVALENT INCREMENTAL COST WHICH COULD BE PAID FOR CLEAN FUEL, CENTS/MM BTU
154
-------
I
m
CD
3
DC
U
in
U
0.
O
UJ
O
OC
I
5
u
FIGURE 7.7
CUMULATIVE AVERAGE CAPITAL COST
WET LIMESTONE SYSTEM APPLIED TO POWER PLANTS
EMISSION STANDARD = 1.2 LBS SO2/MM BTU OF FUEL BURNED
100
90
80
70
60
50
40
30
20
10
NOTE: CONTROL ASSUMED TO
BEGIN WITH LARGEST PLANT
10 20 30 40 50 60 70 80 90 100
% OF TOTAL U.S. CAPACITY UNDER CONTROL
155
-------
O
m
CD
D
(C
O
v>
tr
o
k
o
o
CC
111
a.
O
uj
UJ
1C
o
UJ
>
FIGURE 7.8
CUMULATIVE INCREMENTAL OPERATING COST
WET LIMESTONE SYSTEM APPLIED TO POWER PLANTS
EMISSION STANDARD = 1.2 LBS SO2/MM 8TU OF FUEL BURNED
NOTE: CONTROL ASSUMED TO
BEGIN WITH LARGEST PLANT
5
O
I
10
20 30 40 50 60 70 80 90
% OF TOTAL U.S. POWER GENERATION UNDER CONTROL
100
156
-------
FIGURE 7.9
AVERAGE TOTAL CAPITAL REQUIREMENT FOR INSTALLING
WELLMAN/ALLIED SYSTEM IN EXISTING POWER PLANTS
EMISSION STANDARD = 1.2 LBS SO2/MM BTU OF FUEL BURNED
UJ
LU
CC
a
<
o
UJ
CJ
-------
FIGURE 7.10
AVERAGE UNIT COST FOR INSTALLING WELLMAN/ALLIED
SYSTEM IN EXISTING POWER PLANTS
EMISSION STANDARD = 1.2 LBS SO2/MM BTU OF FUEL BURNED
«/>
in
UJ
1C
Q.
X
Ul
oc
a
UJ
oc
u
a
<
oc
UJ
10
10
20
40 60 80 100 200
PLANT SIZE, MW
400 600 800 1000
2000
158
-------
FIGURE 7.11
INCREMENTAL OPERATING COST FOR
WELLMAN/ALLIED SYSTEM IN EXISTING POWER PLANTS
I
o
UJ
UJ
IT
O
100
90
80
70
60
50
40
30
20
10
9
8
7
6
5
4
i iii i i
j_
i i i i i
10
20
40 60 80 100 200
PLANT SIZE, MW
480 600 800>1000
2000
159
-------
FIGURE 7.12
DEMAND FOR CLEAN FUEL AS ALTERNATIVE
TO STACK GAS SCRUBBING
WELLMAN/ALLIED SYSTEM APPLIED TO EXISTING POWER PLANTS
CD
£C
o
UJ
>
O
50 100 150 200
EQUIVALENT INCREMENTAL COST WHICH COULD BE PAID FOR CLEAN FUEL, CENTS/MM BTU
25O
300
160
-------
o
o
00
D
1C
o
o
8
o
Q.
u
UJ
o
X
UJ
D
O
FIGURE 7.13
CUMULATIVE AVERAGE CAPITAL COST
WELLMAN/ALLIED SYSTEM APPLIED TO POWER PLANTS
100
90
80
70
60
50
40
30
20
10
EMISSION STANDARD = 1.2 LBS SO2/MM BTU OF FUEL BURNED
NOTE: CONTROL ASSUMED TO
BEGIN WITH LARGEST PLANT
10 20 30 40 50 60 70 80 90 100
% OF TOTAL U.S. CAPACITY UNDER CONTROL
161
-------
FIGURE 7.14
CUMULATIVE INCREMENTAL OPERATING COST
WELLMAN/ALLIED SYSTEM APPLIED TO POWER PLANTS
EMISSION STANDARD = 1.2 LBS SO2/MM BTU OF FUEL BURNED
400
1
= 300
m
00
CC
cc
o
p 200
0.
O
LU
DC
O
>
tE
2
O
100
NOTE: CONTROL ASSUMED TO
BEGIN WITH LARGEST PLANT
10 20 30 40 50 60 70 80 90
% OF TOTAL U.S. POWER GENERATION UNDER CONTROL
100
162
-------
FIGURE 7.15
EFFECT OF BOILER CAPACITY ON TOTAL CAPITAL REQUIREMENT
WET LIMESTONE PROCESS APPLIED TO LARGE INDUSTRIAL BOILERS
20
1C
D
CD
5
S
in
2
UJ
CC
5
a
ai
(C
a.
<
O
8% SULFUR COAL
4% SULFUR COAL
2% SULFUR COAL
80% LOAD FACTOR
40% LOAD FACTOR
J_
J_
I I I I I III
3 456789 10 15
BOILER CAPACITY. 108 BTU/HR
20 25
BASIS OF CALCULATION: NO RETROFIT, NO CONTINGENCY, U.S.
GULF COAST LOCATION, END OF 1973 FIGURE, HHV OF COAL
11,000 BTU/LB, 40% EXCESS AIR WITH AN AIR LEAKAGE OF 10%.
163
-------
FIGURE 7.16
EFFECT OF BOILER CAPACITY ON TOTAL CAPITAL REQUIREMENT
WET LIMESTONE PROCESS APPLIED TO SMALL INDUSTRIAL BOILERS
oc
I
5
I
a
UJ
CC
a.
<
O
O
60
50
40
30
20
10
9
8
7
6
8% SULFUR COAL
4% SULFUR COAL
2% SULFUR COAL
80% LOAD FACTOR
40% LOAD FACTOR
2 3456
BOILER CAPACITY. 10? BTU/HR
7 8 9 10
(SEE FIGURE 7.15 FOR BASIS OF CALCULATION.)
164
-------
FIGURE 7.17
EFFECT OF BOILER CAPACITY ON OPERATING COST
WET LIMESTONE PROCESS APPLIED TO LARGE INDUSTRIAL BOILERS
200
D
m
Z
LU
O
CO
o
u
o
\
tr
UJ
Q.
O
20
10
80% LOAD FACTOR
40% LOAD FACTOR
3 4 5 6 7 8 9 10
BOILER CAPACITY, 108 BTU/HR
8% SULFUR COAL
4% SULFUR COAL
2% SULFUR COAL
8% SULFUR COAL
4% SULFUR COAL
2% SULFUR COAL
20
30
(SEE FIGURE 7.15 FOR BASIS OF CALCULATION.)
165
-------
FIGURE 7.18
EFFECT OF BOILER CAPACITY ON OPERATING COST
WET LIMESTONE PROCESS APPLIED TO SMALL INDUSTRIAL BOILERS
CO
5
55
UJ
O
O
O
tc.
UI
400
300
200
80% LOAD FACTOR
40% LOAD FACTOR
20
10
8% SULFUR COAL
4% SULFUR COAL
2% SULFUR COAL
•^ 8% SULFUR COAL
4% SULFUR COAL
2% SULFUR COAL
2 3 4 56789 10
BOILER CAPACITY, 10? BTU/HR
(SEE FIGURE 7.15 FOR BASIS OF CALCUATION.)
166
-------
FIGURE 7.19
EFFECT OF BOILER CAPACITY ON TOTAL CAPITAL REQUIREMENT
WELLMAN/ALLIED PROCESS APPLIED TO LARGE INDUSTRIAL BOILERS
en
00
5
5
I-
2
oc
OL
O
50.
40.
80% LOAD FACTOR
40% LOAD FACTOR
3 4 5 & 7 8 9 10
BOILER CAPACITY, 108 BTU/HR
8% SULFUR COAL
8% SULFUR COAL
4% SULFUR COAL
4% SULFUR COAL
2% SULFUR COAL
2% SULFUR COAL
20
30
BASIS OF CALCULATION: NO RETROFIT, NO CONTINGENCY, U.S.
GULF COAST LOACTION, END OF 1973 FIGURE, HHV OF COAL
11,000 BTU/LB, 40% EXCESS AIR WITH AN AIR LEAKAGE OF 10%.
167
-------
FIGURE 7.20
EFFECT OF BOILER CAPACITY ON TOTAL CAPITAL REQUIREMENT
WELLMAIM/ALLIED PROCESS APPLIED TO SMALL INDUSTRIAL BOILERS
DC
I
5
m
a
o
_l
<
o
8% SULFUR COAL
4% SULFUR COAL
4% SULFUR COAL
4% SULFUR COAL
2% SULFUR COAL
2% SULFUR COAL
80% LOAD FACTOR
40% LOAD FACTOR
2 3 4 56789 10
BOILER CAPACITY, 10? BTU/HR
(SEE FIGURE 7.19 FOR BASIS OF CALCULATION.)
168
-------
FIGURE 7.21
EFFECT OF BOILER CAPACITY ON OPERATING COST
WELLMAN/ALLIED PROCESS APPLIED TO LARGE INDUSTRIAL BOILERS
CD
5
c/5
z
UJ
U
§
o
o
z
UJ
300
200
30
20
10
8% SULFUR COAL
4% SULFUR COAL
8% SULFUR COAL
2% SULFUR COAL
4% SULFUR COAL
2% SULFUR COAL
80% LOAD FACTOR
40% LOAD FACTOR
3 4 5 6789 10
BOILER CAPACITY, 1Q8 BTU/HR
15
20 25 30
(SEE FIGURE 7.19 FOR BASIS OF CALCULATION.)
169
-------
FIGURE 7.22
EFFECT OF BOILER CAPACITY ON OPERATING COST
WELLMAN/ALLIED PROCESS APPLIED TO SMALL INDUSTRIAL BOILERS
UJ
O
O
O
cc
UJ
s
80% LOAD FACTOR
40% LOAD FACTOR
30
20
8% SULFUR COAL
4% SULFUR COAL
8% SULFUR COAL
2% SULFUR COAL
4% SULFUR COAL
2% SULFUR COAL
2 3 4 56789 10
BOILER CAPACITY, 107 BTU/HR
(SEE FIGURE 7.19 FOR BASIS OF CALCULATION.)
170
-------
8. SUBSTITUTE NATURAL GAS PRODUCTION USING A LURGI OXYGEN GASIFIER
8.1 Process Model
9
The model represents a plant producing 250 x 10 Btu/day of pipeline
quality gas having a higher heating value of 970 Btu/SCF. This is
generally considered the standard size plant. At this stage there
is little economic incentive to increase this size for two good
reasons: firstly, most of the equipment already consists of
several parallel trains, and secondly, because of the difficulties
in financing projects well in excess of $300 million.
So the only variations in the size of the units making up the plant
will be produced by different types of coal feed.
8.1.1 Coal Types
There are four basic coal types: lignite, subbituminous,
bituminous and anthracite. All four types have been used as
feed to Lurgi gasifiers. Anthracite can be dismissed as a
feedstock for SNG because of its scarcity. There are, however,
large deposits of lignite, subbituminous and bituminous coal
in the U.S.
These four coal types can best be categorized by the dry ash
free carbon content. Here, dry means free from all water,
not only surface moisture.
% carbon in
dry ash free coal (PCARB)
lignite 65-73
subbituminous 73-77
bituminous 77 - 91
anthracite 91 - 96
171
-------
A mid-range composition for each type is:
% Dry Ash
free basis (DAFB) Lignite Subbituminous Bituminous Anthracite
C 69 75
H 55
O 24 18
N 11
S 11
84
5.5
7
1.5
2
93
3
2
1
1
100 100 100 100
These coals, when received from the mine, will have the follow-
ing range of water and ash content.
% of coal
as received Lignite Subbituminous Bituminous Anthracite
Water (PH2O) 28-40 16-28 4-16 2-4
Ash (PASH) 4- 8 5-12 6-14 6-18
Both the water and the ash content can vary a few percent up
or down throughout the same coal field.
Figure 8.1 gives the higher heating values in Btu/lb of dry
ash free coal against the percentage carbon on a dry ash free
basis. This graph has been derived from information on several
coals of each type (23, 24, 25).
The sulfur contents of the lignites and Subbituminous coals
are generally low, about 1% d.a.f.b.; however, the sulfur
contents of the bituminous coals can be as high as 8% d.a.f.b.,
although they are usually less than 4% d.a.f.b.
8.1.2 Coal, Oxygen and Steam Requirements for the 5NG Plant
Figure 8.2 shows the dry ash free coal, oxygen and steam
requirements in million Ib/hr for a Lurgi oxygen gasification
172
-------
q
plant producing 250 x 10 Btu/day of pipeline quality gas of
higher heating value 970 Btu/SCF. These graphs have been
produced from available designs and published information
(26, 27, 28). They are, of course, simplified linear representa-
tions of the real situation. The difference between the coal
feed to the gasifier and the total coal requirement of the
plant is the fuel to the furnace producing HP steam and power.
8.1.3 Electric Power and High Pressure Steam Requirements for
the SNG Plant
The following bases were used to establish the electric
power and high pressure steam requirements:
1. In this model no major drivers are powered by electri-
city. The total electric power requirement for in-
struments , small drivers not powered by steam and
lighting is about 68 megawatts.
2. The SNG compressor is powered by low pressure steam
generated by the gasifier and waste heat recovery.
This also covers all other low pressure process
steam requirements and a few small drivers.
3. The air and oxygen compressors are powered by HP
steam with vacuum condensation (1100 psig, 825°F->-3
psia requires 10 Ib/hr of steam/KWH). A plant pro-
ducing 0.5 million Ib/hr of oxygen requires 116
megawatts.
4. The Lurgi process units require 0.3 million Ib/hr
of HP steam.
5. The methanator generates about 1.4 million Ib/hr
of HP steam.
173
-------
6. The 550 psig steam for the gasifier is provided by
expanding 1100 psig steam and generating some of the
68 megawatts power requirement (50 Ib/hr of steam/KWH)
I
7. The rest of the 68 megawatt power is provided by
expanding the 1100 psig steam to 3 psia.
8.1.4 Sample Calculation of Plant Total Coal Requirement
The total coal feed to the SNG plant is calculated below for
two coals: a mid-range lignite (PCARB = 69%) , and a mid-range
bituminous (PCARB = 84%).
