EPA-650/2-75-027-B
March 1975
Environmental Protection Technology Series
e n 1
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EPA-650/2-75-027-b
FLUIDIZED BED COMBUSTION
PROCESS EVALUATION
(PHASE I - RESIDUAL OIL GASIFICATION/DESULFURIZATION
DEMONSTRATION AT ATMOSPHERIC PRESSURE)
VOLUME II - APPENDICES
by
D.L. Kcairns, R.A. Ncwby , E.J. Vidl. E.P. O'Neill.
C.H. Peterson, C.C. Sun, C.D. Buscagha, and D.H. Archer
Westinghouse Research Laboratories
Beulah Road, Churchill Borough,
Pittsburgh, Pennsylvania 15235
Contract No. 68-02-0605
ROAP No. 21ADB-009
Program Element No. 1AB013
EPA Project Officer: P.P.Turner
Control Systems Laboratory
National Environmental Research Center
Research Triangle Park, North Carolina 27711
Prepared for
U.S. ENVIRONMENTAL PROTECTION AGENCY
OFFICE OF RESEARCH AND DEVELOPMENT
WASHINGTON, D. C. 20460
March 1975
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EPA REVIEW NOTICE
This report has been reviewed by the National Environmental Research
Center - Research Triangle Park, Office of Research and Development,
EPA, and approved for publication. Approval does not signify that the
contents necessarily reflect the views and policies of the Environmental
Protection Agency, nor does mention of trade names or commercial
products constitute endorsement or recommendation for use.
RESEARCH REPORTING SERIES
Research reports of the Office of Research and Development, U.S. Environ-
mental Protection Agency, have'been grouped into series. These broad
categories were established to facilitate further development and applica-
tion of environmental technology. Elimination of traditional grouping was
consciously planned to foster technology transfer and maximum interface
in related fields. These series are:
1. ENVIRONMENTAL HEALTH EFFECTS RESEARCH
2. ENVIRONMENTAL PROTECTION TECHNOLOGY
3. ECOLOGICAL RESEARCH
4. ENVIRONMENTAL MONITORING
5. SOCIOECONOM1C ENVIRONMENTAL STUDIES
6. SCIENTIFIC AND TECHNICAL ASSESSMENT REPORTS
9. MISCELLANEOUS
This report has been assigned to the ENVIRONMENTAL PROTECTION
TECHNOLOGY series. This series describes research performed to
develop and demonstrate instrumentation, equipment and methodology
to repair or prevent environmental degradation from point and non-
point sources of pollution. This work provides the new or improved
technology required for the control and treatment of pollution sources
to meet environmental quality standards.
This document is available to the public for sale through the National
Technical Information Service, Springfield, Virginia 22161.
Publication No. EPA-650/2-75-027-b
11
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ABSTRACT
This volume contains appendices resulting from the design,
evaluation, and experimental work carried out by Westinghouse to develop
an atmospheric-pressure fluidized bed residual oil gasification process
for power generation. The process, conceived by Esso Research Centre,
Abingdon, England, as the chemically active fluidized bed (CAFB) process,
produces a clean, low heating-value fuel gas for firing in a conventional
boiler. The integrated process, previously operated successfully in a
750 kw pilot plant unit, has demonstrated the ability to meet environ-
mental emission standards for sulfur oxides, nitrogen oxides, and
particulates.
Work carried out under this contract was directed toward a
commercial demonstration of the process. A preliminary design and cost
estimate for a 50 MWe atmospheric-pressure fluidized bed residual oil
gasification/desulfurization demonstration plant was completed for the
50 MWe unit No.12 at the Manchester Street Station, Narragansett
Electric Co., Providence, RI. The design and cost estimate provide
sufficient detail to proceed with the detailed design and construction.
Boiler performance with hot, low heating-value gas is assessed. Experi-
mental programs were carried out to select limestone sorbents for the
plant and to provide design and operating criteria on spent sorbent
processing. The environmental impact of the process is assessed and
areas which require further development identified. Market projections
are presented for the process. A 200 MWe commercial plant design and
cost estimate were prepared. Capital and operating costs are presented
for commercial plants with capacities from 50 to 500 MW . Detailed
c
design and construction of the 50 MW demonstration plant utilizing
low-grade, high metal content fuels is recommended if completion of the
critical experimental tests identified and review of fuel projections
demonstrate process operability and fuel availability.
iii
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PREFACE
The Office of Research and Development (ORD) of the United States
Environmental Protection Agency (EPA) has organized and is sponsoring a
fluidized bed fuel processing program. Its purpose is to develop and
demonstrate new methods for utilizing fossil fuels to produce electrical
energy from utility power plants which meet environmental standards. These
methods should:
Meet environmental goals for sulfur dioxide (SO.), nitrogen
oxide (NO ), ash, smoke emissions, trace element emissions,
and wastes
Utilize fuel resources efficiently
Compete economically with alternative means for meeting
environmental goals.
Westinghouse Research Laboratories, under contract to ORD, is carrying out
a program to evaluate, develop, and demonstrate fluidized bed gasifica-
tion/desulfurization of residual fuel oil for power generation. This
two-volume report describes work performed from May 1973 to December 1974
under contract 68-02-0605. The work carried out during this period is
based on tasks-set forth by EPA which were completed by Westinghouse under
previous contracts. The results from these prior tasks on fluidized bed
gasification of residual oil were published in a three-volume report,
"Evaluation of the Fluidized Bed Combustion Process," in November 1971
under contract No. CPA 70-9 and as oart of a four-volume reoort, Volume TV
"Fluidized Bed Oil Gaslfication/Desulfurization," in December 1973 under
contract 68-02-0217. The previous work on residual oil gasification included:
Conceptual design and cost estimate of atmospheric-pressure
fluidized bed residual oil gaslfication/desulfurizatlon
system for utility newer generation
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Assessment of the effectiveness and economics of an atmospheric-
pressure fluidized bed residual oil gasification/desulfuriza-
tion system
Identification of a project team to demonstrate fluidized
bed residual oil gasification/desulfurization for power
generation
Identification of a boiler unit to carry out the demonstration
plant program
Evaluation of pressurized residual oil gasification/desulfur-
ization for combined cycle power generation
Provision of technical consultation and assistance on the
fluidized bed fuel processing program.
Tasks carried out under contract 68-02-0605, which are reported in
this two-volume report, have included:
Preparation of a preliminary design and cost estimate for
the 50 MW fluidized bed residual oil gasification/desulfur-
ization demonstration plant. The design is for the 50 MW
unit No. 12 at the Manchester Street Station, Narragansett
Electric Company, Providence, Rhode Island. The design and
cost estimate provide sufficient detail to proceed with the
detailed design and construction of the demonstration plant.
Evaluation of the demonstration plant boiler performance.
The boiler performance with the hot, low heating-value gas
is assessed for the demonstration plant. Modifications and
associated costs to maintain the performance are identified.
Evaluation and selection of the fluidized bed gasification
process and design options. Alternative design and operating
parameters are identified and evaluated. Recommendations
are made for the demonstration plant and for commercial plants.
Development of process data. Experimental tests were carried
out to identify candidate limestone sorbents; to obtain data
on sulfur removal, sorbent regeneration, and spent stone
processing; and to assess the environmental impact of
spent stone.
vi
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Determine market for fluidized bed residual oil gasification/
desulfurization. Factors are identified which will determine
the market potential. Market projections are made based on
a review of these factors.
Prepare commercial plant design and cost estimate. A 200 MW
commercial plant design is prepared on the basis of the
demonstration plant design work. A factored cost estimate
is made on the basis of the demonstration plant estimate.
Prepare assessment of environmental impact. Solid, liquid,
and gaseous emissions are identified and assessed to deter-
mine their environmental impact. Resource utilization is
also reviewed.
Identify Development requirements. Areas which require
further development are identified and programs projected to
obtain the necessary information to carry out the program.
Volume I contains a summary of the work performed and program
recommendations. Volume II contains appendices which provide supplemen-
tal information and detailed back-up to support the summary and recommen-
dations.
vii
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TABLE OF CONTENTS
APPENDICES
Page
A. Market Data 1
B. Preliminary Design Process Description 13
C. Demonstration Plant Site 27
D. Fluidized Bed Oil Gasification Demonstration Plant Design
Basis 35
E. Design Manual 63
F. Boiler Modifications 163
G. Environmental Impact 169
H. Commercial Plant Design 177
I. Initial Design Study 191
J. Evaluation and Selection of Reaction System Process and
Design Options 213
K. Spent Stone Processing Options 221
L. Dry Sulfation Experimental Program 271
M. Slurry Recarbonation Experimental Program 307
N. Acid Sulfation Experimental Program 325
0. Dry Recarbonation and Sintering Experimental Programs 345
P. Sulfur Recovery System Design and Evaluation 361
Q. Limestone Selection 377
R. Thermogravimetric Studies on the Sorbtion of Sulfur by
Lime 397
S. Spent Limestone Disposition 407
ix
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LIST OF FIGURES
Page
APPENDIX A
1. United States Fossil Generation Overview 11
APPENDIX B
1. Reaction System 15
2. Stone Processing System (Dry Sulfation) 19
APPENDIX C
1. Manchester Street Station Operating Floor Plan 29
2. Manchester Street Station Basement Floor Plan 31
3. Manchester Street Station Basement Floor Plan 33
APPENDIX D
1. Gasifier Fundamentals 36
2. Regenerator Fundamentals 41
3. Stone Processing Design Options 57
4. Estimated Fines Size Distribution from Boiler prior to
Existing Cyclone 61
5. Existing Cyclone Efficiency 61
APPENDIX E
1. Schematic Flow Diagram and Material Balance Base Case 79
2. Schematic Flow Diagram Material Balance Turndown Case and
Low Air/Fuel Ratio Case 81
3. Reaction System Oil Gasification-Desulfurization CAFB 85
4. Fuel Oil and Waste Stone Handling Systems 87
5. Equipment Arrangement Plan, Lower Level 105
6. Equipment Arrangement Plan, Upper Level 107
7. Equipment Arrangement, Section AA 109
8. Equipment Arrangement, Section BB and CC 111
9. Site Plan 113
10. L-l Gasifier and L-2 Regenerator 119
xi
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List of Figures (Continued)
Page
11. L-3 Absorber 121
12. L-l, L-2, L-3 Design Data 123
13. One Line Diagram 157
APPENDIX H
1. Reaction System Oil Gasification-Desulfurization CAFB 181
2. Schematic Flow Diagram and Material Balance Base Case 183
APPENDIX I
1. Initial Design Study - Reaction System 195
2. Initial Design Study - Stone Processing System 197
3. Initial Design Study - Gasifier Design 205
4. Initial Design Study - Regeneration Design 207
APPENDIX K
1. Logic Diagram for Dispoal of Spent Limestone 230
2. Oil Gasification: Spent Stone Disposal Processes 232
3. Oil Gasification, Nonregenerative Mode-Option NDN-3:
Oxidation plus Carbonation 234
4. Oil Gasification, Nonregenerative Mode-Option NDN-4:
Silica Sintering 235
5. Oil Gasification, Nonregenerative Mode-Option NWS-1:
Wet Carbonation 238
6. Oil Gasification, Nonregenerative Mode-Option NWS-2:
Wet Sulfation 239
7. Oil Gasification, Nonregenerative Mode-Option NDS-6:
Steam Recarbonation , 241
8. Oil Gasification, Regenerative Mode-Option RDN-1:
Dry Sulfation 242
9. Oil Gasification, Regenerative Mode-Option RDS-1:
Direct Disposal 244
10. Oil Gasification, Regenerative Mode-Option RDS-2:
Dry Oxidation 245
xii
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List of Figures (Continued)
Page
11. Oil Gasification, Regenerative Mode-Option RDS-3:
Oxidation plus Carbonation 246
12. Oil Gasification, Regenerative Mode-Option RDS-4:
Silica Sintering 247
13. Oil Gasification, Regenerative Mode-Option RDS-5:
Dead-Burning 248
14. Oil Gasification, Regenerative Mode-Option RWS-1:
Wet Carbonation 250
15. Oil Gasification, Regenerative Mode-Option RWS-2:
Wet Sulfation 251
16. Oil Gasification, Regenerative Mode-Option RWN-1:
Lime Scrubbing 252
17. Oil Gasification, Regenerative Mode-Option RDS-6:
Steam Re Carbonation 254
APPENDIX L
1. The Dry Sulfation Process 272
2. Rate Criteria for Sulfation of Waste in TG Experiments
at 0.5% S02 274
3. The Sulfation of Fine Particles of Calcined Limestone 276
4. First-Order Model for Fluidized Bed Sulfation of CaO 277
5. The Sulfation of Spent Sorbent from the CAFB Regenerator Bed 280
6. The Rate of Sulfation of Finely-Ground Spent Stone 283
7. Schematic Diagram of the 25 mm Quartz Reactor Test System
for Dry Sulfation 285
8. SO. Absorption Profiles of Dry Sulfation
Runs in the 25 mm Quartz Reactor 289
9. SO. Absorption Profiles for Dry Sulfation Runs at 870°C
(1000°F) 290
10. SO. Absorption Profiles for Dry Sulfation Runs at 870°C
(1000°F) 291
xiii
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List of Figures (Continued)
Page
11. Comparison of Dry Sulfation Runs 293
12. Percent of CaO Utilization as a Function of the Total S02
Feed, Showing the Effect of Bed Temperature 294
13. Percent of CaO Utilization of the Total SO Feed, Showing
the Effects of Particle Size and Gas Flow Rate 295
14. Calculated S02 Absorption Profile for Dry Sulfation
Reaction 298
15. Calculated S02 Absorption Profile for Dry Sulfation
Reaction at 870°C (1600°F) 299
16. Comparison of the Calculated Dry Sulfation Reaction Curve
and the Experimental TG Data 301
17. Fixed Bed Experimental Results 304
APPENDIX M
1. Schematic Diagram for the Slurry Recarbonation System 312
2. Neutralization Curve of 10 ml Slurry Recarbonation Run 6
Filtrate with 0.1 N HC 1 319
3. Slurry Settling Curve for Run 25 320
4a. A Typical Optical Microphotograph of the Slurry Recarbonated
Product from Run 22 (200 x) 323
4b. A Typical Electron Scanning Microphotograph of the Slurry
Recarbonated Product from Run 22 (2000 x) 323
APPENDIX N
1. Schematic Diagram for the Acid Sulfation System for Spent
Stone Processing 329
2a. An Optical Microphotograph of the Acid-Sulfated Product
from Run ASS (200 x) 339
2b. An Electron Scanning Microphotograph of the Acid-Sulfated
Product from Run ASS (500 x) 339
3. Slurry Settling Curve for Acid Run AS19 Immediately Following
the Sulfation Reaction 340
xiv
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List of Figures (Continued)
4. Slurry Settling Curve of AS19 after 4-Day Settling and
Reshaking to Form the New Slurry 342
5. Comparison of Photomicrographs of the Acid Sulfation
Products from (a) ASS and (b) AS19 (200 x) 343
APPENDIX 0
1. Recarbonation of Calcium Oxide 349
2. Rate of Recarbonation of Residual CaO in Spent Sorbent 352
3. CO. Requirement for CaO Recarbonation 354
4. Effect of Bed Height on CO. Requirement for Recarbonation
of CaO 355
APPENDIX Q
1. Location of Limestone Supplies in Relation to Providence,
Rhode Island 380
2. Location of Limestone Deposits in Eastern United States 381
3. Repeatability of Stone Sulfidation Kinetics 384
4. Grain Structure of Candidate Stones for the NEES Plant 386
APPENDIX R
1. Sulfidation of a Large Particle of Calcium Limestone 402
2. Comparison of Esso CAFB Oil Gasification Desulfurization and
the Projection of Fluidized Bed from the Thermogravimetric
Data 404
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LIST OF TABLES
Page
APPENDIX A
1. Imports of Residual Fuel Oil by Percent Sulfur Content
by Country of Origin: Jan.-Dec. 1973 4
2. Refinery Production of Residual and No. 4 Fuel Oil by
Sulfur Content by P.A.D. and Refinery Districts:
Jan.-Dec. 1973 6
3. Imports of Residual Fuel Oil by Percent Sulfur Content
by Country of Origin: Jan.-July 1974 8
4. Refinery Production of Residual and No. 4 Fuel Oil by
Sulfur Content by P.A.D. and Refinery Districts:
Jan.-July 1974 10
5. United States Fossil Fuel Generation Survey 12
APPENDIX C
1. Power Plant Design and Operating Characteristics 28
APPENDIX D
1. Fluidized Bed Oil Gasification and Demonstration Plant 38
2. Fluidized Bed Oil Gasification Process Fuel-Gas
Characteristics 39
APPENDIX E
1. Reaction System-Process Equipment 152
2. Fuel-Gas Piping System 154
3. Stone Processing System (Dry Sulfation) 155
APPENDIX G
\
1. Summary of Environmental Performance of Chemically Active
Fluidized Bed (CAFB) Pilot Plant 169
2. Environmental Impact Comparison 173
xvii
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List of Tables (Continued)
Page
APPENDIX H
1. Comparison of Utility Requirements for CAFB and Limestone
Slurry Scrubbing 188
APPENDIX I
202
1. Equipment List of Initial Design Study
2. Cost Breakdown for the Initial Study Demonstration Plant 203
3. Summary of Major System Costs in the Initial Study
Demonstration Plant 204
4. Summary of Major Component Costs in the Initial Study
Reaction System 209
5. Demonstration Plant Modified Costs Projection 211
APPENDIX J
1. General Process Options for the Reaction System 215
2. Reaction System Operating Options 216
3. Equipment Design Options for the Reaction System 218
APPENDIX K
1. Solubility of Calcium Compounds in Water 224
2. Estimate of By-Product Stone Production 228
3. Demand for Construction Material 229
4. Oil Gasification - Comparison of Spent Stone Disposal
Processes 259
5. Realization from Gypsum Landfill 261
6. Base Equipment Cost Comparison 263
APPENDIX L
1. Projected S02 Retention in Fluidized Beds of CAFB Spent
Sorbent 281
2. TG Experiments Probing Conditions for the Dry Sulfatlon
Process 282
3. A Summary of the Dry Sulfation Tests 287
4. Program for Numerical Prediction of S0£ Retention 302
xviii
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List of Tables (Continued)
Page
APPENDIX M
1. Slurry Recorbination Test Results 314
I | _
2. Ca and S Determination for the Slurry Recarbonation
Reactions 316
I | _
3. Chemical Analyses of Ca , S , and Cationic Impurities
for a Typical Slurry Recarbonation Test 318
4. Particle Size Distribution of the Recarbonated Product
from Slurry Recarbonation Run 25 321
APPENDIX N
1. A Summary of the Acid Sulfation Experiments 332
++ = =
2. Chemical Analyses of Ca , S , SO^ , and Cationic Impurities
for Acid Sulfation Reactant, Products, and Filtrates 336
3. Rate of Filtration and pH of the Washing Solutions for the
Acid Sulfation Reaction Mixture ASS 337
APPENDIX 0
1. TG Runs for Spent Sorbent Disposal 347
2. Exothermic Hydration of Waste Lime in 20 ml Water 357
3. Temperature Rise on Hydration of Spent Sorbent CAFB Run 7
Regenerator Bed Material 359
APPENDIX P
1. Demonstration Plant 363
APPENDIX Q
1. Limestones Tested by Esso 377
2. Limestone Specification 378
3. Candidate Materials - Composition 382
4. Conditions for 51 mm (2 in) Unit Attrition Index Tests 387
5. Comparison of Attrition Index from 51 mm (2 in) Unit and
Results Noted by Esso 388
6. Summary of Candidate Stones for the NEES Plant 392
7. Respondents to Sampling Request 394
xix
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List of Tables (Continued)
Page
APPENDIX R
1. TG Results on Sulfidation of Lime 399
APPENDIX S
1. Sulfur Resources in the United States 411
2. Phosphorus Reserves 413
xx
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LIST OF ABBREVIATIONS
A/F
BCURA
BaS0
CAFB
CaO
Ca(OH)2
CaS
Ca(SH3)2
CaSOn
CE
CH
CH,
1/2
CO
co2
Consol
COS
cs2
EPA
GJ
HDS
kg
kPa
kv
kWh
m
air/fuel ratio
British Coal Utilization and Research Association
barium sulfate
calcium carbonate
chemically active fluidized bed
calcium oxide
calcium hydroxide
calcium sulfide
calcium bisulfite
calcium sulfite
Combustion Engineering
residual hydrocarbon content of gasifier stone
methane
methanol
carbon monoxide
carbon dioxide
Consolidated Coal Company
carbonyl sulfide
carbon disulfide
Environmental Protection Agency
9
giga joule (10 ), S.I. unit of energy
hydrodesulfurization
hydrogen sulfide
sulfuric acid
kilogram
kilopascal, S.I. unit of pressure
kilowatt
kilowatt hours
meter
xxi
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MgCO, - magnesium carbonate
MgO - magnesium oxide
ml - milliliter
mm - millimeter
m/s - meters per second
MW - megawatt
NEES - New England Electric Systems
NO- - nitrogen dioxide
02 - oxygen
OCR - Office of Coal Research
ORD - Office of Research and Development
ORM - Office of Research and Monitoring
Pa - pascal, S.I. unit of pressure
PTG - purge transport gas
rms - root mean square
SiO_ - silicon dioxide
S02 - sulfur dioxide
S0_ - sulfur trioxide
SRE - sulfur removal efficiency
SWEC - Stone and Webster Engineering Corporation
TG - thermogravimetric
xxii
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ACKNOWLEDGMENTS
o
The results, conclusions, and recommendations presented in this
volume represent the combined work and thought of many persons at
Westinghouse and the Office of Research and Development (ORD), and
personnel at New England Electric System (NEES). Other ORD contractors
have freely shared with us their ideas and results of their research and
development effort. Westinghouse subcontracted Stone and Webster
Engineering Corporation (SWEC) to prepare the preliminary design and cost
estimate.
In particular, we want here to express our high regard for and
acknowledge the contribution of personnel at Westinghouse Research Lab-
oratories , New England Power Service Company, Esso Research Centre (England),
and the personnel at ORD who have directed the overall fluidized bed
residual oil gasification program and who have defined, monitored, and
supported the efforts of Westinghouse and others on the program.
Mr. F. P. Turner, Chief of the Advanced Process Section, has served as
project officer on our work. Numerous enlightening and helpful discussions
have been held with Mr. Turner; with section members S. L. Rakes and
D. Bruce Henschel; and with R. P. Hangebrauck, Chief of the Demonstration
Projects Branch. Mr. S. K. Batra and Mr. R. Gendreau of New England Power
Service Company contributed to the demonstration plant design, provided
data on the Manchester Street Station, and participated in process
evaluation and selection. Dr. J. W. T. Craig of Esso Research Centre
(England) participated in the preliminary design study from August 1973
through February 1974. Dr. Craig was effective in integrating the Esso
pilot-scale experience and results into the demonstration plant design.
He also participated in the investigations and evaluation of spent stone
processing systems. Mr. John P. Malloy of SWEC served as project manager
for preparation of the preliminary design and cost estimate. Mr. Harry
xxiii
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Robinson, SWEC, was the lead process engineer. Allied Chemical Corporation
provided technical and economic information on sulfur recovery. Mr. W. D.
Hunter of Allied coordinated the effort. Many persons at the Westinghouse
Research Laboratories have contributed to the program. Dr. T. K. Gupta
and Dr. M. Gunasekaran are initiating work to identify and develop market
applications for the spent stone. Mr. R. Brinza and Mr. W. F. Kittle
assisted in the collection of data.
xxiv
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APPENDIX A
MARKET DATA
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APPENDIX A
MARKET DATA
This appendix provides additional data on the availability of
petroleum residues and the fuel utilization in fossil-fired steam power
plants. An overview of the market is presented in Volume I, section IV.
AVAILABILITY OF PETROLEUM RESIDUES
Recent data on the availability of heavy fuel oils by sulfur
level from the Bureau of Mines are appended. Summary tabulations are in-
cluded for January to December 1973 and January to July 1974. Vacuum
tower bottoms or other high metal content fuels are the most likely to be
economically attractive for the chemically active fluidized bed (CAFB)
process (see Volume I, sections I and IV). The utilization of these low-
grade petroleum residues is, unfortunately, unclear.
Commercial processes are being developed to desulfurize fuel
oil from many crude oils. The cost of desulfurizing high-sulfur, high
metals content residual oil fractions, however, is currently too high to
permit the production of a fuel economical for power generation. Thus,
residues from low-quality crudes are more attractive, economically, for
power generation utilizing the^fluidized bed residual oil gasification/
desulfurization process. The price of lower quality residues, therefore,
could be substantially below the price of higher grade, low-to-medium
sulfur residual oils.
Estimates of the potential availability of low-grade petroleum
fractions (for example, vacuum tower resid-- 1050°F+) are approximately
one million barrels per day. This is equivalent to approximately
30,000 MW of power generation capacity. Estimates for high metals con-
tent crude oil reserves are in the order of 15 billion barrels. The use
of this fuel is unclear. Several processing alternatives have been
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proposed, including blending with low-sulfur and/or low-metals residual
oils. What quantity of the low-grade fuel would be available for utility
use with the CAFB will depend on numerous factors, including the fuel re-
quirements of other countries, crude oil availability, fuel distribution
problems, alternative process options available. If half the potential
low-quality fuel is available for utility use, this would represent
roughly 15,000 MW or 25 percent of the present oil-fired power generation
capacity. Figure A-l and Table A-5 summarize the utility power genera-
tion market.
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U. S. DEPARTMENT OF THE INTERIOR
BUREAU OF MINES
WASHINGTON. D. C. 2O24O
Rogers C. B. Morton, Secretary
Thomas V. Falkle. Director
For information call Paul Chapman
Telephone: 703 557-OUte
Fuel Oils by Sulfur Content. Monthly
AVAILABILITY OF HEAVY FUEL OILS BY SULFUR LEVELS
December 1973
New supply of residual fuel oil, including No. k fuel oil, In the United
States in December totaled 91-0 million barrels according to the Bureau of Mines,
United States Department of the Interior. Of this total, domestic refinery production
accounted for 36.5 million barrels and imports for 5U.5 million barrels. In this
report new supply excludes domestic crude oil used for direct burning as fuel. Import
data shown in this report are import data reported to the Office of Oil and Gas under
the oil Import program. Imports reported In the Bureau of Mines Monthly Petroleum
Statement include bonded imports of fuel oils used for vessel bunkering and fuel oil
used by the military offshore in addition to the imports included in this publication.
RESIDUAL FUEL OIL
Domestic refinery production of residual fuel oil in December totaled 35.9
million barrels of which low-sulfur oils (1.0 percent or less sulfur content by weight)
amounted to 17.1 million barrels or 1*7.7 percent of the total. The West Coast Re-
fining District produced 37.9 percent of the total low-sulfur residual fuel oil, fol-
lowed by the Indiana-Illinois Refining District with 16.$ percent. Compared to 1972,
domestic refinery production during 1973 increased by 21.2 percent. Imports of low-
sulfur fuel oil in December, amounting to 2U.9 million barrels, represented U9.2
percent of the total 50.7 million banels of residual oil imported during the month.
Imports from Venezuela accounted for 38.3 percent of the total imports of low-sulfur
residual fuel oil, followed by the Netherlands West Indies with 22.5 percent of the
total. Total imports of residual fuel oil during 1973 increased by 39*1 million
barrpln, or U.9 pervert, «v»r 1972. The Ce^rr"! Atl?".tic Stat?p received J6.1 percent
of the total Imports of low-sulfur residual fuel oil during December, followed by the
New England States with 28.0 percent.
NO. k FUEL OIL
Domestic refinery production of No. It fuel oil In December totaled 0.6
million barrels, including O.k million barrels of low-sulfur oils. Compared to 1972,
domestic refinery production during 1973 decreased by 31-0 percent. During the month
Imports of No. 1| fuel oil totaled 3.8 million barrels of which low-sulfur oils amounted
to 2.8 million barrels or 7^.5 percent of the total. Imports for 1973 increased by
I*.3 million barrels, or 13.7 percent, over 1972.
Stocks of heavy fuel oils held by refining and pipeline companies on
December 31, 1973 amounted to 53.5 million barrels of residual fuel oil and 3.U
million barrels of No. 1* fuel oil. Of the total residual and No. U fuel oil stocks,
25.3 million barrels and 2.7 million barrels, respectively, contained 1.0 percent
or less sulfur.
Table Ik. -Refinery Production of Residual and No. It Fuel Oil by Sulfur
Content, by P.A.D. and Refinery District: 1972
Table 15 . -Imports of Residual Fuel Oil by Percent Sulfur Content, by
Country of Origin: 1972
Table 16. -Imports of Residual Fuel Oil by Percent Sulfur Content, by
States: 1972
Table 17. -Imports of No. It Fuel Oil by Percent Sulfur Content, by
Country of Origin: 1972
Table 18 . -Imports ol No. k Fuel Oil by Percent Sulfur Content, by
States: 1972
Table 19 . -Stock? of Residunl and No. U Fuel Oil Held by Refineries and
Bulk Terminals by Sulfur Content by T.A.D. anJ Refinery
Districts: December 31, 1972.
Prepared April 2, 197!* in the Division of Fossil Fuels - Mineral Supply.
-------
Table A-l
IMPORTS OF RESIDUAL FUEL OIL BY PERCENT SULFUR CONTENT
BY COUNTRY OF ORIGIN: JAN-DEC
(Thousands of Barrels)
1973
DISTRICT
COUNTRY
P.A.D. I
NEW ENGLAND
CANADA
MEXICO
BAHAMAS
BRITISH W.I.
TRINIDAD
N.W.I.
VENEZUELA
BRAZIL
SWEDEN
ENGLAND
NETHERLANDS
BELGIUM
FRANCE
WEST GERMANY
SOVIET UNION
SPAIN
ITALY
GREECE
RUMANIA
GHANA
NIGERIA
VIRGIN ISLANDS
TAT Al
CENTRAL ATLANTIC
CANADA
BAHAMAS
BRITISH W.I.
TRINIDAD
N.W.I.
COLOMBIA
VENEZUELA
ECUADOR
PERU
BRAZIL
SWEDEN
NORWAY
ENGLAND
NETHERLANDS
BELGIUM
FRANCE
WEST GERMANY
SOVIET UNION
SPAIN
ITALY
GREECE
RUMANIA
SAUDI ARABIA
ALGERIA
LIBYA
GHANA
NIGERIA
ANGOLA
VIRGIN ISLANDS
LOWER ATLANTIC
MEXICO
BAHAMAS
BRITISH W.I.
TRINIDAD
N.W.I.
