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

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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.

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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.

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        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:

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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

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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

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     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

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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

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    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

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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

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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

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         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

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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

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    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

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                               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

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                           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

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                              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

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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

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     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

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 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

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                             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

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                                                               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.

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                                                               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.

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               FIGURE 4.3

LOCATION FACTORS FOR  SELECTED CITIES
                                                                         1.82
                     MISSOURI
                      KANSAS
                        CITY
                         1.37

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I-J
01
                                                         FIGURE  4.4


                                            AVERAGE LOCATION FACTORS BY STATE
                                                                                                                 1.3
                                              Y / /i / / QKLAHOHf A
               > 1.75

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                              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

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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

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                                      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

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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

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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

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                                               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

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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

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    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

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                                 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

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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

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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

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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

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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

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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

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     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

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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

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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

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                             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

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                                           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

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                                                    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

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                                    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

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                   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

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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

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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

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     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

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                               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

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                                      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

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  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

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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

-------
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

-------
                                                            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

-------
                                   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

-------
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

-------
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

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                               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

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                              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

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                              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

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                               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

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                          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

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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

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                                     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
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     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
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cc
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        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

-------
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5
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                                 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

-------
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                               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


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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
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      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
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     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
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                   50             100             150            200


            EQUIVALENT INCREMENTAL COST WHICH COULD BE PAID FOR CLEAN FUEL, CENTS/MM BTU
                                                                              25O
                                                                                          300
                                          160

-------
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                               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
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p  200
0.
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tE
2

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   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


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CD

5
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                                      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
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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

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Z
LU

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CO
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     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
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2
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     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

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U




§
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     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

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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

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                                            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

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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

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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

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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

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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

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                                        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

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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

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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

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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

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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

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         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

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       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

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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

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                     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

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                                         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

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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

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                                                                 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

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                         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

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                             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

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                     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

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                           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

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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

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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

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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

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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

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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

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                                 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

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                                                               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

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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

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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

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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

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                            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

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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

-------
                  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

-------
13.  APPENDICES
         257

-------
                            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

-------
                       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

-------
                         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

-------
                      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

-------
                         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

-------
                                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

-------
                           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

-------
                           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

-------
                            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

-------
                             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

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                         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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