LIGNITE BITUMINOUS
megawatts megawatts
Total electric power required 68 68
Power Generated by expanding gasifier
steam 39 50
Net power required 28 18
million Ib/hr million Ib/hr
HP steam required to generate net power 0.28 0.18
HP steam for expansion to gasifier 1.94 2.48
HP steam required for air and oxygen
compressors 0.84 1.39
HP steam required for Lurgi process units 0.30 0 .30
Total HP steam requirement 3.36 4.35
HP steam generated by methanator 1.40 1.40
Net HP steam requirement 1.96 2.95
million Btu/hr million Btu/hr
Furnace Duty 2520 3790
Furnace Liberation 2900 4360
Heat provided by burning tar, etc. 1180 1500
(8% of heat input for lignite and
10% of heat input for bituminous)
Heat provided by coal 1720 2860
174
-------
DAF DAF
million Ib/hr million Ib/hr
Coal required to provide this heat 0.15 0.19
Coal feed to furnace including 0.17 0.21
12% extra for stack gas scrubbing
Coal feed to gasifier 1.28 1.01
Total coal requirement of the plant 1.45 1.22
9
The total dry ash free coal requirement of a 250 x 10 Btu/day
SNG plant is given by:
TDAFC = 1.51 - 0.0156 (PCARB-65) million Ib/hr
9
The total "as received coal" requirement of a 250 x 10 Btu/day
SNG plant is given by:
TCOAL = 100 TDAFC/(100-PH20-PASH) million Ib/hr
8.2 Cost Model
8.2.1 Major Equipment Costs, E
The SNG plant has been considered as 12 units (Fig. 8.3).
Section Solid Handling
Number or Chemical Processing Unit
1 S Coal Preparation and Handling
2 S Fines Agglomeration
3 S Coal Gasification
4 C - Shift Conversion and Gas Cooling
5 C Gas Purification by the Rectisol
Process
6 C Methane Synthesis
7 C SNG Compression
8 C The Oxygen Plant
9 C The Phenosolvan Unit
10 C Furnace Stack Gas Scrubbing and
Plant Sulfur Recovery
11 C Utility Plant
12 C Other Offsites
175
-------
The following equipment costs were developed using in-house and
published data (26, 27) , updated to the end of 1973 and adjusted
to the U.S. Gulf Coast basis. The standard relationships given
in the General Cost Model between major equipment costs, E,
other material costs, M, and direct construction labor costs,
Gulf Coast, L, were used in cross-checking the available
information.
Section 1 - Coal Preparation and Handling
Raw coal from storage is crushed and classified in this section.
The larger size fraction (about 70%) is sent to the gasifiers.
Some of the fines generated during crushing are burned in the
furnace, the remainder are sent to the fines agglomeration
unit. No costs are included for equipment delivering coal
from the mine. Coal is assumed to be delivered to the plant
storage and the delivery costs included in the cost of the coal.
It is also assumed that the ash is removed back to the mine and
the cost of this disposal is included in the cost of the coal.
The plant requires more lignite feed than bituminous coal feed,
however, the lignite crushes more easily. Therefore, it has
been assumed that there is no variation of E with coal type.
El = 2,100 M$
Section 2 - Fines Agglomeration
Variations in coal feedrate to the plant and to the furnace
mean the coal flow to the fines agglomeration unit decreases
as the carbon content of the coal increases. These quantities
were determined, the equipment cost variation calculated as
the 0.6 power of the size and the cost simplified to a linear
equation.
D2 = 5,000 - 100 (PCARB-65) M$
176
-------
Section 3 - Coal Gasification
The number of gasifiers required depends on the quantity of
the coal feed, the slagging properties of the coal and the
reactivity of the coal. Although the coal feed to the gasifier
decreases with the increase in carbon content so does the coal
reactivity, the net effect is that more gasifiers are required
for the highest rank coals. This, of course, is a complicated
effect, which has been simplified as best possible in the
following cost equation.
E3 = 14,800 + 160 (PCARB-65) M$
Section 4 - Shift Conversion and Gas Cooling
In this section, the H2/CO ratio of the crude gas is adjusted
by the shift reaction:
CO + H20 -> CO2 + H2
to 3.0, which is the stoichiometric ratio for methanation.
The crude gas is cooled before the purification unit. No sig-
nificant cost variations could be determined with carbon content.
E4 = 4,500 M$
Section 5 - Gas Purification by the Rectisol Process
This unit removes CO , H_S and naphtha from the gas before
methanation. There is a small increase in cost as the sulfur
content of the coal increases. This variation has been expressed
as a linear equation.
E5 = 13,000 + 200 PSULF M$
PSULF is the percent sulfur in the dry, ash free coal.
177
-------
Section 6 - Methane Synthesis
In this section CO and H in the treated gas are converted to
CH. by the reaction:
CO
The final C02 absorption is also included in this section.
No significant variation in cost could be determined with
carbon content.
E6 = 5,500 M$
Section 7 - SNG Compression
This section is SNG compression for delivery to the pipeline.
The costs include the compressors, steam turbine drivers and
the vacuum condensers .
E7 = 3,000 M$
Seqtion 8 - The Oxygen Plant
This section produces the oxygen feed to the gasifiers. The
oxygen requirements of the plant increase with increasing
carbon content of the coal. The cost variations have been
determined as a function of the 0.6 power of size and expressed
as a linear equation.
E8 = 9,700 + 160 (PCARB-65) M$
Section 9 ^ The Phenosolvan Unit
The unit handles all the gas liquor which has been condensed.
Here the main objective is to remove the water before the
phenols and tars can be routed to the furnace. No significant
178
-------
cost variations can be determined in general terms.
E9 = 1,800 M$
Section 10 - Furnace Stack Gas Scrubbing and Plant Sulfur
Recovery
The section includes a stack gas scrubbing and SO,, regeneration
unit on the furnace if this is required. It contains the
sulfur recovery unit for the whole plant. It has been assumed
that 80% of the sulfur entering the plant emerges as sulfur
by-product. An equation has been derived which contains a
term for the variation of stack gas scrubbing costs. These
were computed as the 0.6 power of size and expressed as a linear
function. This term is effectivley zero when the coal contains
little sulfur. The equation also has a term for the sulfur
recovery unit. Even a coal with a sulfur content of 0.1%
requires a small sulfur recovery unit.
E10 = 1,250 PSULF + 1065 (TDAFC.PSULF)°'6 -250 M$
Section l^L - The Utility Plant
The utility plant supplies the power and HP steam for the
plant. It is made up of 3 areas, the boiler plant, the power
plant and processing of the fuel gas and tar.
The boiler plant increases in size as the carbon content of
the coal increases. The variations in cost were computed as
the 0.8 power of size and expressed as a linear function. The
rest of the unit was assumed to be independent of coal type
and has a major equipment cost of $4,500 M.
Ell = 13,800 + 200 (PCARB-65) M$
179
-------
Section 12 - Other Offsites
Offsites other than the utility plant have been grouped together
as one section. The major items are storage facilities, ser-
vice systems, electrical distribution, sewers and waste disposal,
site preparation, plant buildings and mobile equipment.
No meaningful variation with coal type could be derived.
E12 = 14,000 M$
8.2.2 Total Net Annual Operating Cost
The total net annual operating cost, AOC, is the total cost
of operating the plant less the credits from the sale of by-
products. It does not include return of capital, payment
of interest on debt or income tax on equity returns.
This model conforms exactly to the format in the General Cost
Model, which is fully explained in Section 4.
The total net annual operating cost is, therefore, given by:
AOC = 0.078 TPI + 2 TO-CO (1 + F) + ANR
The total number of shift operators for the SNG plant can be
assumed to be 300.
The annual cost of raw materials less by-product credits
has been simplified and is given by -
ANR = ACOAL + ACHEM - ASULF
The annual cost of catalysts and chemicals, ACHEM, is assumed
constant -
180
-------
Annual Cost
M$
Shift catalyst 40
Methanator catalyst 60
Methanol 500
Isopropyl ether 200
H S04 120
NaOH 400
Activated Carbon 50
Lime 10
Na2C03 20
Process Water 200
ACHEM = 1,600 M$
The annual cost of the coal feed to the plant is given by
ACOAL = CCOAL ' TCOAL x 24 x SD
2,000 x 1,000
where CCOAL is the unit cost of coal as received at the site
in $/ton and SD is the number of days the plant is on stream
per year .
The equation reduces to:
ACOAL = 12 CCOAL . TCOAL . SD M$/Yr .
The credit per year for the sale of sulfur, ASULF, is
given by:
ASULF = CSULF x 0.8 x TDAFC x PSULF x 24 X SD Mc/Yr
2,000 x 100 x 1,000
181
-------
where CSULF is the unit credit for sulfur in $/ton. It has
been assumed that 80% of the sulfur in the coal feed to the
plant is recovered. The equation reduces to
ASULF =0.1 CSULF . TDAFC . PSULF . SD M$/Yr.
8.2.3 Total_Plant Investment, Total Capital Required and
Total Annual Production Cost
This model conforms exactly to the General Cost Model and so
the Total Plant Investment, TPI, at different locations can
be derived from the graph of C vs. F in Section 4.
TPI = C-E
Sections 1, 2 and 3 of the SNG plant are classified as solids
handling and the remaining 9 sections as chemical processing.
The Total Capital Required, TCR, is given by the TCR equation
in the General Cost Model.
TCR ^ 1.21 TPI + 0.8^ TO -CO (1 + F) +0.4 ANR
The Total Annual Production Cost, TAG, is also obtained from
the General Cost Model.
^i
TAG = 0.225 TPI + 2.1 TO-CO (1 + F) + 1.04 ANR
8.2.4 Calculation of Costs for Three Types of Coal in Three
Different Locations
Example 1
Location: New Mexico F = 1.3
182
-------
Coal Details: Subbituminous PCARB = 77%, PSULF = 1%
PH20 = 17%, PASH = 17%
CCOAL = $3/ton
Other Information;
The plant is on stream for 93% of the year, SD = 340 days.
The by-product sulfur credit CSULF = $5/ton. The Gulf
Coast Operating labor costs CO = $7/hour.
Derived Information;
Scale up factor to give TPI obtained from General Cost Model
C = 2.63 solid handling, Sections 1 to 3
C = 3.56 Chemical handling, Section 4 to 12
TDAFC =1.32 million Ib/hr, Figure 8.2
Using the major equipment equations shown in Table 8.1 and
the above values of C, the following costs were calculated:
183
-------
Section E M? TPI M$
1
2
3
4
5
6
7
8
9
10
11
12
2,100
3,800
16,720
4,500
13,190
5,500
3,000
11,620
1,800
2,157
16,200
14,000
5,523
9,994
43,974
16,020
46,956
19,580
10,680
41,367
6,408
7,679
57,672
49,840
315,693
TPI = 315.693 million
TCOAL = 1.32/0.66 =2.0 million Ib/hr
ACOAL = 12 x 3 x 2 x 340 = M$ 24,480
ASULF = 0.1 x 5 x 1.32 x 1 x 340 = M$ 224
ANR = 24,480 + 1,600 - 224
= M$ 25,856
TCR = 1.21 x 315,693 + 0.8 x 300 x 7 (1 + 1.3) + 0.4 x 25,856
TCR = $396.195 million
TAG = 0.225 x 315,693 + 2.1 x 300 x 7 (1+1.3) + 1.04 x 25,856
TAG = $108.064 million
The Annual Gas Production
AGP = 250,000 x 340 = 85.0 million MMBtu/year
The gas cost = 108.064/85.0
= $1.27/MMBtu
184
-------
Location; Wyoming F = 1.3
Coal Details; Subbituminous PCARB = 74%, PSULF 0.12%
PH20 = 30%, PASH = 5%
CCOAL= $3/ton
Other Information:
SD = 340 CSULF = $5/ton
Derived Information;
C = 2.63 Sections 1 to 3
C = 3.56 Sections 4 to 12
TDAFC =1.37 million Ib/hr
CO = $7/hour
Section
1
2
3
4
5
6
7
8
9
10
11
12
TPI =
EM$
2,100
4,100
16,240
4,500
13,024
5,500
3,000
11,140
1,800
260
15,600
14,000
$304 .03 million
TPI M$
5,523
10,783
42,711
16,020
46,365
19,580
10,680
39,658
6,408
926
55,536
49,840
304,030
TCOAL = 1.37/0.65 =2.1 million Ib/hr
ACOAL = 12 x 3 x 2.1 x 340 = M$ 25,704
ASULF = 0
ANR = 25,704 + 1,600
= M$ 27,304
185
-------
TCR = 1.21 x 304,030 + 0.8 x 300 x 7 (1+1.3) + 0.4 x 27,304
TCR = $382.66 million
TAG = 0.225 x 304,030 + 2.1 x 300 x 7 (1 + 1.3) + 1.04 x 27,304
TAG = $106.95 million
AGP = 85.0 million MMBtu/year
The gas cost = 106.95/85.0
= $1.26/MMBtu
Example 3
Location: Illinois F = 1.7
Coal Details: Bituminous PCARB = 78%, PSULF = 5.6%
PH20 = 14% , PASH = 15%
CCOAL = $6/ton
Other Information:
SD = 340 CSULF = $5/ton CO = $7/ hour
Derived Information;
C = 2.88 Sections 1 to 3
C = 3.95 Sections 4 to 12
Section E M$ TPI M$
1
2
3
4
5
6
7
8
9
10
11
12
2,100
3,700
16,880
4,500
14,120
5,500
3,000
11,780
1,800
10,270
16,400
14,000
186
6,408
10,656
48,614
17,775
55,774
21,725
11,850
46,531
7,110
40,566
64,780
55,300
386,729
-------
TPI = $386.729 million
TCOAL = 1.31/0.71 =1.85 million Ib/hr
ACOAL = -12 x 6 x 1.85 x 340 = M$ 45,288
ASULF = 0.1 x 5 x 1.31 x 5.6 x 340 = M$l,247
ANR = 45,288 -t- 1,600 - 1,247
= M$ 45,641
TCR = 1.21 x 386,729 + 0.8 x 300 x 7 (1 + 1.7) + 0.4 x 45,641
TCR = $490.73 million
TAG = 0.225 x 386,729 + 2.1 x 300 x 7 (1 + 1.7) + 1.04 x 45,641
TAG = $146.39 million
AGP = 85.0 million MMBtu/year
Gas Cost = 146.39/85.0
= $1.72/MMBtu
187
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8.2.5 The Influence of Coal Type, Coal Cost, Percentage Sul-
fur and Plant Location on Gas Cost
The graphs on the following pages were derived using the model
to investigate the influence of coal type, coal cost, percentage
sulfur and plant location on the cost of the SNG. The following
general observations can be made after examining these figures:
1- The location of the plant and the coal cost have the
largest effect on the gas cost. If the plant is built
in a high construction cost area or the coal price
is high, SNG costs can become unnecessarily high.
2. The sulfur content of the coal has a secondary effect.
The difference in gas cost between a high and a low
sulfur coal is 8 to 12. .C/MMBtu depending on the
plant location.