COLOMBIA
VENEZUELA
BRAZIL
ARGENTINA
ENGLAND
NETHERLANDS
BELGIUM
ITALY
LIBYA
MALAGASY REPUBLIC
VIRGIN ISLANDS
O-.bO
7,042
3
298
_
5,869
8.706
8,354
196
132
1,069
166
380
ZSO
-
280
236
2,394
155
466
_
117
9,270
j.c tan
* 3 1 JDU
S.022
28,850
_
8.356
27,688
32,545
-
886
1.052
435
«
280
1.126
1.067
575
117
1,222
3,020
18,013
462
173
4,293
8.944
_
2.876
-
_3fti5»3^
IBS. 543
24
-
522
502
-
237
-
_
_
_
_
_
681
_
1.966
F'ERCFNT SI
.bl-1.00
4.532
-
5,237
148
8,761
15,777
5,950
-
-
242
231
1,100
284
132
220
135
4,382
-
259
140
209
/*T . T^O
4ft f JTF
B73
3,041
205
9.057
14.241
-
5.921
-
133
368
-
250
437
1,111
1.060
109
_
249
314
4.966
_
_
_
-
388
-
3,162
45,864
-
333
-
-
21,568
-
12,425
239
_
110
244
no
1,052
486
-
27
36,655
LFUR CONTFK
1.01-2.00
122
-
_
-
2.817
1.116
5.034
-
-
-
-
-
_
_
-
-
_
-
1.072
1 ft » 1 Al
1 U 1 Dl
840
351
-
6.093
3.814
919
12.376
130
_
-
_
_
128
-
«
_
_
_
-
_
_
SO
119
-
-
140
6.095
31.057
-
2.961
_
2,545
4.096
240
23.281
_
_
_
_
_
_
3
606
33.732
Y
OVER 2.00
4.992
75
506
_
153
1.8B5
10.468
-
-
-
-
-
517
_
_
297
-
-
"-
-
170
1 O A A?
5t559
_
M
1.356
11.581
-
24.592
-
-
-
-
_
37
-
54
_
_
206
_
.
.
-
-
-
-
6.722
50.106
561
2.485
131
690
31.485
-
43.568.
-
152
a
_
_
12.966
91 .646
TOTAL
16,688
78
6.040
168
17,600
27,482
29.A06
196
112
1.310
397
1.680
535
650
499
370
7.073
155
666
tvt
?57
10.720
i pa \t*9
IcCt i**c
12.?96
32.763
205
26.B62
57,325
919
75.633
130
1.019
1.470
435
250
717
2.401
2.127
AB5
171
1.470
3.333
23.116
462
173
50
4.612
8.944
388
2.H76
140
54,<;oi
312. 571
S61
5.P06
131
3.757
57.651
240
79.-510
239
152
117
?44
170
1.052
1.167
3
13.199
163.999
-------
Table A-l(Continued)
DISTrtlCT
CODNTSY
P.A.O. II
CANADA
VENEZUELA
LlbYA
NIGERIA
P.A.O. Ill
MEXICO
BAHAMAS
TRINIDAD
N.H.I.
COLOMBIA
VtNLZUELA
ECUADOR
NETHERLANDS
BELOIUX
LIBYA
NIGERIA
VIRGIN ISLANDS
TOTAL. »
P.A.D. V
CANADA
N.H.I.
VENEZUELA
ECUADOR
PERU
ITALY
GREECE
SAUDI ARABIA
BAHRAIN
MALAYSIA
INDONESIA
HAWAII1
U.S.. TOTAL
COUNTRY TOTALS
CANADA
MEXICO
BAHAMAS
BRITISH H.I.
TRINIDAD
N.H.I.
COLOMBIA
VENEZUELA
ECUADOR
PERU
BRAZIL
ARGENTINA
SHE DEN
NORWAY
ENGLAND
NETHERLANDS
BELGIUM
FRANCE
WEST GERMANY
SOVIET UNION
SPAIN
ITALY
GREECE
RUMANIA
SAUDI ARABIA
BAHRAIN
MALAYSIA
INDONESIA
ALGERIA
LIBYA
GHANA
NIGERIA
ANGOLA
MALAGASY REPUBLIC
VIRGIN ISLANDS
HAWAII1
"FKCENI SULfUH CON TEN
o-.io
It 654
_
-
_
1.654
124
-
-
-
-
77
-
201
-
594
2,350
IS5
432
524
213
62B
317
4,332
-
9,542
244, 286
13,718
127
29.172
-
14,747
37.468
_
43.485
155
1.317
1,248
-
567
-
1,349
1.368
1,446
825
117
1,501
3.255
20.930
830
639
628
-
317
4.332
4.293
9.625
-
2.993
-
_
47.813
-
.Sl-I.OO
993
587
198
186
1.964
-
385
609
.
-
»
_
333
504
474
2. 303
-
-
-
_
-
16
16
_
-
32
134.557
6.398
-
8.996
353
17,819
52,195
_
24,882
_
133
607
-
250
788
1.586
2,662
394
132
468
448
10,401
-
-
16
_
16
-
1,188
647
800
-
_
3.377
-
1.01-,!. 00
1.345
374
-
_
1.719
-
-
20
60
355
113
_
_
-
547
26
-
-
_
-
-
567
486
_
384
1<464
78,680
2,333
-
3,312
-
11,455
9,047
1,219
41,419
243
_
_
-
-
128
_
.
*
_
_
_
617
_
486
_
119
_
_
_
140
3
7,772
384
T
OVER 2.00
116
654
-
770
-
11
244
.
783
_
.
_
-
371
!>bOA
wo
-
>
166
_
.
_
-
50
5
-
_
.
221
163.213
10,666
636
2.991
131
2.209
45.195
.
BO. 230
_
_
_
152
.
a
37
_
.
571
_
503
_
_
50
5
_
.
.
.
_
.
_
19.828
TOTAL
4,108
1.615
198
IB*
f>.107
124
385
11
P73
60
1,136
113
77
332
504
474
371
4- 4.CQ
t **3~
26
594
2.516
155
432
524
713
l.?60
5
A19
4.332
1R4
11,759
2620.736
33.116
763
44,472
484
46,?30
143,925
1,?I9
190.017
398
1.450
1.855
152
S67
?50
2.145
3.119
4,109
1,?19
821
1.470
3,704
31.R35
A30
639
1.311
5
A19
4.332
4.412
10.A13
*47
3.793
140
3
78,790
384
SOURCE I OFFICE OF OIL AND GAS Table 5.
NOTE..DATA MAY NOT ADD TO TOTALS SHOWN BECAUSE OF INDEPENDENT ROUNDING
1 FOREIGN TRADE 7ONE
8 INCUDES CRUDE OIL FOR DIRECT BURNING AS FUEL- CANADA, 1.916. INDONESIA.
-------
Table A-2
REFINERY PRODUCTION OF RESIDUAL AND NO.4 FUEL OIL BY SULFUR CONTENT
BY P.A.D. AND REFINERY DISTRICTS: JAN-DEC 1973
(Thousands of Barrels)3
P.A.I). AND REFINERY
DISTRICTS
EAST COAST TOTAL
IND.-ILL. .............
DIST. III. TOTAL
TEXAS INLAND ..........
TEXAS GULF ............
ARK. -LA. INL. .........
RESIDUAL FUEL OIL
PERCENT SULFUR CONTENT
O-.SO
11.743
11.363
4BO
*9H5
29
111
*845
12,790
2.151
4,150
lt,0<<9
1.496
944
B24
70.348
'96,690
.51-1.00
15.834
IS. 056
778
30.368
1.615
26.433
927
* 1.393
* 26.1462
1.383
11.972
* 12, 722
384
1
2.451
7.385
* 82, 500
1.01-2.00
16.112
10.350
S.762
25.952
143
20.518
2.488
2.803
*9.927
l.bB3
4.650
* 1.755
1.718
221
3,323
47.528
* 102,61)2
OVER 2.00
8,569
8.569
13.815
6.632
4.500
2.683
39.276
57B
32.303
5.229
1.166
3.266
7.639
72.565
TOTAL
52.258
45.238
7.020
71.120
1.758
53,612
8.026
7.724
88.U55
5.695
53.075
23.755
4.764
1.166
9.864
132.900
35U.597
NO. 4 FUEL OIL
PERCENT SULFUR CONTENT
O-.SO
153
153
2,040
209
307
1.524
1.220
76
517
5
4on
214
406
1.058
4, 877
.51-1.00
_
1,441
1.401
40
4B5
51
389
45
196
2.122
* KEVISLD.
1.01-2.00
205
205
153
153
82
82
200
640
OVEH 2.00
108
108
2?2
2?2
214
214
20
100
664
TOTAL
466
466
3.P54
1.763
529
1.S64
2.001
127
906
50
704
214
fl?2
1.158
d.303
READS;
AUGUST 1973;
DISTRICT II. TOTAL
OKLA.-KANS.
U.S. TOTAL
SEPTEMBER 1973:
DISTRICT III, TOTAL
LA. CUI.F
U.S. TOTAL
0-0.50
.51-1.00
SHOULD READ:
.51-1.00
32
70
7,367
51-1.00
1,661.
878
5.77U
1.586
77
5.385
1.01-2.00
627
_
7.052
75
63
7,360
.51-1.00
1.592
806
5.702
1.593
8U
5,392
1.01-2.01
1.99
72
7.121.
aMineral Industry Surveys, U.S. Department of the Interior,
Bureau of Mines, Washington, D.C., December 1973, Table 3.
-------
U. S. DEPARTMENT OF THE INTERIOR
BUREAU OF MINES
WASHINGTON. D. C. 2O24O
Rogers C. B. Morton, Secretary
Thomas V. Falkie, Director
For information call Stephen K. Patterson
Telephone: 703 557-0442
Fuel Oils by Sulfur Content, Monthly
AVAILABILITY OF HEAVY FUEL OILS BY SULFUR LEVELS
July 1974
New supply of residual fuel oil, including No. 4. fuel oil, in the United
States in July totaled 75.3 million barrels according to the Bureau of Mines, United
States Department of the Interior. Of this total, domestic refinery production ac-
counted for 33.4 million barrels and imports for 41.9 million barrels. In this report
new supply excludes domestic crude oil used for.direct burning as fuel. Import data
shown in this report are import data reported to the Office of Oil and Gas under the
oil import program. Imports reported in the Bureau of Mines Monthly Petroleum State-
ment include bonded imports of fuel oils used for vessel bunkering and fuel oil used
by the military offshore in addition to the imports included in this publication.
RESIDUAL FUEL OIL
Domestic refinery production of residual fuel oil in July totaled 32.7 million
barrels of which low-sulfur oils (1.0 percent or less sulfur content by weight) amounted
to 17.4 million barrels or 53.3 percent of the total. The West Coast Refining District
produced 29.7 percent of the total low-sulfur residual fuel oil, followed by the Texas
Gulf Refining District with 20.4 percent. Imports of low-sulfur fuel oil in July amount-
ing to 24.0 million barrels, represented 58.9 percent of the total 40.8 million barrels
of residual oil imported during the month. Imports from the Virgin Islands accounted
for 31.0 percent of the total imports of low-sulfur residual fuel oil, followed by the
Netherlands West Indies with 21.9 percent of the total. The Central Atlantic States
received 62.6 percent of the total imports of low-sulfur residual fuel oil during July
followed by the New England States with 16.6 percent.
NO. 4 FUEL OIL
Domestic refinery production of No. 4 fuel oil in July totaled 0.7 million
barrels, including 0.6 million barrels of low-sulfur oils. During the month imports
of No. 4 fuel oil totaled 1.1 million barrels of which low-sulfur oils amounted to
1.0 million barrels or 86.1 percent of the total.
STOCKS
Stocks of heavy fuel oils held by refining and pipeline companies on July 31,
1974 amounted to 59.8 million barrels of residual fuel oil and 4.4 million barrels of
No. 4 fuel oil. Of the total residual and No. 4 fuel oil stocks, 31.7 million barrels
and 3.9 million barrels, respectively, contained 1.0 percent or less sulfur.
Prepared October 31, 1974 in Division of Fossil Fuels - Mineral Supply.
-------
Table A-3
IMPORTS OF RESIDUAL FUEL OIL BY PERCENT SULFUR CONTENT
BY COUNTRY OF ORIGIN: JAN-JULY 1974
(Thousands of Barrels')
DISTRICT
COUNTSY
P.A.O. I
NEW ENGLAND
CANADA
BAHAMAS
BfMTISH W.I.
TRINIOAO
N.W.I.
VENEZUELA
ENGLAND
NETHERLANDS
FRANCE
ITALY
GREECE
RUMANIA
NIGERIA
ANGOLA
VIRGIN ISLANDS
CENTRAL ATLANTIC
CANAOA
MEXICO
BAHAMAS
BRITISH W.I.
TRINIDAD
N.W.I.
COLOMBIA
VENEZUELA
PERU
BRAZIL
ENGLAND
NETHERLANDS
BELC-I'J"
FRANCE
WEST GERMANY
SPAIN
PORTUGAL
ITALY
GREECE
RUMANIA
IRAN
U. APA3 EMIRATES
INDONESIA
ALGERIA
GHANA
NIGEPIA
VIRGIN ISLANDS
LOWER ATLANTIC
MEXICO
BAHAMAS
TRINIDAD
N.W.I.
COLOMBIA
VENEZUELA
NETHERLANDS
BELGIUM
WEST GERMANY
PORTUGAL
ITALY
VIRGIN ISLANDS
TOTAL
0-.50
2.960
504
-
4,649
5,237
2,614
251
-
112
622
101
-
98
-
2,053
19.200
2,771
_
8,187
100
6,366
17.626
-
15,457
144
466
27S
_
-
286
-
376
_
4,872
-
-
-
_
141
239
-
l,32fl
20,32'.
78.958
_
-
-
_
-
.
-
-
-
-
-
_
-
PEPCCNT SU
.i.1-1.00
557
795
150
5,949
7,645
2,054
_
173
-
361
-
182
-
102
-
17,966
196
-
3,625
-
1,154
7,258
-
4,178
-
-
149
352
124
. -
-
_
-
698
-
202
-
12
-
-
261
-
5,619
23i827
-
541
_
8,981
-
5,580
170
-
177
-
539
_
15.S87
LfUIJ CONJEN
1.01-2.00
_
_
308
1,003
264
2,369
_
-
_
-
_
-
-
2,299
6,243
635
-
-
-
4,093
1,159
502
6,745
-
-
-
_
152
368
-
-
50
360
-
-
120
-
10.641
24.824
-
2,724
8*7
2,477
617
8,197
120
-
-
64
-
903
15,950
T
OVER 2.00
2,774
354
-
399
1.071
6,251
_
222
-
-
-
-
-
-
-
11,071
4,054
108
2,765
127
939
7,053
-
11,293
-
-
329
331
-
277
-
798
135
-
54
_
-
-
124
1,968
30.376
103
4,551
54
17,464
-
19,503
125
320
236
-
241
7.675
50.273
TOTAL
6,?90
1.652
458
12,000
14,?17
I3,?f)7
?51
396
112
Qfl3
101
182
98
102
4,. 152
54,4*1
7,656
108
14,577
227
12, 552
33,096
502
37,673
144
466
4?4
682
475
?«6
?77
376
152
6,736
135
202
104
372
141
?39
390
1,4«52
38,552
157. QH5
103
7,R16
Q01
28, "22
617
33,280
416
320
413
64
780
P.S78
62.210
-------
Table A-3(Continued)l
DISTRICT
COUNTRY
P.A.D. II
CANADA
VENEZUELA
P.A.D. Ill
BAHAMAS
BRITISH W.I.
TRINIDAD
N.W.I.
COLOMBIA
VENEZUELA
RUM AN I A
VIRGIN ISLANDS
P.A.D. V
N.W. I.
VENEZUELA
ECUADOR
PERU
ITALY
SAUDI ARABIA
MALAYSIA
INDONESIA
VIRGIN ISLANDS
HAWAIIAN F.T. ZONE
COUNTRY TOTALS
CANADA
MFXICO
BAHAMAS
BRITISH W.I.
TRINIDAD
N.W.I.
COLOMBIA
VENEZUELA
ECUADOR
PERU
BRAZIL
ENGLAND
NETHERLANDS
BELGIUM
FRANCE
WEST GERMANY
SPAIN
PORTUGAL
ITALY
GREECE
RUMANIA
IRAN
SAUDI ARABIA
U. ARAB EMIRATES
MALAYSIA
INDONESIA
ALGERIA
GHANA
NIGERIA
ANGOLA
VIRGIN ISLANDS
HAWAIIAN F.T. 70M
PF-RCFNT SULFUR CONTENT
0-.50
1,335
-
1.33S
149
-
_
72
_
_
?21
_
3.964
-
691
282
79
20
3,584
1.102
9,721
109.436
7,066
_
8,690
249
11.015
22.864
22,108
-
835
466
526
-
-
399
_
376
_
5,776
101
-
_
79
20
3.724
239
1.425
_
23.479
.51-1.00
847
341
1,188
853
221
363
130
-
-
246
93^
2,746
_
-
150
-
-
-
-
_
ISO
61,864
1,600
5,813
371
7,466
24,013
-
12,152
150
-
149
696
124
-
177
-
-
1,598
-
630
-
-
12
-
-
261
-
102
6,550
1.01-2.00
79
R
87
-
-
30
525
?0
575
37
23
-
-
300
244
403
1,007
48,686
714
2,7?4
308
5,943
3,937
1,149
17.868
-
-
120
-
-
-
216
368
-
-
50
300
360
244
-
120
-
14,266
OVER 2.00
60
60
229
-
234
57
-
569
7IB
1.B07
-
-
-
59
-
59
93,646
6.888
211
7,899
127
1.626
25,645
-
37,615
-
-
677
671
-
513
-
-
1,039
135
-
54
-
59
-
124
10,362
"
TOTAL
2,3?2
149
2,671
1 «0fl2
370
597
187
30
1,166
?46
1.^70
5.349
37
3.9P7
150
691
?R2
379
323
3,584
1.505
"
10,937
1313,632
16,268
?1 1
25,127
1,055
26,050
76,459
1,149
89,74>
150
835
466
675
1,493
795
399
689
376
?16
8,780
?37
f.30
104
379
372
323
3,7?4
?39
380
1.550
102
54,658
SOURCE: OFFICE OF OIL AND GAS Table 5.
NOTE..DATA MAY NOT ADD TO TOTALS SHOWN BECAUSE OF INDEPENDENT ROUNDING
1
INCLUDES CRUDE OIL FOR DIRECT BURNING AS FUEL: CANADA, 1,274; IRAN, 50; TRINIDAD, 241;
UNITED AHAB EMIRATES, 360, VENEZUELA,1,708..
-------
Table A-4
REFINERY PRODUCTION OF RESIDUAL AND NO. 4 FUEL OIL BY SULFUR CONTENT
BY P.A.D. AND REFINERY DISTRICTS: JAN-JULY 1974
.(Thousands of Barrels)3
P.A.D. AND REFINERY
DISTRICTS
DIST. IV
RESIDUAL FUEL OIL
PERCENT SULFUR CONTENT
o-.so
6011
6.110
201
2.411
1.868
150
393
11*951
1.397
6*056
1.837
1.958
703
1.158
32.063
53.874
.51-1.00
11.079
10.881
198
15.260
It 045
12.882
616
717
22.633
1.472
10.892
10.045
53
171
1.564
2.903
53.439
1.01-2.00
8.391
4.608
3.783
11.838
B.1B4
1.387
2*267
8.581
890
4.932
1.584
975
200
1.476
31.414
61.700
OVER 2.00
7.282
7.282
fl.175
5.189
1*795
1*191
22*924
306
17*930
3*346
1*342
2.356
4.413
45. ISO
TOTAL
33*063
28*881
4*182
37*684
1*045
28.123
3*948
4*568
66*089
4,065
39,810
16*812
4,328
1*074
6*554
70*773
214,163
NO. 4 FUEL OIL
PERCENT SULFUR CONTENT
O-.SO
505
214
291
393
225
150
18
617
6
215
3S9
37
417
487
2*419
.51-1.00
20
20
547
456
89
2
726
335
247
. us
26
30
.
1,323
1.01-2.00
261
169
92
13
13
68
68
167
53
562
OVER 2.00
_
137
137
164
164
62
383
TOTAL
786
383
403
1.090
818
252
20
1.575
6
335
462
709
63
614
622
4,687
Industry Surveys, U.S. Department of the Interior,
Bureau of Mines, Washington, D.C., July 1974, Table 3.
-------
HYDRO/DIESEL
NOTE; EACH STATE IS SHADED TO REPRESENT MORE THAN 50% OF ITS FOSSIL GENERATION
BY FUEL TYPE.
Figure A-l-United States fossil generation overview
-------
Table A-5
UNITED STATES FOSSIL FUEL GENERATION SURVEY3ib
State
Alabama
Alaska
Arizona
Arkansas
California
Colorado
Connecticut
Delaware
District of Columbia
Florida
Georgia
Hawaii
Idaho
Illinois
Indiana
Iowa
Kansas
Kentucky
Louisiana
Maine
Maryland
Massachusetts
Michigan
Minnesota
Mississippi
Missouri
TOTAL
Gas Oil
(Hydro & diesel)
30
16682
375
1609
365
573
75 9859
112 318
935
-
1209
46
117
1484
3654 1269
381
1021
3647
427
202
686 1231
562
Coal
3206
-
700
1600
1220
633
497
2961
1331
4064
-
11959
7535
393
540
2996
1412
867
7851
2634
200
4973
State
Montana
Nebraska
Nevada
New Hampshire
New Jersey
New Mexico
New York
North Carolina
North Dakota
Ohio
Oklahoma
Oregon
Pennsylvania
Rhode Island
South Carolina
South Dakota
Tennessee
Texas
Utah
Vermont
Virginia
Washington
West Virginia
Wisconsin
Wyoming
Gas Oil
(Hydro)
(Diesel
275
317
2406
334
6689
16
2
3153
479
1605
215
98
18843 418
305
(Hydro & Diesel)
1427
87
-
95
(Diesel)
23,704 59,961
Coal
-
-
234
5883
206
7278
7873
113
19196
164
8134
62
1489
116
15124
-
561
6022
-
1173
182
131.382
^Information extracted from Electrical World Directory of Electrical Utilities, 1971-1972.
In MW.
-------
APPENDIX B
PRELIMINARY DESIGN PROCESS DESCRIPTION
INTRODUCTION
The chemically active fluidized bed (CAFB) demonstration plant
is a facility which will demonstrate system operability, environmental
performance, boiler performance, and process retrofit capability. It will
also provide a projection of the economics of an ultimately commercial
oil gasification process. The demonstration plant has been designed with
sufficient flexibility to enable the exploration and determination of
optimum operating conditions and operating procedures.
The demonstration plant consists of five major processing systems:
Reaction
Stone processing
Limestone handling
Fuel-oil handling
Waste/by-product stone handling.
The general function of the reaction system is to generate a low heating-
value fuel gas for utilization by the boiler. The stone processing system
employs a dry sulfation process to convert the spent sorbent into a form
suitable for disposal or by-product utilization. The remaining three systems
store and carry the raw materials (residual oil and limestone) and the waste
or by-product materials (sulfated limes) which are utilized in the process.
REACTION SYSTEM
The reaction system boundaries and components are shown on
Figure B-l. The reaction system consists of
A gasifler
A regenerator
Solids transport components
«
Particulate removal components
13
-------
Air and stack-gas compression components
Fuel-gas handling and combustion components
Heat exchange components
A purge and transport gas (PTG) system.
General
The process logic is as follows: Residual fuel oil is fed to
the gasifier at 12,700 kg/hr (28,000 Ib/hr) along with limestone at
1,070 kg/hr (2,355 Ib/hr), air at 32,100 kg/hr (70,604 Ib/hr) and stack
gas for temperature control at 30,800 kg/hr (67,020 Ib/hr). Oil cracking,
partial oxidation, and hydrogen sulfide (H.S) absorption proceed at 880°C
(1616°F) in the gasifier to generate 76,000 kg/hr (167,175 Ib/hr) of low
heating-value fuel gas. The reaction mechanisms involved in the gasification
1 2
step have been described in detail previously. ' The hot fuel gas flows
to the four boiler burners through four separate refractory-lined pipes, where
combustion is completed for steam generation. The sulfided lime (^ 6 wt %
calcium sulfide [CaS]) is transported pneumatically to the regenerator
vessel at 31,400 kg/hr (69,419 Ib/hr). The calcium sulfide stone and the
carbon (^ 0.3 wt % CH.._) are reacted with preheated air at a temperature
of 1075°C (1970°F) to generate calcium oxide (CaO), calcium sulfate, sulfur
dioxide (SO.) (8 mole %), and water vapor. Regenerated sorbent is recir-
culated to the gasifier at 32,400 kg/hr (71,615 Ib/hr). The sulfur dioxide
gas at 3,980 kg/hr (8,778 Ib/hr) and a regenerated stone purge stream at
297 kg/hr (655 Ib/hr) are removed from the reaction system for processing
in the stone processing system (dry sulfation process). Regenerator
1 2
chemistry has also been previously described in detail. '
Solids Transport Components
Solids transport components are utilized to carry stone between
the gasifier and regenerator, to transport the regenerator stone purge
stream to the stone processing system, to recycle gasifier fines collected
in the gasifier cyclones to the gasifier and regenerator, and to transport
fines collected in the boiler cyclone and the recycled stack-gas baghouse
14
-------
Dwg. I67IB21)
Limestone
from
Limestone Handling
System
Steam
Fuel Oil
from '
Waste Stone
to Stone
Processing System
Oil Handling
System
Processing
System
Fines to Stone
Processing
Fines to
Stone Processing
Fines to
Stone
t Processing
S02-Gas System
to Stone
I. D. Fan
Recycle Gas
from Stone
Processing System
Gasifier
Regenerator
Solids Transport
Particulate Removal
Air and Stack Gas Compression
Fuel Gas Handling
Heat Exchange
Purge & Transport Gas
Figure B-1-Reaction system
-------
to the stone processing system. Stone transport between the gasifler and
regenerator Is carried out using a pulsed transport scheme developed in
the Esso Research Centre, Abingdon, England (Esso) pilot plant.
The transport ducts are of rectangular cross sections 12.7 cm high, 73.7
cm wide, (5 in. high, 29 in. wide), with a lateral leg about 0.92 m (3 ft)
long set at a 30 degree angle to the vertical and a horizontal leg about
0.31 m (1 ft) long set at the base. The whole transport system is
formed in a single refractory block connecting the gasifier and regenerator
together. Solids in the gasifier enter the lateral leg of the transport
duct at a depth of about 0.92 m (3 ft) in the bed. Transport gas is
pulsed to the transport leg through a pipe located at the junction between
the lateral and horizontal legs periodically to fluidize and defluidize
the leg and,thus, periodically permit solids flow. The rate of circula-
tion is controlled by the pulse frequency and the pulse duration. The
time-average transport gas rate is about 264 kg/hr (580 Ib/hr). Solids
return to the gasifier from the regenerator by an identical mechanism.
The circulation rate of solids is set to control the regenerator
temperature.
The stone purge stream is transported by means of a smaller pulsed
transport leg operated by the same principle described above. About 2.7 kg/hr
(6 Ib/hr) of transport gas is required for this leg, which is used to
control the gasifier bed depth.
Cyclones, which are set with insufficient elevation to transport
collected fines by gravity, utilize a commercial pulsed transport system
(Sturtevant Engineering Company Ltd., London, England) consisting of a
lock hopper/control valves arrangement which permits long-distance
transport of solids as plugs. The operation of this system in the Esso
3
pilot plant has been described. Four refractory-lined Sturtevant
systems are used to return gasifier fines from the four gasifier cyclones.
Each system handles a maximum of 780 kg/hr (1,720 Ib/hr) of fines,
requiring about 6.8 kg/hr (15 Ib/hr) of transport gas. Two transport
systems return fines directly back to the gasifier,while the other two
transport fines to the entrance of the gaslfier-to-regenerator transport
16
-------
duct. This arrangement permits very fine particles to be removed from
the gasifier rather than to be circulated continuously through the
cyclone-gasifier circuit. The fines are added at the transport duct
entrance so that they can mix with the bed material before going to the
regenerator and not disrupt regenerator performance. The boiler stack
cyclone Sturtevant system carries about 136 kg/hr (300 Ib/hr) of fines
and the recycle stack-gas baghouse about 7.7 kg/hr (17 Ib/hr).
Particulate Removal Components
Conventional cyclone and filter equipment is utilized to control
partlculate loadings in gas streams for environmental and compression
equipment protection. Four parallel refractory-lined cyclones are
located in the four fuel-gas lines to collect elutriated bed material
from the hot fuel gas before it is combusted. Each of the cyclones
handles a maximum of 19,000 kg (42,000 Ib) of fuel gas per hour and a
maximum of 820 kg (1,800 Ib) of fines per hour. The cyclones are elevated
so that the total length of hot fuel-gas piping is minimized.
A pair of refractory-lined cyclones is arranged in series to
collect fines elutriated from the regenerator. The primary cyclone is
designed to return fines of a coarse size (larger than 300 microns in diameter)
to the regenerator. The secondary cyclone collects the remaining fines
and removes them from the reaction system. Its purpose is to act
as an elutriation point for fines which would otherwise be trapped in
the system.
A bag filter is situated at the point of regenerator sulfur
dioxide gas-release from the reaction system for a final cleaning of the
gas prior to its processing in the stone processing system. Fines
collected are also transported to the stone processing section. Stack
gas recycled for gasifier temperature control also requires a baghouse
filter to remove fines before the stack gas is compressed. This baghouse
handles 30,800 kg/hr (28,000 Ibs/hr) of stack gas which has been cooled
to temperatures less than 260°C (500°F) by a water spray. Fines collected
17
-------
are transported to the stone processing system. The stack gas includes
a 5,670 kg/hr (12,500 Ib/hr) recycle stream from the stone processing
system.
Air and Stack Gas Compression Components
The major reaction system blower supplies air and stack gas to
the gaslfier at a rate of 62,600 kg/hr (138,000 Ib/hr). The exit pressure
produced by this blower is 161 kPa (23.3 psia). The regenerator air
blower supplies 3,860 kg/hr (8,500 Ib/hr) of air at 165 kPa (24.0 psia)
for regenerator operation. Recycled stack gas is given a 4.6 kPa (0.7 psi)
boost prior to passing through the stack-gas baghouse.
Fuel-Gas Handling Components
Four parallel refractory-lined pipes carry the hot fuel gas
from the gasifier vessel to the four boiler burners. All pipe bends and
transition regions are smoothed to prevent turbulent mixing regions
which lead to carbon/fines deposits. The entrances to the cyclones are
points where deposits are likely to form, which causes an increase in
system pressure drop. The demonstration plant is equipped with a steam/
air injection point at the cyclone inlet to remove deposits and permit
continuous gasifier operation.
Heat Exchange Components
Heat exchange equipment is utilized in the reaction system for
the purposes of start-up heating of the gasifier and regenerator and for
cooling the regenerator sulfur dioxide gas prior to its processing in
the stone processing system.
The sulfur dioxide gas from the regenerator is cooled from 1075°C
(1970eF) to 315°C (600°F) in a waste heat boiler to raise 1033 kPa (150 psig)
steam. The 315°C (600°F) sulfur dioxide gas is then used to preheat the
regenerator air stream. The resulting 149°C (300°F) sulfur dioxide gas
stream passes through a baghouse for fines collection.
18
-------
Fines from
Reaction
System
Recycle Gas
to
Reaction System
YY Functional Components
(D Absorber
© Stone Pulverizer
Gas Compression
Particulate Removal
Heat Exchange
Waste Stone
from the
Reaction
System
Sulfate Stone
to Waste Stone
Handling System
S02-Gas
/from the
Reaction
System
Figure B-2-Stone processing system (dry sulfation)
-------
Purge and Transport Gas Component
A small portion of the recycled stack gas (about 635 kg/hr
[1400 lb/hr]) is used for instrument purge gas and for solids transport
in the demonstration plant with stone streams which cannot be transported
with air. The purge and transport gas is withdrawn after recycling in the
stack-gas baghouse, is compressed to about 345 kPa (35 psig), and is
cooled to about 38°C (100°F). The gas cooler is followed by a water
separator to eliminate the corrosive water vapor in the gas. The small
amount of separated water contains sulfur dioxide and is recycled to the
gasifier,where it is combined with the fuel-oil feed stream. The purge
and transport gas is then distributed to numerous components in the plant,
the major user being the pulsed stone circulation component.
Gasifier and Regenerator Vessels
The gasifier and regenerator vessels are conventional refractory-
lined fluidized bed vessels with carbon-steel shells. The gasifier has
an inner diameter of about 5.95 m (19.5 ft). The design basis bed depth
is 1.22 m (4 ft), and the freeboard height is about 5.5 m (18 ft). The
air distributor is of a refractory construction to allow a start-up temperature
below the grid of 760°C (1400°F). Fuel oil is injected into the gasifier
bed at a level of about 15.2 cm (6 in) above the grid through three
independent oil rings surrounding the vessel which distribute oil to
2
about 60 horizontal injection tubes, with one injection point per 0.47 m
2
(5 ft ) of bed cross section. A single vertical baffle wall is placed
in the bed and the regenerated stone inlet to the gasifier is situated
on the side of the baffle opposite to the gasifier stone outlet to
guarantee that the general circulation pattern will not bypass the bulk of
the bed. The transport leg outlet is set at a 0.92 m (3 ft) bed
depth, and the inlet from the regenerator is set at a 15.2 cm (6 in) bed
depth. Four fuel-gas outlets are arranged in the top of the gasifier.
The regenerator design is similar to that of the gasifier. The
regenerator vessel has a 1.37 m (4.5 ft) diameter and is designed on the
basis of a 1090°C (2000°F) operating temperature. No flow baffles are
20
-------
utilized in this regenerator. In addition to the circulating stone inlet
and outlet, the regenerator has a waste stone drain line located in the
wall at about a 0.92 m (3 ft) bed depth.
Control
The major control points in the reaction section of the
demonstration plant are the gasifier temperature and bed depth, and the
regenerator temperature and bed depth. The gasifier temperature is
controlled at about 870°C (1600°F) by the rate at which stack gas is
recycled to the unit. The demonstration plant can also utilize steam
or water injection as a method for gasifier temperature control. The
t.
gasifier bed depth is controlled by the rate at which stone is withdrawn
from the regenerator through the waste stone line. The pressure
differential across the gasifier bed controls the transport gas pulsing
rate to this withdrawal leg.
The regenerator temperature is controlled by the rate of stone
circulation between the gasifier and regenerator. The transport gas
pulsing rate to the gasifier-to-regenerator transport leg is controlled
by the regenerator temperature measurement. The regenerator bed depth
is controlled by the regenerator-to-gasifier transport leg pulsing rate.
An emergency quench system also exists for the regenerator in case of a
temperature overshoot.
The performance of the gasifier and regenerator is not too
sensitive to these control points so the control system response may
be slow over a large operating range. Other important performance
variables, such as the gasifier sulfur removal efficiency and the regener-
ator product-gas sulfur dioxide and carbon dioxide contents, will be
monitored and controlled manually by setting the fresh limestone feed rate,
the air-to-fuel ratio, the regenerator air rate, and the set points for
the automatically controlled parameters.
Turndown
The demonstration plant is designed to operate stably at
21
-------
specified levels of performance down to a partial load of 25 percent of
full plant capacity. The stability of the reaction system is limited by
the minimum operable fluidization velocity in the regenerator and gasifier.
The minimum operable fluidization velocity of these units is one-half of
the full capacity design velocity, so a turndown of the reaction system to
50 percent of full capacity can be carried out by a simple linear reduction
by 50 percent in the capacity of all process streams. Further reduction
in the plant output must be carried out by increasing the gasifier air/
fuel ratio to such a degree that the 0.92 m/sec (3 ft/sec) minimum operating
fluidization velocity is maintained.
The minimum operable regenerator fluidization velocity of 0.92 m/sec
(3 ft/sec) is maintained by operating the regenerator with high excess air.
This procedure may reduce the regenerator performance (product-gas sulfur
dioxide concentration reduced and rate of calcium sulfate generation increased)
but will not disrupt the plant performance significantly.
Turndown of the other components in the reaction system is
facilitated by conventional practices.
Flexibility
The reaction system in the demonstration plant is designed on
the basis of the best existing pilot-plant data, but the demonstration
plant Is also designed with the flexibility of an experimental installation
to reflect the uncertainties involved in the scale-up of the pilot-plant
information, the integration of the process with a full-scale power plant,
the turndown performance of the plant, and operating procedures;
and also to allow the testing of many operating and design options. The
capacity and complexity of all the components in the reaction system
are increased by this requirement for flexibility.
STONE PROCESSING SYSTEM
The boundaries and components of the stone processing system are
defined in Figure B-2. Dry sulfation has been selected as a means to
process the waste stone from the reaction system in the demonstration plant.
22
-------
The dry sulfation stone processing system consists of an absorber com-
ponent, a stone pulverizer component, gas-compression components, particu-
late removal components, and heat exchange components*as is shown in
Figure B-2.
General
The dry sulfation process contacts the waste stone from the
reaction system and the boiler cyclone with the sulfur dioxide product-
gas from the regenerator to produce a highly sulfated by-product stone
and a low concentration sulfur dioxide gas stream which is recycled to
the gasifier. This contacting is carried out at about 870°C (1600°F) in
a fluidlzed bed absorber. In order to enhance the rate of absorption and
the degree of sulfation of the stone, the stone is pulverized while hot
to an average size of about 75 microns before being contacted with the
sulfur dioxide. The kinetics of this reaction have been described in
Appendix L. The sulfated by-product stone from the absorber is cooled
and transported to the waste stone handling system.
Stone Pulverizing Component
The 1075°C (1970°F) waste stone from the regenerator is pulver-
ized by jet grinding in a commercial grinding device (Fluid Energy
Processing and Equipment Co., Jet-0-Mizer). A single pulverizer
(refractory lined) handles about 435 kg/hr (960 Ib/hr) of stone, producing
material with an average diameter of about 75 microns. About 1130 kg/hr
(2,500 Ib/hr) of air at 517 kPa (75 psla) is required for the jet
pulverizing operation.
Gas-Compression Components
Three blowers are utilized in the stone processing system:
A compressor for stone pulverization
An air blower for the absorber, which provides air tempera-
ture excess oxveen in the absorber
23
-------
A blower which handles the sulfur dioxide gas stream from
the reaction system.
The pulverizer compressor requirements are 1130 kg/hr (2500 Ib/hr) air at
517 kPa (75 psia). The absorber air requirements are about 1500 kg/hr
(3,300 Ib/hr) air at 159 kPa (23 psia). The sulfur dioxide blower
compresses about 3630 kg/hr (8,000 Ib/hr) of gas from 97 kPa (14.0 psia)
to 155 kPa (22.5 psia).
Particulate Removal Components
High efficiency refractory-lined cyclones are located at the
pulverizer and absorber exits. The pulverizer cyclone removes about
95 percent of the pulverized solids and transports them to the absorber
by gravity drainage. The absorber product-gas is combined with the
pulverizer air stream; and the combined streams pass through the absorber
cyclone, where 95 percent of the fines are removed and drained to the
waste stone cooler.
Heat Exchange Components
Sulfated stone from the absorber and fines drained from the
absorber cyclone are cooled and conveyed in a Holoflite screw conveyor
at a rate of about 1360 kg/hr (3000 Ib/hr). These solids are handled
by the waste stone handling system.
A conventional oil burner is utilized to heat the absorber air
for start-up and for absorber turndown where the rate of heat generation
in the absorber will not maintain the 870°C (1600°F) absorber temperature.
Absorber
The absorber is a fluidized bed reactor having a design very
similar to the gasifier vessel. Its inner diameter is about 6.1m
(20 ft). The bed depth is 1.22 m (4 ft), and the freeboard height is
about 5.5 m (18 ft). No baffle flow restrictions are utilized in
the absorber because of the long stone residence time (^50 hr). A
24
-------
refractory distributor design similar to that of the gasifier distributor
is used. A single stone inlet point handles the pulverized stone from
the pulverizer cyclone and the fines streams from other points in the
plant. A single stone outlet removes the sulfated stone from the
absorber. The outlet is in the form of an overflow wire which controls
the bed height. The fluidization velocity in the unit is about 0.12 ra/sec
(0.40 ft/sec), which is many times the minimum fluidization velocity of
the pulverized material. For this reason turndown of the absorber is
facilitated simply by reducing the fluidization velocity proportionally
to the plant output.
SUPPORT SYSTEMS
The limestone handling systems, the fuel-oil handling system,
and the waste stone handling system function to support the two major
processing systems in the demonstration plant which have been described.
These three support systems are all of a conventional nature and are not
described in any detail. The limestone handling system receives, stores,
and feeds limestone at a controlled rate of about 1135 kg/hr (2500) Ib/hr
to the gasifier. The fuel-oil handling system receives, stores, and
feeds high-sulfur residual oil to the gasifier at 12,700 kg/hr (28,000 Ib/hr).
The waste stone handling system transports and stores waste stone at
about 1360 kg/hr (3,000 Ib/hr).
25
-------
REFERENCES
1. Craig, J.W.T., G. Moss, J.H. Taylor, D.E. Tisdall. Sulfur Retention
in Fluidized Beds of Line under Reducing Conditions.(Proceedings of
the 3rd International Conference on Fluidized Bed Combustion, EPA,
Hueston Woods, Ohio, 1972.) p. 379.
2. . Study of Chemically Active Fluid Bed Process for Sulfur
Removal during Gasification of Heavy Fuel Oil. Interim Report for
1974, EPA Contract No. 68-02-0300.
3. . Final Report for 1974, EPA Contract No. 68-02-0300.
26
-------
APPENDIX C
DEMONSTRATION PLANT SITE
-------
APPENDIX C
DEMONSTRATION PLANT SITE
The 46 MW unit, No. 12, at the Manchester Street Station of
Narragansett Electric Company, Providence, R.I., was selected by New
England Electric System for the demonstration plant program. Power plant
layout drawings and an elevation drawing of unit No. 12 are shown in
Figures C-l, C-2, and C-3. A summary of the plant design and operating
characteristics is presented in Table C-l.
27
-------
Table C-l
POWER PLANT DESIGN AND OPERATING CHARACTERISTICS
Unit No. 12
Manchester Street Station
Capacity
Heat rate
Steam conditions
Burners
Number
Capacity
Fuel
Operation
46 MW
1.48 MJ/kWh
8446 kPa (1225 psi)/510°C (950eF)
6
84.4 GJ/hr (80 x 106 Btu/hr)
coal 1949-1967
gas 1967-1970
oil 1970-present
daily cycling
on-line between 8 and 50 MW
16-20 hr/day
28
-------
PAGE NOT
AVAILABLE
DIGITALLY
-------
PAGE NOT
AVAILABLE
DIGITALLY
-------
O-IIS-9
COMPANION SHtCTl
Pt+" JfoSmT A Xt*-S
os- Plor*
' O ft 4
frENERAL SECTION THUQUGM UNIT, ^11 LOQKMfr WEST
Figure C-3-Manchester Street Station basement floor
CLKTMC CO
MANCHESTER ST. STATION
POWEH PLANT
SECTIONS
-------
APPENDIX D _
FLUIDIZED BED OIL GASIFICATION
DEMONSTRATION PLANT DESIGN BASIS
-------
APPENDIX D
FLUIDIZED BED OIL GASIFICATION
DEMONSTRATION PLANT DESIGN BASIS
The design basis for the demonstration plant is tabulated
under four general divisions - overall plant considerations and reaction,
stone processing, and support sections. Explanations for the listed
bases are included as notes.
OVERALL PLANT CONSIDERATIONS
Plant - New England Electric System (NEES) Manchester Street
Station, boiler No. 12
Plant size - * 50 MW
Plant sulfur removal efficiency - 90%
Plant turndown ratio - 4:1
Plant points of sulfur (S) emissions - from boiler stack
only; all other sulfur-containing streams are recycled to
gasifler.
Boiler excess air - 6%
Stack gas temperature - 149°C (300°F)
Boiler cyclone (existing) efficiency - 85%
Fines elutriated to boiler - all calcium sulfide (CaS) in the
fines are converted to calcium oxide (CaO) and sulfur dioxide
(SO.): all carbon on fines is burned off; all fines pass
through the boiler.
RFACTION SECTION ~ MAJOR COMPONENTS
Gasifier
Chemistry - See Figure D-l
Operating conditions
35
-------
1
Limestone
^
^
Oil
i Fuel
Gas
to
Boiler
j^^ ^^^^"^^ .^^b.
^
871° C
(1600° F)
Reducing
h Pnrlr i
ng
Oxidizing '^-
fT>. . .-Oi QI i
%l ivxj KXXI K//vl K//O| K/xl K?
-------
- Bed temperature - 880°C2 (1616°F)
3
- Gas space pressure - 17.9 kPa (2.6 psig)
A
- Maximum bed depth - 1.22 m (4 ft)
- Superficial fluidization velocity - 1.83 m/s (6 ft/sec)
Flow rates and composition
- Fuel oil rate - 12,700 kg/hr (28,000 Ib/hr)
- Fuel oil composition - see Table D-l
- Fuel oil temperature - 93°C (200°F)
- Air rate - air/fuel ratio of 20% of stoichiometric
- Limestone feed rate - 0.5 to 1.5 x stoichiometric,
normal being stoichiometric based on fuel-oil/sulfur feed rate
- Limestone composition - see Table D-l
- Limestone size = 300-3175 microns (u)
- Fuel-gas composition - see Table 072°
- Sulfided lime composition - 3 wt% S, 0.3 wt% S as calcium
sulfate (CaSO^)
- Sulfur removal efficiency - 90% overall plant
- Fines elutriation rate - 3,175 kg/hr (7,000 Ib/hr)9
- Elutriated fines composition - same as bed
- Stack gas for temperature control - ^35,000 kg/hr (77,200 Ib/hr)
- Stack-gas composition - based on 6% excess air to
boiler burners
- Metals retention - 100% vanadium (V), 75% nickel (Ni), 36%
sodium (Na)
Operation
- Temperature control - stack-gas recycle
- Turndown ratio - 4/1
- Turndown method - adjustment of air/fuel ratio to minimize
12
velocity reduction during turndown
- Burnout - provision for continuous steam/air injection
into gasifier cyclone entrances/or flue gas burnout during
13
shutdown
- Staging - no special considerations for horizontal staging
of gasifier
37
-------
Table D-l
FLUIDIZED BED OIL GASIFICATION DEMONSTRATION PLANT
Raw Materials and Product Gas
Material
Composition,
wt %
Metals content,
ppm
Limestone - BCR 1359
CaO
MgO
Fe2°3
A12°3
co2
S (total)
V
Na
Ni
54.1
0.60
0.75
0.09
0.31
44.0
0.12
50
< 20
30
Fuel Oil3, No. 6
S
C
H2
°2
W2
V
Na
Ni
I
2.6 (3.0 max.)
85.9
11.0
0.1
0.4
442
30
52
14
300 SSU at 60°C (140'F) viscosity;
f n
41.8 x 10 kJ/m (150,000 Btu/gal) gross (high) heat value)
12.8° API gravity
38
-------
Table D-2
FLUIDIZED BED OIL GASIFICATION PROCESS
FUEL-GAS CHARACTERISTICS
Composition
Mole
%a
Weight
%b
CO
57.0
10.3
9.7
10.0
6.5
5.8
0.5
0.2
Condensable tars, soot
58.70
9.55
14.60
0.28
2.31
4.96
0.41
0.37
5.61 C + 0.92
0.20
1.99
Actual product analysis.
Estimated chemical analysis.
39
-------
- Start-up - external direct heating air heater for 788°C
(1450°F) start-up air
Design factors
- Freeboard - 6.1 m (20 ft)
- Baffles - single vertical baffle
- Number of fuel-gas outlets - 4 (one per burner)
- Stone inlet location - *v 127 mm (5 in) above distributor
- Stone outlet location - 0.916 m (3 ft) above distributor
- Distributor - single plate of refractory construction
(Harbison-Walker design)
. 12.4 kPa (1.8 psi) maximum AP at maximum flow rate
. 38.2 ro/s (125 ft/sec) efflux velocity at maximum
flow rate
2 2
. 10.75 chimney per m (1 chimney per ft )/4 holes per
chimney
. Start-up temperature of 788°C (1450°F)
- Fuel-oil injection - horizontal pipes with one Injection
.46
19
2 2
point per 0.465 m (5 ft ) set at 127 mm (5 in) above
distributor
- Gaslfier elevation - minimum required for construction
and maintenance (to minimize fuel-gas piping length)
20
- Number of gasifier modules - 1.
Regenerator
21
Chemistry - see Figure D-2
Operating conditions
- Bed temperature - 1080°C (1976°F)22
- Gas space pressure - 0-2488 Pa (0-10 in) H.O
23 i
less than gasifier pressure
- Maximum bed depth - 1.22 m (4 ft)
- Superficial fluidization velocity - 1.83 m/s (6 ft/sec)
Flow rates and compositions
40
-------
Dwg. 6238A99
Regenerated
Stone
S02 Gas
Sulfided
Stone
1093°C
(2000° F)
CaS + 3/2
CaS +20,
2 ^± CaO
^± CaS0
\
Waste
Stone
Air
Figure D-2-Regenerator fundamentals
41
-------
- Product-gas composition - 8 mole % SO., 6%
carbon dioxide (CO.), 0.05 percent oxygen (0 ), 2%
24
KyO, N
25
- Regenerated stone composition - CaO, CaS, CaSO,, inerts
- Fines elutriation rate - 191 kg/hr (420 lb/hr)26
- Fines composition - same as bed
- Air - preheated by exchange with regenerator SO. product
Operation
- Temperature control - stone circulation rate, with
27
emergency quench for temperature overshoot
- Turndown method - reduction in regenerator SO. concentration
when maximum velocity reduction is reached by excess air/stack
28
gas
29
- Staging - no special considerations for horizontal staging
- Start-up - utilize gasifier start-up heater
Design factors
- Freeboard - 5.5 m (18 ft)
29
- Baffles - none
- Stone inlet location - > 127 mm (5 in) above distributor
31
- Stone outlet location - 0.916 m (3 ft) above distributor
- Distributor - (Harbison-Walker)
. 13.8 kPa (2 psi) pressure drop at maximum flow rate
.38.2 m/s (125 ft/sec) efflux velocity18
- Regenerator elevation - same as gasifier
32
- Number of regenerator modules - 1
- 'Regenerator primary cyclone - remove 80% of fines >
300y and returns to regenerator
- Regenerator secondary cyclone - high efficiency (95%)
with fines drained to pulverizer.
Solids Circulation System
Flows
- Solids rate - up to 45,360 kg/hr (100,000 lb/hr)3A
42
-------
o
- Carrier gas - about 80.3 dm /S (170 scfm) per pulser
(maximum)
Operation
36
- Turndown - by flue gas pulsing rate
- Constraints - solids cannot sit in presence of oxygen
due to the possibility of the occurrence of particle
agglomeration
Design factors
- Valves - camflex type
- General - fit into cylindrical gasifler and regenerator
38
vessel designs as refractory block (Harbison-Walker)
- Gasifier/regenerator separation - about 0.305 m (1 ft)
39
,39
- Transfer duct height - 127 mm (5 in)39
- Transfer duct width - 645 mm (29 in)'
- Flue gas distribution for pulsing - 508 mm (2 in)
I.D. pipe with 803 mm (3.16 in) holes spaced at 25.4 mm
(1 in) intervals and located horizontally across duct
at duct base.
40
- Thermal expansion - regenerator mounted on flexible base.
Gasifier Cyclones and Fine Return - Design Factors
Number of cyclones - 4
Cyclone overall efficiency - 90%
Cyclone elevation - set to minimize fuel-gas piping length
Cyclone fines return - Sturtevant Pulse Phase Powder Conveyor
Number of Sturtevant Systems - 4 (I/cyclone)-
Fines return distribution - two return fines to the gasifier,
two return fines to the entrance of the gasifier-to-regenerator
41
stone transport line
42
Fines return system cycle time - about three minutes
43
Valves - use high-temperature valves for Sturtevant system.
43
-------
Fuel Gas Ducts - Design Factors
Design velocity - increase from pilot plant basis of 18.3 tn/s
0^60 ft/sec) to 45.8 m/s (150 ft/sec)
Length - minimum required
Shape - smooth entrances, bends, and so on, to reduce turbu-
lence and resulting deposits at these points
44
Valves - eliminate valves in fuel-gas lines.
Air and Stack-Gas Recycle Circuits
Pressure drop - minimize losses by proper design of blower
controls, air lines, and so forth (13.8 kPa [2 psi] drop
45
estimated between gasifier blower exit and gasifier plenum)
Blower - single gasifier blower for air and recycled stack gas.
Other Items
Burners - 4
Start-up heater
Baghouse - handles stack-gas recycle and recycled absorber
46
product-gas fines sent to absorber
47
Purge transport gas (PT6) blower - assumes 310 kPa (45 psia)
exit pressure
Stack-gas recycle fan
Waste heat boiler - cools regenerator SO.-gas to 315°C (600°F)
Regenerator air preheater - regenerative heating to cool
regenerator S02-gas from 315°C (600°F) to 149°C (300°F)
PTG cooler and water separator
Limestone feeding system - based on 1.5 x stoichlometric
STONE PROCESSING SECTION MAJOR COMPONENTS - DRY SULFATION CONCEPT48
Spent Stone Pulverizer
Operating conditions - maximum temperature - 500°C (932°F)
Flow rates
44
-------
- Spent stone from regenerator drain and second-stage
cyclone flow to pulverizer by Esso pulsed flow
- Spent stone size 10 mesh (see Regenerator)
- Pulverized stone size - 95% 900u (170 mesh), 5% 640p
(230 mesh); average size - 75y (200 mesh)
- Pulverizer air rate - air at 515 kPa (74.7 psia); based
on pulverizer requirements
Operation
- Turndown - reduced pulverizing efficiency as velocity is
reduced
- Temperature control - steam injection can be used to
49
avoid high-temperature conditions
Design Factors
- Pulverizer type - Jet-0-Mlzer, Fluid Energy Processing
and Equipment Company, Hatfield, Pennsylvania
- Number of units - 1
- Transfer of stone from regenerator to pulverizer - by
Esso solids transport scheme
2 ?
. About 25.8 cm (4 in ) duct cross-sectional area
3
. About 1.415 dm /s (3 cfm) transport gas required.
Pulverizer Air Compressor - Operating Conditions
Inlet pressure - 100 kPa (14.6 psia)
Outlet pressure - 515 kPa (74.7 psia)
Inlet temperature - ambient.
Pulverizer Cyclone - Design Factors
Type - high efficiency (95%); conventional design
Number - 1
Cyclone drain - into absorber by dense sta'ndleg
Pulverizer gas - to inlet of absorber cyclone.
45
-------
Absorber
Operating Conditions
- Bed temperature - 870°C (1608°F)51
- Maximum bed depth - 1.22 m (4 ft)52
- Superficial fluidization velocity - 0.122 m/s (0.40 ft/sec)
Flow rates and composition
- Regenerator SO.-gas rate - see Regenerator
- Regenerator SO.-gas composition - see Regenerator
54
- Absorber air rate based on temperature control
- Pulverizer stone rate - see Pulverizer
- Pulverizer stone composition - see Regenerator
- SO. absorption efficiency - 90%
- Fines elutrlation rate - 68 kg/hr (150 lb/hr)55
- Fines particle size - < 30p
- Fines composition - same as absorber bed material
- Sulfated stone particle size - same as pulverizer
product stone
Operation
- Temperature control - by excess air control or fuel
combustion
- Turndown method - velocity reduction
- Start-up - fuel combustion flue gas into distributor
Design Factors
58
- Type of reactor - fluid bed
- Staging (horizontal) - none
- Staging (vertical) - none; space available for second-
stage absorber if necessary
- Number of units - I59
- Freeboard height - 6.1 m (20 ft)55
- Baffles - none
- Pulverized stone inlet - set at absorber wall at maximum
bed depth level of 1.22 m (4 ft)61
- Sulfated stone outlet - single overflow outlet set at
46
-------
wall opposite stone inlet at 1.22 m (4 ft) above distri-
em
62
butor; dense phase drain to stone cooler
- Distributor
. Single refractory construction
. 13.8 kPa (2 psi) pressure drop based on maximum flow
. 38.2 m/s (125 ft/sec) efflux velocity - see Gasifier
7 2 63
. 172/m (16 chimneys per ft )/8 holes per chimney
. Start-up temperature of 788°C (1450°F).
Absorber Cyclone
Operating conditions
- Temperature - 870°C (1608°F) (absorber gas inlet)
- 500°C (932°F) maximum (pulverizer gas inlet)
Flow rates - see Absorber and Pulverizer
Design factors
- Type - high efficiency, such as Aerodyne
- Number - 1 (depending on sizes available)
- Efficiency - removes 98% > lOy
- Cyclone drain - into stone cooler by dense phase standleg.
Absorber Air Blower
Operating conditions
- Inlet pressure - 100 kPa (14.6 psia)
- Outlet pressure - 155 kPa (22.5 psia)
- Inlet temperature - ambient
Flow rate
- Air rate - set by temperature control of absorber and
excess oxygen requirement for reaction.
Absorber SO^ Blower
Operating conditions
- Inlet pressure - 98 kPa (14.2 psia)
47
-------
- Outlet pressure - 155 kPa (22.5 psia)
- Inlet temperature - 149°C (300°F)
Flow rate - see Regenerator.
Regenerator SO -Gas Bag Filter
Operating conditions - inlet temperature - 149°C (300°F)
Flow rates - see Regenerator
Design factors - type: Reverse jet cleaning Sly Dynaclone
Air/cloth ratio - 2.5:1.
Stone Cooler/Conveyor
Operating conditions
- Stone inlet temperature - 870°C (1608°F)
- Stone outlet temperature - 150°C (302°F)
Flow rates and compositions - see Absorber.
SUPPORT SECTIONS
Equipment for waste stone handling and for fuel oil storage is
of a conventional nature and is based on normal limestone and fuel oil
requirements.
ASSUMPTIONS FOR ENERGY AND MATERIAL BALANCES
Common assumptions have been made for generating energy and
material balances having sufficient accuracy for a preliminary design
task. The assumptions include:
No heat losses
Steady-state with no accumulations
Air to blowers is at 25°C (77°F) and bone dry
Purge and transport gas streams have been neglected
The pulsed nature of solids transport systems have been
neglected by assuming steady flow behavior
48
-------
When possible, data and correlations generated by Esso experimental
programs have been utilized to carry out balancesfor example, for fuel-gas
compositions and gasifier heat generation rates. Fines carried over in the
gas streams have been included in the balances because they are an important
64
operating and environmental consideration.
49
-------
NOTES
1. Gasifier Chemistry
Due to the endothermic nature of the calcination reaction, the
limestone feed rate must be continuous rather than periodic.
The calcium sulfate (CaSO.) produced in the regenerator is com-
pletely reduced to calcium sulfide (CaS) in the gasifier. A
low level of residual CaSO. circulates between the gasifier
and regenerator.
The fuel oil cracks immediately to liberate hydrogen sulfide
(H.S), carbonyl sulfide (COS), and carbon disulfide (CS.),
the bulk being H S.
An important balance exists between the air/fuel ratio and
gasifier temperature in terms of the rate of carbon laydown
and burn-off in the oxidizing region. Lower air/fuel ratios
require higher gasifier temperatures to Increase the carbon
burn-off rate. s
An important balance exists between stone carbon level and
sulfur removal efficiency. Too low a carbon level permits
SO- to be released in the oxidizing region once the stone
carbon layer is burned off. Once released, SO. is difficult
to reabsorb in the bed.
The addition of water or steam to the gasifier has an important
effect on the sulfur removal efficiency. The sulfur removal
efficiency may be reduced because of thermodynamic effects,
reduced kinetic driving force, or reduced bed-carbon level.
Operation at lower air/fuel ratios (< 20 percent), where
higher carbon levels result, may permit the use of water or
steam for temperature control.
2. The bed may be required to operate over a range of temperatures from
850 to 900°C (1562 to 1652°F) depending on the air/fuel ratio.
3. Reflects initial Stone and Webster Engineering Corporation (SWEC)
design study.
4. Bed depths down to 0.61 m (2 ft) may be utilized. Deeper beds may
improve sulfur removal efficiency and reduce the rate of carbon
50
-------
build-up in the fuel-gas lines due to increased bed surface for carbon
laydown or cyclone inlet scouring due to increased elutriation of oarticles.
5. Minimum operable superficial velocity * 0.915 m/s (3.0 ft/sec). Velocities
lower than 1.83 m/s (6 ft/sec) will result at lower air/fuel ratios than
20% of stoichiometric at full capacity.
6. A minimum air/fuel ratio of ^15% gives adiabatic operation.
Includes oxygen supplied by stack-gas recycle. Twenty percent has
been demonstrated and should be the maximum required.
7. A maximum of 1.5 x stoichiometric permits faster bed depth build-up
and operation with higher sulfur fuel oils.
8. An example of an experimentally determined composition. Composition
will change with operating conditions and process design.
9. Based on recent pilot plant data (run 8) the elutriation rate could
be twice as great as this amount (90.7 kg/hr [i, 200 16/hr] in pilot
plant). Because the freeboard in the demonstration design is about
twice that of the pilot plant, the 3180 kg/hr (7,000 Ib/hr) figure is
considered reasonably conservative.
10. Pilot runs have indicated fines compositions such as 6 wt % CaSO^ and
2% carbon. Because of the high carbon content of the gasifier fines,
care must be taken In returning all or part of them to the regener-
ator, if this is done as a batch, because of the high heat load.
11. The amount of stack gas required for temperature control depends
greatly on the air/fuel ratio. This amount is based on 20%
air/fuel ratio. This amount is also sufficient for burnout. The
options of water, steam, or indirect cooling will not be considered
in the preliminary design for temperature control.
12. Only a 50% reduction in gasifier velocity is permitted if
operating at 1.83 m/s (6 ft/sec) in order to maintain fluidizatlon.
By increasing the air/fuel ratio (and therefore the stack-gas recycle
rate) as the fuel is turned down, a high gasifier velocity can be
maintained. The limiting factor on how high an air/fuel ratio can be
utilized may be the heating value of the fuel gas produced during
turndown.
51
-------
13. Recent pilot plant information indicates that deeper beds (up to
1.22 m [4 ft] reduce the rate of carbon build-up in the cyclone
inlets due to increased bed surface for carbon laydown or increased
solids rates through the cyclones. Water and steam injection into
the bed also reduces carbon build-up. Steam/air injection into the
cyclone entrance has not yet been demonstrated. The 200-hour operating
times on the pilot plant have been scaled to ^2000-hour operating
time between burnouts on the demonstration plant.
14. Horizontal staging of the gasifier has been proposed as a method
of improving gasifier sulfur removal performance. Because of the
increased complexity, this option should not be utilized unless it is
demonstrated that sufficient sulfur removal cannot be achieved
without staging.
15. Esso Research Centre, Abingdon, England (Esso) has proposed utilizing
a vertical baffle connected to a central core of circular cross-section
in the gasifier to permit staging, reduce distributor structural
problems, and reduce the lengths required for fuel injectors. This
advanced design is not required for the preliminary design study.