3. For a given location, sulfur content and coal cost,
the cost of gas decreases as the percentage carbon
in the coal increases, i.e., gas cost is less for
bituminous coal than lignite.
4. All of the curves shown in the four figures do not
represent possible real situations. For example,
it is highly unlikely that low sulfur bituminous
coal will be available at $3/ton and even more un-
likely that it would be available in an area with
a location factor around 1.0.
5. The most attractive real situations appear low sulfur
subbituminous coal in areas like New Mexico and Wyom-
ing where there is a possibility that coal could be
purchased for around $3 or 4/ton including re-land-
scaping strip mines. Here a gas price of between
$1.2 and 1.4/MMBtu (1973 plant costs) appears realis-
tic .
188
-------
6. An underestimation of 10% in the total plant in-
vestment would mean an underestimation of the gas
cost by 8 to 11 C/MMBtu depending on the value of
TPI.
189
-------
TABLE 8.1
Summary of Major Equipment Cost Equations
Substitute Natural Gas Production
M$
El = 2,100
E2 = 5,000 - 100 (PCARB-65)
E3 = 14,800 + 160 (PCARB-65)
E4 = 4,500
E5 = 13,000 + 200 PSULF
E6 = 5,500
E7 = 3,000
E8 = 9,700 + 160 (PCARB-65)
E9 = 1,800
0.6
E10 = 1,250 PSULF + 1,065 (TDAFC.PSULF) -250
Ell = 13,800 + 200 (PCARB-65)
E12 = 14,000
190
-------
FIGURE 8.1
HIGH HEATING VALUE OF VARIOUS RANKS OF COAL
16,000 - -
15.000 --
14,000 - -
13,000 • -
12,000 • -
11,000 • -
10,000 - -
HIGH
HEATING
VALUE
BTU/LB
DRY ASH
FREE COAL
NOTE
DRY, HERE, MEANS
FREE FROM ALL H2O
NOT ONLY SURFACE
MOISTURE
LIGNITE
SUBBITUMINOUS
BITUMINOUS
ANTHRACITE
50
60 70 80 90
% CARBON DRY ASH FREE BASIS
100
REFERENCES: 1. 1972 KEYSTONE COAL INDUSTRIAL MANUAL, MCGRAW HILL.
2. PERRY, H., "CHEMICAL ENGINEERING HANDBOOK," 4TH EDITION.
3. LOWRY, H., "CHEMISTRY OF COAL UTILIZATION," WILEY.
191
-------
2.8 T
2.6
FIGURE 8.2
DRY ASH FREE COAL, OXYGEN AND STEAM
REQUIREMENTS FOR A 250 X 109 BTU/DAY
LURGI SNG PLANT
2.4
MM LB/HR
TO GASIFIER
2.2
STEAM TO
GASIFIER
2.0
1.8 - -
1.6 - -
1.4
1.2
1.0
0.8
TOTAL DRY ASH FREE
COAL REQUIREMENT
DRY ASH FREE
COAL TO GASIFIER
0.6
OXYGEN TO
GASIFIER
0.4
0.2 - -
LIGNITE
SUBBITUMINOUS
BITUMINOUS
50
60
70 80
% CARBON DRY ASH FREE BASIS
90
100
192
-------
FIGURE 8.3
LURGI SNG PROCESS FLOW DIAGRAM
RAW
COAL
FEED
UJ
70%
COAL
PREPARATION
AND
GRINDING
30%
FINE
AGGLOMERATION
LURGI
GASIFIERS
AIR
OXYGEN
PLANT
1
STEAM &
POWER
GENERATION
11
SHIFT
CONVERSION
& COOLING
SULFUR
REMOVAL &
RECOVERY
10
PURIFICATION
C02 + H2S
REMOVAL
TARS. PHENOLS ETC.
PHENOL
SOLVAN
UNIT
CLEAN STACK GAS
SULFUR BY-PRODUCT
METHANATION
SNG
COMPRESSION
SNG TO
, PIPELINE
FUEL TO STEAM
& POWER GENERATION
OTHER
OFFSITE
12
-------
FIGURE 8.4
EFFECT OF LOCATION FACTOR ON GAS COST
SUBBITUMINOUS COAL % CARBON 74 DAFB
% SULFUR 0.1 DAFB
MINE-MOUTH COAL COST 6 $/TON
4.5 $/TON
3 $/TON
GAS COST
$/MM BTU
1.1
1.0
COAL
S/TON
1.8 1.9 2.0
LOCATION FACTOR F
194
-------
FIGURE 8.5
EFFECT OF LOCATION FACTOR ON GAS COST
BITUMINOUS COAL % CARBON 78 DAFB
% SULFUR 6 DAFB
2 DAFB
MINE-MOUTH COAL COST $6/TON
$3/TON
2.0
1.9 • -
1.8
1.7
1.6
1.5
1.4
1.3
1.2
GAS COST
$/MM BTU
COAL COST
$6/TON
% SULFUR
6%
2%
6%
2%
1.1
10
1.0 1.1 1.2 1.3 1.4 1.5 1.6 1.7 1.8 1.9 2.0
LOCATION FACTOR F
195
-------
FIGURE 8.6
EFFECT OF CARBON CONTENT OF COAL ON SNG COST
% SULFUR 0,1,2,4,6
MINE-MOUTH COAL COST $6/TON
$3/TON
LOCATION FACTOR F = 1.5
2.0 -._ GAS COST
$/MM BTU
1.9
1.8
1.7
1.6
1.5
1.4
1.3 - -
1.2 - -
1.1
1.0
60
% SULFUR DAFB
70 80
% CARBON DAFB
90
COAL COST S6/TON
COAL COST $3/TON
100
196
-------
FIGURE 8.7
EFFECT OF CARBON CONTENT OF COAL ON SNG CAPITAL COSTS
MILLION $
500 -.-
480 • -
460 • -
440 4-
420
400 -4-
380
360 4-
340
320 J_
300
60
LOCATION FACTOR F = 1.5
% SULFUR
DAFB
MINE-MOUTH 6
COAL COST
S/TON 3
$/TON
70 80
6%
TOTAL
CAPITAL
REQUIRED
1%
90
TOTAL
PLANT
INVESTMENT
100
% CARBON DAFB
197
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9. SOLVENT REFINED COAL PRODUCTION
9.1 Process Appraisal
The Stearns-Roger Corporation's Report of July 1969, prepared for
the Pittsburg and Midway Coal Mining Company was chosen as the
tj.sis for the development of this model (30, 32). Their design
was examined by MWK Research and Engineering Development Department
and found to be unrealistic in some areas. A number of process
modifications have been made and these are discussed below.
Their design imports a quantity of natural gas equivalent to
the heat content of 10% of the total solvent refined coal product.
This gas is used for the hydrogen plant feed and for fuel. Under
present conditions this is basically unsound. Although the
price of natural gas is still low at present, it is more flexible
and has a lower sulfur content than the solvent refined coal
and is potentially a more, valuable product. The process has,
therefore, been modified to use the light oil and hydrocarbon
by-product streams as fuel and hydrogen plant feed.
In the Stearns-Roger Design, a fluidized combustor was included
to burn off the carbon remaining on the ash. An examination was
made by MWK to determine the economics of burning off this carbon.
The total plant investment of the carbon burn-off section was $9
million and the total annual production cost of the electricity
produced by the plant was $2.1 million. The electricity could
only be sold for $1.3 million, so this area of the plant would
operate at a loss. Instead, the dry ash is conveyed to storage
after stripping off wash solvent. The steam, which would have
been generated by burning-off the carbon, produced approximately
the excess power which was to be sold. The chance of selling
carbon-ash as by-product is considered slim. In fact, it is more
realistic to provide an annual operating cost for the disposal
of ash. The ash could be dumped as land fill where coal is mined,
assuming the plant is built adjacent to the coal mine.
198
-------
The process conditions were changed so that 2% more fuel was
produced at the expense of producing 2% less solvent refined coal
and the plant is now in energy balance. The hydrogen consumption
and losses were estimated to be about 60% more than the figure
given in the Stearns-Roger design and the feed to the hydrogen
plant was adjusted accordingly.
The plant size was increased so that the solvent refined coal
Q
production rate is 250 x 10 Btu/day (higher heating value).
This enables direct comparison to be made with the SNG
process.
The revised plant flow rates are given below:
Tons/Day
Raw Coal Feed to Plant 13,600
Solvent Refined Coal Product 7,834
(Equivalent HHV of SRC = 250 billion
Btu/day)
Sulfur by-product (LT/D) 300
Cresylic acid by-product 170
The revised heat requirements and fuel production are given below:
Consumption MMBtu/hr
Dissolver preheaters 1,770
Vacuum flash preheater 450
Wash solvent splitter heater 410
Ash residue drying 70
Power generation 50
Hydrogen plant fuel 680
Hydrogen plant feed 1,150
Miscellaneous 70
4,650
Production
Fuel Gas 3,260
Light Oil burned as fuel 1,390
4,650
199
-------
If the coal feed has about 4% sulfur DAFB, then the solvent
refined coal product would usually contain not less than 1%
sulfur. (According to information from EPA, the sulfur content
of solvent refined coal could be reduced to as low as 0.4% with
process modification at additional cost.) This particular design
produces a solid refined coal, which can be pulverized or sold
as briquettes, or a liquid product depending on the need cf the
plant.
9 . 2 Process Description
The process model for the solvent refined coal process has been
developed as 9 sections (Fig. 9.2). A brief description is therefore
given for each of these sections.
Section 1
Raw coal from storage is first crushed and then processed through
a secondary grinder to reduce the coal particles to less than
1/8 of an inch. The resulting coal fines pass through a flash
dryer to remove the moisture content.
Section 2
The coal together with the solvent and hydrogen are passed through
preheaters and dissolvers. The coal dissolves in the solvent in
the presence of hydrogen at 1000 psig and 825 °F. The dissolution
of coal involves hydrogenation and depolymerization. The remaining
undissolved material consists of the ash content of the coal. This
section also includes the hydrogen compressors.
Section 3
The ash residue from the coal is separated from the solvent by
rotary filters at 150 psig and 600°F. The ash portion is transfered
to the ash drying section for further solvent recovery and on to
storage.
200
-------
Section 4
The solvent and light oils are recovered by a series of flash
separators followed by vacuum distillation. The overheads are
further distilled to recycle the solvent and produce a light
fuel oil. The vacuum tower bottom product is the liquid refined
coal. This section also includes a cresylic acid recovery unit.
Section 5
In this section the liquid coal product is solidified and trans-
fered to storage.
Section 6
This section generates makeup hydrogen for the process by steam
reforming the light oil stream.
Section 7
The fuel gas and light fuel oil are treated to remove hydrogen
sulfide and sulfur compounds. The hydrogen sulfide goes to the
sulfur recovery unit, which produces a saleable by-product.
Section 8
The steam and power generation plant is fired by fuel gas and
light fuel oil.
Section 9
This section includes other offsites: the cooling water system,
water treatment and general plant buildings.
201
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9.3 Cost Model
9.3.1 Totaj^ Plant Investment
The costs presented in the Stearns-Roger Report (30) were
updated to the end of 1973 and adjusted using the factors
listed in the General Cost Model to represent the increased
plant size. For this model no estimating has been carried
out by MWK and the costs have therefore only been presented
as plant investments for each area. The cost of each section
has been examined and approximate adjustments have been made
to the estimates in those areas whose cost appeared to be low.
There is general agreement with the Steams-Roger's cost
in the following areas: raw coal preparation, ash filtration
and drying, product solidification, hydrogen plant, fuel
treatment and sulfur recovery, and steam/electricity generation.
However, it was felt that the other areas, viz., the preheater/
dissolver units, solvent/light oil/cresylic acid recovery
and the general off site units were on the low side.
The preheater/dissolver units are large items operating at
high pressures and temperatures . Much of the material of
construction is stainless steel. The dissolver design is
complex and not very well defined as yet. The plant invest-
ment of this area has therefore been increased over the ad-
justed Steams-Roger's cost by $10 million.
The solvent/light oil/cresylic acid recovery units are
relatively complex, large units largely constructed of stain-
less steel. The plant investment for this area has been
increased by $15 million over the Stearns-Roger figure.
It was felt that $30 million for other offsite units was
more likely than the adjusted Stearns-Roger Figure of under
$10 million.
In making these adjustments to the Stearns-Roger costs, it
202
-------
should be stressed that no detailed estimating was done by
MWK and these changes should be regarded as approximate only.
The Total Plant Investment for a plant producing 250 x 10
Btu/day of solvent refined coal is given below.
Total Plant Investment (TPI)
Section
Number
1
2
3
4
5
6
7
8
9
Section
Description
F=1.0
M$
Coal preparation (solid handling section)10 , 000
Preheater/dissolvers
Ash filtration, drying and disposal
Solvent/light oil/cresylic acid
recovery
Product solidification/handling and
storage
Hydrogen plant
Sulfur removal from fuels and sulfur
recovery
Steam and power generation
Other offsites
40 ,000
15 ,000
30,000
10,000
10,000
10,000
10,000
30,000
165,000
If F = 2.0, the value of TPI is $215 million.
9.3.2 Total Net Annual Operating Cost, Total Capital
Requirement and Total Annual Production Cost
The total net operating cost, AOC, is the total cost of
operating the plant less the credits from the sale of by-
products. It does not include return of capital, payment
of interest on debt or income tax on equity returns and
is given by:
AOC = 0.078 TPI + 2TO.CO(1+F) + ANR
203
-------
The total number of shift operators for a plant producing
9
250 x 10 Btu/day of solvent refined coal is estimated to
be 200.
The annual cost of raw materials less by-product credits
(ANR) is given by:
ANR = ACOAL + ACHEM - ASULF - ACRES
The annual cost of coal feed to the plant (ACOAL) is given by:
ACOAL = CCOAL-TCOAL-SD M$/Year
where CCOAL is the cost of coal in $/ton, TCOAL is the total
raw coal feed to the plant in Mton/day and SD is the number
of days the plant is on stream per year.
The annual cost of catalyst and chemicals, ACHEM, can be
assumed to be 500 M$ for the plant considered. The annual
credit for the sale of sulfur, ASULF, is given by:
ASULF = CSULF'TSULF-SD M$/Year
where CSULF is the unit credit for sulfur in $/LT and TSULF
is the sulfur production rate in MLT/day.
The annual credit for the sale of cresylic acid is given by:
ACRES = CCRES-TCRES-SD M$/Year
where CCRES is the unit credit for cresylic acid and TCRES
is the production rate of cresylic acid in Mton/day.
For the present plant using a 4% sulfur coal (DAF) , TSULF
is estimated to be 300 LT/day and TCRES, 170 tons/day.