16. The pressure drop across a 1.22 m (4 ft) deep bed (maximum depth) is
about 13.8 kPa (2 psl). The distributor pressure drop should be
^ 25 percent of the bed pressure drop for good distribution at the
lowest operating velocity, which is one-half of the maximum velocity
at 4:1 turndown. At 0.915 m/s (3 ft/sec) the distributor pressure
drop then equals 3.45 kPa (1/2 psl) and at 1.83 m/s (6 ft/sec) it
will be 13.8 kPa (2 psi). This has been reduced to 12.4 kPa (1.8 psi)
for structural reasons.
17. An efflux velocity too high may cause excessive particle attrition
at the air jets.
18. The start-up temperature should be high enough to prevent recarbonatlon of
the lime bed and defluidization caused by increased bed density.
19. Spacing of fuel injection points based on pilot plant tests. Recent
tests with a single vertical fuel injector having adjustable height
indicate that the elevation of the fuel injection point above the
52
-------
distributor has little effect on gasifier performance, except as it
affects the active bed height.
20. Larger installations (> 100 MW) may use multiple gaslfiers which will
simplify turndown. Small units could use multiple gasifiers,1if
optimization studies indicate an economic advantage.
21. Regenerator chemistry
Carbon deposited on the stone surface is combusted before
any stone regeneration takes place, according to batch results.
Carbon combustion on the particle surface permits the stone
to heat up to the bed temperature before the competing
reactions listed in Figure D-2 can progress. Thus, the formation
of CaO and SO. is favored. Pilot plant operation, with low
carbon levels on the stone, have provided poor regenerator
performance (i.e., high CaSO, and low SO.).
Stagnant regions in the regenerator, or regions where air
leaks into the regenerator during shutdown, lead to large
agglomerated deposits of stone.
22. Regenerator temperature > 1100°C (2012°F) should be avoided.
23. Based on pilot plant experience. Pressure arranged to draw fuel
gas into regenerator rather than SO.-gas into gasifier to maximize
sulfur removal. No seal is required between the vessels.
24. With the dry sulfation reaction scheme, lower SO. levels from the
regenerator can be tolerated. This additional degree of freedom may
permit operating the gasifier with improved performance. Also, the
SO.-fuel from the regenerator may be changed during turndown without
reducing the plant performance. Thus, no SO. recycle system is
required as was the case in the previous SWEC design with sulfur
recovery.
25. Small amounts of inerts and V, Ni, Na from the fuel oil are also
present.
26. Based on fuel-gas dust loading - very conservative estimate.
27. Required circulation rate depends on stone carbon level and
selectivity of the regenerator to CaO formation rather than CaSO,.
A maximum stone carbon level exists above which the regenerator
53
-------
temperature cannot be controlled by circulation rate (^ 1 wt %).
28. Some fuel combustion may be required in the regenerator at turndown
greater than 50 percent to maintain the bed temperature. Two
regenerators could be used if this mode of turndown were difficult.
29. Esso has proposed staging of the regenerator to improve
performance. Would require a single baffle and bed temperature
gradient measurement for control.
30. The diameter of the regenerator used in the initial SWEC design is
somewhat small. This factor will not affect the resulting cost
projection sufficiently to require resizing the regenerator.
31. For bed levels less than 0.915 m (3 ft) the normal splash rate of
particles is sufficient for solids circulation.
32. With multiple gasifier modules a single regenerator vessel per
gasifier would probably be used. Two regenerators could be used
in the demonstration plant if regenerator turndown were expected to
be a problem.
33. The regenerator acts as an elutriator to eliminate fines from the
system which would otherwise be circulated continuously through the
regenerator-cyclone circuit.
34. Normal pilot plant operation would indicate a circulation rate of
about 24,950 kg/hr (55,000 Ib/hr). This excess factor accounts for
operation under unusual circumstances, with higher sulfur fuel oils,
or with limestones having higher inert contents, and so on.
3
35. Based on pilot plant transport gas rate of 1 dm per 0.16 kg (1 cf
per 10 Ibs ) solids transferred.
36. Full-scale cold modeling will be carried out prior to designing
detailed solids circulation system.
37. Stone circulation with flue gas rather than nitrogen as the transport
gas is due to be demonstrated in run 9.
38. Esso has suggested an advanced design which integrates the
pulsed stone circulation system into the gasifier and regenerator
designs. The cost difference is not expected to be significant.
39. Based on direct pilot plant scale-up for rate of 24,950 kg/hr
54
-------
(55,000 Ib/hr) Large duct dimensions should reduce wall effects
present in pilot plant system, making the projected design capable
of higher transport rates.
40. Alternatively, expansion joints between the vessels and the
refractory block could be utilized.
41. The fines are split in this manner to permit gasifier fines to enter
the regenerator (which acts as an elutriator) and be removed from
the system. All of the gasifier fines are not sent to the regener-
ator, because of their high carbon content and the resulting high
heat load. The fines which are to be sent to the regenerator are
placed in the entrance of the gaslfier-to-regenerator stone transport
line to mix the fines with the gasifier bed material being transported,
so that hot spots will not develop periodically in the regenerator. The
fines will be pulled into the regenerator only when the circulation
system Is being pulsed.
42. Surge vessels may be required in the fines return circuits carrying
material to the regenerator in order to provide a smooth return rate
and reduce regenerator fluctuations. This is not done in the pilot
plant and is not to be considered in the preliminary design.
43. High-temperature butterfly valves supplied by the Dally Engineering
Valve Company, Pittsburgh, Pennsylvania are being used in the Consoli-
dation Coal Company (Consol), Rapid City. South Dakota, coal gasification
olant at 870 to 982°C (1600 to 1800°F) and M.033 kPa ('v-lSO psi). The
FMC Princeton Plant has used a high-temperature slide gate knife valve
(Alovco 310 SS gate valve) custom made at 870°C (1600°F). The designs
for that valve are available.
The alternative of cooling the recycled fines and striping the line
to prevent tar condensation has been considered, but the high-temperature
valve is more advantageous.
44. The valves would not function long because of carbon build-up. Also,
methods to limit build-up or burnout during operation are being
actively investigated, making valves unnecessary. The legal and
insurance aspects of valves in the fuel-gas lines must be explored.
55
-------
45. The initial SWEC design was based on a 2.44 m (8 ft) maximum bed
depth and used
Bed pressure drop - 27.6 kPa (4 psi)
Distributor pressure drop - 27.6 kPa (4 psi,)
Losses between blower and plenum - 41.3 kPa (6 psi).
46. Water injection into the recycled absorber gas may be required at
reduced stack-gas recycle rates in order to limit the baghouse
temperature to 260°C (500°F).
47. If the recycle flue-gas oxygen content becomes too great due to the
presence of the recycled absorber product gas, then the purge and
transport gas may have to be taken from the stack gas upstream of
the point at which the streams combine. The PTG stream would then
require its own baghouse prior to compression.
48. A number of process modifications have been proposed. These options
deal with how to reduce the stone particle size enough to permit
SO. absorption to a high level of efficiency with high stone
utilization (^75u); how to place the reactants into the absorber;
and how to handle the products from the absorber.
Pulverization may be by mechanical grinding after the stone has
been cooled to about 260°C (500°F), (Option e) by pneumatic
grinding of the hot stone (Jet-o-Mizer), (Options a-d), or bed
jet pulverizing directly in the absorber (Option f). These
options are shown in the accompanying simple flow diagrams.
- With external pulverization, stone may be carried to the absorber by
dilute pneumatic transport using the pulverizing air as a carrier
gas (Option a) or may be separated from the pulverizing gas in a
cyclone and sent to the absorber by dense phase standleg (Options b-d).
The pulverizer air may be utilized in the absorber (Options a, b, d)
or combined directly with the absorber gas product (Option c).
- The absorber product may be sent directly to the stack if SO.
and particulate contents are low enough, may be cooled and sent
through a bag filter before going to the stack if SO is low
enough, or may be recycled to the gasifier by combining it with
the recycled stack gas before the stack-gas baghouse.
56
-------
Jon a
in
Hot
Regenerator
Waste_
Stone
Dilute Pneumatic Transport
of Pulverized Stone
Cyclone
Fines
Option b
Hot
Regenerator
Waste
Stone
Option c
Hot
Regenerator
Waste .
Stone
Pulverizing
Air
Absorber
Product
Regenerator
S0n- Gas
Sulfated
Waste
Stone
Regenerator
Sulfated
Waste
Stone
Figure D.-3 -Stoneprocessing design options
Option d
Hot
Regenerator
Waste . je
Stone Pulw
Regen
so2-
r&
t
jrizer
| Pulverizi
Option e
Hot
Regenerator
Waste r Sti
Cot
Option f
me
Her
Hot
Regener
Wast
Ston
erator \
Gas
B
Stone
ng
Mechanical
Crusher
rl<
1
ator
B Absor
e
t
faste
Heat fr
oiler
In eater) ~
p| Cyclone .
1
1 Fines
Absorber Sulf
-x v
AiJ
r(§
1
Fuel^j
Air
yclone
Fines
"r \
1
Pulverizing Air
Air
t
clone]
Fines
er
Absorber
*" Product
Regenerator
S02 - Gas
\ Sulfat
>vWa$U
Ston
Bag 1
Filter 1 1
Absorber
" Product
Blower fc
ated
Waste
^^C^tone
.Absorber
Product
Regenerator
S02 - Gas
s. Sulfated
XWaste
Stone
ed
1
e
-------
Referring to the flow diagrams:
Option a: Combines simple pulverization system and no need for a
cyclone with the problem of distributing the gas-solid
suspension over the absorber cross-section.
Option b: Combines simple pulverization and cyclone separation of
pulverized stone with problem of distributing the separated
pulverized (dusty) air over the absorber cross-section.
Option c: Retains simple pulverization and cyclone separation but
eliminates the problem of pulverizer air distribution.
The exit oxygen content of the absorber product is reduced
from about 8 mole % in Options a, b, d to about 4 mole %
which should still give sufficiently good kinetics in
the absorber.
Option d: Eliminates the problem of pulverizer air distribution but
increases the capacity of the SO.-gas compression system
to permit utilization of the pulverizer gas.
Option e: Eliminates absorber distribution problems. May give superior
control of pulverizing quality. Requires fuel combustion
to maintain 870°C (1600°F) absorber temperature.
Option f: Simple system, but in-bed quality of pulverization is
questionable. May be feasible because of long stone
residence time (^50 hr) but may require extensive
development.
Option c has been selected for the demonstration plant preliminary
design. Small amounts of fuel may be combusted to raise the absorber
oxygen level, if required.
49. Resulting temperature with required pulverizing air rate. No lower
limit on this temperature as long as hydrating conditions are not
approached.
50. Conservative estimate is 991 dm (35 cf) air per 0.4536 kg (1 lb)
stone pulverized.
51. Based on thermodynamic and kinetic information.
58
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52. A gas residence time of three seconds should be sufficient for 90
percent absorption of SO.. This corresponds to about a 0.458 m (1.5 ft)
bed depth.
53. The minimum fluidization velocity of this material is about 0.00305 m/s
(0.01 ft/sec) Fluidization studies at room temperature have indicated
poor fluidization quality of this material, but at high velocities
mixing is fast with particle sticking at the wall. High-temperature
fluidization quality may be better (Esso finds that their
materials flow much better when hot). The maximum allowable ratio of
U/U , is about 80 based on the average particle size.
mt
54. Based on bed temperature control. Fuel may be burned to heat air
and increase the bed oxygen content. Fuel will be required for turndown
to maintain the absorber temperature.
55. The particle elutriation rate due to bubble splash should be very small
with a 6.1 m (20 ft) freeboard height, especially if very shallow
beds are used 0.61 m (^2 ft). At the operating velocity of 0.122 m/s
(0.40 ft/sec) the gas will carry out by pneumatic transport any
particles less than about 15 microns in diameter. The estimated rate
is set at 10 percent of the stone feed rate.
56. Assumed that the fines are extremely reactive. The fines may be
recycled to the absorber or combined directly with the stone waste.
If the fines are not highly reacted, they may have to go to a
separate sulfation stage.
57. The range of operable velocities is very large with very fine
particles, but in order to reduce distributor pressure drop, the
absorber air rate can remain high during turndown.
58. The absorber could also be an entrained bed reactor (or fast
fluidized bed), in which case the quality of fluidization would be
unimportant; the absorber size and possibly pulverization could be
carried out in the reactor.
59. Multiple units could be used if economic considerations indicated
an advantage.
60. Because of long stone residence time (about 50 hours) in the absorber,
59
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no baffles should be required, since the mixing rate of solids will
greatly exceed the bed net velocity.
61. Set arbitrarily to maximize inlet and outlet separation distance.
62. For the option chosen (Option c) the distributor is of a conventional
design. Options a and b require unusual distributor designs because
of the particulate content of the gas. For Option a, a disengaging
zone located at the base of the absorber having a coarse distributor
could be used, or a series of nozzles located around the circumference
of the absorber. Option b could use a nozzle concept or an independent
tubular distributor located above the refraction distributor plate.
63. For shallow bed operation with fine particles, dead spots at the
distributor should be avoided. Corresponds to 25.4 nun (1 in)
chimneys set on square lattice with 50.8 mm (2 in) separation.
64. An estimate of the size distribution of fines passing through the
boiler is shown in Figure D-4. The quantity of fines passing through
the boiler has been estimated from the following basis:
Direct scaling of pilot-plant elutriation rates to the base
design conditions indicate a fines rate of 3180 kg/hr (7000
Ib/hr) elutriated from the gasifier. This is a pessimistic
assumption since the demonstration plant freeboard will be
much greater than the pilot-plant freeboard.
Gasifier cyclone efficiencies of 95percent (optimistic) yields
a rate of 159 kg/hr (350 Ib/hr) of fines passing to the boiler.
Assuming all of the fines entering the boiler pass through the
boiler, then a dust loading of 159 kg/hr (350 Ib/hr) or 0.3 kg/GJ
(0.7 lb/10 Btu) are carried on to the existing stack-gas cyclone.
The existing stack-gas cyclone efficiency is indicated in
Figure D-5 for fly ash. If the overall cyclone efficiency with
the fines elutriated from the CAFB is 85 percent, then the stack
dust loading would be 0.043 kg/GJ (0.1 lb/10 Btu)which would
meet the federal particulate standard of 0.43 kg/GJ (0.1 lb/10 Btu).
At a low air/fuel ratio 0\»17%) and with a shallow gasifier
bed the elutriation rate from the gasifier could be an order
60
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1000
800
600
400
200
100
«. 80
I 60
§
, 40
£20
£
10
8
6
20 40 60
% Less Than Stated Size
100
Curve 674341-A
Standard Tubular Unit
Thermal Duplex
Design 6 Tubular Collector
Fly Ash Specific Gravity 2.2-
Dust Concentration 2.0 GR/ cf
Flue Gas Temperature 300° F
Resistance Primary Unit 2" W. G.
Resistance Secondary Unit 4" W. G.
6 8 10 12 14 16 18 20 22 24 26 28 30 32 34
Size of Dust-Microns
Figure D-5-Existing cyclone efficiency
Figure D-4-Estimated fines size distribution from boiler
prior to existing cyclone-Rate= 350 Ib/hr
-------
of magnitude less than with the base design conditions, or about
318 kg/hr (700 Ib/hr). A gasifier cyclone efficiency of 95 percent
would yield 16 kg/hr (35 Ib/hr) of fines passing to the boiler.
If all this material passes through the boiler and an 85 percent
efficiency is representative of the existing stack-gas cyclonet
then a dust loading of 0.004 kg/GJ (0.01 lb/10 Btu) would result
from the power plant stack.
62
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APPENDIX E
DESIGN MANUAL
E
-------
APPENDIX E
DESIGN MANUAL
This design manual was prepared as Phase I of a contract with
Stone & Webster Engineering Corporation (SWEC). The contract between
Westinghouse Electric Corporation and SWEC provided for the completion
of a preliminary design and cost estimate for the 50 MW fluidized bed
oil gasification demonstration plant with dry sulfation of waste limestone.
SWEC completed the work on December 5, 1974.
63
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DESIGN MANUAL
INDEX
1. Description of Process Elements
1.1 Chemistry and Basis of Design
1.2 Process Alternatives
1.3 System Performance Assessments
2. Utilities, Raw Materials and Products
2.1 Feeds and Products
2.2 Utilities and Fuel Summary
3. Flow Diagrams
3.1 Material Balance Flow Diagrams
3.2 Process Flow Diagrams
4. Functional Description of System Operation
5. Equipment Arrangement Drawings
6. Equipment Drawings and Specifications
6.1 Equipment List
6.2 Vessel Drawings
6.3 Equipment Specifications
6.4 Summary of Equipment Rooms
6.5 Summary of Equipment Costs
7. One Line Diagram
8. Economics
8.1 Base Conditions and Assumptions
8.2 Investment Cost Summary
8.3 Operating Cost Summary
64
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1. DESCRIPTION OF PROCESS ELEMENTS
1.1 Chemistry and Basis of Design
The basic process used in the oil gasification-desulfurization unit
was developed by Esso Petroleum Company, Research Division, Abingdon,
Berkshire, England, and Esso Research and Engineering Company, Florham
Park, New Jersey. Data for most of the design was provided by bench
scale batch tests and one megawatt pilot plant continuous runs of 200
to 500 hours duration carried out at Abingdon.
The Esso CAFB (chemically active fluidized bed) process vaporizes and
thermally cracks heavy fuel oil at 1,616 F by partial oxidation in a
fluidized bed. In this reducing atmosphere the sulfur compounds in the
fuel are converted to t^S (hydrogen sulfide). The fluidized bed material
is CaO (calcium oxide, quicklime) which reacts with the gasous t^S and
traps the sulfur as solid CaS (calcium sulfide). The reactor in which
this is carried out is called the Gasifier. The cracked gas formed,
with oil vapors and combustion products, is fed hot to special burners
installed in the boiler. About 90 percent of the sulfur in the original
fuel is removed from the gas. Design in this area is based on pilot
plant results.
Air for the partial combustion is fed to the Gasifier along with re-
cycled flue gas from the boiler stack. The recycled flue gas coming
in cool and leaving at reactor temperature removes excess heat of
combustion and regulation of its quantity controls the Gasifier temp-
erature. In the Gasifier some carbon is laid down on the stone and the
amount varies with the ratio of air to fuel. The rate of carbon lay
down determines the minimum feasible air/fuel ratio. The unit is de-
signed to operate normally at an air/fuel ratio of 20% of stoichio-
metric, with capability of 25%. Air/fuel ratios below 17 percent have
been demonstrated in the pilot plant.
Make-up of solids for the fluidized bed is in the form of crushed lime-
stone sized 1/8" to 300 microns, approximately 600 microns average size.
The CaCf>3 (calcium carbonate, limestone) becomes calcined to CaO in the
Gasifier giving off C02 (carbon dioxide). The term stone is used below
to refer to the solids in the system at any point whether limestone or
quicklime or sulfided or sulfated lime. In an ideal system, no stone
make-up would be required but pilot plant continuous runs showed that
the CaO loses its activity (ability to pick up sulfur) so stone purge,
plus loss make-up, at a rate of 20 to 28 tons per day of limestone is
required. 28 tons per day is coincidentally the stoichiometric rate
of calcium through-put with reference to sulfur to be removed from
the fuel.
65
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The stone in the Gasifier picks up 3 to 5 wt% sulfur as CaS and is
circulated at high rate to a second fluidized bed reactor called the
Regenerator. Here it is contacted with preheated air which rapidly
burns the carbon off the stone and more slowly reacts with CaS until
essentially all the oxygen is consumed. The reaction of CaS is in
two steps, first forming CaSOo (calcium sulfite) which is unstable
at reaction temperature and, secondly, decomposing to regenerated
CaO and SC>2 gas (sulfur dioxide). An alternate second step is
further oxidation of CaSC-3 to stable solid CaSO, (calcium sulfate),
which is encouraged by low sulfur to oxygen ratio. The percentage
of CaS going to CaO, referred to CaS reacted both ways is defined
as the Selectivity of regeneration, and selectivity is found to
vary directly with sulfur content of the stone.
For each mol of CaS going to CaO, 1/2 mol of oxygen is absorbed
from the gas into the stone while one mol of S02 gas joins the
associated nitrogen. For each mol of CaS going to CaSO, , two moIs
of oxygen are absorbed into the stone leaving only associated nitrogen
in the gas. Hence the 862 concentration in the gas depends on the
selectivity of regeneration, and at 100% selectivity (with no carbon
present) the concentration would be a maximum of 15% S0« in nitrogen.
At design selectivity of 80%, and with carbon to yield 6% C02 in the
gas, the SC>2 concentration is 8%.
If the input of air (oxygen) is reduced, the quantity of CaS reacted
will be reduced, relative to sulfur pick-up by the stone, so the
sulfur content of the stone will increase. That increase gives higher
selectivity, and a higher portion of smaller total reaction goes to
S02- The net result is the same S02 output with decreased CaSO^,
decreased nitrogen from the air, and hence higher SOg concentration.
Conversely, if air input is increased, total reaction increases, total
sulfur decreases, selectivity decreases, more oxygen goes to CaSO.,
and more nitrogen is left to dilute the SO2 to lower concentration.
The limit of the effect of reduced air, in causing increased SO2 and
decreased CaSO^, is reacted when the sulfur loading on the stone be-
comes high enough to interfere with pickup of sulfur in the Gasifier.
Then the required amount is not delivered to the Regenerator and more
than allowable sulfur goes to the boiler and stack. The operator ad-
justs the flow controlled air input for constant SOo percentage in the
Regenerator off gas at a target value deduced from analysis of gas and
stone samples.
The reactions in the Regenerator are exothermic and the temperature is
controlled by varying the stone circulation rate. The stone enters at
1616 F and leaves at 1970 F, and a massive but practical circulation
rate is capable of carrying away a large amount of sensible heat.
Sulfur and carbon levels on the stone tend to vary inversely with cir-
culation rate, so that change in sensible heat exceeds change in
66
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reaction heat and control is obtained. Control could become weak if
carbon on the stone were excessive, and in this case the air/fuel ratio
in the Gasifier must be increased to reduce carbon lay down.
Design parameters for the Gasifier-Regenerator System have been well
established by the one megawatt pilot plant operations. In the event
of failure of normal means of temperature control for either, steam
and water injection is provided for emergency use.
At this point in the processing the sulfur which was in the fuel oil
has been transferred to a hot gas stream leaving the Regenerator in the
form of SO2 at about 8 vol% concentration. A net purge stream of re-
generated stone, principally CaO, is also withdrawn from the Regenera-
tor. These two streams plus fines collected at various points are fed
to a third reactor called the Absorber, to recombine the sulfur and the
net stone. For this purpose, the stoichiometric rate of stone, or
higher, is' required and the waste stone produced by the process can be
essentially CaSO. (calcium sulfate) containing all the sulfur removed
by the process. The reversal of the decomposition reaction occurring
in the Regenerator is to be obtained in the Absorber by operating at a
lower temperature, 1600F, such that the CaO and the S02 can combine to
make CaS03 and with an excess of air to oxidize that to CaSO^. Near
completion of this reaction, at flowrates close to stoichiometric, is
to be aided by reducing the particle size of the net stone leaving the
Regenerator to 60 micron average in a proprietary air jet pulverizer,
before feeding to the absorber.
The temperature in the Absorber is controlled by admitting excess air
to carry off excess heat, when operating near full capacity. At re-
duced rates of throughput, the heat of reaction may be insufficient
to maintain temperature, and in this case additional heat can be sup-
plied by combustion of low sulfur No. 2 oil at the gas inlet to the
reactor.
The stone leaving the Absorber constitutes the waste solid from the
process which has been converted primarily to CaSC>4 with 6% residual
CaO. It is essential that CaO be small and that CaS be completely
absent, since CaO releases heat and CaS is hydrolyzed to H~S gas by
water and atmospheric Q^. In the order of 1% CaS is present in the
Regenerator stone, and it is expected that this residual CaS will be
completely oxidized in the Absorber so that an l^S nuisance will not
occur after dumping the waste stone.
The design of the absorber, the estimation of its effectiveness in
absorbing S0_, and the prediction of the waste stone composition to be
produced from it, are projections from a small amount of data. The
design parameters have not been proven in the Pilot Plant at this
writing.
67
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In an earlier draft report, prepared for Westinghouse in March, 1974,
the regenerative process was used, with the 502 £as ^rom t*ie Regen-
erator being converted to elemental sulfur in a process unit to be
designed by Allied Chemical Company. However, the sulfur conversion
unit would cost about 3 million dollars and a lower cost process was
desired.
Also, for the plant as designed in the earlier draft report, the waste
stone was taken from the regenerator to disposal. The stone at that
point Is primarily CaO with small percentages of Ca SO^ and CaS. The
CaS created an even greater hazard in disposal than CaO. When con-
tacted by rainwater containing C02 the CaS is hydrolized and releases
H2& gas which is six times as poisonous as hyrogen cyanide. The
characteristic rotten egg odor is not sensed at high concentrations,
but dilute gas gives odor nuisance down to five parts per billion.
Laboratory work indicated that the stone could be sintered or "dead
burned" at temperature over 2200F to eliminate CaS and render the CaO
unreactive toward water. However, for lack of design data, uncer-
tainty as to the long term permanence of the sintering inertness, and
fuel debit of the sintering furnace, this method was not adopted.
The method proposed to avoid the heat, poison and odor effects was to
slurry the stone in water in the plant so that released heat would
form steam and released H~S would be directed back into the process.
To render the stone completely unreactive the slurry was contacted
with flue gas so that the stone would be reconverted to CaCO- which is
inert and harmless. Although the stone could be filtered and dried,
the use of a slurry system in the plant was considered undesirable.
It is possible that the fluid bed Absorber now proposed will be
roughly the same cost as will the slurry system formerly proposed.
An additional change which has been made is in the method of handling
the stone flow between the units. In the March, 1974 design, stone
was transferred by Fluidizing technique used in petroleum plants in
fluid coking units called a "dense phase riser system". Such a sys-
tem imposes minimum elevations at the base of the reactors since
vertical seal loops must be provided below the reactor level. In the
Esso pilot plant at Abingdon transfer in essentially horizontal manner
was accomplished by using inclined throughs or conduits with gas pulse
to maintain the flow of solids. In the present design it is postulated
that scale up of this method by a factor of 50 is feasible. This pos-
tulate is to be tested by cold mock-up of the stone flow system before
the demonstration unit is erected.
68
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1.2 Processing Alternatives
Some alternatives to the processing scheme outlined above are avail-
able. One would be to operate the second reactor at conditions to
retain the sulfur on the stone as CaSO^ rather than to regenerate
CaO. The stone would feed directly through the system without re-
cycling to the Gasifier. However, the maximum allowable sulfur pick-
up by the stone is indicated by batch tests to be 16 weight percent
which represents about 33 mole percent conversion of the stone. There-
fore, the amount of stone throughout would have to be three times the
stoichiometric rate or 84 tons per day, in order to attain 90 percent
sulfur removal.
In the above mode, called Once-Through, the waste stone to be disposed
of would be two-thirds CaO and one-third CaSO^ on a mole basis (54
weight percent CaSO^). Its disposal as land fill would represent a
hazard due to the fact that the reaction between CaO and water liberates
a large amount of heat. Given the presence of sufficient water, steam
could be generated explosively in a'large mass of solids, or with defi-
cient water, temperature sufficient to ignite flamables have been dem-
onstrated, the theoretical maximum being about 1200F. Furthermore,
pilot plant work on a Once-Through process has not been carried to a
point to provide adequate design for such a system.
1.3 System Performance Assessments
Gasification
The pilot plant work provides a sound basis for anticipation of
90 percent or better removal of sulfur at level of 2.6 percent in the
fuel. The depth of the bed of fluidized lime in the demonstration unit
is more than double that in the pilot plant. This should provide
adequate safety factor to compensate for degree of grossness in air
and fuel distribution caused by practical mechanical designs and to
allow for modest recycle of SO2 from the Absorber.
The most probable source of difficulty is the formation of solid
deposits of coke and lime dust on surfaces exposed to the hot
fuel gas. It is believed that the rate of increase of depth of
deposit is constant with time for given operating conditions. The
effect of larger scale is beneficial in this respect since a depth
which could shut down the pilot plant by restriction of small
passages would have little effect in the large passages of the
50 MW unit. Nevertheless, rates of deposit as high as 1/2" in 200
hours have been reported, and it will be necessary to maintain
operating conditions giving the much lower but unquantified rates
which were also reported. No provision is made for decoking in
operation, and complete shutdown of the gasifier is required.
69
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On occasion an accumulation of solid particles too large to
fluidize has been noted in the pilot plant. These were eliminated
by rapid drainage of solids from the bottom of the bed. A bottom
drain has been provided for the 50 MW gasifier also, but its
effectiveness is not readily predictable.
Regeneration
The pilot plant work provides a sound basis for anticipation of
sulfur rejection at 8% SC>2 in the off-gas and loading of sulfur
on the stone sufficiently low to provide pickup in the gasifier.
The stone circulation rate required for temperature control is
calculated to be proportionately higher than in the pilot plant.
This should tend toward lower sulfur loading, or better than
90% pickup in the gasifier, or reserve capacity to carry Absorber
S02 leakage.
The rate of carbon lay-down on the stone could be a limiting factor,
since it reduces the effectiveness of temperature control by stone
circulation and the S02 concentration margin is provided by in-
creasing air/fuel ratio to the gasifier, but this decreases Btu
value of the fuel gas. On this account, the running of vacuum
residual oils should not be accorded undue optimism; there is no
continuous pilot plant backup, although batch unit tests were made.
The concentration of 8% S02 in the Regenerator off-gas is probably
not critical to the performance of the S02 Absorber. However, it
would be important if conversion to elemental sulfur were desired,
since 8% is a borderline concentration which might require supple-
mental methane consumption in the conversion unit. Every effort
should be made to meet the design concentration. The pilot plant
has shown values both higher and lower than 8% in various runs.
SO. Absorber
At this writing, no pilot plant data on the Absorber reaction is
at hand, and there is no basis for comment on the Westinghouse
projection of 947. reaction efficiency of CaO and S02.
The particle size required for near completeness of reaction in the
Absorber is not established. The smaller the better for this purpose,
obviously. However, the particle size of the waste stone may have
important effects on the stability of land fill, and in that respect
larger particles are preferable.
While some residual CaO is desirable in the waste solids, it is
essential that CaS be completely reacted in the Absorber. Adequate
test of this point must be made before the 50 MW unit is built.
70
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Gas - Solids Separations
The material balance streams indicating fines content of gas
streams seem reasonable estimates, for the most part predicated
on 95% cyclone efficiency. There is no way to make accurate
predictions without knowledge of the quantity and particle size
distribution of the solids entering the cyclones. It is to be
expected that these values will be quite different in the 50 MW
unit than in the pilot plant, and extensive efforts to get more
pilot plant data do not seem warranted. The pilot plant cyclones
were designed for 90% efficiency on assumed solids loading, and
results were adequate. With an appropriate cyclone design, the
solids effluent will vary with the attrition characteristics
of the stone being fed.
Solids Transport
The transport of solids for controlled circulation between
Gasifier and Regenerator is a vital function whose scaleup
from pilot plant size is regarded with less certainty than
other factors. If this should fail, it is difficult to
conceive of feasible modifications short of elevating the
major reactors. Full scale demonstration is essential.
However, the arrangement as specified does appear to have
good prospects of functioning.
The Sturtevant systems for fines transport have been proven
in all-metal systems at temperatures up to 750F, but there
is no experience with slug flow in refractory lined pipe.
In case of inoperability, substitution of Incolloy pipe is
a possible modification.
71
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2. UTILITIES, RAW MATERIALS AND PRODUCTS
2.1 Feeds and Products
The principal feed to the system is high sulfur residual fuel oil, a
No. 6 oil containing 2 1/2 to 3 percent by weight of sulfur. High
sulfur fuel oil is received by barge, held in the storage tank, and
fed to the unit as required by the boiler. The design feed rate is
28,000 Ib/hr or 514 million Btu/hr. The specifications of the oil
used for design calculations are as follows:
API Gravity 12.8 (8.166 Ib/gal) @ 60F
Viscosity 300 SSU @ 140F
Higher Heating Value 150,000 Btu/gal
Compositions: Carbon
Metals
Hydrogen
Sulfur
Nitrogen
Oxygen
content : Vanadium
Nickel
Sodium
Iron
85.9 wt%
11.0
2.6
0.4
0.1
442 ppm
52 ppm
30 ppm
14 ppm
The ultimate product sought by use of this system in conjunction with
a power boiler, is, of course, clean stack gas. The projected per-
formance of the system will yield a stack gas containing about 140
ppm (Vol) sulfur dioxide and 0.80 Ib/million Btu particulates.
The principal direct product of the system is low Btu fuel gas, with
90% of fuel sulfur removed. The calculated higher heat of combustion
of the gas is 199 Btu/scf, referred to 77F; however, since the gas is
delivered to the boiler at 1616F, it has an appreciable sensible heat
content of 38 Btu/scf, therefor effective heating value of 237 Btu/scf.
The gas ignites spontaneously in air and burns with a luminous flame.
The gas has a solids content estimated at 1.25 grains/scf and also
tends to deposit coke on contact surfaces, particularly at points of
high turbulence. At turndown rates from 50% to 25%, the gas becomes
progressively leaner in combustibles with increased nitrogen and CO
content, and 25% oil rate is the lowest design condition. The normal
and turndown compositions of the gas are tabulated below.
Component Normal wt% Turndown (25%) wt%
Hydrogen 0.44 0.12
Methane 2.07 1.41
Ethylene 3.31 0.82
Heavier HC 4.41 1.33
Hydrogen Sulfide 0.05 0.03
72
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Carbon Monoxide
Carbon Dioxide
Water Vapor
Nitrogen
Heat of Combustion 199 88
Sensible Heat (ai6000F 38 34
Heating Value (HHV,77°F) 237 Btu/scf 122 Btu/scf
Secondary feeds to the system are air and limestone. The design is
based on a limestone containing 98.1 wt% calcium carbonate and 0.1 wt%
sulfur, the balance inert. Normal feed rate is 2355 Ib/hr or 28 tons/
day, presized at 1/8" to 300 microns. If the commercial grade of 1/8"
to dust is fed, the weight rate will be somewhat larger in amount of
the quantity of ineffectively used fines content.
The secondary product of the system is the waste stone, primarily as
calcium sulfate. The calculated rate is 2920 Ib/hr or 35 tons/day.
To the degree that any excess fines in the limestone feed are collected
and sulfated in the Absorber, the stated feed need not be increased.
The waste stone will be 75 microns and smaller in size, averaging 60
tricrons.
An auxiliary feed to the system is low sulfur No. 2 fuel oil, presently
the principal fuel of the Manchester St. Station. This feed is used
only for startup heating and for Turndown Case operation of the
Absorber.
73
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2.2 Utilities and Fuel Summary
A. FUEL OIL - LBS/HR Intermittent Design Continous
1. F.O. (High Sulfur)
L-l Gasifier 28.000.00
Total 28,000.00
2. F.O. (Low Sulfur)
F-l Start-up Heater 13,000 13,000 0
F-3 Absorber Heater 110 120 0
Total 0
B. PROCESS WATER-GPM
Service Water
50PSIG. 65"F
G-18 Absorber 5 4.4
Cyclone (Exit)
R-9 Absorber Air Blower 8 O'1
L-l Gasifier 14 17 0
Total 4.5
C. COOLING WATER-GPM
1. Service Water
50PSIG. 65"F
G-27 Solid Waste/Cooler Conveyor 150.00 120.0
T-3 Purge & Transport 40 32.0
Gas (PTG) Cooler
T-7 PTG Blower Suction Cooler 10 6.0
R-8 GA/FG Blower L.O. Console 100.0
Total 258.0
D. STEAM - LBS/HR.
150PSIG. 366"F
1. Import 14,000 14,000 0
(From Battery Limits)
2. Produced
M-3 Steam Drum 4,000 4,000
Total Continuous 4,000
3. Used
T-5 0.1 Storage tk 1,480
Suction Heater
T-6A 0.1 Oil Transfer 1.635
Line Heater Total 3,115
Decoking Lines 150 200 0
From Gasifier
Cyclones
74
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Intermittent
Continuous
3. Used (Cont'd)
Process
L-l Gasifier
Heating H-L
Control Building
Heating H-2
Laboratory
15,000
none
none
15,000
none
none
0
0
885
E. BOILER FEED WATER-GPM
M-3 Steam Drum
F. ELECTRIC POWER-KW
1. 2300 V. 3 PH. 60 Hz
R-8 Gasifier Air/Flue Gas Blower
2. 600 V. 3 PH. 60 Hz
Gl-1 Filter Receiver Air lock
Gl-2 Pressure System Air blower
Gl-4 Vac. System Air blower
Gl-5 Baghouse Rotary airlock
G2-1 Waste Stone Rotary airlock
G2-4 Vacuum System blower
G2-5 Filter Receiver airlock
G2-6 Pressure System blower
G2-8A Silo Rotary airlock
G2-8B Silo Rotary airlock
G-4A Screw Conveyor
G-4B Recycle flue gas airlock
G-26 Limestont otar valve
G-27 Solids waste cooler/conveyor
P-4A Oil transfer pump
P-4B Spare
R-3 Purge & transport
Gas (PIG) blower
R-4 Flue Gas
Booster Fan
R-6 Regenerator
Air blower
R-9 Absorber
Air blower
R-10 Absorber S02
Blower
R-ll Pulverizer
Air Blower
H-l Control Room Bldg.
A. cond. unit
10
8
Connected HP
3000.0
0.75
75.0
60.0
0.75
0.75
30.0
1.0
20.0
0.5
0.5
3.0
1.0
0.5
1.0
5
5
75.0
100.0
100.0
200.0
75.0
150.0
150.0
150.0
Spare or
Intermittent-KW /HR Continuous
KW /HR
2141.0
0.6
62.00
50.00
0.6
0.4
5.0
0.02
2.18
1.80
0.02
0.6
25.00
1.2
16.6
0.4
0
0.25
1.2
0.4
1.2
5.0
0
119.0
48.0
134.0
41.0
241.0
101.0
5.0
75
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Spare or
Connected HP Intermittent-KMH Continuous-KWH
H-2 Laboratory Bldg. 5.0
A. Cond.unit
Lighting & instruments
(assumed average) 50.0
3. 120 V. 1 PH. 60 Hz
G-3A Limestone Weigh feeder 1.0 0.8
G-3B Limestone Weigh feeder 1.0 0.8
G-5 Gasifier Air Screen 1/6 0.1
G-6 Regenerator 1/6 0.1
Air Screen
G-17 Pulverizer-cyclone 0.1 0
G-19 Bag filter-rapper 0.1 0
G-24 Pulverizer air blower- 1/6 0.1
Air screen
G-25 Absorber air blower - 1/6 0.1
Afir screen
G-34 Air screen intake for G2-3 1/6
G-35 Air screen intake for G2-6 1/6 0.1
TOTAL 2943.17
G. PLANT AIR - SCFM
100 PSIG. Ambient
Sturtevant Valves 200*0
G-23A-H, G-30A/B, G-31A/B
H. INSTRUMENT AIR - SCFM
100 PSIG, 20°F.
G-4 Recycle Flue Gas Baghouse
G-19 Bag Filter
76
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3. FLOW DIAGRAMS
3.1 Material Balance Flow Diagrams
Drawing No. 6ID
Schematic Flow Diagram and
Material Balance Base Case
Drawing No. 6IE
Schematic Flow Diagram
Material Balance
Turndown Case and
Low Air/Fuel Ratio Case
77
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AVAILABLE
DIGITALLY
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AVAILABLE
DIGITALLY
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PAGE NOT
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DIGITALLY
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DIGITALLY
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4.FUNCTIONAL DESCRIPTION OF SYSTEM OPERATION
Fuel Oil Feed
The present normal fuel of the Manchester St. Station is No. 2 grade
low sulfur distillate oil, which has replaced natural gas, which had
replaced coal, the original design fuel. For the Oil Gasification -
Desulfurization CAFB Demonstration Unit on No. 12 Boiler, No. 6 Grade
Residual Fuel, of about 2.6 wt% sulfur content, will be received at
the existing barge unloading facilities and sent by a new branch line
to the new 30 day storage tank Q-l. Process Flow Diagram 12418.01-61C
shows the new oil feed equipment.
/
Tank Q-l is 104' diameter by 40' high and is provided with a flame ar-
restor on the vent, level indicator, overflow, and dike. In Q-l, a
steam heated U-tube bundle of 1350 sq. ft. surface, T-5 Suction Heater,
warms the oil to 150F for pumping from Q-l by positive displacement
Fuel Oil Pump P-4A or spare P-4B. P-4A & B are gear pumps of 72 gpm
capacity at 35 psig discharge pressure. A duplex strainer G-14 is pro-
vided in the pump suction line, and oil can be recirculated back to Q-l
from P-4 discharge via a back pressure controller. All piping for No. 6
oil is steam traced to maintain fluidity when flow is stopped.
Oil is pumped from P-4 through T-6 A&B Fuel Oil Heaters of 560 sq.. ft.
surface each. T-6 A&B warm the oil to a controlled temperature of 200F
with 150 psig steam, and thence to process. The pumping rate is higher
than the design process rate, and continuous recirculation from the
point of use to Q-l is maintained.
For startup of the reactors, for experimental or contingency operations,
and for Turndown case operation of the Absorber L-3, a supply of No. 2
low sulfur oil is required. This is drawn from the existing oil supply
of the No. 12 Boiler and may be fed to startup Heater F-l, to Absorber
Heater F-3, or to Gasifier L-l via the normal No. 6 oil controls de-
scribed below, and shown on Dwg. No. 12418.01-61A. The No. 6 oil feed
is Material Balance Stream No. 4 (MBS-4) of quantities shown on Dwgs.
Nos. 12418.01-61D & -61E.
No. 6 oil from the heaters T-6 A&B is admitted to the Gasifier L-l at a
flow controlled recorded rate with a master FRC set by the operator.
The FRC resets three slave FIC's feeding a three segment ring main on
the gasifier. The FIC's can have different offsets put in so that oil
rate to three zones in the gasifier may be graduated. Multiple injec-
tion nozzles on each segment of the ring main distribute the fuel oil
over the cross section of each zone of the fluidized bed.
The master FRC also puts out a signal of measured oil feed rate to a
ratio controller which regulates the quantity of air supplied to the
gasifier, as described under subheading 3.2.2. The system is designed
to be operable over an oil feed rate range of 25% to 115% of the normal
89
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rate, which is 28,000 Ib/hr. If desired, the master FRC could be made
responsive to power demand on No. 12 Boiler; however, for purposes of
the Demonstration Unit, it is deemed preferable that the master FRC
be set by the operator.
Air and Recycle Gas Feed
Air (MBS-1) is drawn from atomosphere by Gasifier Air/Flue Gas Blower
R-8 through Gasifier Air Screen G-5. G-5 has a glass fiber screen,
automatically renewable by electric moter drive. R-8 is a centrifugal
compressor of 46,700 CFM capacity and 23.3 psia discharge.
The air is sent to Gasifier L-l along with recycle flue gas (MBS-3)
which is also fed to R-8 suction. The combined gas (MBS-2) enters L-l
below a refractory arch distribution grid which is provided with multi-
ple holes to distribute the gas evenly throughout the cross-section of
the fluidized bed.
The air rate (MBS-1) is normally regulated to obtain 207. of the stoichio-
metric amount of oxygen for complete fuel oil combustion. This is
referred to as air/fuel ratio of 20%, although correction is made by
the operator for oxygen content of the recycle flue gas so the air
rate is not exactly 20%. Provision is also made for operation at
lower air/fuel ratio equivalent to 17% of stoichiometric oxygen. The
flow rate and oxygen content of the combined gas are continuously re-
corded to provide a basis for operator adjustment of the automatic
controls described below. Under the Turndown conditionsbetween 25%
and 50% of normal fuel rate, the air/fuel ratio will be higher than
normal, resulting in decreased heating valve of the fuel gas as shown
under 3.1.3.
The air rate is controlled by an FRC throttling the air intake to R-8.
The index of the air FRC is pneumatically reset normally by a ratio con-
troller which receives an oil flow rate signal from the master FRC on
the fuel oil feed. The ratio to be maintained is constant and normally
would not be changed over long periods of operation, the primary opera-
tor control being the oil FRC. The ratio setting would be changed
slightly if a change in oil feed stock altered carbon lay down. The
ratio setting would be altered significantly for operation, for ex-
ample at 17% air/fuel. Ratio control of air FRC set point is main-
tained over the fuel oil feed range from maximum to approximately 50%
of normal, which will result in total gas flow from R-8 to L-l varying
over about the same range.
The air FRC index can also be reset by the FRC on total gas flow, but
the signal is normally blocked. The purpose of the latter is to main-
tain a minimum flow of gas to L-l for adequate fluidization of the
solids bed, which is approximately 50% of normal flow, and the total
gas FRC is set permanently at that value. When air flow demand by the
oil ratio control falls below the demand by the total gas FRC, a se-
90
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lector relay is actuated which blocks the ratio control signal and
allows the total gas FRC signal to take over the reset function. The
total gas FRC remains in control of the air FRC reset so long as its
demand is higher than that of the oil ratio controller.
The effect of varying air flow on total gas flow is roughly double the
change in air itself. Since all of this air goes for fuel combustion
in the Gasifier, a change in rate produces a change in heat release.
The recycle flue gas also fed to R-8 is throttled by an FRC reset by
the Gasifier temperature controller. Therefor a change in air rate
results in a similar change in recycle flue gas rate to compensate
for the change in heat release and maintain constant Gasifier tem-
perature.
Recycle flue gas (MBS-39) is drawn from the outlet of the existing
boiler stack cyclones, upstream of the existing induced draft fans,
by new Flue Gas Booster Fan R-4. R-4 has a capacity of 22,800 CFM
at suction conditions of -14" wg and 300F, and discharges at +6" wg
(MBS-40) to Recycle Flue Gas Bag House G-4. Off-gass from the Absorber
Cyclone G-18 (MBS-38) joins the recycle flue gas between R-4 and G-4,
constituting about 20% of the gas flow to G-4 and raising its tempera-
ture to 350F (MBS-3).
Recycle Flue Gas Bag House G-4 is a reverse gas Jet type dust collector
with approximately 6000 sq. ft. of Nomex fabric bags, for a capacity of
30,000 cfm gas flow at 5:1 gas to cloth ratio. Its function is to re-
move fine particulates for protection of downstream compressor and the
Gasifier distributor grid. The solids collected (MBS-42) are auto-
matically transferred to Absorber L-3 and become part of the waste stone
(MBS-30), along with parallel transfer of solids from the stack cyclones
(MBS-41).
Virtually all of the gas from G-4 passes to R-8 at the rate allowed by
the recycle flue gas FRC (MBS-3). About 27. of the gas is used for
pur^e and transport gas (PTG) as described under 3.2.8, and this quan-
tity is not included in the Material Balance Streams. Delivery of gas
by R-4 is thus restricted to the foregoing less the quantity displaced
by the independently controlled Absorber off-gas.
In starting up the unit from a cold condition, it is necessary to warm
the Gasifier L-l to HOOF before any oil is fed; including low sulfur
oil. The bed must be calcined and fluidized before high sulfur oil is
fed. If the starting bed is fresh limestone, it must be heated well
above 1250F in order to calcine it. For heating purposes, part or all
of the gas to L-l may be passed through Startup Heater F-l and heated
to 1400F by direct combustion of low sulfur No. 2 oil, at a maximum
rate of 22 million Btu/hr. Gas lines downstream of F-l are refractory
lined. Heated air and flue gas from F-l is also piped to Regenerator
L-2 for warmup.
91
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Warmup must be done slowly to avoid damage to refractory. Temperature
indicating controllers are provided in the plenum chambers of L-l and
L-2. Regulating valves throttle the heated gas to be combined with cool
gas which bypasses F-l. The temperature setting may then be increased
manually at a rate to be specified by the manufacturer of the refrac-
tory, probably 100 degrees per hour. Temperature of the gas at F-l
exit should be held constant by the heater controls, or may, if
necessary, be manually controlled.
Limestone Feed
Limestone of the proper grade and size, which is 1/8 in. to 300 mu
will be received by rail cars of approximately 200 tons capacity. The
rail car unloading system (G-l) is shown on drawing No. 12418.01-61C,
and is designed to unload a 200 ton railroad car in less than one
shift, or about six hours, which entails a flow rate of approximately
67,000 Ib per hour. Normal flow rate to the process (MBS-5) is 2,355
Ib per hour and one rail car of limestone should last about seven days.
The railroad car unloading system will be a packaged solids handling
system of the push-pull type. The limestone will be drawn from the
railroad hopper cars through unloading Kit Gl-6 by a vacuum system into
Filter/Receiver Gl-7 with suction provided by Vacuum System Air Blower
Gl-4. The solids separate in the vacuum receiver and drop through
Rotary Airlock Gl-1 to a pressure system. Pressure System Blower Gl-2
drives the limestone up about 200 ft to Baghouse Gl-3. The solids then
drop through Rotary Airlock Gl-5 into the existing coal bunkers above
the boiler, which will be converted to limestone use. The bunker is
designated Q-4 and this portion of the flow is shown on drawing
No. 12418.01-61A.
The limestone to be fed to the unit flows by gravity from the converted
coal bunker to an automatic weigh feeder. Limestone Weigh Feeders
G-3A and B of which one is to be operating and one spare. The design
weight rate of limestone feed (MBS-5) is set by the operator and the
weigh feeder drops the limestone into a feed pipe to flow by gravity
to Gasifier L-l. Since L-l is operating under positive pressure and
contains hot flammable gas, a seal is necessary between the gasifier
and the weigh feeder. This is provided by Limestone Star Valve G-26.
Gas purge is used to G-26 and to a point just down stream of G-26 to
prevent the contact of hot gas with G-26 itself. The capacity of G-26
is well above any expected feed rate and a level indicator and switch
keeps it in on-off operation according to the presence or absence of
solids in a small hopper which is part of the weigh feeder G-3.
Limestone feed enters the gasifier at one of the inlets returning
solids to the Gasifier from one of the four cyclones G-8A, B, C, D
described below.
92
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Gasifier Operation
The principal reactor in the system is Gasifier L-l which contains the
fluidized bed of CaO solids (stone) wherein partial combustion of the
fuel oil occurs and sulfur is trapped. The Gasifier is 19 ft-4 in.
ID of the refractory lining. The depth of the fluidized bed is 4 ft
and the clear space from the top of the bed to the upper bend line
is 14 ft. The fluidized bed rests on a distributor composed of re-
fractory brick, each with a hole for gas passage, and the distributor
is integral with a refractory arch. The depth of the gasifier below
the distributor is that required to accommodate the supporting arch,
and the gas inlet to the plenum chamber formed by the space beneath
the arch. Air and flue gas from blower R-8 (MBS-2) enter the plenum
chamber and pass through the distributor holes to fluidize the solids
bed.
The oxygen content of the air/flue gas mixture entering the Gasifier
is continuously monitored by a recording analyzer, and the flow rate is
also recorded. VIth this information, the operator can set the ratio
controller which normally regulates the ratio of air to oil being fed.
On turndown operation below 50% of normal, the air/flue gas flow is
automatically held at the mechanical minimum required to fluidize the
bed of solids. The high sulfur No. 6 fuel oil feed (MBS-4) is injected
into L-l through 36 injection nozzles grouped 12 on each of three seg-
ments of the ring main around the Gasifier. Controls for the gas and
oil inlet rates are described in Sections 4.2.1 and 4.2.2. Connections
are provided for the injection of steam and water into the oil fuel
lines to the Gasifier dowstream of the fuel oil controls. These are
for use in the event that the Gasifier temperature cannot be controlled
and tends to rise excessively.
The temperatures in the Gasifier bed are measured at 12 points at
various elevations and circumferential locations in the bed and these
12 points are carried on a single recorder. There is also temperature
measurement of the gas in the plenum chamber beneath the refractory
arch, and of the gas in the upper part of L-l immediately below the
gas outlet. This last is used for purposes of control, by connecting
to the TRC which resets the FRC on recycle flue gas to R-8.
The pressure in Gasifier L-l will be as is developed for the fuel gas
to pass through the exit cyclones, through the transfer lines to the
burners and through the burners into the boiler firebox. At normal
firebox draft, pressure in the Gasifier is estimated to be 2.6 psi
gage, and L-l is designed for a maximum pressure of 10 psi gage.
There are four fuel gas outlets (SUM=MBS-6) in parallel of size 42 in
OD steel by 30 in ID refractory, each having a refractory lined cyclone,
Gasifier Cyclones G-8A, B, C, and D. These remove 95% of the solids
(MBS-8) which are entrained from the gasifier by the fuel gas. The
fuel gas tends to deposit coke on all solid surfaces and particularly
93
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at points of high turbulence. The location of the most severe coking
is at the cyclone inlets, and to monitor this, there is a pressure
differential indicator connected between the top of the gasifier and
the outlet of each cyclone. There is a separate transfer line for
fuel gas from the outlet of each cyclone to one of the four burners
which are to be installed in the existing boiler. Although overall
pressure difference between boiler and L-l for the four paths must be
equal, significant changes can occur in cyclone pressure drop, abso-
lute, relative to the overall, and relative to each other. An eight
point temperature recorder is provided with one thermocouple at the
outlet of each cyclone and one at the inlet of each burner. These
are for use in following the progress of the decoking of the cyclones
and lines, for which air and steam connections are provided.
Each of the four cyclones has a Sturtevant system for returning fines
to the gasifier (SUtt=MBS-8). The Sturtevant systems are described in
Section 4.2.7. There is a chunk trap on the outlet of each cyclone
upstream of the Sturtevant conveyor system consisting of a coarse
strainer in the line and a plug valve through which chunks can be re-
moved. The fines from each cyclone are returned through separate in-
lets on the gasifier, two being located on each side of the baffle
described below. Thereby, half of the fines (MBS-9) enter the bed
with the fresh and regenerated stone and half of the fines (MBS-10)
are deposited at the outlet to flow immediately to the regenerator
with the main flow of sulfided stone (MBS-12).
Connections are necessary between Gasifier L-l and Regenerator L-2
for the circulation of the stone in both directions (MBS-11, MBS-10+12).
The stone passages must be shallow, wide and short, and L-2 is much
smaller in diameter than L-l, which results in passages having compound
curves and opening into L-l very close together. The passages for the
stones will be cast into a monolithic block which extends through the
refractory walls of L-l and L-2, and this monobloc will be encased in
a gas-tight steel shell. Since the stone inlet and outlet of the gasi-
fier are adjacent, a baffle is provided between them of twice the
height of the bed and extending to the center of the gasifier, so that
the stone flow will not short-circuit. The baffle is formed of high
cr-ni alloy which will have adequate strength at the design temperature
of 1,710 F, to be supported primarily from the wall and not by the re-
fractory arch. Stone flow rate through each passage is regulated by
providing a short horizontal run near the outlet which requires pulsed
injection of gas to aid flow, and by making the pulse rate automaticly
variable. The prime function of the circulation is to send sulfided
stone to the Regenerator for removal of its sulfur content, and to
return for more sulfur pickup. However, the flow is actually controlled
for a secondary purpose described in the next section which demands a
rate more than sufficient for sulfur transport.
The performance of the gasifier in terms of sulfur removal frcm the
fuel is determined by a continuous recording analyzer on the stack.
At 90 percent removal of sulfur from the fuel containing 2.6 weight
94
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percent, the concentration of SC^ in the stack gas should be about
140 ppm. If the removal of sulfur is not satisfactory then oil feed
must be decreased or limestone feed must be increased in order that
the stone in the bed be purged more rapidly and its activity maintained
at a higher level. If the circulation rate to and from the regenerator
becomes very high with associated difficulty in controlling the re-
generator temperature this probably indicates excessive laydown of
carbon on the stone and the remedy for this is an increased ratio of
air to fuel.
The density of fluidized bed is measured by finding differential pres-
sure across two points a known vertical distance apart in the bed.
The level of the bed is measured by taking the differential pressure
from the bottom of the bed to a point well above the bed and comparing
this to the bed density. The level of the bed is controlled by regula-
ting the rate of withdrawal of stone from the system (MBS-30) via the
regenerator (MBS-30) as described in Section 4.2.5.
A condition has been noted in the pilot plant gasifier wherein activity
of the stone decreases seriously due to formation of large and heavy
particles which tend to stratify in the bottom of the bed. To main-
tain satisfactory operation, it is necessary that these be removed
rapidly. A connection with plug valve is provided at the bottom of the
bed, piped to the inlet of Pulverizer G-20 (see 4.2.5) through which a
fraction of the bed can be drawn fairly rapidly. Normal flow from
Regenerator L-2 to G-20 should be interrupted during this procedure.
Regenerator Operation
Regenerator L-2 receives a massive circulation of stone from Gasifier
L-l, burns the carbon off the stone, converts the CaS partially to CaO
and S02 and partially to CaSO,. The stone is recirculated to the
Gasifier. The circulation back and forth passes through a monobloc
connection between the two reactors as described below. Regenerator
L-2 is 4 ft-6 in. ID of refractory lining and is mechanically an ap-
pendage of the Gasifier which is about four times larger.
The Regenerator is fitted with a refractory distributor grid for the
fluidized bed, containing multiple holes for passage of gas to the
bed. L-2 is supplied with air by Regeneration Air Blower R-6 equipped
with an air screen G-6. R-6 is a positive displacement axial flow
machine of 2500 cfm capacity from 14.6 to 24.0 psia pressure. The air
(MBS-13) is throttled by a flow controller on the blower suction, the
rate being set by the operator. Air from the discharge of R-6
(MBS-14) is preheated in Regenerator Off-gas/Air Exchanger T-l to a
temperature of 452 F. From T-l, the air (MBS-15) passes to a chamber
beneath the refractory grid of L-2. The reactions of L-2 are exothermic
and the temperature is controlled at 1,970 F. The temperature is nor-
mally controlled by adjusting the circulation rate of stone between L-l
and L-2, whereby the excess heat of reaction is carried away as sensible
heat of the stone circulation.
95
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The off-gas from the Regenerator L-2 (MBS-16) contains SO at about 8%
concentration and CO. at about 6%, by volume. The off-gas passes
through two cyclones in series. The Regenerator Primary Cyclone, G-9,
returns fines (MBS-18) to the Regenerator by gravity. The Regenerator
Secondary Cyclone, G-10, sends the fines (MBS-20) to join the net stone
purge stream (MBS-19) flowing from L-2 to Pulverizer G-20.
The gas from Secondary Cyclone G-10, (MBS-21) still at 1,970 F, passes
through Waste Heat Boiler B-l, where the temperature is reduced to
600 F (MBS-22), and the heat is used to generate steam at 150 psig.
Boiler feed water is used as make-up for the Steam Drum M-3, and
approximately 4,000 Ib/hr of steam are produced. About 3,000 Ib/hr
are normally used for preheating the fuel oil feed in T-5 and T-6A&B.
A small excess of steam is sent to the boiler service steam system,
normally, or steam may be drawn from that system when usage is higher
than normal.
The off-gas from B-l (MBS-22) passes through Regenerator Off-Gas/Air
Exchanger, T-l, where it preheats incoming air and is reduced itself
to 300 F (MBS-23). The off-gas then passes to Bag Filter G-19, for
removal of fines (MBS-24) before forwarding to Absorber L-3. The
off-gas leaving G-19 (MBS-25) is throttled to maintain the back pres-
sure on the Regenerator, L-2. The throttling is controlled by a PDRC
which maintains a differential of 0.3 psi between L-l and L-2.
In Regenerator L-2, the amount of sulfur which is released to the off-
gas as SO, must come to balance with the amount of sulfur removed from
the fuel gas in the Gasifier, L-l. The amount of sulfur as CaS which
is converted to CaSO,, rather than being rejected to off-gas, has a
very significant effect on the quantity of heat released in L-2 as does
the amount of carbon which is oxidized in L-2. At the sulfur levels
and carbon levels on the stone experienced in the pilot plant, there
should be no difficulty in controlling the temperature of L-2 by ad-
justment of the stone circulation rate.
The circulating stone passes through inclined passages in the monobloc
connection between L-l and L-2 mentioned above. The lower end of each
passage is horizontal for a short distance which tends to impede the
flow. A gas distributor pipe lies across the horizontal section and timed
pulses of gas are injected to speed the flow of solids. By altering
the time rate of pulsing, the flow rate of the stone can be controlled.
The flow of sulfided stone from Gasifier to Regenerator (MBS-10+12) is
controlled by a TRC in the gas space of the Regenerator which alters
the pulsing rate. The flow from L-2 back to L-l (MBS-11) is similarly /
regulated by the level controller of the Regenerator.
In the event that modulation of the rate of stone circulation is unable
to maintain the temperature at 1,970 F, and it begins to rise, a por-
96
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tion of the air/flue gas mixture from the discharge of blower R-8 is
admitted to the plenum chamber below the refractory grid on automatic
control by the same IRC which regulates the stone rate. It is con-
sidered that the nature of the reactions occurring in Regenerator L-2
is such that a tendency to overtemperature would be naturally self-
limiting. In any case, provision is made for manual injection of
steam and water to L-2 as auxiliary temperature regulation, if required.
L-2 is provided with a density recorder which measures the differential
pressure across a fixed depth of bed. There is also a level controller
and alarm which functions by measuring the differential pressure be-
tween the bottom of the bed and a point well above the bed. The level
controller acts to regulate the rate at which the stone circulation is
returned from L-2 to L-l (MBS-11).
The quantity of stone in the fluidized bed in the Gasifier is roughly
35 tons while that in the regenerator is about two tons. The stone
circulation in each direction between the two is over half a ton per
minute, which would seem to place a severe requirement on level regu-
lation. However, the circulation flow rates are strongly influenced
by difference in level between Gasifier and Regenerator and the levels
tend toward a stable balance.
A net purge stream of stone (MBS-19) is taken from Regenerator L-2
for further processing as waste stone. The limestone make-up to the
system (MBS-5) is fixed by the operator's setting of Weigh Feeder
G-3, so the purge must be withdrawn at a rate to maintain constant
system inventory. The withdrawal of stone from the system via L-2 is
regulated by the level controller on Gasifier L-l. Since this is the
larger body of stone in the system and the bed level is not subject
to rapid fluctuation, a steady rate of net stone flow can be obtained.