204
-------
The Total Capital Required, TCR, including interest on
construction capital, startup costs and working capital is
given in the General Cost Model as:
TCR =1.21 TPI + 0.8 TO-CO (1+F) +0.4 ANR
The Total Annual Production Cost, TAG, including the return
of capital, payment of interest and income tax on equity
return is given by:
TAG = 0.225 TPI +2.1 TO-CO (1+F) +1.04 ANR
9.3.3 Calculation of Costs of Solvent Refined Coal
Location Factor: F = 2.0
Coal Details:
Bituminous PCARB = 78%, PSULF =3.8%
Plant Details;
By-Products:
250 billion Btu/day of SRC
Coal Feed rate 13,600 tons/day
On Stream for 340 days/year
Total number of shift operators TO = 200
300 tons/day of sulfur at $5/ton
170 tons/day of cresylic acid at $100/ton
Example 1
Coal Cost:
$3/ton
ACOAL = 13.600 x 240 x 3 = M$ 13,900
ASULF = .300 x 340 x 5 = M$ 500
ACRES = .170 x 340 x 100 = M$ 5,700
ACHEM = Cost of catalysts and chemicals = M$500
The cost of raw materials and chemicals less by-product credits
is given by:
205
-------
ANR = 13,900 + 500 - 500 - 5,700
ANR = M$8,200
TAG = 0.225 x 215,000 + 2.1 x 200 x 7 (1+2) + 1.04 x 8200
TAG = $65.7 million/Yr
Cost of SRC = $65.7 x 106/(340 x 250,000 MMBTU)
= $0.77/MMBtu
Example 2
Coal Cost; $6/ton
ACOAL = M$27,800
ANR = 27,800 + 500 - 500 - 5,700
ANR = M$22,100
TAG = 0.225 x 215,000 + 2.1 x 200 x 7 (1+2) + 1.04 x 22,100
TAG = $80.2 million/Yr
Cost of SRC = $80.2 x 106/(340 x 250,000 MMBtu)
= $0.94/MMBtu
Figure 9.1 illustrates the variation in the cost of solvent
refined coal with location factor, for coal costs of $3 and
$6/ton. The costs of substitute natural gas produced by a
plant of the same size using the same bituminous coal feed-
stock are also given for comparison.
An underestimation of 30% in the total plant investment would
result in an underestimation of SRC costs by 18 to 24*/MMBtu
depending on the location. Because of the limited state of
development of the SRC process, the order of accuracy of the
TPI estimate is only about 30%.
9.4 Conclusions
Solvent refined coal can be produced more cheaply than SNG. It is,
however,, a solid fuel containing normally about 1% sulfur when
produced from a 4% sulfur coal (DAF) . According to proprietary
206
-------
information from EPA, SRC with as low as 0.4% sulfur can be
produced with process modification and increase in cost. Presumably
this can be done by increasing the hydrogenation pressure in the
dissolvers. However the main area that needs to be improved
seems to be the ash filtering section. This section is probably
the most costly as well as troublesome in operation.
Another aspect of investigation for the production of SRC appears
to be the market. It is basically an expensive, low sulfur, ash
free fuel not suitable for direct use with gas turbines. Since
the process can also be geared towards specialized refinery type
products, the question arises as to whether this would be a more
worthwhile direction than the production of a solid fuel.
207
-------
FIGURE 9.1
COMPARISON OF SOLVENT REFINED COAL
AND SUBSTITUTE NATURAL GAS COSTS
BITUMINOUS COAL FEED, 3.8% SULFUR
$/MM BTU __
1.7
1.6 - -
SUBSTITUTE NATURAL GAS
SOLVENT REFINED COAL
MINE-MOUTH
COAL COST
S6/TON
$3/TON
MINE-MOUTH
COAL COST
$6/TON
$3/TON
0.5
1.0 1.1
LOCATION FACTOR F
208
-------
FIGURE 9.2
SOLVENT REFINED COAL PROCESS FLOW DIAGRAM
USED WITH
PLANT
STEAM
& POWER
GENERATION
FUEL GAS & OIL
to
O
HYDROGEN
PLANT
RAW
COAL
FEED
COAL
HANDLING
& GRINDING
LIGHT OIL
SLURRY
PREHEAT
DISSOLVER
RECYCLED
HYDROGEN
GAS
SULFUR
REMOVAL &
RECOVERY
ASH
FILTERING
DRYING
SULFUR
BY-PRODUCT
SOLVENT
AND LIGHT
OIL RECOVERY
OTHER
OFFSITES
PRODUCT
SOLIDIFICATION
SRC
CRESYLIC ACID
BY-PRODUCT
-------
10. THE COMBINED GAS TURBINE - STEAM TURBINE POWER PLANT USING A
LOW BTU LURGI GASIFIER
10.1 Introduction
A new coal fired conventional steam cycle power station without
stack gas cleaning can achieve a heat rate as low as 8,600 Btu/
net kwh which is equivalent to an overall cycle efficiency (based
on the higher heating value of the coal) of almost 40%. If this
power station was designed with, say, a Wellman/Allied stack gas
scrubbing system and was burning coal with 4% sulfur DAFB, the
fuel required would be increased by more than 11%, using the present
design. The heat rate would be increased to more than 9,600 Btu/
net kwh and the overall cycle efficiency (net power/coal input)
reduced to 35.5%. If sulfur dioxide emission controls are to
be imposed on new power stations there is clearly a great deal
of incentive to investigate alternatives to the conventional
steam cycle power plant with stack gas scrubbing. One such alter-
native is a coal gasification plant supplying clean, hot, low
Btu fuel gas under pressure to a combined gas turbine - steam
turbine power generation unit.
There are many coal gasification processes in various degrees
of development. The Lurgi coal gasification unit using either
air or oxygen is, however, well established on a commercial
scale. For this reason, in these studies, the Lurgi coal gasifi-
cation unit, using air, has been selected to provide the low Btu
fuel gas. Emphasis has been given to designs which are possible
at present, although calculations have been made for future
cycles not limited by the gas turbine inlet temperature.
When this work was started it was felt that a gas turbine inlet
temperature of 1700°F was the highest allowable design temperature for
base loaded plants. Recent discussions (December 1973) with General
210
-------
Electric Company (39) reveal however, that they now have marketed
a base load air compressor/gas turbine/generator unit producing
55 net megawatts, operating with an inlet temperature of 1950°F
and an inlet pressure of 150 psia. To some extent this makes
cycles 1 and 2 (to be defined later) already outdated; however,
the results and discussion on them are included for general interest.
These power plants are obviously base loaded, since the Lurgi
gasifiers cannot be shut down other than for maintenance. The
gas turbine must therefore be capable of in excess of 8000 hours
operation per year. It is also absolutely necessary that the in-
let gas and air have practically zero alkali metal impurities.
A special water wash, free of sodium and potassium ions is there-
fore needed after the hot potassium carbonate purification plant.
A gas turbine inlet temperature of 1600°F obviously puts some
limitations on the design of the power cycle, since the exhaust
temperature is less than 1000°F (gas turbine exhausts at 16 psia).
The most widely used power plant steam cycle has 1000°F, 2400 psia
steam with reheat to 1000°F after expansion down to 500-600 psia
(33, 38). This steam cycle is illustrated in Figure 10.1 and is
essentially the steam cycle used in all the combined cycle studies,
although in these, it is clearly necessary to preheat the water with
the flue gas. In a conventional power plant, this is done by bleed-
ing steam from several pressure levels in the turbines and preheating
the combustion air with the flue gas. The steam cycle illustrated in
Figure 10.1 is not the most efficient possible, since steam turbine
efficiencies of 88% (isentropic work to electrical power) have been
used. However, this is intended to be a comparative study and in the
combined power cycle work, gas turbine, steam turbine and air com-
pressor efficiencies of 88% were used. Again it must be emphasized
that the combined cycles reported in this work should not be con-
sidered as finalized designs, but taken as illustrations of the
possibilities of the cycles under these stated conditions.
It is clear that with gas turbine inlet temperature of 1600°F, the
steam cycle cannot receive heat only from the gas turbine exhaust,
211
-------
since its temperature is too low (Fig 10.16).
One method of achieving a sufficiently high temperature for the
superheat and reheat tubes in the steam turbine section is to divide
the fuel gas into 2 streams. One stream is combusted with excess
air to give an inlet temperature of 1600°F and expanded to 16 psia
and less than 1000°F. The second fuel gas stream is expanded to 16
psia and then mixed with the first gas turbine exhaust in a combustor.
The combustion temperature will then be in excess of 1200°F and
suitable for supplying heat to the reheat and superheat tubes.
This cycle was investigated and found to produce less power than
cycle 1 (illustrated in Figure 10.3). It is also more complicated
and requires an extra turbine. The results for this alternative
have therefore not been reported here.
Another area of work for which the results are not reported
should be mentioned if only to avoid further study. There is
a great temptation to assume that conventional centrifugal air
compressors with intercoolers to reduce the power consumed would
be used in a combined power cycle. This is not so. In fact,
much higher overall cycle efficiencies are achieved by using
axial flow compressors with no intercoolers. This fact has been
reported in a few publications but is still not widely recognized.
The use of centrifugal compressors rejects well over 1000 MMBtu/hr
of heat to cooling water. A design utilizing axial flow compressors
keeps the heat of compression in the power cycle.
10.2 The Lurgi Gasification Plant
Before discussing details of the power cycles, a brief description
of the low Btu Lurgi gasification plant is useful. Approximately
20 Lurgi gasifiers are necessary to generate the fuel gas for a 1000
megawatt combined cycle power plant. The air enters the gasifiers
at about 600°F, preheated by compression to 320 psia, and the steam
enters at about 460°F from an intermediate pressure level in the
steam power cycle. For the case studied a bituminous coal with
212
-------
the following analysis was used:
Wt % DAFB % As Received
C 78.0 Water 15
H 5.5 Ash 15
O 11.0
S 4.0
N 1.5 HHV = 14,200 Btu/lb DAF
100.0
The coal feed to the gasifiers was taken as 710,000 Ib/hr DAF.
The significance of this number is that it is the coal feed to
a new conventional steam cycle power plant fitted with a Wellman/
Allied stack gas scrubbing system and generating 1000 megawatts
net power. The Wellman/Allied process produces a sulf-ur by-product
and is therefore compatible with the low Btu Lurgi design. The
steam cycle is as shown in Figure 10.1 and its heat rate would
be 9000 Btu/net kwh without the stack gas scrubbing unit. So
in the power cycle studies a combined cycle which produces net
power over 1000 megawatts is more efficient than the conventional steam
cycle with Wellman/Allied stack gas scrubbing. (If the power
plant were fitted with a Wet Limestone system, the net power pro-
duced would be 1074 megawatts. Combined cycles producing power
in excess of this amount would be more efficient than a conventional
plant with Wet Limestone scrubbing.). The conventional steam cycle
without stack gas scrubbing would achieve 1127 net megawatts with
this coal feed rate.
The composition of the crude gas after the tars have been removed is
approximately (34):
Mole % (Dry Basis)
C02 14.5
CO 15 .6
H2 21'5
CH4 6.0
N2 41'7
H S 0. 7
f £
100.0
213
-------
The gas has an additional 26% of water. It is purified by a hot
carbonate scrubbing system. The gas leaves the absorber at 260°F
and is reheated to 300°F by heat exchange with gas leaving the
gasifier. The gas is saturated by the water wash for alkali metal
removal. The reboiler heat required by the stripper tower of the
hot carbonate system is supplied by cooling the tar-free gas enter-
ing the purification system from 500°F to 260°F.
The gas entering the combined power cycle has approximately the
following composition:
Mole % (Dry Basis)
C02 10.65
CO 16.5
H2 22.5
CH4 6.3
N2 44.0
H2S 0.05
100.00
The gas is saturated with water at 300°F and 250 psia. The total
wet flow rate is 163,300 Ib moles/hour. All tars, etc. are recycled
to the gasifiers after being removed in a water wash tower.
A heat balance around the Lurgi gasifiers and purification system
is given in Figure 10.2.
10.3 Description of Cycles Studied
The cycles which were studied are illustrated in Figures 10.3-10.6.
Cycle 4 (Figure 10.6) is a future cycle. Cycles 1 to 3 (Fig. 10.3-
10.5) are the cycles which are possible now. The differences between
them are in the positions of the preheat, superheat and reheat
tubes in the steam generators. The differences between cycle 3 and
cycle 4 is the gas turbine inlet pressure.
214
-------
In cycle 3, gas turbine A has an inlet pressure of 100 psia and
inlet temperatures greater than 1900°F. The exhaust temperatures
from turbine A are therefore greater than 1200°F (see Figure 10.16)
and suitable for providing heat to the whole steam cycle (see Fig.
10.14). Thus, no heat is given up to the steam cycle by the high
pressure gas at the higher temperatures (the superheat & reheat
tubes in Cycle 1, and the superheat tubes in Cycle 2) and the cycle
is potentially more efficient. Another illustration of this
principle is given by comparison of cycles 1 and 2 (Figures 10.3
and 10.4). These are different from cycle 3 in one way. The
inlet temperature to gas turbine A is less than 1900°F and the
exhaust temperature is too low to provide heat for the superheater
and the reheat tubes of the steam cycle. Cycle 1 is not as effi-
cient as cycle 2 (see Figure 10.7) at the same gas turbine inlet
temperatures because the steam system of cycle 2 removes less heat
from the high pressure hot gas. Cycle 1 has both the superheater
and the reheat tubes heated by the gas before the inlet to turbine
A, whereas cycle 2 only has the superheater in this location.
Examination of Figure 10.7 will also reveal another interesting
point. Cycle 3 generates less power than cycle 2 when the inlet
temperature to turbine A is less than 1960°F. The reason for this
is that the steam cycle with 1000°F superheat and 1000°F reheat is
not efficient when the gas entering the steam generator is less than
1250°F gas inlet temperature. The pinch point between the cooling
curves results in the slope of the gas cooling curve being much less
than that of the water preheat curve. This means that the two curves
diverge and the stack gas temperature is 350°F, which is obviously
undesirable, especially in view of the large amounts of excess air
in the flue gas. At 1300°F gas inlet temperature the gas cooling
curve is almost parallel to the water preheat curve and a stack gas
temperature of 230°F is possible.
Cycles 1, 2 and 3 are identical up to the point where the gas is
combusted. The clean fuel leaving the Lurgi unit at 300°F and 250
psia is preheated to 420°F by two heat exchanges with the hot com-
pressed gasifier air. The air itself is cooled to 600°F. The fuel
gas at 250 psia and 420°F is then expanded in turbine B to 105 psia
215
-------
and 270°F providing a power generation of 59 megawatts. The expanded
fuel gas is then combusted with a quantity of air which is in excess
of the stiochiometric requirements. The quantity of air which is
used produces the control of the combustion temperature (see Figure
10.9). The air temperature leaving the axial flow compressor is
580°F.