Stone from L-2 (MBS-19) flows to Pulverizer G-20 through an inclined
pipe, followed by a short horizontal run where the solid flow rate is
controlled by the rate of gas pulsing to the horizontal section. It
is necessary that this horizontal run be very short, and also necessary
that it receive connections for fines from Regenerator cyclone G-10
(MBS-20) and for heavy particle drainage from Gasifier L-l. Due to
the requirements of the physical dimensions, this multiple connection
is fabricated of a high Cr-Ni alloy capable of withstanding the re-
generator temperature of 1,970 F, rather than of refractory lined pipe.
Further treatment of the waste stone is discussed in under 3.2.6.
The operation of L-2 is monitored by continuous recorded analysis of
the off-gas (MBS-25) downstream of G19, for S02, CO* and 02< Since
the air entering L-2 may contain recycled flue gas from R-8, analysis
is provided at that point also. Comparison of the two analyses to-
gether with the air rate yields the results of the reaction in L-2.
Samples of the stone are analyzed periodically also.
97
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For start-up purposes, hot air and flue gas mixture from Start-up
Heater F-l are introduced with normal air into the Regenerator plenum
chamber below the refractory grid. A TIC is provided in the chamber
to throttle this hot gas stream and assist the operator in the gradual
warmup required to avoid damage to the refractory.
SP2
Ahsorher Ooeration
The function of Absorber L-3 is to recombine the sulfur removed from
the fuel and converted to SO^ in the Regenerator off-gas with the net
stone purge also leaving the Regenerator, in order that the waste solids
disposed of may consist primarily of non-reactive calcium sulphate
(CaS04).
The stone as it is drawn from the Regenerator (MBS-19) is joined by
fines from Regenerator Secondary Cyclone (MBS-20) and the mixture is
of 1/8 in. to dust size range. In order to enhance completeness of
the reaction in the Absorber, the stone is first pulverized to 88 microns
and smaller with an average size of 60 microns. Since the stone is at a
temperature of 1,970 F, a pneumatic pulverizer is used rather than mec-
hanical grinding. The stone flows to pulverizer G-20 which is supplied
with air (MBS-32) by Pulverizer Blower R-ll, having an air screen G-24
on the suction. R-ll is a positive displacement machine of 600 cfm
capacity with 14.6 psia suction and 75 psia discharge. A flow indica-
ting controller is provided with orifice in R-ll discharge line and
valve throttling a recycle line from discharge to suction of R-ll,
which compensates for any volume flow mismatch between G-20 and R-ll,
or for any blockage of air flow through G-20.
In the pulverizer G-20, the stone is reduced in size by the kinetic
energy of jetted air (MBS-33), and contact between the stone and the air
in G-20 gives a resultant temperature of 794 F. The total air and stone
(MBS-34) is blown upward through a stainless steel line to Pulverizer
Cyclone G-17, which is located at an elevation that will allow separated
solids (MBS-35) to flow by gravity into Absorber L-3. In addition to
the pulverized stone separated in G-17, Absorber L-3 also receives fines
at the same inlet, from the existing stack cyclones (MBS-41), from the
Recycle Flue Gas Bag House (MBS-42) and from Bag Filter G-19 (MBS-24).
The air separated in G-17 (MBS-36) passes forward to join the gas efflu-
ent from the Absorber (MBS-31) at the inlet of Absorber Cyclone G-18.
Absorber L-3 is a refractory lined carbon steel vessel 20 ft diameter
inside of the refractory. It is provided with a refractory brick dis-
tributor grid, integral with an arch, similar to that in gasifier L-l.
The bed of fluidized solids is maintained at 3' depth by an overflow
weir, and the upper bend line is 10* above the nominal bed level. The
Regenerator off-gas to be fed is picked up (MBS-25) from Bag Filter
G-19 by Absorber S02 Blower R-10 and sent to the plenum chamber in L-3
below the refractory grid. R-10 is a centrifugal blower of 3500 cfm
capacity, 14.0 psia suction and 22.5 psia discharge. The flow is
throttled for back pressure at the outlet of G-19 as described in the
98
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preceding section, and the rate is as produced by the Regenerator.
For operation in the turndown mode, and also for start-up purposes,
Absorber Heater F-3 is provided, mounted directly on Absorber L-3 inlet
nozzle, through which the Regenerator off-gas from R-10 as well as re-
action air for L-3 are passed. Under conditions where the reaction
heat in L-3 is not sufficient to maintain the normal temperature of
1,600 F, F-3 is fired with low sulfur fuel oil under temperature control
of a TIC with thermocouple in the plenum chamber. Under 25 percent
turndown condition, the usage of low sulfur oil is approximately 110 Ib
per hour.
In Absorber L-3, two principal reactions are carried out. The regener-
ated stone is primarily CaO, entering addition reaction with SO to
give CaS03, and air sent to L-3 oxidizes the CaS03 to CaSO^. Also any
residual sulfide on the regenerated stone is expected to be completely
oxidized. Air for L-3 is provided by Absorber Air Blower R-9, a posi-
tive displacement axial flow blower, with Air Screen G-25, of 870 cfm
capacity at 14.0 psia suction, 22.5 psia discharge. The air flow is
throttled on the suction of R-9 by an FIG which is reset by a TIC mea-
suring the temperature in L-3. A small amount of water is also neces-
sary for the reaction and service water is injected between R-9 and F-3
on flow control.
The fluidized solids in L-3 form a bed of 3 ft depth. A density re-
corder is provided measuring the differential pressure across a known
depth of bed. The total depth of the bed is recorded by measuring the
differential pressure from a point just above the refractory grid to a
point well above the bed. The reacted stone is removed by gravity over-
flow from L-3 and no level control as such is used. The reacted stone
overflowing from L-3 drops into Solid Waste Cooler Conveyor G-27, and
is joined by fines separated in Absorber Cyclone G-18. These gravity
flow lines are purged with a small flow of air which is tapped from
the discharge of Pulverizer Blower R-ll. G-27 is helical ribbon cooler
manufactured by Joy Manufacturing Company and service water flows through
passages in the helix. The solids are cooled to 200 F in G-27 and fall
from the outlet to be picked up by the waste stone system described in
the next section.
Off-gas from absorber L-3 consists primarily of nitrogen, C0» and some
excess oxygen, with a small amount of SO.. The gas leaves L-3 at
1,600 F and is mixed with the pulverizer air from cyclone G-17 giving
a resultant temperature of 1,435 F. This gas passes through Absorber
Cyclone G-18 where entrained fine solids are separated. The pipes en-
tering and leaving G-18 are refractory lined due to the temperature of
the gas. A short distance from G18, service water is injected into the
gas to reduce its temperature to 500 F and from that point a carbon
steel line is used. This gas joins recycle flue gas entering Recycle
99
-------
Flue Gas Bag House G-4, as described in Section 3.2.2, so any unreacted
S02 leaving Absorber L-3 will return to Gasifier L-l where it is re-
duced to H2S and picked up by the stone again.
Waste Stone and Fines Handling
Waste stone leaving the Waste Stone Cooler Conveyer G-27 will be picked
up by the Waste Stone Transport and Loading System item G-2. This will
be a package solids handling unit of the push/pull type. Waste stone
from G-27 passes through a rotary air lock, G2-1 into a vacuum line,
carrying air drawn through a screen G-34, by air blower G2-4. The air
and solids pass into filter/receiver G2-3, where the air separates and
passes through blower G2-4. Solids from G2-3 pass through a rotary
air line G2-5 into a pressure blowing system. For this system, the
pressure system blower G2-6 picks up air through a screen G-35 and
blows it to filter/receivers G2-7A and B. A diverter valve is used to
pass the solids and air to one or the other of these receivers. The
receivers G2-7A or B are located above the two existing ash silos now
designated Q-2 and Q-3. The solids flow by gravity from the receivers
through rotary air locks G2-8A and B, into the silos. The existing ash
silos are already set up for gravity loading of trucks to remove the
solids from the plant. The set up can load a 10 ton truck in about 20
minutes, and five truckloads per day must be removed on a five day
week basis.
At some points in the process it is necessary to transfer fines from
the bottom of the cyclone to a point higher than the bottom of the
cyclone. For this purpose a fine solids transport method developed
in Britain and provided commercially by Sturtevent Engineering Co. Ltd.
is suggested by Esso Research. The system consists of the hopper re-
ceiving the fines by gravity with shut off valves above and below the
hopper. When the hopper is full the upper valve is closed and the hop-
per is pressurized to about 20 psig with transport gas. The lower valve
is open and the solids flow out. In order to maintain flow through the
exit line, the solids coming out are divided into discreet slugs by an
air knife which is a device for injecting the air on a pulse basis into
the line. Normal operation is controlled from an automatic timer panel.
These devices are used under the Gasifier Cyclones G-8A, B, C, and D,
which remove fines from the fuel gas make and transport these fines
back into the Gasifier as described in Section 4.2.4. Similar systems
but much smaller in size are used to take fines from the existing stack
cyclones and from the Recycle Flue Gas Bag House G-4, to the inlet of
the Absorber L-3.
100
-------
Puree and Transport Gas
At several points through the system a gas feed under pressure is re-
quired for the transport of solids. Also, at many points through the
system, instrument connections and other small connections must be
purged to prevent their being plugged with solids. In several of these
places it is not possible to use air since the air might react with hot
solids, and a small portion of the recycled flue gas is used for this
purpose.
About two percent of the flue gas leaving Recycle Flue Gas Bag House
G-4 is taken for purge and transport gas, referred to as PTG. The
gas at about 350 F passes first through PTG Blower Suction Cooler T-7,
where it is cooled to about 200 F. The water dew point of the gas is
about 150 F, and condensate would be corrosive due to the presence of
C02 and SC^. Also, the gas passes next to a blower and therefore the
outlet from T-7 is held well above the water dew point. This gas is
picked up by Purge and Transport Gas Blower R-3. R-3 is a positive
displacement axial flow blower with a discharge pressure of 35 psig.
The compressed gas passes through PTG Cooler, T-3 for cooling to about
100 F. This is well below the dew point of the gas and some water will
condense. The gas and water pass through PTG Water Separater M-4 where
the water is knocked out and sent on level control into the Gasifier.
The water is of very small quantity, and is introduced into the oil
feed to the gasifier. The gas from M-4 passes through PTG filter G-7
for the removal of mist and any fine particulates remaining.
From this point, the PTG is distributed as required to the control
panels of the six Sturtevant solids transport systems described in
Section 3.2.7. This gas also supplies the pulsing systems used for
the circulation of stone from Gasifier to Regenerator and from Regenera-
tor to Gasifier and for the net stone coming from the Regenerator to the
Pulverizer. The gas will also be piped in manifold to all points re-
quiring gas purge of instrument connections.
101
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5. EQUIPMENT ARRANGEMENT DRAWINGS
103
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PAGE NOT
AVAILABLE
DIGITALLY
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AVAILABLE
DIGITALLY
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AVAILABLE
DIGITALLY
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AVAILABLE
DIGITALLY
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AVAILABLE
DIGITALLY
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6. EQUIPMENT DRAWINGS AND SPECIFICATIONS
6.1 EQUIPMENT LIST
Item Number
Service
Drive
Acct. B - Boilers
B-l
Waste Heat Boiler
Acct. F - Heaters
F-l
F-2, A-D
F-3
Start-up Air Heater
Hot Fuel Gas Burners
Absorber Heater
Acct. G - General Equipment
G-l
Gl-1
Gl-2
Gl-3
Gl-4
Gl-5
Gl-6
Gl-7
G-2
G2-1
G2-3
G2-A
G2-5
G2-6
G2-7 A&B
G2-8 A&B
G-3 A&B
G-4
G4A
G4B
G-5
G-6
G-7
G-8 A-D
G-9
G-10
G-14
Fresh Limestone Unloading System
Filter Receiver Airlock Motor
Pressure System Air Blower Motor
Baghouse
Vac. System Air Blower Motor
Baghouse Rotary Airlock Motor
Railcar Unloader
Filter/Receiver
Waste Stone Transport and Loading
System
Waste Stone Rotary Airlock Motor
Filter/Receiver
Vacuum System Blower Motor
Filter Receiver Rotary Airlock Motor
Pressure System Blower Motor
Silo Filter Receivers
Silo Rotary Airlocks Motors
Limestone Weigh Feeders Motors
Recycle Flue Gas Baghouse
Screw Conveyor Motor
Recycle Flue Gas Airlock Motor
Gasifier Air Screen Motor
Regenerator Air Screen Motor
Purge & Transport Gas Filter
Gasifier Cyclones
Regenerator Primary Cyclone
Regenerator Secondary Cyclone
Duplex Oil Filter
115
-------
Item Number Service Drive
Acct. G - General Equipment (Cond't.)
G-17 Pulverizer Cyclone
G-18 Absorber Cyclone
G-19 Bag Filter Motor (Rl)
G-20 Jet Pulverizer
G-21 A-D Fines Lock Hoppers (Rl)
G-22 A-D Air Knife & Transport System
G-23 A-D Fines Lock Hopper Valves-Inlet
G-23 E-H Fines Lock Hopper Valves-Outlet
G-24 Pulverizer Air Blower-Air Screen Motor
G-25 Absorber Air Blower-Air Screen Motor
G-26 Limestone Star Valve Motor
G-27 Solid Waste Cooler/Conveyor Motor
(Service Water)
G-28 Stack Cyclone Lock Hopper
G-29 Baghouse Lock Hopper
G-30 A&B Stack Cyclone Lock Hopper
Valves - Inlet and Outlet
G-31 A&B Baghouse Lock Hopper
Valves - Inlet and Outlet
G-32 Stack Cyclone Lock Hopper Air Knife
& Transport System
G-33 Baghouse Lock Hopper Air Knife
& Transport System
G-34 Air Screen Intake for G2-3 Motor
G-35 Air Screen Intake for G2-6 Motor
Acct. H - Buildings (These to be in existing Building)
H-l Control Room (Rl)
(Air conditioned) Motor (Rl)
H-2 Laboratory Motor
(Air conditioned)
Acct. L - Reactors
L-l Gasifier
L-2 Regenerator
L-3 Absorber
Acct. M - Drums
M-3 Waste Heat Boiler Steam Drum
M-A PTG Water Separator
116
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Item Number
Service
Drive
Acct. P - Pumps & Drivers
P-4 A&B
Oil Transfer Pump and Spare
Motors
Acct. Q - Storage Tank
Q-l
Q-2
Q-3
Q-4
High Sulfur Oil Storage Tank
Ash Silo (Existing-16'0 to be modified)
Ash Silo (Existing-22'0 to be modified)
Limestone Storage Bunker
(Existing- to be modified)
Acct. R - Compressors & Drivers
R-3
R-4
R-6
R-8
R-9
R-10
R-ll
PTG blower (Z-in series) Motors (Rl)
Flue Gas Booster Fan Motor
Regenerator Air Blower Motor
Gasifier Air/Flue Gas Blower Motor
Absorber Air Blower Motor
Absorber SO. Blower (2 in Parallel)Motors (Rl)
Pulverizer Air Blower Motor
Acct. T - Heat Exchangers
T-l
T-3
T-5
T-6 A&B
T-7
Regenerator Off-Gas/Air Exchanger
Purge and Transport Gas (PTG) Cooler
Oil Storage Tank Suction Heater
Oil Transfer Line Heaters
PTG Gas Blower Suction Cooler
(Rl)
117
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6.2 Vessel Drawings
L-l Gasifier & L-2 Regenerator
L-3 Absorber
L-l, L-2, L-3 Design Data
Fuel Oil Storage Tank
118
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PAGE NOT
AVAILABLE
DIGITALLY
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PAGE NOT
AVAILABLE
DIGITALLY
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PAGE NOT
AVAILABLE
DIGITALLY
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1
2
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7
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11
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14
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17
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29
18
11
18
18
80
81
82
83
84
85
88
81
88
at
40
41
42
43
44
4ft
48
41
48
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-------
STONE & WEBSTER ENGINEERING CORPORATION
CLIENT
LOCATION
SUMMARY
TYPE OF UNIT .!
/ $ as
*,*, />//?
SHEET NO.
EST /9OB NO
DATE '
ISSUE NO /
J .
ITEM NO »
DESCRIPTION
DUTY MM BTU/HR
HOT Tl TZ
COLO T2 Tl
MTD 8F
U SERVICE RATE
TOTAL SURFACE Pf*"
NO. OF SECTIONS
ARRANGEMENT
SIZE
TYPE
GPM OF WATER
LBS/HR STEAM
NO. A HP OF MOTORS
SIDE
MEDIUM
ESTIMATED AP
DESIGN PRESSURE
DESIGN TEMPERATURE
MATERIAL - SHELL
MATERIAL - CHANNEL
MATERIAL - TUBES
MATERIAL - TUBE SHEET
NO. OF TUBES
SIZE X BWG X LENGTH
TUBE PITCH
ESTIMATED WEIGHT(LBS)
COST
VENDOR
DIRECT LABOR MH'S
J5-/4M-3
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366 346
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-------
6.3 Equipment Specifications
for Items in Accounts
B. F. G. M. P. R & T
127
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STONE & WEBSTER ENGINEERING CORPORATION
LOCATION
TYPE OF UNIT
DESCRIPTION
DUTY MM BTU/HR
HOT Tl TZ
COLO T2 Tl
MTD "F
U SERVICE RATE
ffoui £afe
TOTAL SURFACE
NO. OF SECTIONS
ARRANGEMENT
SIZE
TYPE
GPM OF WATER
LBS/HR STEAM
NO. Ok HP OF MOTORS
SIDE
MEDIUM
ESTIMATED AP
DESIGN PRESSURE
DESIGN TEMPERATURE
P**t« -1*4 Z>jf
MATERIAL SHELL
MATERIAL - CHANNEL
MATERIAL - TUBES
MATERIAL - TUBE SHEET
NO. OF TUBES
SIZE X BWG X LENGTH
TUBE PITCH
ESTIMATED WEIGHT(LBS)
4«PV#X. SfOP
COSY
VENDOR
DIRECT LABOR MH'S
F- /
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00
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WESTINGHOUSE ELECTRIC CORP.
FLUID BED/GAS OIL DESULFURIZATION
DEMONSTRATION PLANT
PROVIDENCE, RHODE ISLAND
SWEC J.O. NO. 12418.01
EQUIPMENT DUTY SPECIFICATION
OCT. 18, 1974
PAGE: 1 of 1
ABSORBER HEATER - ITEM F-3
One in-line burner for No.2 fuel oil for direct heating of air fed
by R-9 and 4% SO gas fed by R-10 to mixed gas outlet of 1200 F. F-3
similar to F-l and consisting of oil atomizing nozzle in a refractory
lined combustion chamber with connections for air inlet (estimated
size 8") Process gas inlet from R-10 (estimated size 8") at 300°F.
Normal intermittent oil flow 110 Ib/hr, heat release approx. 2MM BTU/Hr.
The following is a preliminary sketch for estimating, layout and report
purposes
129
STONE & WEBSTER
-------
WESTINGHOUSE ELECTRIC CORP.
FLUID BED/OIL DESULFURIZATION
DEMONSTRATION PLANT
PROVIDENCE, RHODE ISLAND
SWEC J.O. NO. 12418.01
EQUIPMENT DUTY SPECIFICATION
OCTOBER 18, 1974
PAGE: 1 of 15
LIMESTONE UNLOADING SYSTEM - ITEMS G-l SYSTEM
(See Dwg. 12418.01-61C)
System is designed to unload a 200 ton railroad car in less tha 1
shift (6 hours). This rate will be approximately 67,000 Ibs/hr.
The system is designed to convey over a 100' lift and 400 lineal
feet of tubing with 4 elbows. A pull-push system will be employed.
The material will be pulled out of the railroad car under a vacuum
and then sent to the existing coal bunker under pressure. System
will include: railcar unloading kit, 6" dia. tubing and fittings,
2-Filter/Receivers, 2 Rotary Airlocks with 3/4 HP motors, Vacuum
blower with 60 HP motor, Pressure blower 75 HP motor, required
silencers, flexible hoses, and inlet air filter.
Similar to systems manufactured by Ducon, Shick, etc.
130
STONE a WEBSTER
-------
WESTINGHOUSE ELECTRIC CORP.
FLUID BED/OIL DESULFURIZATION
DEMONSTRATION PLANT
PROVIDENCE, RHODE ISLAND
SWEC J.O. NO. 12418.01
EQUIPMENT DUTY SPECIFICATION
OCT. 18, 1974
PAGE: 2 of 15
WASTE STONE SYSTEM - ITEMS; G-2 SYSTEM
(See Dwg. 12418.01-61C)
System is designed to pick-up all the waste stone produced and convey
the material at a rate of 4,500 Ibs/hr., to either of the two (2)
existing ash silors. The system is a vacuum pressure type, and re-
flects a lift of 100 feet, and includes 600 lineal feet of tubing, 4
Rotary airlocks, 3 filter/receivers, 1 Diverter Valve, 1-30 HP Vacuum
Blower, 1-20 HP Pressure Blower, Lot of 4" tubing, fitting elbow
etc., silencers, flex hoses, manifolds, and inlet air filters.
Similar to systems manufactured by Ducon, Shick, etc.
131
STONE 81 WEBSTER
-------
WESTINGHOUSE ELECTRIC CORP.
FLUID BED/OIL DESULFURIZATION
DEMONSTRATION PLANT
PROVIDENCE RHODE ISLAND
SWEC J. 0. NO. 12418.01
EQUIPMENT DUTY SPECIFICATION
OCT. 18,1974
PAGE: 3 of 15
LIMESTONE WEIGHT FEEDERS-ITEM; G-3 A/B
Belt type weight feeder to meter between 750 and 4500 #/Hr. with an
accuracy of + % %. Weigh measurement accomplished by load cell.
Feeder is complete with 1 Hp DC motor, totalizer, digital rate setter,
HI-low alarm contacts, zero speed switch, external signal acceptor,
complete internal controls, all in a totally enclosed dust cover.
Unit is constructed of carbon steel with 304SS skirtboards.
Similar to Acrson Model 203-105Z.
132
STONE at WEBSTER
-------
WESTINGHOUSE ELECTRIC CORP.
FLUID BED/OIL DESULFURIZATION
DEMONSTRATION PLANT
PROVIDENCE, RHODE ISLAND
SWEC J. 0. NO. 12418.01
EQUIPMENT DUTY SPECIFICATION
OCT. 18,1974
PAGE: 4 of 15
RECYCLE FLUE GAS BAG HOUSE ITEM: G-4
Reverse air jet type dust collector to filter 30,000 ACFM of Flue
gas to 350 F and + 6" W.G. Unit sized at a 5:1 air to cloth ratio
(approx. 6,000 ft. of filtration area). Unit is furnished complete
with solenoid valves, air header, internal catwalks, external ladder
and catwalks, bracing, support legs, Nema IV electricals, trough (s),
screw conveyor (s), and air lock (s). Unit is constructed of carbon
steel with, Nomex bags.
Similar to Flex-Kleen Model 84UDC-608
133
STONE a WEBSTER
-------
WESTINGHOUSE ELECTRIC CORP.
FLUID BED/OIL DEULSURIZATION
DEMONSTRATION PLANT
PROVIDENCE RHODE ISLAND
SWEC J. 0. NO. 12418.01
EQUIPMENT DUTY SPECIFICATION
OCT.18, 1974
PAGE: 5 of 15
GASIFIER AIR SCREEN ITEM: G-5
Vertical Roll type, with automatic renewable media to filter 20,200
cfm of atmospheric air. Filter complete with 1/6 Hp motor drive
mechanism, internal controls, timers, etc. Unit shall be of carbon
steel construction with a glass fiber media.
Similar to American Air Filter Model Roll-0-Matic 5.93
REGENERATOR AIR SCREEN ITEM; G-6
Same description as G-5 except 2450 cfm of air is to be filtered.
Similar to American Air Filter Model Roll-0-Matic 3-50
134
STONE a WEBSTER
-------
WESTINGHOUSE ELECTRIC CORP.
FLUID BED/OIL DESULFURIZATION
DEMONSTRATION PLANT
PROVIDENCE, RHODE ISLAND
SWEC J. 0. NO. 12418.01
EQUIPMENT DUTY SPECIFICATION
OCT. 18,1974
PAGE: 6 of 15
PTG FILTER ITEM G-7
Inline cartridge type filter to remove all particles larger than 1
micron from 140 ACFM of gas at 100°F & 30 psig. Housing and cartridge
to be 316SS, replaceable filter sock to be polypropylene cloth. Particles
wet with corrosive steam condensate containing S02& C02.
Similar to Twistloc Model 100-1.
135
STONE a WEBSTER
-------
WESTINGHOUSE ELECTRIC CORP.
FLUID BED/GAS OIL DESULFURIZATION
DEMONSTRATION PLANT
PROVIDENCE, RHODE ISLAND
SWEC J. 0. NO. 12418.01
EQUIPMENT DUTY SPECIFICATION
OCT.18, 1974
PAGE: 7 of 15
GASFIER CYCLONES ITEM; G-8 A -D
Four Cyclones in parallel to handle 670 ACFS each at 1600°F and 2.6
psig. Internal design conditions 1710°F, 10 psig unit will be of
Carbon Steel construction with 5" of refractory lining on all inside
walls. Units to be 7'-ll"OD x 33'-0" high.
Similar to Ducon Size 1025 Type VM700-L5
REGENERATOR PRIMARY CYCLONE ITEM; G-9
Cyclone to handle 148 ACFS at 1960°F & 2.1 psig. Internal design
conditions 2012°F, 10 psig unit will be of carbon steel construction
with 7" of refractory lining on all inside walls. Unit to be 4«-6"OD
x 19'-0" high.
Similar to Ducon Size 225 Type VM700-L7
REGENERATOR SECONDARY CYCLONE ITEM; G-10
Cyclone is in series with G-9 and designed exactly the same as G-9.
136
STONE ft WEBSTER
-------
WESTINGHOUSE ELECTRIC CORP.
FLUID BED/GAS OIL DESULFURIZATION
DEMONSTRATION PLANT
PROVIDENCE, RHODE ISLAND
SWEC J. 0. NO. 12418.01
EQUIPMENT DUTY SPECIFICATION
OCT.18, 1974
PAGE: 8 of 15
SUCTION LINE DUPLEX OIL FILTER ITEM; G-14
Manual twin basket type oil filters. Carbon Steel housing with stainless
steel baskets. Switch-over accomplished by a manually operated plug
valve. Cover to be of quick opening type. Unit sized for a maximum
pressure drop of less than 1 psi.
Similar to Zurn Industries Model 4" 590 FPS.
137
STONE a WEBSTER
A
-------
WESTINGHOUSE ELECTRIC CORF.
FLUID BED/GAS OIL DESULFURIZATION
DEMONSTRATION PLANT
ROVIDENCE, RHODE ISLAND
SWEC J.O. NO. 12418.01
EQUIPMENT DUTY SPECIFICATION
OCT. 18, 1974
PAGE: 9 of 15
PULVERIZER CYCLONE ITEM; G-17
Cyclone to handle 25 ACFS at 794°F & 1.0 psig. Internal design
conditions 850 F, 10 psig. Unit of 304 S.S. construction to be
I1-11" O.D. x 9'-6" high.
Similar to Ducon size 40 VM700-L3
ABSORBER CYCLONE ITEM; G-18
Cyclone to handle 200 ACFS at 1436°F & 1.4 psig. Internal design
conditions 1700°F, 10 psig. Unit of carbon steel construction with
5" of refractory lining.
20'-0" high
Similar to Ducon size 330 VM700-L5
Unit is 4'-0" O.D. x
138
STONE a WEBSTER
-------
WESTINGHOUSE ELECTRIC CORP.
FLUID BED/GAS OIL DESULFURIZATION
DEMONSTRATION PLANT
PROVIDENCE, RHODE ISLAND
SWEC J.O. NO. 12418.01
EQUIPMENT DUTY SPECIFICATION
OCT. 18, 1974
PAGE: 10 of 15
BAG FILTER ITEM G-19 lincluding components G-19A and G-19B)
Reverse air jet type dust collector to filter 2625 ACFM of Regenerator
off gas at 300°F and - 8.5 WC (Design 3300 ACFM). Unit sized at a 5:1
air to cloth ratio (approx. 640 ft2 of filtration area). Unit is fur-
nished complete with solenoia valves, air header, external ladder and
cage, four support legs, NEMA IV electrical components, and air lock
(1/6 HP). It is carbon steel construction with nomex bags, size 13'-2"
high x 5'-4" square.
Similar to Flex-Kleen Model - 84 RA-64.
139
STONE Ot WEBSTER
-------
UESTINGHOUSE ELECTRIC CORP.
FLUID BED/GAS OIL DESULFURIZATION
DEMONSTRATION PLANT
PROVIDENCE, RHODE ISLAND
SWEC J.O. NO. 12418.01
EQUIPMENT DUTY SPECIFICATION
OCT. 18, 1974
PAGE: 11 of 15
FINES LOCK HOPPER - ITEMS G21A-D
o
Carbon steel shell, use Sturtevant -25ft size with 4" refractory lining.
Inlet nozzle 6" ID refractory.
Outlet nozzle 4' ID refractory.
Valves - Item G-22-A-D - 2" nominal bore.
Valves - Item G-23A-D - assume 2" air operated knife gate valves with
10" OD size companion flanges.
Valves - Item G-23 E-H - assume 2" air operated knife gate valves
with 10" flanges.
3
C.S. Stack cyclone lock hopper - Item G-28 Sturtevant - 3ft
3
C.S. Bag house lock hopper - Item G-29 Sturtevant - 3ft
G-30 A and G-31 A - 8" Carbon steel air operated knife gate valves.
G-30 B and G-31-B - 1-1/2" Carbon steel air operated knife gate valves.
G-32 and G-33 - 1-1/2" nominal bore
140
STONE a WEBSTER
-------
WESTINGHOUSE ELECTRIC CORP.
FLUID BED/GAS OIL DESULFURIZATION
DEMONSTRATION PLANT
PROVIDENCE. RHODE ISLAND
SWEC J. 0. NO. 12418.01
EQUIPMENT DUTY SPECIFICATION
OCT. 18, 1974
PAGE: 12o£ 15
PULVERIZER BLOWER AIR SCREEN ITEM: G-24
Same as G-6 except 700 cfm air
Similar to American Air Filter Model
Roll-0-Matic 3-50
ABSORBER BLOWER AIR SCREEN ITEM; G-25
Same as G-7 except 940 cfm air
Similar to American Air Filter Model Roll-0-Matic 3-50
141
STONE ft WEBSTER
-------
WESTINGHOUSE ELECTRIC CORP.
FLUID BED/GAS OIL DESULFURIZATION
DEMONSTRATION PLANT
PROVIDENCE, RHODE ISLAND
SWEC J.O. NO. 12418.01
EQUIPMENT DUTY SPECIFICATION
OCT. 18, 1974
PAGE: 13 of 15
LIMESTONE STAR VALVE - ITEM G-26
One star valve for feeding limestone of 1/8" to dust size at a rate
governed by G-3 A or B between 750 and 4,500 Ib/hr. The star valve
serves as a gas lock to prevent escape from L-l Gasifier, normal
pressure 2.6 psig, design 10 psig. The star valve and downstream
line will have purge gas supply available at 100F and 25 psig, to
prevent entrance of hot gas into the line. Motor size - %HP.
Similar valve by Beaumont - Birch Co.
142
STONE a WEBSTER
-------
WESTINGHOUSE ELECTRIC CORP.