Cycle 4 differs from cycle 3 in that in this case, the combustion
air is compressed to 250 psia and 920°F by an axial flow compressor.
Turbine B is therefore not required. Turbine A expands the combusted
gas from 250 psia to 16 psia. Again cycle 4 does not produce an
efficient steam cycle until the gas turbine exhaust temperature
exceeds 1250°F which corresponds to an inlet temperature 2430°F.
In all four cycles the steam to the gasifier was taken from the
steam cycle after the HP turbine. It was expanded down to the
gasifier pressure to produce about 6 megawatts.
10.4 Discussion of Results
In view of the statements by General Electric Corporation that
their present 55 net megawatt unit can operate base loaded with gas
turbine inlet temperatures of 1950°F (39), it appears reasonable
to claim that cycle 3 with an inlet temperature of 2000°F is the
best available cycle for a plant designed in the next two years.
It is possible that the best intermediate pressure between turbine
A and turbine B may be higher than 100 psia. In fact the net power
out of the air compressor/turbine A unit is higher for 150 psia
than for 100 psia, but the higher pressure means using a steam cycle
with superheat and reheat temperatures lower than 1000°F. In short
the most suitable design at present appears to be close to cycle 3,
but more detailed work and discussion with machine vendors would be
required to produce an optimized design.
MWK's calculations show the combined cycle using a Lurgi gasification
system produces about the same power as the conventional steam
cycle plant without stack gas scrubbing. It is true that the steam
cycle shown in Figure 10.1 is not the most efficient available,
216
-------
but the same is probably true of the combined cycle work. The
efficiency ratio of (isentropic work, to megawatts consumed or
generated) has been assumed to be 88% for the air compressors,
the gas turbines and the steam turbines. It is likely that the
steam turbine and air compressor efficiencies could be higher than
88% and there are indications that 88% is possible for the gas
turbines.
Thus, the overall Lurgi gasification plant and generation unit
appears capable of efficiencies higher than 38% (net power generated
divided by the higher heating value of the coal input). On occasion,
Lurgi publications have expressed the overall efficiency in terms
of the lower heating value of the coal. This produces a figure
which is approximately 2 points higher, i.e., 38% (HHV) is equi-
valent to 40% (LHV).
The power (other than compressors) used by the plant itself was
taken as 30 megawatts for the combined cycle and Lurgi plant, and
60 megawatts for the conventional steam cycle. The difference
is due to the reduced boiler feed water pump power requirements
and the fact that no induction fans are needed.
10.5 Cost Model
The design which has been used for the cost model is cycle 3 with
a gas turbine inlet of 2000°F.
Megawatts
Gas turbine power = 1230
Air compressor power consumed = -605
Net generation = 625
Steam turbine power = 540
Total power generation = 1165
Auxiliary power requirement = -30
Total net generation = 1135
217
-------
The plant will be scaled down to a 1000 net megawatt size (1030
megawatts including auxiliary power) . For this size unit:
Net gas turbine/air compressor power = 1030 coc. _
T T f c X D^3 — 3 D .3
HDD
Steam turbine power = 1030 r^n _
T -^f r- X DfrU fl / f
1165 low
Auxiliary power requirements -30
Total net generation 1000
Table 10.1 was prepared by examination of published Tennessee
Valley Authority plant data (36, 37, 38). Information about the
Bull Run Plant provided most of the data. The plant investment was
updated to end of 1973 and brought to Gulf Coast cost. The size of
the plant was adjusted to 1000 net megawatts. Figure 10.18 was pre-
pared after examination of several TVA units. The cost of the
power generating units was found to vary with the 0.8 power of
plant size.
As shown in Table 10.1, the conventional steam cycle power station
total plant investment is $200 million (Gulf Coast, end of 1973).
A conventional power plant fitted with Wellman/Allied stack
gas scrubbing and generating 1000 net megawatts would require
an incremental boiler plant investment of $10 million and an
extra $50 million for the scrubbing plant. Total plant investment
for 1000 net megawatt station with Wellman/Allied stack gas scrubbing
is $260 million.
Table 10.1 and the Lurgi SNG model were used to generate Table 10.2.
The cost of the air compressor/gas turbine/generator unit was
firmed to a certain extent by General Electric Company's approxi-
mate cost for their 55 megawatt unit, which was between $4 and
$4.5 million, not installed.
218
-------
The conclusion appears to be that the conventional steam cycle
power plant without stack gas scrubbing would cost less than the
equivalent combined cycle unit with a Lurgi gasifier, which in
turn appears to cost less than the conventional station with the
Wellman/Allied stack gas scrubbing system. The combined cycle
power plant requires less coal feed than the conventional steam
power plant fitted with stack gas scrubbing.
The costs and efficiencies shown here indicate an incentive to
develop gas turbines which can handle higher inlet temperatures
and pressures, thus making cycle 4 possible. The combined cycle
power plant with a Lurgi gasification unit merits a more detailed
technical and cost examination.
219
-------
TABLE 10.1
Cost: of a 1000 Net Megawatt Conventional Power Station
Total Plant
Investment, TPI % of TPI
(MS)
Boiler plant equipment 102,000 51
Turbo-generator unit 46,000 23
Land and structures 26,000 13
Accessory electrical equipment 14,000 7
Transmission plant 8,000 4
Miscellaneous equipment 4,000^ 2_
200,000 100
Notes:
1) Coal-fired plant
2) Costs are Gulf Coast, end of 1973
3) No stack gas scrubbing
220
-------
TABLE 10.2
COST OF A 1000 NET MEGAWATT COMBINED CYCLE POWER PLANT
TOTAL PLANT
INVESTMENT, TPI
Power Generation
Boiler plant including extra capacity for
Lurgi gasifier steam. (No coal handling,
ash handling or draft equipment needed) 38,000
Steam-turbo generators 25,000
Air compressor/gas turbine/generator 50,000
Land and structures 26,000
Accessory electrical equipment 14,000
Transmission plant 8,000
Miscellaneous equipment 4,000
165,000
Lurgi Plant
Coal preparation and handling 3,500
Fines agglomeration 8,500
Gasification 23,500
Gas Purification 6,500
Sulfur control 7,000
Offsites 16,gpQ
65,000
Total for gasification unit and power station = $230 million
NOTE; Costs are Gulf Coast, end of 1973
221
-------
FURNACE
1
1
I
H
H
\
1521 _.
1306
M ^^s,,
^4-9 ^^
P 3000
T 550
H 550
H
<
I
^~~~
M 45-9
h
P 1150
T 792
H 1378
G
1 1
1
P 550
T 620
H 1306
F
FIGURE 10.1
STEAM BALANCE FOR 1000 MEGAWATT POWER PLANT
j (2400 PSIA 1000°F, 1000°F CYCLE)
P 475
T 1000
H 1521
M 35-2
\
\
K
i
P 180
T 750
H 1401
E
515
MEGAWATTS
i
M
37-8
r^
P 50
T 475
H 1272
D
. •
M
62-7
r^i
P 16
T 288
H 1187
C
i
— —
M
87-3
•^
— —
P
T
H 1
B
.. •
M
112-1
^=*
1
A
B
C
D
E
F
G
H
J
K
FLOW
MM LB/HR
4.420
0.213
0.273
0.334
0.456
0.718
0.843
7.257
5.696
5.240
FRACTION
OF A
1.0000
0.0482
0.0618
0.0756
0.1032
0.1624
0.1907
1.6419
1.2888
1.1865
BFW PUMP 30
LD. FANS 10
OTHER AUX20
60
H
545
-60
485
NET
MEGAWATTS
H 1018
8% WATER
VACUUM
CONDENSER
1-1/2" Hg
92°F
4
155
103
1
P 0.74
T92
H60
A
P PSIA
T °F
H BTU/LB
M MEGAWATT/MM LB/H
FURNACE DUTY
7850 MM BTU/HR
FURNACE LIBERATION
9000 MM BTU/HR
FURNACE EFFICIENCY
87%
TURBINE EFFICIENCY
88%
T 550
T 460
T 360
T 270
T 200
T 140 T 92
HEAT RATE 9000
BTU/NET KWH
-------
FIGURE 10.2
HEAT BALANCE AROUND LURGI LOW BTU GASIFIER
COAL 710,000 LB/HR DAF
HHV
10,082 MMBTU
N)
U)
FEED
WATER
AIR
600°F
665,000 LB/HR
ENTHALPY 0
65,700 MPH
232 MMBTU
GASIFICATION
CRUDE GAS
RECYCLE
TAR ETC
COOLING PURIFICATION
REHEAT &
RESATURATED
HEAT LOSS
C. IN ASH
350 MMBTU
150 MMBTU
500 MMBTU
FUEL GAS
163,313 MPH (WET)
300° F
250 PSIA
SATURATED
WITH WATER
SATURATED GAS
LEAVING HjS
STRIPPER
SEN. HEAT 290
LAT. HEAT 830
HHV GAS 8574
IMPORTED STEAM
400,000 LB/HR
480 MMBTU
SEN. HEAT
LAT. HEAT
HHV H2S
30 MMBTU
370 MMBTU
200 MMBTU
600 MMBTU
9654
COAL
STEAM
AIR
10,082 MMBTU
480 MMBTU
232 MMBTU
10,794 MMBTU
PURIFIED GAS
STRIPPED 0/H
HEAT & C. LOSS
OUT
9,694 MMBTU
600 MMBTU
500 MMBTU
10,794 MMBTU
-------
FIGURE 10.3
ADVANCED POWER CYCLE
POWER GENERATION SECTION (CYCLE 1)
250 PSIA. 300°F
CLEAN FUEL
GAS TURBINES:
ISENTROPIC WORK TO
MEGAWATTS GENERATED 88%
GAS FROM
GASIFICATION UNIT
1 EXHAUST GAS
T TO STACK
MVVWWNMr
STEAM PREHEAT & BOILING
M = MEGAWATTS/MMLB/HR OF STEAM
STEAM TURBINES:
ISENTROPIC WORK TO
MEGAWATTS GENERATED 88%
224
-------
FIGURE 10.4
ADVANCED POWER CYCLE
POWER GENERATION SECTION (CYCLE 2)
250PSIA, 300°F
CLEAN FUEL
GAS TURBINES:
ISENTROPIC WORK TO
MEGAWATTS GENERATED 88%
GAS FROM
GASIFICATION UNIT
COMPRESSED AIR
TO GASIF!ERS
EXHAUST GAS
TO STACK
M = MEGAWATTS/MMLB/HR OF STEAM
M= 112.1
STEAM TURBINES:
ISENTROPIC WORK TO
MEGAWATTS GENERATED 88%
225
-------
FIGURE 10.5
ADVANCED POWER CYCLE
POWER GENERATION SECTION (CYCLE 3)
AIR COMPRESSOR
250PSIA,300°F
CLEAN FUEL GAS
FROM GASIFICATION
UNIT
250PSIA. 420°F
COMPRESSED
AIR TO AIR
GASIFIER
105 PSIA, 580°F
AIR
100 PSIA
^
t
/Ov
03
GAS
BURNER
270°F
105 PSIA
GAS
TURBINE
.A
14 PSIA, 80°F
AIR
COMPRESSOR
EXHAUST GAS, 16 PSIA
GAS TURBINES:
ISENTROPIC WORK TO
MEGAWATTS GENERATED 88%
STEAM TURBINES:
ISENTROPIC WORK TO
MEGAWATTS GENERATED 88%
M = MEGAWATTS GEN./MMLB/HR
OF STEAM
226
-------
FIGURE 10.6
ADVANCED POWER CYCLE
POWER GENERATION SECTION (CYCLE 4)
AIR
COMPRESSOR
COMPRESSED
AIR TO
GASIFIERS
250PSIA, 300°F
CLEAN FUEL GAS
FROM GASIFICATION
UNIT
250PSIA, 920 F
AIR
14 PSIA. 80 F
AIR
COMPRESSOR
EXHAUST GAS 16 PSIA
GAS
BURNER
GAS TURBINE:
ISENTROPIC WORK TO
MEGAWATTS GENERATED 88%
STEAM TO
AIR GASIFIER
M = MEGAWATTS GEN./MMLB/HR
OF STEAM
STEAM TURBINES:
ISENTROPIC WORK WORK TO
MEGAWATTS GENERATED 88%
1-1/2" Hg. 92°F
227
-------
1340 -r-
1320 --
1300 --
1280 --
1260
1240
1220
FIGURE 10.7
EFFECT OF GAS TURBINE INLET
TEMPERATURE ON OVERALL
POWER GENERATION OF
ADVANCED POWER CYCLES
CVCLE 4
1. COAL FEED: 710,000 LB/HR
2. STEAM CYCLE: 2400 PSIA/1000°F/1000°F
3. CONVENTIONAL STEAM PLANT WITH NO STACK GAS SCRUBBING GENERATES 1127 MW
4. CONVENTIONAL STEAM PLANT WITH WET LIMESTONE SCRUBBING GENERATES 1074 MW
5. CONVENTIONAL STEAM PLANT WITH WELLMAN/ALLIED SCRUBBING GENERATES 1000 MW
<
o
1200 - -
1180 --
CYCLE 3
(T
LU
1160 --
I
Q_
1140 --
1120 --
1100
1080 - -
1060 - -
1040 - -
1020 - -
CYCLE 2
CYCLE 1
1000
1000
1400
1800
2200
2600
3000
GAS TURBINE INLET TEMPERATURE, °F
228
-------
46 --
FIGURE 10.8
EFFECT OF GAS TURBINE INLET
TEMPERATURE ON OVERALL
CYCLE EFFICIENCY OF
ADVANCED POWER CYCLES
CYCLE 4
t
44
OVERALL CYCLE EFFICIENCY IS
DEFINED AS NET ELECTRICAL
POWER PRODUCED DIVIDED BY
HEAT INPUT OF COAL
STEAM CYCLE: 2400 PSIA/1000°F/1000°F
CYCLE 4
42 --
CYCLE 3
40
o
ID
O
HI
UJ
_i
O
O
_l
-I
oc
01
o
38 . -
CYCLE 2
CYCLE 1 —
CYCLE 2
36 •-
CYCLE 1
CYCLE 3
34 --
32 . _
NOTE: A CONVENTIONAL STEAM
CYCLE POWER PLANT EFFICIENCY
IS ALWAYS DEFINED USING THE
HHV OF THE COAL, HOWEVER,
LURGI HAVE ON OCCASIONS USED
THE LHV IN THEIR PUBLICATIONS.