FLUID BED/GAS OIL DESULFURIZATION
DEMONSTRATION PLANT
PROVIDENCE, RHODE ISLAND
SWEC J. 0. NO. 12418.01
EQUIPMENT DUTY SPECIFICATION
OCT.18, 1974
PAGE: 14 of 15
SOLID WASTE COOLER CONVEYOR- ITEM G-27
Normal Feed Rate: 2292 Ib/Hr. solids, size 75 micron to dust
Normal Heat Duty: .968 MM BTU/hr.
Design for 150^ of capacity and duty
^W
37.;
c*o
CdS
c*s^.
£*«Vfs
341.
Note;
Solids feed in and out by gravity, independently.
If feasible, use constant speed drive set for
Hell-flow to handle maximum solids rate.
If necessary to regulate drive to maintain level of solids,
suggest skin thermocouple on cover which will be responsive
to approach or recession of hot solids level, TC to vary speed.
143
STONE a WEBSTER
-------
WESTINGHOUSE ELECTRIC CORP.
FLUID BED/GAS OIL DESULFURIZATION
DEMONSTRATION PLANT
PROVIDENCE, RHODE ISLAND
SWEC J.O. NO. 12418.01
EQUIPMENT DUTY SPECIFICATION
OCT. 18, 1974
PAGE: 15 of 15
AIR SCREEN INTAKE FOR G-2-3 Item G-34
Same as G-
Simllar to American Air Filter Model - Roll-o-matic
AIR SCREEN INTAKE FOR G2-6 Item G-35
Same as G-
Similar to American Air Filter Model - Roll-o-matic
144
STONE a WEBSTER
-------
WESTINGHOUSE ELECTRIC CORP.
FLUID BED/GAS OIL DESULFURIZATION
DEMONSTRATION PLANT
PROVIDENCE, RHODE ISLAND
SWEC J.O. NO. 12418.01
EQUIPMENT DUTY SPECIFICATION
OCT. 18, 1974
PAGE: 1 of 1
PTG WATER SEPARATOR ITEM: M-4
A centrifugal type water separator to remove 200 Ibs/hr of water
from 205 acfm of gas at 100 F and 40 psig.
Similar to unit manufactured by Dyna-therm.'
145
STONE ft Wes9TER
-------
WESTI MG»MOU-:-E'
CLIENT
UOCATION
TYPE OF "-" OIL. GAtL>lF"lCATiON
STONE 8c WEBSTER ENGINEERING CORPORATION
PUMP SUMMARY
SHEET NO.
EST ./JOB NO. .
, OC.T \ 19 7A
ISSUE NO. .
K
ON
Q'TY |
KRVICG
a
a.
1 DRIVER I
L
TURBINE |
0
In
8
i
ARRANGEMENT & TYPE
MANUFACTURER » MODEU
SIZE | STAGES | RPM
CAP.EA.(GPM)
SUCTION(PSIG)
TEMP F
NP9H REQ.(FT)
DISCH .(PSIG)
PROD. S.G
MATERIAU
F .S EACH
INCU DRIVER
WEIGHT (UBS)
MANUFACTURER & MODEU
STEAM
DRIVER RPM
QTY |
CONDITIONS
GEARED TO RPM
HP | WR
WEIGHT OF TURBINE(S) (UBS)
MANUFACTURER 8r TYPE
DRIVER RPM
Q'TY [ HP
GEARED TO RPM
ENCU | VOUTS
WEIGHT OF MOTOR(S) (UBS)
PUMP (S )
TURBINE (S )
MOTOR ( S )
GEAR (S )
FREIGHT
TOTAU COST \/Q
| DIRECT UABOR MH'S
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-------
STONE & WEBSTER ENGINEERING CORPORATION
CLIENT
LOCATION .
COMPRESSOR SUMMARY
SHEET NO. .
EST./JOB NO. .
TYPE OF UNIT
OIL
OC.T-/ / "
ISSUE NO. .
APAA
ITEM NO.
QTY |
SERVICE
COMPRESSOR 1
DRIVER
TURBINE
K
O
0
b
s
TYPE
MANUFACTURER ft MODEL
SIZE | RPM
CAP.EA.(CFM) NO OF STAGES
SUCTION PSIA DISCH PSIA
TEMP * F MOL. WT.
NORMAL HP DESIGN HP
MATERIAL
F.a. EACH INCL. DRIVER
AUXILIARIES INCLUDED
WEIGHT (LBS )
MANUFACTURER Ck MODEL
STEAM CONDITIONS
DRIVER RPMJGEARED TO RPM
QTY | HP | WR
WEIGHT OF TURBINE(S) (LBS)
MANUFACTURER
DRIVER RPM GEARED TO RPM
Q'TYJ MP ENCL | VOLTS
WEIGHT OF MOTOR(S) (LBS)
COMPRESSOR (S) f£)
TURBINE (S)
MOTOR ( S ) f2j
GEAR ( S )
CONSOLE ( S )
FREIGHT
TOTAL COST
R-3
1 CZ. IN Seftieo)
PTfc BLOWER
CENTRiFud Ai
SujuoVAJtr /./*\C?//P
1
S"SO Z.
IH,b HSO
2. SO "2.8.69
/ 2. ^
3/<& 55 (2]
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I L
DIRECT LABOR MH'S
-------
STONE & WEBSTER ENGINEERING CORPORATION
CLIENT
LOCATION PROVI&EMC.E.
TYPE OF "-r Oil- £ AS \ PlgAT lOAJ
COMPRESSOR SUMMARY
SHEET "**
EST./JOB M
/ 2. «4 t ft- O
ISSUE NO. .
BV
ITEM NO.
QTY |
SERVICE
COMPRESSOR
DRIVER
TURBINE
K
0
S
k
8
TYPE
MANUFACTURER 8t MODEL.
SIZE | RPM
CAP.EA.ICFM) NO OF STAGES
SUCTION PSIA DISCH PSIA
TEMP F MOL. WT .
NORMAL H= DESIGN H=>
MATERIAL.
F.S. EACH INCL DRIVER
AUXILIARIES INCLUDED
WEIGHT (LBS )
MANUFACTURER flk MODEL
STEAM CONDITIONS
DRIVER RPM GEARED TO RPM
QTY | H=> | WR
WEIGHT OF TURBINE(S) (LBS)
MANUFACTURER
DRIVER RPM GEARED TO RPM
QTv| H=> EltCL | VOLTS
WEIGHT OF MOTOR(S) (LBS)
COMPRESSOR (S)
TURBINE (S)
MOTOR (S )
GEAR (S)
CONSOLE (S)
FREIGHT
TOTAL COST
R-1
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DIRECT LABOR MH'S
-------
LOCATION .
STONE & WEBSTER ENGINEERING CORPORATION
EXCHANGER SUMMARY
SHEET NO.
/
OF
/
TYPE OF UNIT fej
&*eM£asJ2jJL
DATE _
ISSUE NO
NO J2,4rigi._0_/
. io)nP/74.
1 E
DESCRIPTION
DUTY MM BTU/HR
HOT Tl TZ
COLD T2 Tl
MTD eF
U SERVICE RATE
TOTAL SURFACE - pf**
NO. OF SECTIONS
ARRANGEMENT
SIZE
TYPE
GPM OF WATER
LBS/HR STEAM
NO 8. HP OF MOTORS
SIDE
MEDIUM
ESTIMATED AP
DESIGN PRESSURE
DESIGN TEMPERATURE
MATERIAL SHELL
MATERIAL - CHANNEL
MATERIAL - TUBES
MATERIAL - TUBE SHEET
NO. OP TUBES
SIZE X BWG X LENGTH
TUBE PITCH
ESTIMATED WEIGHT(LBS)
COST
P"y O I&U t"
VENDOR'
DIRECT LABOR MH'S
T"_ /
1 ~ 1
feeeHffafo-S
W-fai/S/
iSXflttinqe I/
0.67
600 3oo
4-ao 100
80 _.
2.7
34oo
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/
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C . ^- X« /a^X
O
-------
6.4 SUMMARY OF EQUIPMENT ROOMS
A. H-l Control Room & Model Shop
Control room size; 10'-0" W by 20«-0"L by 15'-0"H
B. H-2 Laboratory
1. Space for following type of tests and apparatus
a) Viscosity, density, sulfur content, carbon, conrad-carbon,
vanadium, nickel, sodium, water, asphaltenes, calcium, carbon dioxide,
inerts
b) Sieve analysis, atomic adsorption, wet extraction, muffle furnace, gas
chromatograph.
c) Racks for samples - 24 hour basis
2. Size - 10'-0"W by 20'-0"L by 15'-0"H
C. Switchgear Room
Size; 10'-0"W by 15'-0"L by 15'-0"H
D. Analyzer Shelter
Size; 10'-0"W by 15'-0"L by 8'-0"H
6.5 SUMMARY OF EQUIPMENT COSTS
Table E-l summarizes the equipment costs for the reaction system
of the demonstration plant. Table E-2 gives a breakdown of the fuel-gas
piping system cost. The dry sulfation system equipment costs are listed in
Table E-3.
151
-------
Table E-l
REACTION SYSTEM-PROCESS EQUIPMENT
Item
I Material ($) I Labor ($) I Total (S) I
Basis and Comments
Gasifier (L-l)
Regenerator (L-2)
222,000
41,000
Gasifier Cyclones (G-8 A-D) 213,000
Gasifier Fines Recycle . 188,100
System (G-21 A-D),
(G-22 A-D), (G-23).
(G-23 A-D), (G-23 E-H)
Regenerator Primary Cyclone 23,800
and Trickle Valve (G-9),
(G-9A)
Regenerator Secondary 23,800
Cyclone and Trickle Valve
(G-10) and (G-10A)
Gasifier Air/Flue Gas 208,750
Blower (R-8) and Console
and Screen (G-S)
Regenerator Air Blower (R-6) 39,230
and Screen (G-6)
Fuel-Gas Piping System 521,415
(see Table E-2)
13,400 235,400 20Z A/F ratio; 6 ft/sec velocity
20' 8" diameter; 23" vertical
matIs. - shell and Internals - 82,500
refractory lining - 78,700
distributor - 66,250
900 41 ,'900 8Z S02; 6 ft/sec velocity
6' dlam. x 22' vertical
matls. - shell - 17,800
refractory lining - 15,340
distributor - 6,895
1,800 214,800 20Z A/F ratio; 4 units; 670 scfs each;
7'11" O.D. x 33' vertical
(30Z o.c. on volume)
matls. -
shell (est. from SWEC) - 168,000
refractory - 38,000
5,800 193,910 4 Sturtevant Systems; 25 ft volume
(i> 40 mln capacity) per lock hopper;
valves from Dally Eng. Co.
matls. - lock hoppers, refractory,
air knife valves and
connecting piping - 108,000
inlet butterfly valves - 38,000
outlet ball valves - 34,000
450 24,250 8Z SO,; 148 acfs; (151 o.c. on volume)
4' 6" O.D. x 19' vert.
matIs. -
shell (est. from SWEC) - 17,000
refractory - 4,240
trickle valve - 2.500
450 24,250 Identical to primary cyclone
20,300 229,050 20Z A/F ratio, 46,000 ft3/mln; 23.3 pala
2585 hp. c.s.i 79.70 $/hp
(4Z o.c. on volume/15Z on hp)
1.800 41,030 2500 ft3; 24.0 psla; 145 hp, c.s.,
266 $/hp (30Z o.e. on volume) 38Z o.c.
on hp) net 80Z o.c. on h.p.
150,000 671,415 (design velocity < 100 ft/sec;
150 ft/sec velocity would reduce
diameter by 21Z)(Plpe layout not
optimum)
152
-------
Table E-l(Continued)
Item | Material ($) I
Regenerating Off-Gas System
Waste heat boiler and
steam drum' (B-l), (M-3)
Regenerator off-gas/air
exchanger (T-l)
Bag filter (G-19),
32,000
17,500
8,800
Labor (S) I
700
500
300
Total ($) |
32,700
18,000
9,100
Basis and Comments
3.39 MM Btu/hr; 3930 Ib steara/hr at
366°F; c.s.
0.67 MM Btu/hr; c.s.
2525 acfm; 300° F
wrapper (G-19A), and
discharge valve (G-19B)
Flue-Gas Recycle System
Stack cyclone lock 10,200
hopper (G-28), valves
(G-30A, G-30B, G-32),
and connecting piping
Recycle flue gas bag- 55,600
house (G-4), screw
conveyor (G-4A), and
airlock (C-4B)
tfaghouse lock hopper 10,200
(G-29), valves (G-31A-B,
G-33) and connecting
piping
Flue gas booster fan 10,300
-------
Table E-2
FUEL-GAS PIPING SYSTEM
Ul
Item
Shop Fab. Pipe
48"
46"
90° Elbows
45° Ells
Tees
Reductions
Flanges - 48"
Flanges - 46"
Blind Flanges
Welds
Testing Fab.
Bolt-up
Hangers and Supports
Design Allowance
Refractory
Additional Structures
TOTAL COST
| Material ($)
2,700
15,330
11,440
4,200
21,200
9,215
59,160
79,200
8,100
70,820
6,850
57,600
51,900
50,000
73,700
521,415
| Labor ($) | Total ($)
12,700 15,400
57,000 72,330
11,440
4,200
21,200
9,215
59,160
79,200
8,100
70,820
6,850
9,400 9,400
43,700 101,300
14,100 66,000
50,000
13,100 86,800
150,000 671,415
| Basis and Comments
35 ft; gasifier to cyclone; C.S.
210 ft; cyclones to burners; C.S
3; 46"; C.S.
1; 46"; C.S.
4; 46"; C.S.
4; C.S.
34; C.S.
45; C.S.
3; 46"; C.S.
79; C.S.
i
Matls. at 20% of above items
15% allowance
Estimated from SWEC costs
Includes 10% design allowance
-------
Table E-3
STONE PROCESSING SYSTEM (Dry Sulfation)
Item
T Material (?) \ Labor ($) I Total ($) I
Basis and Comments
in
Ln
Absorber (L-3)
241,000
Absorber Cyclone (G-1B) 24,000
and Trickle Valve (G-18A)
Jet Pulverizer (G-20) 34,200
Pulverizer Cyclone (G-17) 7,600
and Trickle Valve (G-17A)
Absorber Air Blower (R-9) 33,130
and Screen
Absorber S02 Blower 67,000
Pulverizer Air Blower 37 730
(R-ll) and Screen (G-24)
13,500 254,500 0.4 ft/sec; 21' 4" diam. x 23' vert.
MatIs. - shell - 72,700
refractory - 76,200
distributor - 87,050
300 24,300 200 ACFS at 1436°F;
4' diam. x 20' vert.
(35Z o.c. on volume)
tiatls. -
shell (est. from SWEC) - 18,000
refractory - 3,400
trickle valve - 2,500
560 34,760 870 scfm at 75 psia, 932°F
1500 Ib stone/hr. (45Z o.c. on gas)
Fluid Energy Co., Hatfield, Pa.
300 7,900 25 ACFS at 794°F;
I1 11" O.D. x 91 6" vert.
(27Z o.c. on gas volume)
Matls. -
shell (est. from SWEC) - 5,000
trickle valve - 2,500
1,200 34,330 870 cfm; 22.5 paia; 45 hp
722 $/hp (15Z o.c. on gas/67Z*o.c.
on hp)
1,700 68,700 3500 cfm; 22.5 psla; 316 S.S.;
262 hp; 255 $/hp (15Z o.c. on
gas/lSZ o.c. on hp)
1,800 39,530 600 cfm; 75 psia; C.S.; 110 hp
337 $/hp (5Z over capacity on
compressor/37Z o.c. on motor bp)
Solid Waste Cooler/
Conveyor (G-27)
Absorber Heater (F-3)
TOTAL PROCESS EQUIPMENT
47,500
2.200 49,700 1.5 x stoichiometric stone rate
560
13.160 2 MM Btu/hr; 1200°F exit temp.
22,120 526,880
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7. ONE LINE DIAGRAM
156
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PAGE NOT
AVAILABLE
DIGITALLY
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8. ECONOMICS
8.1 Base Conditions and Assumptions - 50 MW Demonstration Plant
(a) Retrofit to New England Power Service Co. No. 12 Boiler, Manchester St.
Station, Providence, R. I.
(b) Cost of removal of No. 10 and No. 11 Boilers to provide space is not in-
cluded, and cost of refractory lined 42" piping for hot fuel gas is based
on four burners located together in near side of No. 12 boiler. Tangen-
tial firing will require at least 50% more 42" piping.
(c) Boiler house, overhead coal bunker, ash silos, railroad siding, rail car
spotting equipment, oil barge unloading facilities are all existing and
usable.
(d) Steam, boiler feed water, service cooling water, electric power, plant
air, instrument air, and low sulfur fuel oil are all available by con-
necting to existing supplies at no cost.
(e) Continuous firing of present normal fuel at 5% minimum of boiler capacity;
50 MW power requires 422,000 Ib/hr steam. The hot fuel gas is estimated
to raise 430,000 Ib/hr. The boiler capacity is 450,000 Ib/hr.
r
(f) Fuel oil feed-wt-off has sufficiently fast response to boiler fuel re-
jection signals so no need for hot fuel gas line to have cut-off and
relief facilities; if the latter were required, significant cost and
major practical problems would be entailed.
(g) Design parameters and effectiveness of Absorber reaction at 94% minimum
efficiency (not more than 6% of SO- and of CaO unreacted) will be tested
and proven in Pilot Plant.
(h) Complete reaction of residual CaS in regenerated stone occurs in Absorber,
(so that waste stone is unreactive toward natural waters) will be tested
and proven in the Pilot Plant.
(i) Suitable limestone supply 300 miles by rail. Suitable waste solids
disposal 30 miles by truck.
(j) Horizontal pulsed flow for controlled stone circulation will be tested
and proven in a full scale cold flow model.
(k) Air jet pulverizing of regenerated stone for feed to Absorber will be
tested and proven in manufacturer's test equipment.
(1) The 50:1 scaleup of processes proven in Abingdon Pilot Plant is feasible
and appropriate.
159
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8.2 Investment Cost Summary - 50 MW Pilot Plant. Total Project
Item
PROCESS EQUIPMENT
A. Towers
B. Boilers, S- heaters
F. Process furnaces
G. General equipment
L. Reactors
M. Drums
Q. Storage tanks
P. Pumps (including drives)
R. Compressors (including drives)
S. Stacks
T. Heat exchangers
TOTAL PROCESS EQUIPMENT
PROCESS MATERIALS
C. Piping
D. Structures
E. Electrical
H. Buildings
J. Civil
K. Instruments
N. Insulation
N. Paintingb
Refractory lining
TOTAL PROCESS MATERIALS
TOTAL DIRECT COST (DM +DL)
DISTRIBUTABLE ACCOUNTS
VI. Insurance (excluding all risk)
V2. Federal and states taxes
X. Temporary construction facilities
Y. Field office (Including insuiance <
Z. Construction tools and equipment
0. Other distributable items
TOTAL DISTRIBUTABLE
SUBTOTAL-COST OF WORK
U. INDIRECT ACCOUNTS
Engineering
Design
Other headquarters office
Taxes headquarters payroll
Overhead allowance
TOTAL INDIRECT (HEADQUARTERS OFFICE)
TOTAL, PRESENT DAY PRICES (BARE COST)
Fee, Escalation, and Contingency
Order of Accuracy
"From original SWEC table.
Subcontract.
(Material
32,
151,
775,
504,
355,
3,
429,
58,
2,309,
1,048,
169,
158,
24.
65,
129,
54,
87,
60,
1,798,
4,107,
ind taxes)
f$)| Labor (S)
000 600
200 5,100
000 45,900
000 27,800
500 200
500 35,500
100 300
500 30,100
700 1,800
500 147,300
000 545,000
800 48,800
800 221,400
500 25,400
800 110,100
200 54,600
000
900
000
000 1,005,300
500 1,152,600
I Total ($)
32,600
156,300
820,900
531,800
700
391,000
3,400
459,600
60,500
2,456,800
1,593,000
218,600
380,200
49,900
175,900
183,800
54,000
87,900
60.000
2,803,300
5,260,100
850,000
850.000
6,110,100
1,250,000
1.250.000
7,360,100
Not included
15*
160
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8.3 Operating Cost Summary
Unit costs used are as provided by NEES (New England Electric Systems) basis
1976; - Internal Power - 23 Mils/KW Hr; No. 6 Fuel Oil - $2.25/MM BTU (HHV);
Capital Charges - 20 Percent of Bare Cost/Yr; Maintenance (including con-
struction tools, supplies and labor) - 5 percent of bare cost/yr.
Annual costs assumed on 7000 hr/yr stream operation. Other costs estimated
by SWEC.
Bare Cost (Accuracy 15%) $7,360,100.00
Direct Operating Costs
Power $1,472,020.00
Operating Labor 304,500.00
Limestone Supply 76,250.00
Solid Waste Disposal 40,880.00
Maintenance 368,005.00
Total 2,261,655.00
Indirect Operating Costs
Capitol Charges 1,472,020.00
Total Direct & Indirect Costs 3,733,675.00
Fuel Operating Cost 8,100,000.00
Total Operating Costs 11,833,675.00
161
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APPENDIX F
BOILER MODIFICATIONS
F
-------
APPENDIX F
BOILER MODIFICATIONS
An initial evaluation of the interfaces between the demonstra-
tion plant boiler and the fluidized bed residual oil gasification process
was carried out. Combustion Engineering Power Systems Services (CE) pro-
vided certain engineering services designed to establish parameters for
firing the high-temperature, low heating-value gas in the CE steam gene-
rator (unit No. 12) at the Providence, Rhode Island plant of the New
England Electrical Services (NEES) (see Appendix C). Westinghouse pro-
vided CE with a design basis taken from conservative CAFB performance
estimates (20 percent air/fuel ratio; high stack-gas recycle requirement).
Little information was provided on the probable range of performance cap-
able by CAFB. A summary of their study and an assessment of the requirements
for the demonstration plant are presented.
SUMMARY OF THE CE ENGINEERING EVALUATIONS
An analysis of current performance data and latest equipment
capabilities reveals that high-temperature, low heating-value gas can be
fired in this unit. The maximum steam generating rate would be limited
to 180,440 kg (400,000 Ib) steam per hour with existing equipment because
of induced draft (I.D.) fan limitations. To achieve a 204,120 kg
(450,000 Ib) per hour steaming rate new I.D. fans must be supplied.
New burners will be required to fire the low heating-value gas.
After careful consideration, the CE tangential firing system was chosen
to fire this fuel. Tangential firing burners would be installed in the
four corners of the unit. The burners would be arranged with a center
No. 6 oil compartment in which the flame scanner (supplied with the burner
control system) would be located. Adjacent to this oil compartment, a CE
152 mm (6 in) side eddy plate No. 2 oil igniter would be installed.
163
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Above and below the center oil compartment, compartments would be provided
for low heating-value gas firing. Within these compartments, Type 309
stainless steel burners would be installed with their manifolds. Mani-
folds would terminate immediately adjacent to the windbox.
If the tangential firing system is selected, it will be neces-
sary to plug the existing burner openings when low heating-value gas is
fired, and it is expected that a 48-hour shutdown will be required to
change over to low heating-value gas firing.
When firing the high-temperature gas, higher exit gas tempera-
tures are expected in the economizer area,based on the conservative
design basis used by CE. The economizer supports must be modified to
provide insurance against failure. Inasmuch as the condition of the
lower economizer bank is questionable, a replacement is recommended.
PREDICTED PERFORMANCE
When firing low heating-value gas having an analysis
projected for the 20 percent air/fuel ratio case
see Appendix E:
- Steam load - 204,120 kg (450,000 Ib/hr)
- Unit efficiency - 82.5 percent
- Superheater steam temperature - 510°C (950°F)
When firing No. 6 Oil:
- Steam load - 204,120 kg (450,000 Ib/hr)
- Superheater steam temperature - 510°C (900°F)
MATERIALS AND EQUIPMENT AND COST ESTIMATE
The following material and equipment would be required to imple-
ment the recommendations made by CE:
Four (4) tangential burners for low heating-value gas
and No. 6 oil firing. All material in contact with
low heating-value gas is to be 309 stainless steel.
Burners to be complete with windbox, air dampers, and
so on.
164
-------
Four (4) No. 2 oil side eddy plate self-proving ig-
niters rated at 2.11 MJ/hr (2 M Btu/hr), one (1) per
corner, located adjacent to oil compartments
Four (4) igniter air fans, one (1) per corner, com-
plete with butterfly valve, air piping
Two (2) scanner air fans with flapper arrangement for
back-up complete with air piping
Furnace Safety Supervisory System hardware
Bent tube inserts for the tangential burner openings
in the sidewall pressure parts as close as possible
to the corners
Bent tube inserts for the side igniters
Wall boxes and mountings for the side igniters
Refractory, insulation, casing, and so on, for the
burner openings, igniter openings, closing up other
openings, patching up affected furnace areas, and so
forth
Hot air ductwork from existing hot air source to
burner windboxes, including air foils in each side for
air measurement
Insulation, and so on, for hot air ductwork
Removable refractory plugs for existing six (6) hori-
zontal burners, to be used for protection when these
burners are not in use during tangential firing
Economizer supports for entering gas temperature of
704°C (1300°F)
Two (2) I.D. fans complete with motor drives
Transition gas ductwork for new I.D. fans
Insulation, and so on, for gas ductwork and I.D. fans
One (1) set lower bank economizer elements (included
since condition of existing elements is unknown)
The budget price associated with this scope of supply is
$612,000. The erection effort associated with the installation of the
165
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above listed material, to include both removal and installation, is esti-
mated at $300,000.
The following materials and actions would be required to com-
plete the modifications. The above prices do not cover this work scope.
Low heating-value gas piping to burner including gas
control valves and shutoff dampers for each low heat-
ing value gas compartment (two [2] per corner)
Burner piping material
Relocation of boiler feedwater piping
Relocation of platforms and stairways
Relocation of any other equipment
Foundations, structurals, and so forth, for I.D. fans
Any gas-cleaning equipment to meet pollution codes
Electrical material and labor for wiring burner con-
trols
Combustion control material and erection labor.
It is estimated the modifications can be completed in 18 months.
ASSESSMENT
The study carried out by CE represents an initial evaluation of
the boiler requirements. The study adopts the point of view that the
plant should be modified to produce the maximum boiler design capacity,
given certain assumptions about the gasification process. While this
perspective is of interest, it may not be necessary or desirable to adopt
this point of view for a demonstration plant. The study was also influ-
enced by the basis which CE used for the gasification process. Certainly,
some modifications would be required to adapt the existing boiler to its
role in a low heating-value gas demonstration plant. The new economizer
supports, burner plugs, and dual fuel safeguard system are examples of
those required adaptations. In Westinghouse's opinion, however, the need
for tangential firing, plus wall and tube modifications for that type of
firing, is marginal. They are certainly the kind of modifications that
would be required to maintain maximum boiler efficiency, but this is not
166
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a requirement of the demonstration plant. It appears that a substantial
saving could be made by using the four-burners-in-one-wall approach to
boiler firing, and with fewer potential problems than with longer, more con-
voluted fuel-gas pipes. If this approach caused some loss in boiler ef-
ficiency, it would not be significant to the overall program. In addition,
it should be noted that Stone and Webster Engineering Corporation (SWEC)
included in their costs new air ducts and wind boxes for the new gas
burners. Two factors indicate that a new I.D. fan may not be required
for the demonstration plant:
The design basis for the plant is 195,000 kg steam/hr
(430,000 Ib steam/hr) (equivalent to 50 MW) and not
204,000 kg/hr (450,000 Ib/hr).
The base design conditions are the most conservative
conditions, and it is probable that lower flue gas
volumes will exist in reality (by using a 17 percent
air/fuel ratio, or water injection for gasifier tem-
perature control, for example'). Thus, the existing
I.D. fan may be sufficient but, in any case, should
not be replaced until the demonstration plant program
has progressed to a point indicating its need.
In summary, further review of the boiler modification study is
required before an optimal decision can be made regarding material and
equipment requirements. Heat and material balance data on low heating-
value gas firing, specific equipment costs, trade-off reviews on tangen-
tial burners over side burners, projection of boiler performance with and
without new I.D. fans, and other related engineering data must be devel-
oped and reviewed.
167
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APPENDIX G
ENVIRONMENTAL IMPACT
-------
APPENDIX G
ENVIRONMENTAL IMPACT
The fluidized bed oil gasification process, in which limestone
or dolomite removes the sulfur from fuel gas during the gasification
process, has been developed to permit the utilization of high-sulfur
residual fuel oil with conventional boilers by producing a low-sulfur fuel
gas. The process can be operated as a once-through limestone sorbent
system, a sorbent regeneration/sulfur recovery system, or a sorbent re-
generation system without sulfur recovery by capturing the sulfur-rich
gas from the regenerator with the spent stone. The spent stone from each
system alternative can be processed to minimize the environmental impact
of the waste stone for disposal or to provide material for potential market
utilization (Appendices K-N).
Under contract to the U.S. Environmental Protection Agency (EPA),
Westinghouse is carrying out laboratory support work on the sulfur removal
system, and Esso Research Centre, Abingdon, England (Esso) is carrying out
pilot-scale tests to investigate sulfur removal. A summary of the Esso
pilot-scale experimental results is shown in Table G-l. The demonstration
plant has been designed on the basis of these results.
Table G-l
SUMMARY OF ENVIRONMENTAL PERFORMANCE OF
CHEMICALLY ACTIVE FLUIDIZED BED (CAFB) PILOT PLANT
Sulfur Removal Potential - Up to 95%
NOx Control Potential - Normal 270 ppm reduced to 155 ppm
Limestone Makeup Rate - 0.5 to 1.5 times stoichiometrlc
Removal of Metals - 100% fuel vanadium, 75% nickel,
40% sodium
Fuel Oils Utilized - Atmospheric residual (vacuum bottoms
on batch unit tests)
Processing Steps Demonstrated - Gasification and regeneration
169
-------
The gaseous sulfur dioxide (SO.) and particulate sulfate emission
from the 50 MW oil gasification demonstration plant has been investigated.
New federal Source Performance standards for fossil fuel-fired steam
generators of greater than 263.75 Gj/hr (250 million Btu/hr) heat input
are 0.34 kg/GJ (0.8 lb/106 Btu) for sulfur dioxide and .043 kg/GJ
(0.1 lb/10 Btu) for particulates. For the oil gasification demonstration
plant, the sulfur dioxide output is projected to be 0.15 kg/GJ (0.35 lb/10
Btu) and is safely within the limit. This is based on 90 percent sulfur
removal for the 2.6 percent sulfur residual oil for the demonstration
plant. The total particulate emission is projected to be 0.0043 to
0.043 kg/GJ (0.01 to 0.1 lb/10 Btu) which meets the Performance Standard
of 0.043 kg/GJ (0.1 lb/106 Btu). (See Appendix D).
Since the particulate emission from the oil gasification plant
consists mainly of calcium oxide (CaO), the environmental impact of calcium
oxide particulate emission is also of interest. According to the Analytical
Handbook of Toxicity, calcium oxide is not toxic, but is irritating to the
o
eye and skin and respiratory system. Its Threshold Limit Value is 5 mg/m .