LHV OF COAL
HHV OF COAL
30
1000
1400 1800 2200
GAS TURBINE INLET TEMPERATURE, °F
2600
3000
229
-------
FIGURE 10.9
COMPRESSED AIR FLOW TO
GAS BURNER VS. COMBUSTION TEMPERATURE
700
600 - -
oc
X
e/5
UJ
500 - -
400 - -
CYCLE 1, 2 & 3
a
UJ
CC
300 - -
CYCLE 4
200 --
100 - -
4-
1000 1400 1800 2200 2600
COMBUSTION TEMPERATURE IN GAS BURNER, °F
3000
230
-------
FIGURE 10.10
TURBINE EXHAUST GAS FLOW VS.
GAS TURBINE INLET TEMPERATURE
700 -,-
600 - -
cc
I
o
5
n
o
I
500 • -
400 - -
CYCLE 4
CYCLE 1
LU
O
300
200
100 • -
1000
1400
1800
2200
2400
2800
GAS TURBINE INLET TEMPERATURE, F
231
-------
1400-T
1300--
FIGURE 10.11
TURBINE POWER GENERATION AND AIR
COMPRESSOR REQUIREMENTS FOR CYCLE 3
1200. _
GAS TURBINES POWER GENERATION
1100- -
1000- -
cc
LU
I
a.
900- -
800- -
700-
\
STEAM TURBINES POWER GENERATION
600- -
500- -
400--
300- -
200--
AIR COMPRESSOR POWER REQUIREMENT
1000
1400 1800 2200
GAS TURBINE INLET TEMPERATURE, °F
2600
3000
232
-------
2200 ^_
2100 . _
2000 . _
1900 . _
1800 . _
FIGURE 10.12
TYPICAL STEAM GENERATION CURVE FOR CYCLE 1
BOILER
GAS
COOLING
EXHAUST GAS FLOW = 536,205 MPH
GAS TURBINE INLET TEMPERATURE = 1800°F
GAS TURBINE EXIT TEMPERATURE = 1120°F
STEAM FLOW RATE
3,550.000 LB/HR
3,150.000 LB/HR
TURBINE
HP
I .P. & L.P.
POWER GENERATION
163
; 464
STEAM EXPANSION TO GASIFIER PRESSURE 6
NET POWER GENERATION 633
(400,000 LB/HR OF STEAM TO GASIFIER)
Q.
HI
1100 - _
1000 . _
900 . J
800 . _
700 . _
EXHAUST
GAS
COOLING
600 . _
500 ._
400 . _
300 ._
200 - -
100 . _
STACK 240°F
34567
HEAT EXCHANGE. 109 BTU/HR
233
10
-------
FIGURE 10.13
TYPICAL STEAM GENERATION CURVE FOR CYCLE 2
2200 ^_
2100
2000 ._
1900 ._
BURNER
GAS
COOLING
EXHAUST GAS FLOW = 546,571 MPH
GAS TURBINE INLET TEMPERATURE = 1920°F
GAS TURBINE EXIT TEMPERATURE = 1200°F
STEAM FLOW RATE TURBINES
3.480,000 LB/HR
3,080,000 LB/HR
STEAM EXPANSION TO GASIFIER PRESSURE
NET POWER GENERATION
(400,000 LB/HR OF STEAM TO GASIFIER)
POWER GENERATION
160
454
6
620
LU
tr
3
1200 - -
1100
1000
900
800
700 ._
600 -_
500
REHEAT
SUPERHEAT
EXHAUST
GAS
COOLING
400 ._
PREHEAT
300 . _
200 . _
A 1 1 1-
4567
HEAT EXCHANGE, 109 BTU/HR
234
10
-------
Ill
EC
1-
tr
LU
Q.
LU
1700 __
1600
1500 ._
1400 --
1300 ._
1200
1100
1000
900
800 ._
700 ._
600 ._
500 -_
400 ._
300 -_
200 __
100 ._
FIGURE 10.14
TYPICAL STEAM GENERATION CURVE FOR CYCLE 3 OR 4
EXHAUST GAS FLOW = 451,393 MPH
GAS TURBINE INLET TEMPERATURE = 2450°F
GAS TURBINE EXIT TEMPERATURE = 1600°F
STEAM FLOW RATE TURBINES
3,200.000 LB/HR
2,800,000 LB/HR
POWER GENERATION
147
412
EXPANSION OF STEAM TO GASIFIER PRESSURE 6
NET POWER GENERATION 565
(400,000 LB/HR OF STEAM TO GASIFIER)
234
HEAT EXCHANGE, 109 BTU/HR
235
-------
FIGURE 10.15
APPROXIMATE POWER GENERATION FOR GAS TURBINE A
3000
10,000 LB MOLE/MIN. OF EXHAUST GAS FLOW
88.0% TURBINE EFFICIENCY,'EXHAUSTS AT 16 PSIA
2500 - -
m
DC
Q
LU
tr
LU
2
LU
DC
LU
S
2000 __
1500
1000 ._
GAS TURBINE
INLET PRESSURE
800 PSIA
400 PSIA
300 PSIA
200 PSIA
100 PSIA
500
1000
1500 2000 2500
GAS TURBINE INLET TEMPERATURE, °F
3000
236
-------
2000-r-
IOC PSIA
1760--
FIGURE 10.16
APPROXIMATE EXIT TEMPERATURE FOR
GAS TURBINE A
GAS TURBINE EXIT PRESSURE 16 PSIA
88% TURBINE EFFICIENCY
1500--
1250--
UJ
DC
(-
cc
UJ
Q.
UJ
X
UJ
CD
DC
D
1000--
750--
500--
250 ._
GAS TURBINE
INLET PRESSURE
200 PSIA
300 PSIA
400 PSIA
800 PSIA
1000
1500 2000 2500
GAS TURBINE INLET TEMPERATURE, °F
237
3000
-------
40
38
36
34
32
30
28
FIGURE 10.17
APPROXIMATE ENTHALPIES OF FUEL GAS,
STOICHIOMETRIC FLUE GAS AND AIR
MM BTU/1000 LB MOLES
ENTHALPY
26
24
22
20
18
16
14
12
10
STOICHIOMETRIC
FLUE GAS
FLUE GAS UP TO
1000°F IS SAME AS
STOICHIOMETRIC
FLUE GAS
I
I
I
J
200 400 600 800 1000 1200 1400 1600 1800 2000 2200 2400 2600 2800 3fi
TEMPERATURE. °F
-------
FIGURE 10.18
COST OF CONVENTIONAL COAL-FIRED STEAM POWER PLANT
\u
I
300
250
200
150
100
90
80
70
60
50
40
30
20
10
100
U.S. GULF COAST
END OF 1973
I
I I I I I I I
200 300 400 500 600 800 1000 1500
PLANT SIZE, MW
239
-------
11. PRESSURIZED FLUIDIZED BED STEAM GENERATOR WITH DRY DOLOMITE
INJECTION FOR SO REMOVAL
11.1.Process Appraisal
The fluidized bed combustion process, as evaluated and presented by
Westinghouse in their report for EPA (41) , was chosen as the basis
for the development of this model and their design was examined by
MWK Research & Engineering Development Department. The process
includes a regenerative dolomite system for SO removal. Material
and heat balances prepared by Westinghouse were in agreement with
ours, and we have accepted Westinghouse1s statement of reactions
taking place under the conditions specified by them. However, in
some areas their design was found to be unrealistic.
The regenerative system involves several large units and compressors.
The design of the solids handling systems would present enormous
difficulties. In addition, experimental results show that the re-
generative efficiency falls off markedly with the number of cycles.
The costs for a dolomite regeneration system appear to be unrealistic,
and we do not feel convinced by Westinghouse1 s low figures for an
item which has yet to be constructed commercially. However, in view
of the decision to drop the regenerative scheme and concentrate on
once through dolomite design, the costs and problems of a regenera-
tive system are academic.
To ensure adequate turn down facility for the pressurized fluidized
bed steam generator, it is necessary to use four boiler modules,
which means that the concept becomes less desirable for small plant
sizes and industrial applications.
11.2.Process Description
The vertical pressure vessel of the fluidized boiler has four exchangers,
mounted one above another. The preheater section is at the bottom.
Above this are the evaporator, superheater, and reheater beds. The
240
-------
pulverized coal is introduced at the bottom of the bed, and about
six times the stoichiometric amount of dolomite (limestone) enters
at the top of the bed to react with sulfur in the coal. About
10% excess air is supplied to the fluidized bed, giving a superficial
velocity 6-9 ft/sec and a bed temperature of 1750°F. The entrained
solids are recycled to a carbon burn-up cell in the combustor vessel
itself, which operates at 79% excess air at 2000°F. The overall com-
bustor unit takes 15% excess air and about 1% of the carbon is even-
tually lost without combustion.
The flue gases from the CBC are passed through the second stage of a
particulate separator, before entering the gas turbines. Flue gases
enter the gas turbine at 1600°F and leave at 900°F, generating 490
megawatt. The sensible heat of the turbine effluent is used to
preheat the feed water in two stages. The flue gases enter the
stack at 200°F.
SO
_ evolved during combustion reacts with lime to form CaSO.:
CaCO3 -»•' CaO
CaO + S0_ + 1/2 02 -»•
The regeneration flowsheet is as described by Westinghouse.
Sulfated dolomite from the fluidized bed boiler is converted back to
carbonate by reducing to calcium sulfide and subsequent regeneration
with steam and CO-.
CaS04 + 4CO -»• CaS +
CaS04 + 4H_ -»• CaS + 4H2
-------
and 135 psia. The producer gas is generated by oxidizing coal with
an air-steam mixture. Steam controls the gas temperature and pro-
vides H2 for reduction.
Feed gas for the second stage regenerator is obtained by purifying
CO_ from boiler flue gas. A slip-steam from stack gas is compressed
to 135 psia and absorbed in regenerable hot carbonate. CO_ is then
stripped from the carbonate and cooled to 200°F before recompression
to 180 psia. It is then fed to the H-S generator. Rich H_S leaving
the generator at 1100°F and 165 psia is expanded through a turbine
to 2 atmosphere before sending to the Glaus unit.
242
-------
11.3 Conclusions
A fluidized bed gives better heat transfer and more uniform temperature
distribution. Surface area requirements are reduced by 60-70% due to
the high heat transfer rate. Pressurized fluidized bed operation has
certain other advantages to conventional boilers:
1. It can burn low grade, high sulfur coals
efficiently while conforming to stringent
air pollution control regulations.
2. NO emissions are reduced substantially.
J\,
3. Cycle efficiencies of 35-39% can be
achieved with dry dolomite injection
for SO removal.
Although this design gives a higher gas-steam combined cycle effi-
ciency than a conventional steam cycle with stack gas scrubbing,
there are certain limitations and problem areas for pressurized indus-
trial boilers which need development work.
The major areas which need further consideration are:
1. More stringent particulate removal is needed before the
gas turbine.
2. High temperature and high pressure piping, valving and
ducting, a particulate removal system, plus coal and dolomite
feeding systems have to be used. This is inherently expensive.
3. Turn down is the big problem in pressurized fluidized bed
operation. As shown by Westinghouse data, even by using
four modules there is a discontinuity in the turn down.
4. Nothing definite has been established about regenerative
efficiency of dolomite with time.
11.4 Addendum
After our review was completed, Westinghouse Corporation issued a
243
-------
second set of reports on the evaluation of the fluidized bed combus-
tion process (EPA 650/2-73-D48 a, b, c, and d, December, 1973).
These reports contain information on sorbent requirements for a
once-through sulfur removal system, regeneration system costs, re-
generation system potential, turn down capabilities and development
requirements . The readers are encouraged to refer to this set of
reports for the latest information.
244
-------
TABLE 11.1
to
GAS S TREAT-IS
STREAM
Gl
G2
G3
G4
G5
G6
G7
G8
G9
G10
Gil
TEMP
80
700
700
640
1500
1500
1600
300
200
1100
1000
PRESS
PSIA
14.7
150
150
150
135
116
150
135
19
165
2400/475
FLOW
MPH/HR H2
287,500
278,000
9,500
750
13,173 8.0
10,630 1.4
295,250
34,120
10,050
7,642
295,278
H2°
-
• -
-
100
10.7
21.7
8.8
8.8
63.7
73.3
100
MOLE %
£Q CO2 N, °_2 H2S
- 79 21 -
79 21
79 21
_ _
16.3 7.4 57.1 - 0.5
0.5 5.5 70.9
14.3 74.3 2.6
14.3 74.3 2.6
36.3
16.2 - - 10.5
_ _
-------
TABLE 11.2
SOLID STREAMS
to
£*
CD
STREAM DESCRIPTION
SI Total coal feed to plant
S2 Coal feed to combustor
S3 Coal feed to gas generator
S4 Dolomite make up
S5 Regenerated stone to combustor
S6 Sulfated dolomite to reducer
S7 Spent stone purge
FLOWRATE
Ib/hr
710,000 DAF
662,000 DAF
48,000 DAF
96 ,000
700,000
630,000
81,000
COMPOSITION
WT%
C 78, H 5.5, 0 11.0, S 4.0, N1.5 DAFB
C 78, H 5.5, 0 11.0, S 4.0, N1.5 DAFB
C 78, H 5.5, O 11.0, S 4.0, N1.5 DAFB
80
CaC03/ 20 Mgco
60 CaCO-, 20 CaO, 20 MgO
16 CaS04, 64 CaO, 20 MgO
60 CaC03/ 20 CaO, 20 MgO
-------
TABLE 11.3
Power Generation of an FBC Combined Cycle Plant
GENERATION MEGAWATTS
Net Steam Cycle 916
Gas Turbine 517
Reducer Reactor Effluent Turbine 16
H-S Generator Reactor Effluent Turbine 10
1459
REQUIREMENTS
Air Compressor (including air to producer gas generator) 366
CO- Compressor 12
Flue Gas Slipstream Compressor 36
Auxiliary Power Other Than Steam Cycle 10
Equivalent Power of Steam to Producer Gas Generator
and CO Stripper Reboiler 15
439
Net Power Generation =1020 megawatts.
Overall Cycle Efficiency = 34.4% (HHV of coal)
NOTES: 1). Plant uses regenerative dolomite system for sulfur control
2). Total plant feed is 710,000 Ib/hr DAF bituminous coal
containing 4% sulfur
247
-------
TABLE 11.4
Heat to Steam Cycle
Heat Losses From Combustor % of Coal HHV
Radiation and Convection 1.8
Incomplete Combustion 1.5
Heats of Calcining Reactions etc. 0.4
Heat to bring Dolomite up to 1700°F 0.3
Heat Loss by Hot Ash . 0.4
Heat Loss by Transferences between Combustor
and Regenerator 0 . 6
Total 5.0
HEAT IN WITH RAW MATERIALS MMBTU/HR
HHV of coal to combustor 9,400
Enthalpy of air @ 700°F 1,250
10,650
Less heat losses above -470
Net useful heat 10,180
HEAT OUT WITH COMBUSTED GASES
Sensible @ 1600°F 3,697
Latent Ht. of water 488
4,185
HEAT INTO STEAM CYCLE = 10,180 - 4185
= 5,995 MMBTU/HR
248
-------
TABLE 11.5
Cost of a 1000 Net Megawatt FBC Combined Cycle Plant
POWER GENERATION TOTAL PLANT
INVESTMENT, TPI
(M$)
Coal handling and injection system 15,000
Pressurized corabustor/boiler 50,000
Steam-turbo generator and condensers 44,000
Air compressor/gas turbine/generator 25,000
Land and Structures 14", 000
Accessory electrical equipment 8,000
Transmission plant 4,000
160,000
REGENERATION OF DOLOMITE
Producer gas generator 2,000
CaSO. reducer unit 5,000
H2S generator unit 5,000
Sulfur Recovery 5,000
CO 2 absorber/stripper unit 4,000
Compressor and turbines 4,000
Other offsites and solids transportation 5,000
30,000
Total for dolomite regeneration and power station is $190 million.
NOTES: 1). Plant uses a regenerative dolomite system for sulfur control,
2). Costs are Gulf Coast, end of 1973.
249
-------
C02
COMPRESSOR
12
IGW
G9
C02
ABSORBER-STRIPPER
REBOILER DUTY
175 MM BTU/HR
f
[
FIGURE 11.1
FLOW DIAGRAM FOR FLUIDIZED
PRESSURIZED COAL COMBUSTOR POWER GENERATOR
TO
STACK
G8
36
MGW
STACK
GAS
COMPRESSOR
WATER
180 PSIA
5CO°F
S5
MAKEUP
DOLOMITE!
SPENT
STONE
S4
S7
H2S
GENERATOR
On
O
TO CLAUS PLANT
G10
H2S GENERATOR
EFFLUENT TURBINE
REDUCER
EFFLUENT
TURBINE
STRIPPER
|p REBOILER
3JTACK
100°F
850°F
45 MM BTU/HR
CaS
CaS04
REDUCER
REACTOR
G7
COMBUSTOR
AND CARBON
CLEAN-UP CELL
S6
STEAM
CYCLE
NET
GENERATION
MGW
S2
G2
1600°F
1100
MM BTU/HR
500
MM BTU/HR
8FW PREHEAT
GAS
TURBINE
SI
COAL
G5
S3
PRODUCER
GAS
GENERATOR
T
G3
MGW
AIR
COMPRESSOR
G4
STEAM
ASH
-------
CUMBUSTOR
H 1521
H 1306
P 2400
T 1000
H 1463
M
45 9
P 475
T 1000
H 1521
FIGURE 11.2
STEAM BALANCE FOR PRESSURIZED COMBUSTOR
(THE ENTIRE PLANT RECEIVES 710.000 LB/HR
DAFB OF BITUMINOUS COAL
THIS CORRESPONDS TO THE COAL FEED TO
A CONVENTIONAL 1000 NET MEGAWATT POWER
PLANT FITTED WITH WELLMAN-LORD SCRUBBING)
M
35 -2
P 3000
T 550
H 550
1100 MM BTU/HR
TURBINE
EXHAUST
900°F
T 550
H 550
A
B
E
K
H
J
FLOW
MM LB/HR
4-515
0 220
0 580
4 735
5-315
5-315
STEAM CYCLE
NET GENERATION
916 MEGAWATTS
P 180
T 750
H 1401
500 MM BTU/HR
420°F
T 370
H 343
T 246
H 214
BFW PUMP 25
OTHER AUX J5
40
525
-40
485
NET
MEGAWATT
H 1018
8% WATER
VACUUM
CONDENSER
1-1/2" Hg
92°F
P 4
T 155
H 1103
2-4%
WATER
p 0.74
T 92
H60
P PSIA
T °F
H BTU/LB
M MGW/MM LB/ HR
STEAM COILS DUCT
5995 MM BTU/HR
TURBINE EFFICIENCY
88%
420°F
STACK
GAS
T 140
H 108
T 92
H 60
200°F
-------
FIGURE 11.3
GAS TURBINE EXHAUST COOLING CURVE
GAS FLOW RATE 295.250 LB MOLES/HR
LU
CC
CC
LU
o.
Ill
1000 _
900
800 - -
700 ._
600 .-
500
400 --
300 .-
200
100 . _
STACK EXIT TEMPERATURE 2(1
WATER ENTERS AT 140°F
-I 1 1 \
\ \-
200 400 600 800 1000 1200
HEAT EXCHANGE, MM BTU/HR
1400
1600
1800
252
-------
12. REFERENCES
1. "Steam Electric Plant Factors", National Coal Association,
Twenty-second Edition, December 1972.
2. "Systems Study For the Control of Emissions - Primary
Nonferrous Smelting Industry", For Division of Process
Control Engineering, National Air Pollution Control
Administration, by Arthur G. McKee & Company, San Francisco,
June 1969.
3. "Engineering Analysis of Emissions Control Technology
For Sulfuric Acid Manufacturing Processes", For Division
of Process Control Engineering, National Air Pollution
Control Administration, by Chemical Construction Corporation,
New York, March 1970.
4. "Characterization of Claus Plant Emissions", For Control
Systems Laboratory, National Environmental Research Center,
by Process Research Incorporated, April 1973.
5. "Industrial Growth Forcasts", Task 16 Final Report, by
Stanford Research Institute, Contract 68-02-1308, Sept., 1974
6. "Data File of Nationwide Emissions, 1971", Office of
Air Quality Planning & Standards, U. S. Environmental
Protection Agency, May, 1973.
7. "Energy Scenarios Consumption Considerations", Inter
Technology Corporation, July 11, 1973.
8. Rochelle, G. T., "SO_ Control Technology For Combustion
Sources", Task 6 Final Report, submitted to Control
System Laboratory, EPA, by M.W. Kellogg Company,
Contract 68-02-1308, September 1974.
9. Guthrie, K. M."Data and Technique for Preliminary Capital
Cost Estimating',' Chemical Engineering, March 1969.
10. Engineering News-Record, September 21, 1972.
11. Mendell, Otto,"How Location Affects U. S. Plant -
Construction Costs"»Chemical Engineering, December 21,
1972
12. "Labor Productivity Factors, Contractor Consensus - First
Quarter 1973", M.W. Kellogg Company (confidential).
253
-------
12. REFERENCES (CONT'D)
13. "The Supply - Technical Advisory Task Force - Synthetic
Gas From Coal", Final Report, April 1973.
14. Bauman, H. Carl, "Fundamentals of Cost Engineering in the
Chemical Industry", Reinhold Publishing Corporation, New
York.
15. "Evaluation of SO,, - Control Processes", Task No. 5
Final Report, Submitted to Environmental Protection
Agency, Office of Air Program, by M.W. Kellogg Company,
Contract No. CPA 70-68, October 15, 1971.
16. "A Process Cost Estimate of Limestone Slurry Scrubbing
of Flue Gas", Parts I & II, prepared for Office of Research
and Monitoring, Environmental Protection Agency, by
Catalytic Inc., Contract No. 68-02-0241, January 1973.
17. "Applicability of SO2 - Control Processes to Power
Plants" Task No. 11 Final Report, prepared for Office
of Research and Monitoring, U. S. Environmental Protection
Agency, by M.W. Kellogg Company, Contract CPA 70-68, Nov., 1972
18. "Evaluation of the Controllability of Power Plants Having
a Significant Impact on, Air Quality Standard", Task No. 1
Final Report, prepared for Office of Air & Water Programs,
OAQPS, Environmental Protection Agency by, M.W. Kellogg
Company, February 1974.
19. "SO2 Absorption Efficiencies of the Wet-Limestone Process",
Memorandum, From Derek Shore, M. W. Kellogg Company to
W. R. Schofield, EPA/ORM, March 13, 1973.
20. Cost Estimate For Northern Indiana Public Service Co.
(NIPSCO) Demonstration Plant, by Davy Power Gas and
Allied Chemical Corporation (confidential).
21. Peters, M. S., and K. D. Timmerhaus, "Plant Design and
Economics For Chemical Engineers," McGray-Hill, New York 1968.
22. "Long Range Sulfur Supply & Demand Model", Final Report
submitted to Environmental Protection Agency by Stanford
Research Institute, Contract No. EHSD 71-13, November 1971.
23. 1972 Keystone Coal Industrial Manual.
24. Perry, J. H., Chemical Engineers' Handbook, 4th edition
McGraw-Hill, New York.
25. Lowry, H. H., "Chemistry of Coal Utilization," Supplementary
Volume, Wiley, New York, 1963.
254
-------
12. REFERENCES (CONT'D)
26. M. W. Kellogg Company Report, "Engineering Evaluation
of a Process to Produce 250 Billion BTU/Day Pipeline
Quality Gas", June 1972 (confidential).
27. El Paso Natural Gas Company application to Federal
Power Commission for Burham Coal Gasification Complex
in New Mexico, November 7, 1972
28. "The Lurgi Process - The Route to S.N.G. From Coal"
Presented at the Fourth Synthetic Pipeline Gas
Symposium, by Lurgi Mineraloltechnik GmbH at Chicago,
October 30 - 31, 1972
29. "Solvent Processing of Coal to Produce a De-Ashed
Product", Spencer Chemical Division, Gulf Oil
Corporation, Contract No. 14-01-0001-275 (OCR), 1965
30. "Economic Evaluation of a Process to Produce Ashless,
Low-Sulfur Fuel from Coal" Pittsburg and Midway Coal
Mining Company, Contract No. 14-01-0001-496 (OCR), 1969
31. Brant, V. L. and Schmid, B.K.,"Pilot Plant for De-
Ashed Coal Production", C.E.P. 65, 55 (1969).
32. "Development of a Process for Producing an Ashless,
Low-Sulfur Fuel From Coal" Pittsburg Midway Coal
Mining Company, Contract No. 14-01-0001-496 (OCR),
November 1969
33. Robson, F. L. Giramonti, A. J., Lewis, G. P. Gruber,
G., "Technological & Economic Feasibility of Advanced
Power Cycles and Methods of Producing Non-Polluting
Fuels for Utility Power Stations", United Aircraft
Research Laboratories, 1970.
34. Rudolph, Paul, F. H., "New Fosil-Fueled Power Plant
Process Based on Lurgi Pressure Gasification of Coal,"
Lurgi Mineraloltechnik GmbH, 1970.
35. "Clean Fuel Gas From Coal", Lurgi Mineraloltechnik
GmbH, Fuel Technology Division, October, 1971
36. TVA Steam Plants, Technical Monograph No. 55, Volume
3, 1963.
37. The Paradise Steam Plant, Units 1&2, Technical Report
37, TVA 1964,
255
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12. REFERENCES (CONT'D)
38. The Bull Run Steam Plant, Technical Report, TVA
39. Communication with General Electric Company, December
1973.
40. Foster, A. D. "Gas Turbine Fuels" presented at the
General Electric Gas Turbine State of Art Engineering
Seminar, June, 1971.
41. "Evaluation of the Fluidized Bed Combustion Process",
Submitted to Office of Air Programs, Environmental
Protection Agency, by Westinghouse Research Laboratories,
Pittsburg, Penn. (Contract No. CPA 70-9) Volume I, II,
III, November 1971.
42. "Evaluation of the Fluidized Bed Combustion Process",
Prepared for Office of Research & Development, U. S.
Environmental Protection Agency by Westinghouse
Laboratories, Pittsburg, Penn. (Contract No. 68-02-
0217) Volume I, II, & III, December 1973.
256
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13. APPENDICES
257
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APPENDIX A
General Cost Model Derivations
Interest During Construction (IDC)
Interest during construction represents the cost of interim
financing of a project during the design and construction period.
If the project is financed on borrowed capital, it is a real
cost to the company. If corporate funds are used, it is an
internal charge equivalent to the income which would have been
obtained if the capital had been used for short-term investment
at normal commercial interest rates.
The total interest during construction can be obtained from the
following general equation:
interest = capital x interest rate x time
where "capital" is the total construction cost of the plant, i.e.,
the total plant investment (TPI). The time period referred to
in the equation is a function of the project schedule as it
affects cash flow.
A typical project schedule can be represented as follows:
EH
cn
EH
U
w
^
8
ENGINEERING
&
DESIGNS
(A YEARS)
CONSTRUCTION
(B YEARS)
TOTAL PROJECT
W
EH
W
a
o
CJ
EH
U
EH
CD
2!
O
U
(C YEARS)
258
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APPENDIX A (cont'd)
From the General Cost Model, the cost of engineering and design
is normally about 10% of the total plant investment while construc-
tion costs about 90%. Assuming a uniform cash flow durinq the
engineering and design phase (A yrs) , the interest charge would
be approximately equal to the interest rate applied to the entire
engineering and design cost over one-half the time period (For
interest purposes, a uniform cash flow is roughly equivalent to
a single cash flow at the mid-point of the time period) . When
the engineering and design phase is completed, interest continues
to accumulate until the end of the construction period (C-A years)
Thus, the total interest charge on the engineering and design
costs, I,, is:
Ix = 0.1 TPI x i x |- + 0.1 TPI x i x (C-A)
=0.1 TPI x i [| + (C-A) ]
= 0.1 TPI x i (C-j)
The interest on construction costs, I2, would be approximately
equal to the interest rate applied to the entire cost over one-
half the construction period.
Thus:
I2 = 0.9 TPI x i x
The total interest during construction, IDC, is therefore given
by:
IDC = I, + I
= 0.1 TPI x i (C-^-) +0.9 TPI x i x j
= TPI x i [0.1 (C-|) +0.9 (|)]
259
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APPENDIX A (cont'd)
In order to obtain values for IDC for use in the General Cost
Model, typical project schedules have been assumed as follows:
ABC
(Yrs) (Yrs) (Yrs)
Stack gas scrubbing 1.5 2.5 3.0
SNG, SRC, power plants 2.5 3.5 5.0
Assuming an interest rate of 9%/year and substituting in the
preceding equation, we obtain:
IDC =0.12 TPI . Stack gas scrubbing
IDC =0.18 TPI SNG, SRC, power plants
Interest on Debt and Return on Equity (AIC)
The calculation of AIC is based on the utility method used
by the Synthetic Gas-Coal Task Force of the FPC National Gas
Survey and illustrated in their final report (13). The method
assumes that the total capital required, TCR, is split into a
percentage debt (borrowed capital) and a percentage equity (owned
capital). The debt portion is charged at the commercial interest
rate while the equity is charged at some desired net rate of
return. Depreciation covers return of capital for both the debt
and equity portions of TCR. Interest on debt and return on
equity are calculated over the life of the plant and the average
yearly value is expressed as a fraction of TCR.
At any given time, the book or asset value of a plant equals its
original cost (TCR) less the total accrued depreciation. This
quantity, which represents the debt still outstanding plus the
equity capital yet to be recovered, is the rate base upon which
interest on debt and return on equity are calculated. For any
given year, the average rate base, (BR)., equals TCR less the
260
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APPENDIX A (cont'd)
accrued depreciation at the mid-point of the year. Thus, (AIC) .,
the annual interest on debt plus return on equity for year i, is
calculated as follows:
(AIC) . = (BR) . x fraction debt x fraction annual interest
1 rate
+ (BR) . x fraction equity x fraction annual net
rate of return
The average value of AIC over the life of the plant, expressed
as a fraction of TCR, is:
(AIC,
N
AIC = -=• - TCR
This is the equation which is used in the General Cost Model.
Typical values were assumed for fraction debt, fraction equity,
interest rate, and net rate of return to obtain a numerical value
for AIC for use in the General Cost Model. A sample calculation
of AIC, based on these assumed values, is shown in Table A-l on
a year-by-year basis.
Note that since straight- line depreciation has been used, the
rate base, and therefore the return on rate base, decreases
linearly with time. Thus, AIC alternately could be calculated
by using the average rate base over the plant life. This average
base is equal to TCR less one-half of the total depreciation:
(BR)AV(, = TCR - j (TCR-WKC)
= i- (TCR + WKC)
Substituting the assumed values of TCR and WKC from Table A-l:
(BR)AVG = 1 (100,000 + 3,000)
= $51,500M
261
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APPENDIX A (cont'd)
Using the assumed values in Table A-l for percent debt, percent
equity, interest rate, and rate of return, AIC can be calculated
as follows:
AIC = (0.75 x 0.09 + 0.25 x 0.15) (BR) _.._
AVlj
= 0.105 x 51,500
= $5407 M/Year
It should be noted that plant life does not enter into the cal-
culation. Thus, the expression for AIC used in the General
Cost Model is valid for all types of plants (stack gas scrubbing
units, SNG plants, power plants, etc.).
Federal Income Tax (AFT)
From Table A-l, the average net (after tax) return on equity is
$1,931M. Expressed as a fraction of TCR, this is:
net return = ioo ' TCR
= 0.0193 TCR
Assuming an income tax rate of 48%, the net return represents
52% of the gross (pre-tax) return. The federal income tax is
therefore given by:
AFT = x 0.48
= 0.0193 TCR x -p||
= 0.018 TCR
262
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TAEI.S A-l
LE'CAI.CMI-.Yi'IO:: O!' AfC
Ajs-.:ir.ptions
Total Capital Required (TCP.)
Working Capital (WKC)
Plant Life
2ebt
Equity
Interest rate on <:
PSTUR:; or; RATE BASE
(3 * 4)
-
$ 10,245 M
9,736
9,227
8,717
8,209
7.639
7.190
6.6.71
6.171
5,662
5,153
4,643
4,135
3,625
3,116
2,607
2,097
1,588
1,079
5C9
108,149
5 ,407
c)
o^p-.zf;i.-.r:o:;
•:•:. :s x
ITCR-WX:) )
_
$ 4, 850 M
4.850
4,35'.-
4.35:
4,85C
4, as:
4,850
4,850
4,850
4,850
4,853
4.850
4,850
4,850
4,850
4.850
4,850
4,850
4,850
4,850
97,000
IC
5,407
x TCR - 0.054 TCR
263
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APPENDIX B
WET LIMESTONE PROCESS-CATALYTIC INC. ESTIMATE, 500 MEGAWATT
EQUIPMENT COSTS (MATERIAL & SUBCONTRACTS)
(Inflation index from mid 72 to the end of 73 = 1.08)
Scrubbing System
Mid 1972
End of 1973
4 Venturi Scrubbers 253,000
(380,000 ACFM at the inlet to each)
4 Two-stage TCA1s 840,000
4 Sumps for Venturis & TCA's 254,000
4 Horizontal Chevron Entrainment Separators 531,000
4 Venturi Tanks 85,000
4 TCA Tanks 121,000
4 Venturi Tank Agitators 26,400
4 TCA Tank Agitators 31,000
4 Sets of Venturi Recirculation Pumps 47,600
4 Sets of TCA Recirculation Pumps 91,000
Ammonia Injection System 10,000
Entrainment Separator Recir. Tank 28,500
Entrainment Separator Recir. Pumps 30,900
2,349,400
273,000
908,000
274,000
574,000
92,000
131,000
28,500
33,500
51,500
98,300
10,800
30,800
33,400
2,538,800
Flue Gas Reheat and. Discharge
Mid 1972
End of 1973
Ductwork including dampers etc.
4 I.D. Fans and Motors
(360,000 ACFM, 37 inchw.g., 3000 BHP)
4 Reheater Burner Units
Fuel Oil Tankage and Loading Pump
1,003,200
332,000
156,000
69,270
1,560,470
1,085,000
359 ,000
169,000
75,000
1,688,000
264
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APPENDIX B (con't)
WET LIMESTONE PROCESS-CATALYTIC INC. ESTIMATE, 500 MEGAWATT
EQUIPMENT COSTS (MATERIAL & SUBCONTRACTS)
(Inflation index from mid 72 to the end of 73 = 1.08)
Limestone Hand1ing and Slurry Preparation
Mid 1972
End of 1973
Limestone Silo Conveyor & Stockpile Feeder
Limestone Storage Silo with 3 Cones
3 Limestone Weigh Feeders
3 Tube Mill Wet Grinders
Tube Mill Air Compressor
Tube Mill Surge Tank
Limestone Slurry Transfer Pumps
Limestone Slurry Hold Tank
Limestone Slurry Tank Agititator
Limestone Feed Pumps
57,050
76,000
21,400
550,000
12,500
1,500
2,200
46,300
26,900
2,900
796,750
61,700
82,000
23,100
595,000
13,500
1,600
2,400
50,000
29,100
3,200
861,600
Waste Disposal
Mid 1972
End of 1973
Sumps and Tankage 6,150
Pumps & Drives (inc. process water pumps) 31,840
Separating Pond (250 acres 50 ft. deep) 3,694,000
Pond Located in Cincinnati.)
3,731,990
6,700
34,500
4,000,000
4,041,200
265
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APPENDIX C
WET LIMESTONE PROCESS-CATALYTIC INC. ESTIMATE,
LABOR AND MATERIAL FACTORS
MAJOR EQUIPMENT COSTS
Material
Subcontract
Total
CATALYTIC INC
Part I P65
$
2,925,300
1,819,300
4,744,600
EC chemical process
ES solid handling
Total
This Report
$
3,947,900
796 ,700
4,744,600
FIELD LABOR COSTS (Cincinnati with location factor = 1.53)
$ $
LC chemical process 2,363,000
LS solid handling 212,000
Total 2,575,000 Total 2,575,000
OTHER MATERIAL COSTS
(Piping, instrumentation, electrical, civil etc.)
Total Material 6,218,700 MC chemical process
Maj. Equip. Mat.2,925,300 MS solid handling
Other Material 3,293,400
Total
3,225,700
67,700
3,293,400
LC = 0 .60 EC
LS = 0.27 ES
MC = 0.82 EC
MS = 0.09 ES
266
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APPENDIX D
A SUMMARY OF THE COMMENTS MADE DURING AND AFTER
THE PRESENTATION BY MWK TO EPA, DECEMBER 13, 1973
1. There was agreement with the $1.2 to $1.4/MMBtu cost given for
SNG,but surprise by some to find that this cost did not represent
almost any location at a higher coal cost. There was disbelief
that a contract to sell coal at anything like $3/ton would ever
be signed. It was stressed by MWK that the utility industry
in 1970 paid an average of over $7/ton.
2. Mr. K. Janes expressed the view that costs for solvent refined
coal production while higher than those given in the Stearns-Roger
estimate are still on the low side.
MWK agreed with this and pointed out that the accuracy of
the plant investment was only within + 30% and most probably
the figure would be as underestimate if an actual commercial
design were ever costed.
3. There was considerable interest in the Lurgi combined cycle
presentation and a general agreement with the cycle efficiencies
given in the presentation. The feeling expressed by a few
people was that Lurgi was not as far advanced in the field of
gasification and combined power cycles as Lurgi publications
say. It was stated by Mr. P. Spaite that there have been
considerable technical difficulties with Lurgi's Steag 165 megawatt
unit and the gasification unit was not as reliable as was needed
for a combined cycle. It was felt that reliability and better
control of coal feed rather than cost reduction were the main
reasons for development of other gasifier designs. The opinion
that the hot carbonate purification unit was not proved in this
service was expressed by at least two people. It was pointed out
by MWK that Benfield Corporation felt confident of their design
267
-------
APPENDIX D (con't)
and that Lurgi also had a hot carbonate design. Several people
wondered why the costs for the hot carbonate unit were much
lower than those for the SNG purification unit and the Wellman/Allied
stack gas scrubbing units. It was pointed out by MWK that
the purpose of the Rectisol unit used in the SNG plant was much
different from that of the hot carbonate unit. The rectisol
unit had to remove almost all of the CC- in the gas stream compared
to the hot carbonate's 30%. The main difference was that the accept-
able sulfur level for the methanator catalyst was 0.1 ppm whereas
the hot carbonate left as much as 500 ppm H2S. The Rectisol
unit was more complicated and required a refrigeration unit.
The coal feed to the combined cycle Lurgi unit was approximately
70% of that to the Lurgi SNG plant. Again, the function of the
Wellman/Allied unit was much different in that it handled a flue
gas with a flowrate approximately 2.5 times more (1000 megawatt
plant) than that of the fuel gas from the low Btu Lurgi unit.
The costs of the hot carbonate unit should also be added to the
cost of the sulfur recovery unit before even a broad comparison
could be made to the Wellman/Allied plant.
4. There was interest in the relative recovered energies of the
SNG unit, the SRC unit and the power plants. MWK stated that
the HHV of the product SNG was 58% of the HHV of the input coal
and the corresponding figure was 79% for the SRC units compared
to the best power plant efficiency of 40%. There was strong feeling
by several people present that direct comparison was misleading
and the relative forms of the energy had to be taken into account.
SNG required power for transportation and electrical heating devices
were more efficient than gas heating devices. There was agreement
to some extent by MWK, but it was stated that there was a need
to study transportation costs of the various alternatives and
establish the final efficiency after consumption. MWK also
pointed out that the transportation of low sulfur Wyoming coal
to the Eastern states ought to be a more sensible use of the
coal rather than gasification.
268
-------
APPENDIX D (cont'd)
5. There was agreement with the conclusions expressed about the
Westinghouse fluidized, pressurized combustor with regenerative
dolomite sulfur removal. It was felt by MWK that regeneration
was not technically possible to any consistent level by a commercial
plant of the Westinghouse design. On top of this the costs for
a finallized design could be many million dollars more than those
presented. This view was substantiated by a statement by EPA
that Westinghouse had now dropped the idea of regeneration and
interest was now directed towards the pressurized combustor
with sulfur removal either by a once through throwaway dolomite
system or by stack gas scrubbing. It was stated by MWK that
the stack gas scrubbing alternative would produce overall cycle
erficiencies about the same as a conventional power plant with
staxck gas scrubbing. The only possible advantage would be one of
reduced cost of the pressurized fluidized boiler over the conventional
boiler\. It is obviously in Westinghouse's best interest to establish
whetherXthis is in fact so at the earliest possible time.
It is likely that the once through dolomite alternative would
have a higher overall cycle efficiency than the conventional
plant with Wellman/Allied stack gas scrubbing; it may possibly be
better than the conventional plant with Wet Limestone Scrubbing.
A critical factor may be how many times the stiochiometric flow
of dolomite is required. Waste disposal may be even more expensive
than with the wet limestone process. The main areas for establishing
concrete facts at an early date are the cost of the equivalent
sized pressurized boiler compared to the conventional unit and
the efficiency of sulfur removal at a-stated dolomite flowrate.
269
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TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
i. REPORT ivo.
EPA-650/2-74-098
2.
3. RECIPIENT'S ACCESSIOr»NO.
4. TITLE AND SUBTITLE
Evaluation of R$D Investment Alternatives for
SOx Air Pollution Control Processes
5. REPORT DATE
September 1974
6. PERFORMING ORGANIZATION CODE
7. AUTHOR(S)
D. Shore, J.J. O'Donnell, and F.K. Chan
8. PERFORMING ORGANIZATION REPORT NO.
9. PERFORMING ORG \NIZATION NAME AND ADDRESS
The M. W. Kellogg Company
Research and Engineering Development
Houston, Texas 77046
10. PROGRAM ELEMENT NO.
1AB013; ROAP 21ADE-029
11. CONTRACT/GRANT NO.
68-02-1308 (Task 7)
12. SPONSORING AGENCY NAME AND ADDRESS
EPA, Office of Research and Development
NERC-RTP, Control Systems Laboratory
Research Triangle Park, NC 27711
13. TYPE OF REPORT AND PERIOD COVERED
13. TYPE OF REPORT AND PE
Final; 10/72-12/73
14. SPONSORING AGENCY CODE
15. SUPPLEMENTARY NOTES
16. ABSTRACT
The report presents data on sulfur oxide (SOx) emissions from five major
source groups: utility plants, industrial boilers, non-ferrous smelters, sulfuric
acid plants, and sulfur (Claus) plants. For all source groups studied, the bulk of the
SOx emissions comes from a relatively small number of the largest plants. The
report also includes evaluations of several different sulfur control systems, incl-
uding stack gas scrubbing (wet limestone process and Wellman/Allied system),
substitute natural gas, solvent refined coal, Lurgi gasification with a combined
power cycle, and pressurized fluidized-bed combustion with a combined power
cycle. Process and cost models and/or economics are presented for each system.
Cost models for the stack gas scrubbing processes were applied to existing utility
plants in the U.S. and the results analyzed.
17.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.IDENTIFIERS/OPEN ENDED TERMS C. COSATI Field/Group
Air Pollution
Sulfur Oxides
Cost Effectiveness
Boilers
Electric Utilities
Smelters
Sulfuric Acid
Air Pollution Control
Stationary Sources
Industrial Boilers
Claus Plants
13B
07B
14A
13A
11F
18. DISTRIBUTION STATEMENT
Unlimited
19. SECURITY CLASS (ThisReport)'
Unclassified
21. NO. OF PAGES
288
20. SECURITY CLASS (Thispage)
Unclassified
22. PRICE
EPA Form 2220-1 (9-73)
270
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