The stack exhaust from the oil gasification demonstration plant as
projected from the composition of the stack fines from the Esso pilot
tests is to contain less than 0.026 kg/GJ (0.06 lb/106 Btu) of calcium oxide
based on the fines containing 60 percent calcium oxide and with 0.043 kg/GJ
(0.1 lb/10 Btu) particulate emission. This is equivalent to-approximately
0.032 kg/453 kg (0.07 lb/1000 Ib) effluent gas in the 50 MW demonstration
plant, which is well below the emission standard of 0.1134 kg/453 kg
(0.25 lb/1000 Ib) for the cement industry.
There is evidence that sulfuric acid (l^SO^) aerosols and sulfates
may be more potent irritants than sulfur dioxide. 2~5 The literature on
current knowledge and on emission standards of these species has been
reviewed:
Approximately 95 percent of the sulfur emitted to the
atmosphere from urban sources is in the form of sulfur
dioxide. The primary urban source of sulfate is the
atmospheric oxidation of sulfur dioxide to sulfuric acid,
170
-------
with subsequent neutralization or exchange reactions giving
a variety of sulfates.
Sulfates can have a deleterious effect on human health and
weather, climate, visibility, and vegetation. Reliable no-
effect threshold levels have not been established. Sulfuric
acid and certain sulfates are more potent irritants than
sulfur dioxide.
The irritant response of certain mixtures of sulfates is
probably greater than the sum of the responses of the
individual compounds and is related to the size and number
density of the sulfates.
The major portions of the sulfate particles in the atmosphere
are in the respirable range.
Approximately 20 percent of the sulfate exists as heavy-metal
(zinc, iron, manganese, lead, vanadium, etc.) sulfates, 20 to
30 percent is in the form of ammonium sulfates, and 50 percent
exists as sulfuric acid.
Little is known of the individual sulfate compounds. Amdur
and Corn studied the irritant potency of zinc ammonium sulfate,
zinc sulfate, and ammonium sulfate and found the double salt
sulfate the most and ammonium sulfate the least irritating
of the three compounds. Not all sulfates are irritants in
nature; ferric sulfate is, but ferrous, manganese and calcium
g
sulfates are not.
At eastern urban sites (east of the Mississippi River), sulfate
contributed 10 to 20 percent of the annual average particulate
loading, whereas at western urban sites (west of the Missis-
sippi) , sulfate contributed 5 to 10 percent of the annual average
particulate matter in ambient air.
In 1970, the national average sulfate concentration at urban
3 3
locations was 10.1 yg/m . The nonurban average was 6.3 pg/m .
Current knowledge and available data are inadequate at this
time to provide criteria that might be used as a basis for
standards.
171
-------
Chemical analysis of the stack fines from the Esso pilot
plant indicates that the fines contain 13 percent calcium sulfate (CaSO.).
If this composition holds true for the stack particulate emission from
the demonstration plant, calcium sulfate emission is projected to be
0.0043 kg/GJ (0.01 lb/10 Btu). Although there is no emission regulation
for sulfate particulate at the present time, a research and development
program is outlined by EPA to establish criteria that can be used as a
basis for control standards in the near future. Since calcium sulfate
p
is considered inert, it is not expected to cause any pollution problems
at this level.
Environmental factors for power generation by oil gasification
with a conventional plant are compared with those of a conventional oil-
fired power plant using limestone wet scrubbing in Table G-2. Each of the
processes has some advantages. The major advantage of the once-through limestone
wet-scrubbing process is the status of its development. The oil gasifi-
cation processes have cost advantages and potentially lower waste and
emission levels.
The spent stone from the fluidized bed oil gasification/desul-
furization process is dry and granular (-6.35 mm/-l/4 in). The composition
of the spent stone is primarily calcium oxide and calcium sulfide, with a
small amount of calcium sulfate. The disposal of this solid may be
12
accomplished by a variety of methods. Several processing alternatives
have been developed to convert calcium oxide and calcium sulfide to an
environmentally acceptable form for disposal (Appendix F). Dry sulfation
(Appendix L) and dead-burning (Appendix 0) are examples of dry processing
systems; slurry carbonation (Appendix M) and acid sulfation (Appendix N)
are examples of wet methods investigated for spent stone processing before
disposal or further utilization.
The environmental Impact of any disposed material is a function
of its physical and chemical properties and the quantity involved. Poten-
tial water pollution problems can, in many cases, be predicted by chemical
properties such as solubility, the presence of toxic metals, and pH of
leachates. The disposal of spent stone from the oil gasification processes
172
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Table G-2
ENVIRONMENTAL IMPACT COMPARISON
200 MW conventional oil-fired
power plant with limestone or
lime slurry flue-gas scrubbing
200 MM conventional oil-fired
power plant with CAFB oil
gasification and dry sulfatlon
bf solid product
Plant Emissions
S02 kg/GJ (Ib/mm Btu)
N0x kg/CJ (Ib/mm Btu)8
Total Participates kg/GJ (Ib/mm Btu)
CaO
Solids Mg/day (tons/day)
chemical form
Liquids Kg (tons/day)1*
chemical form
0.15 (0.35)
0.34 (0.80)
0.043 (0.10)
None
None
145 (160)
CaSO., CaSO,, CaCO., inerts (sludge)
145 (160)
Water (sludge)
0.15 (0.35)
0.17 (0.40)
0.043 (0.10)
0.026 (0.06)
0.0043 (0.01)
145 (160)
95Z CaS04, CaO, inerts
None
None
Resource Usage
Limestone Mg (tons/day)
Process water, liters/day (gal/day)
Auxiliary power (ku)
Steam, kg/day (lb/day)d
Land, m (acres)
Ability to utilize atmospheric
residual oils (06 oil)
Ability to utilize vacuum bottoms
with high metals content.
118 (130)
1.1 x 106 (3.0 x 105)
6 x 103 - 8 x 103
3.2 x 105 (7 x 105)
2 x 10*-4 x 10* (5-10) per year
per 3 m (10 ft) depth of pond
Yes
No
118 (130)
1.1 x 105 (3.0 x 10*)
6 x 103 - 8 x 103
None
None
Yes
Yes
*NO emission for gaslfler based on pilot-plant measurements and experience with fluidized bed combustion.
(Taste water based on 50Z water sludge product. Typical sludge properties and land requirements are reported in
J.W. Jones, "Environmentally Acceptable Disposal of Flue Gas Desulfurization Sludges: The EPA Research and
Development Program" and C.N. Ifeadi, U.S. Rosenberg, "Lime/Limestone Sludge Disposal - Trends in the Utility
Industry"; both presented at the EPA Flue Gaa Desulfurization Symposium, Atlanta, Georgia, Nov. 1974.
°Process water and auxiliary power for limestone and lime slurry scrubbing baaed on G.G. McGlamery and R.L.
Torstrick, "Cost Comparisons of Flue Gas Desulfurization Systems," presented at the EPA Flue Gas Desulfurization
Symposium, Atlanta, Georgia, Nov. 1974.
Based on steam for stack-gas reheat.
-------
may create air pollution or odor nuisance, e.g. hydrogen sulfide (H_S),
depending on the amount of calcium sulfide present. Heat may be released
on hydration and carbonation of calcium oxide. Potential water pollution
may be introduced from the runoff leachates caused by the rainfall and
naturally occurring subsurface flow through the landfill site. In order
to predict leachate characteristics of a landfill, it is first necessary
to describe the general features of water movement and geological consid-
eration for this disposal method. Due to the recent surge of ecological
Interest in sanitary landfills utilized for solid waste disposal, there
is an abundance of information available. Enrich's review of research
9 10
in this field presents an overall view of progress on this subject. '
Westlnghouse will be continuing investigations conducted toward the
definition and assessment of this problem; results available to date,
however, are not sufficient to assess fully its extent. Tests on stone
analysis, leaching properties, heat release properties, landfill proper-
ties, and air emission are being planned to obtain this information.
To summarize, the CAFB process compares favorably environmentally
to limestone and lime slurry scrubbers. Sulfur removal capabilities for
the processes appear comparable:
CAFB provides a considerable reduction in nitrogen oxide emissions.
The CAFB process consumes an order of magnitude less process
water than slurry scrubbing processes.
Limestone usage is comparable for the processes at about 1
mole of calcium per mole of sulfur removed from the fuel,
but the CAFB process could potentially reduce this utiliza-
tion to half that level by utilizing a regenerative operation
with sulfur recovery.
The CAFB process produces a dry oroduct with potential market
value rather than a sludge which is difficult to handle and
requires large land areas for disposal ponds.
174
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REFERENCES
1. Lund, H.F. Industrial Pollution Handbook McGraw Hill Co. New York.
2. Summary Report on Suspended Sulfates and Sulfuric Acid Aerosols.
EPA-650/3-74-000. Final Report, 1974.
3. Air Quality Criteria for Sulfur Dioxide. U.S. Department of Health,
Education and Welfare. AP-50. 1969.
4. Air Quality Criteria for Partlculate Matter. U.S. Department of Health,
Education and Welfare. AP-49. 1969.
5. Lewis, T.R., M.O. Amdur, M.D. Fritzland, and K.I. Campbell. Toxicology
of Atmospheric Sulfur Dioxide Decay Products. EPA. Research Triangle
Park, North Carolina. July, 1972.
6. Amdur, M.O. and M. Corn. The Irritant Potency of Zinc Ammonium Sulfate
of Different Particle Sizes. Am. Ind. Hyg. Assoc. J. 24:326-333,
July-August 1963.
7. Amdur, M.O. and D. Underbill. The Effect of Various Aerosols on the
Response of Guinea Pigs to Sulfur Dioxide. Arch. Environ. Health.
16:460-468, April 1968.
8. Sunshine, I. Handbook of Analytical Toxicology. The Chem. Rubber Co.,
Cleveland. 1969.
9. Emrich, G. H. Guidelines for Sanitary Landfills Ground Water and
Percolation. Compost Science, May, 1972. pp. 12-5.
10. Emrich, G.H., G.L. Merrit, R.C. Rhindress. Geocriteria for Solid
Waste Disposal Sites. Program for the 5th Annual Meeting, Northeastern
Section, Geological Society of America, 2(1). 1970. p. 17.
175
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APPENDIX H
COMMERCIAL PLANT DESIGN
H
-------
APPENDIX H
COMMERCIAL PLANT DESIGN
A commercial plant design study was made by Stone and Webster
Engineering Corporation (SWEC) based on direct proration of selected data
for the 50 MW plant described in detail in Appendix E. This appendix will
present summaries of those prorations and the bases for them.
177
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COMMERCIAL PLANT
DESIGN INDEX
1. Base Conditions and Assumptions - 200 MW Commercial Unit
2. Process Flow Diagram
3. Material Balance Flow Diagram
4. Economic Evaluation
A.I Investment Cost Summary
4.2 Operating Cost Summary
4.3 Evaluation
178
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1. BASE CONDITIONS AND ASSUMPTIONS - 200 MW COMMERCIAL UNIT
All aspects of operation of the 50 MW demonstration unit
will function as projected with no more than minor modifica-
tions , giving a sound basis for scale-up
200 MW is an appropriate commercial module size which can be
built as a single train. The gasifier and absorber reactors'
refractory arch distributor grids cannot be built at the ap-
proximately 12.2 m (40 ft) diameter required, so more expen-
sive stainless steel grids must be used; these have an under-
side temperature limitation of 594°C (1100°F), which requires
added provision for charging the gasifier with precalcined
stone for every start-up following a bed removal and requires
that such grids of this size are technically feasible.
Use of regenerator/dry sulfation system with sulfur removal
efficiency 90 percent, turndown ratio 4:1, same boiler effi-
ciency, waste solids suitable for landfill disposal .
Air/fuel ratio of 20 percent, limestone makeup design rate
stoichiometric, suitable limestone supply at 300 rail miles,
waste disposal area at 30 truck miles
Outdoor installation adjacent to boiler; no change in heat
conservation allowed for
Existing boiler using No. 6 fuel oil, with all storage and
handling facilities for oil, also rail-siding and car-
spotting equipment, available
Sixteen gas burners of the same capacity as the 50 MW plant
burners; also eight gasifier cyclones and eight Sturtevant
transfer systems for hot fines
Real estate provided at no charge to the unit
Clear site, with no overhead or underground obstructions;
179
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2 2
soil bearing pressure of 130 kg/m (3000 Ib/ft ) water table
below all foundations, and seismic zone zero
Limestone rate is one (1) times stoichiometric equal 4/1.5
or approximately 2.7 times demonstration plant rate.
2. PROCESS FLOW DIAGRAM
The flow diagram, 61A - Reaction System, has been included here
to indicate the general methods to be used for interconnecting the
equipment in a 200 MW oil gasification system. This flow diagram was
prepared for the 50 MW demonstration plant and carries pipeline sizes
and instrumentation for that smaller-sized unit. The 200 MW plant piping
and reactor vessels would be approximately double the sizes shown on the
flow diagram. One major exception to this would be the fuel-gas lines
and cyclones. Since there would be eight of these In parallel rather
than four as shown on the flow diagram, the size multiplier would be
about 1.4.
3. MATERIAL BALANCE FLOW DIAGRAM
The flow diagram, 61D - Base Case, has been included here to
indicate approximately one quarter of the flow rates required for the
200 MW commercial plant design. As stated in the preceding two sections
of this appendix, there are eight burners in the commercial plant rather
than four as indicated for the demonstration plant on diagram 61 D.
180
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4. ECONOMIC EVALUATION
4.1 Investment Cost Summary - 200 MW Unit (Single Train)
Item
PROCESS EQUIPMENT
A. Towers
B. Boilers, S- heaters
F. Process furnaces
G. General equipment
L. Reactors
M. Drums
Q. Storage tanks
P. Pumps (including drives)
R. Compressors (including drives)
S. Stacks
T. Heat exchanges
TOTAL PROCESS EQUIPMENT
PROCESS MATERIALS
C. Piping
D. Structures
E. Electrical
H. Buildings'*
J. Civil
K. Instruments
N. Insulation
N. Painting*9
Refractory lining
TOTAL PROCESS MATERIALS
TOTAL DIRECT COST (DM*+DL)
DISTRIBUTABLE ACCOUNTS
VI. Insurance (excluding all risk)
V2. Federal and state taxes
X. Temporary construction facilities
Y. Field office (including Insurance c
Z. Construction tools and equipment
0. Other distributable items
TOTAL DISTRIBUTABLE
SUBTOTAL-COST OF WORK
U. INDIRECT ACCOUNTS
Engineering
Design
Other headquarters office
Taxes-headquarters payroll
Overhead allowance
TOTAL INDIRECT (HEADQUARTERS OFFICE)
TOTAL, PRESENT DAY PRICES (BARE COST)
Fee, Escalation, and Contingency
Order of Accuracy
Illaterial
90,
520,
1,800,
1,300,
5,
1,350,
50,
1,720,
150,
6,985,
3,300,
900,
530,
250,
240,
350,
300,
320,
180,
6,370,
13,355,
md taxes)
($)| Labor (S)
000 2,500
000 15,000
000 110,000
000 72,000
000 1,000
000 30,000
000 3,000
000 120,000
000 5,000
000 358,500
000 1,700,000
000 320,000
000 500.000
000
000 450,000
000 150,000
000
000
000
000 3,120,000
000 3,478,500
Total ($)
92,500
535,000
1,910,000
1,372,000
6,000
1,380,000
53,000
1,840,000
155,000
7,343,500
5,000,000
1,220,000
1,030,000
250,000
690,000
500,000
300,000
320,000
180,000
9,490,000
16,833,500
2,500,000
2,500,000
19,333,500
2,250,000
2,250,000
21,583,500
Not Included
25Z
aFrom original SWEC table.
Subcontract.
185
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4.?.Operating Cost Summary
200 MW Commercial Units
A. Single Train Unit
Bare Cost (Accuracy 257.)
Direct Operating Costs
Power
Operating Labor
Limestone Supply
Solid Waste Disposal
Maintenance
Total
Indirect Operating Costs
Capital Charges
Total Direct & Indirect Costs
Fuel Operating Costs
Total Operating Costs
B. Four (4) Train Unit
Bare Cost (Accuracy 20%)
Direct Operating Costs
Power
Operating Labor
Limestone Supply
Solid Waste Disposal
Maintenance
Total
Indirect Operating Costs
Capital Charges
Total Direct & Indirect Costs
. Fuel Operating Costs
Total Operating Costs
$1,900,000.00
304,500.00
305,000.00
163,520.00
1,079,175.00
$1,900,000.00
609,000.00
305,000.00
163,520.00
1,454,000.00
$21,583,500.00
3,752,195.00
4,316,700.00
8,068,895.00
32,400,000.00
40,468,895.00
$29,080,000.00
4,431,520.00
5,816,000.00
10,247,520.00
32,400,000.00
42,647,520.00
186
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4.3 Evaluation
The preceding sections show estimated capital and operating
costs for the chemically active fluidized bed (CAFB) process at 200 MW
capacity. For comparative evaluation purposes, the 200 MW single train
costs of $21.6 million capital and $8 million/year operating costs convert
to $108/kw installed capacity and 5.7 mils/kWh operating cost at 7,000
hr/year. Relative to an oil firing rate of 8,000 bbl/d these convert to
$2,700/bbl/d capital and $3.42/bbl fuel adder costs. It must be noted
that investment costs are present day, while operating costs used the
specified NEES figures on a 1976 basis.
From the study by McGlamery and Torstrick prepared for EPA,
using their capital costs for a 200 MW oil-fired unit and SWEC adjustment
of their 500 MW operating costs, the regenerative wet scrubber with
elemental sulfur by-product shows $76/kw capital and 4.4 mils/kWh operating
cost. Similarly, the limestone scrubbing system shows $65/kw capital
and 3.0 mils/kWh operating cost. Maintenance costs for limestone scrubbers
are high. While CAFB maintenance costs are expected to be lower, no
meaningful projection can be made without large-scale operating data.
Utility requirements are compared in Table H-l for 200 MW oil-fired power
plants. Although the CAFB process consumes about the same auxiliary
2
power as a limestone slurry scrubber, no steam is used in the CAFB process
and an order of magnitude less process water is required.
Low-grade fuel oils (for example, vacuum bottoms) may be avail-
able at significantly lower costs than a No. 6 fuel oil, providing a cost
advantage for the CAFB process. With the present uncertain fuel situation
this potential advantage is difficult to assess, and the availability of
lower-grade fuel oils is unknown.
The cost and availability of hydrodesulfurized fuel oils is
also in a state highly sensitive to world political and economic conditions.
Discussions have been held with Esso Research Centre, Abingdon, England
(Esso) engineers and Westinghouse fuels personnel on hydrodesulfurization.
Costs for installation of a 60,000 bbl/d addition to an existing eastern
United States refinery to produce 0.3 percent sulfur oil with a January
187
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Table H-l
COMPARISON OF UTILITY REQUIREMENTS FOR
CAFB AND LIMESTONE SLURRY SCRUBBING
Annual requirement (7000 hr/yr) for 200 MW
power plant retrofit3
CAFB process
Limestone slurry
f> H
process0 »°
Steam, kg(lb)
0.45 x 108 (1.0 x 108)
(produced)
Process water, liters (gal) 3 x 10 (8 x 10 )
Electricity, kWh
% of power
plant output
4 x 107 - 5.5 x 107
3-4
0.9 x 108 (2 x 108)
(used)
3.3 x 108
4 x 107 - 5.5 x 107
3-4
The variability of the utility requirements for both processes is great.
The source of this energy is combustion of carbon deposited on the lime
and from the conversion of calcium sulfide to calcium sulfate in the
regenerator.
188
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1977 start-up are estimated to be around $3,000/bbl/d of capacity,
including off-sites.
Hydrodesulfurization costs of $600/bbl/d capital and $1.60/bbl
operating are derived from Oil & Gas Journal 8/18/74. The capital
indicated is for refinery installation and does not include capital for
the hydrogen unit, so it might well rise above $2,000/bbl/d for a grass
roots installation. Thus, residual desulfurization costs will vary widely,
depending upon the assumptions made with regard to crude source, distill-
ation cut-point, degree of desulfurization, unit capacity, source of
hydrogen, geographical location, and so on. Nine residual fuel oil
desulfurization processes were available for licensing in 1973. Only one
process, however, (H-Oil, Hydrocarbon Research Inc., 44,000 bbl/d plant
in Kuwait) is in use on a commercial scale with vacuum resid. Desulfur-
ization up to 75 percent is feasible. The use of high-metals feedstocks
is currently limited by practical considerations catalyst life, higher
costs, and so on.
None of the above figures can be considered rigorous, although
the comparison with McGlamery and Torstrick is probably as good as is
possible in present circumstances. It should be noted, however, they were
not prepared on the same basis or in the same period of time. It is
entirely possible that cost comparisons on the same basis and/or construction
and operation of the 50 MW CAFB unit would alter these cost comparisons.
Environmentally, the CAFB process compares favorably with lime-
stone and lime slurry scrubbers (see Appendix G). Sulfur removal capa-
bilities for the processes appear comparable. CAFB provides a considerable
reduction in nitrogen oxide emissions. The CAFB process consumes an order
of magnitude less process water than slurry scrubbing processes. Limestone
usage is comparable for the processes at about 1 mole of calcium per mole
of sulfur removed from the fuel, but the CAFB process could potentially
reduce this utilization to half that level by utilizing a regenerative
operation with sulfur recovery. The CAFB process also produces a dry
product with potential market value rather than a sludge which is difficult
to handle and requires large land areas for disposal ponds.
189
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APPENDIX I
INITIAL DESIGN STUDY
-------
APPENDIX I
INITIAL DESIGN STUDY
An initial design of the 50 MW chemically active fluldized bed
(CAFB) demonstration plant was carried out by Stone & Webster Engineering
Corporation (SWEC) under contract to Westinghouse Research Laboratories.
The basis for the initial design study was the utilization of conventional
catalytic cracking technology developed in the petroleum industry for
the solids transport system, the specification of equipment to provide
extensive system flexibility and reliability, and the inclusion of sulfur
recovery from the spent stone. The study identified the critical design
parameters and cost items in the plant. The initial design basis, process
flow diagrams, equipment design, and cost estimate are presented. These
data provided a basis for evaluation and identification of alternative
design and operating parameters to reduce the plant cost while maintaining
system flexibility and operability. A revised demonstration plant design
(Appendix D) was established on the basis of this evaluation.
GENERAL DESIGN BASIS FOR INITIAL STUDY
The design basis applied in the initial design study is similar
to the demonstration plant design basis eventually selected (Appendix D)
except for the following items.
Reaction System
The initial design basis included items relating specifically to
process flexibility and reliability:
Once-through operation capability in addition to regenerative
operation
2.44 meter (8 ft) bed depth capability
191
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About 50 percent excess capacity for stack-gas recycle
Extremely high pressure drops assumed in air and stack-gas
circuits; 96.5 kPa (14.0 psi) for control, line losses, bed
losses, and distributor losses and equipment design
Low gasifier plenum start-up temperature and lime storage
system for start-up material
Valves in the fuel-gas lines to isolate the fuel liner for
burnout during operation. Fuel-gas lines based on 18.3 m/s
(60 ft/sec) velocity
A solids transport system to circulate solids between the gas-
ifier and regenerator based on known catalytic cracker dense-
phase pneumatic transport technology. Required tall stand-
legs and relatively high-pressure transport gas.
High elevation of gasifier and regenerator vessels and cyclone
fines recycle to reactors by dense-phase standleg.
Stainless steel distributor support plate in gasifier
A stone quench drum used to cool waste solids before they enter
the stone processing system.
The other operating parameters (air/fuel ratio, fluidlzation
velocity, bed temperatures, and so forth) and design parameters (materials
of construction, filter types, operating procedures, and so on) are
essentially the same as those described in Appendix D with only minor
changes or additions.
Sulfur Recovery System
A commercial process to recover elemental sulfur from the
regenerator product gas was specified. The process (Allied Chemical) is
described in Appendix 0.
Stone Processing System
The stone processing system selected for the initial design
study was a slurry carbonation process which removed carbon dioxide (CO.)
192
-------
from the plant stack gas to contact the regenerator waste stone and produce
a dry calcium carbonate product. The hydrogen sulfide (H.S) produced in
the process is recycled to the sulfur recovery system.
Limestone Handling and Feeding System
The design basis for the limestone handling and feeding system
used in the initial study is four times the stoichiometric rate. The
limestone feed system is a dense-phase standleg system feeding limestone
to the stone circulation system.
Waste Stone Handling System
The design basis for the waste stone handling system is a four-
times-stoichiometric limestone feed rate.
Fuel Oil Handling System
The design basis for the fuel oil handling system used in the
initial design study is the same as that used in the eventual design
(Appendix D).
PROCESS DESCRIPTION
Flow diagrams for the reaction system and the stone processing
system (slurry carbonation) as utilized in the initial design study are
shown in Figures 1-1 and 1-2.
Reaction System Description
The same processing functions and processing steps are carried
out in the initial demonstration plant process shown in Figure 1-1 as in
the selected process described in Appendix B. The major process differences
are in the solids transport components. The initial design study uses
gravity flow to return cyclone fines to the reactors, while the selected
design uses pulsed flow fines return equipment.
193
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The major solids transport component, which circulates solids
between the gasifier and regenerator, incorporates conventional,
commercially demonstrated fluidized solids transport design principles
which have been successfully utilized in such well-known applications as
fluid coking units, many manometer-balanced fluid catalytic cracking
units, the chlorination reactors of titanium dioxide (TiCL) chloride plants,
and so on.
In the initial design study process diagram shown in Figure 1-1
the sulphided solids in the gasifier are withdrawn by means of a fluidized
withdrawal well which also serves to generate static pressure in the same
manner as the standpipe below the well. The direction of flow of the
dense-phase solids at the base of the standpipe is reversed by means of a
short-radius bend. The solids travel up a dense-phase lateral transport
line sloped at forty degrees above horizontal and are then vertically
lifted by means of recycled flue gas up into the regenerator bed; this
lift section of the line is termed a dense-phase riser since it can operate
over a very wide range of densities for example, from about 160.2 kg/m
3 33
(10 Ib/ft ) to about 561 kg/m (35 Ibs/ft ), depending upon the relative
quantity of lift gas and solids entering the line.
Slurry Carbonation System
The purpose of slurry carbonatlon is to eliminate calcium sulfide
(CaS) from the waste stone and to convert the hygroscopic quicklime or
alkaline slaked lime to the neutral and unreactive carbonate form (limestone),
thereby rendering the waste innocuous for disposal.
Calcium oxide (CaO) reacts rapidly with water to form calcium
hydroxide (Ca[OH]_ or slaked lime) with evolution of heat. A small amount
of water (for example, rain) on a mass of calcium oxide could theoretically
create a temperature of 650°C (1200°F), and has, in fact, been observed
to ignite paper. Given an excess of water, the heat is released by
vaporization. Calcium sulfide reacts with water to a very slight degree
forming calcium hydroxide and hydrogen sulfide. Calcium sulfide reacts
completely with carbonic acid (H2C03), occurring by solution of carbon
194
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dioxide (CO.) in water, and the products are calcium carbonate (CaCO,)
and hydrogen sulfide.
Referring to Figure 1-2, the carbonation reaction is carried out
in three stages, in drums M-5 A/B and M-7 A/B, having a residence time of
one hour each. Essentially all of the reaction occurs in the first stage,
M-7 A/B, and the other two stages are aimed at converting any residual
unreacted calcium sulfide, which could cause an odor problem or health
hazard in the vicinity of the stone dump.
Carbon dioxide for the reaction is supplied by flue gas drawn
from the baghouse, G-4, by blower R-7 and delivered to the carbonation drums.
Precise control of the gas flow is not required, and hand valves should
be set to divide about 5 percent of the gas to M-7 A/B, 10 percent to
M-5B, and 85 percent to M-5A. The only requirement is that the slurry
be kept saturated with carbon dioxide at all points. The design gas flow
contains 300 percent of the estimated carbon dioxide to be reacted. The
gas streams leave the drums at 60°C (140°F) to 82°C (180°F), carrying water
evaporated by the heat of reaction, and flow together to the inlet of
carbonation off-gas scrubber A-l.
Each drum has an agitator whose main function is to keep the
stone in suspension, with the secondary function of aiding gas dispersion.
The larger particles of stone are 3.2 mm (1/8 in) diameter and settle
very rapidly. The slurry flows by gravity from M-7 A/B to M-5B to M-5A,
with each outflow throttled by the level in the drum.
The slurry is pumped from M-5A to centrifuge G-12 by P-1A or B.
The rate is set by the level in M-2A, with override for excess or
deficiency to keep the centrifuge in operable range. Liquor from the
centrifuge flows by gravity to surge drum M-6,,and provision of return
flow from M-6 via pump P-2A or B to G-12 or M-5A should be included. The
main flow from P-2A or B is to the quench drum M-l under flow control as
needed for suitable slurry quality. M-6 provides water surge for the
system and has a level recorder to facilitate inventory regulation, as
discussed later.
199
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Stone cake from the centrifuge drops into a rotary drier, G-ll,
where it is indirectly heated by 1035 kPa (150 psi) steam coils. Moisture
from G-ll is vented to the existing stack cyclones which run at a negative
pressure of about 3480 Pa (-14 in) water gauge. Dried stone is conveyed
to the waste stone disposal railcar loading facilities. The stone should
be dried to a free-flowing consistency yet be damp enough to avoid dusting.
Centrifuge operation can be adjusted to aid in reaching this result.
Carbonation off-gas from the four drums M-5 A/B and M-7 A/B at
60 to 82°C (140 to 180°F) and steam from the quench drum M-l at 93 to 104°C
(200 to 220°F) enter the base of carbonation off-gas scrubber A-l. The
lower six trays are a cooling and condensing section to retain water in
the system. Heat is removed by circulating the bottoms liquid through
cooler T-4 and back to tray 6. It is to be expected that the gas and
steam will contain entrained slurry droplets, so the bottoms liquid will
be water containing some fine solids, as well as be saturated with carbon
dioxide plus dissolved hydrogen sulfide and oxygen. (The flue gas contains
about one percent oxygen from boiler excess air. It also contains about
40 ppm sulfur dioxide, but this will be absorbed by the stone in the slurry
as calcium sulfite). The upper four trays of A-l provide final cleanup
of any solid fines entrained from tray 6 liquid by washing with makeup water.
A minimum of 71.54 dm (5 gal) per minute of makeup water should be fed
to the top tray at all times for washing, and up to 715.4 dm (50 gal) per
minute may be fed if desired for water balance purposes.
The scrubbed off-gas from A-l will still contain 0.2 to 0.3
percent hydrogen sulfide and must not be released from the system. The
gas is sent to the suction of R-2 where it is combined with the recycle
flue gas to the gasifier, constituting about one-fourth of the total.
The hydrogen sulfide content is recovered by the gasifier solids, along
with the sulfur from the fuel.
The sulfur recovery system is described in Appendix 0. The
support systems (residual fuel oil handling, limestone handling, and
waste stone handling) are identical in description to the support systems
described in Appendix B.
200
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PROCESS EQUIPMENT
An equipment list indicating the equipment components considered
in the initial design study is presented in Table 1-1. The two major
reaction vessels, the gasifier and regenerator, are described in Figures 1-3
and 1-4. The two fluidized bed reactors are very similar in design to the
gasifier and regenerator described in Appendix D, except for the construction
of the distributor plates and the design of the solids Inlets and outlets.
Comparable drawings for the selected demonstration plant gasifier and
regenerator are included in Appendix E.
COST SUMMARY FOR THE INITIAL DESIGN STUDY
A breakdown of the costs estimated for the initial demonstration
plant study is shown in Table 1-2. The breakdown refers to the equipment
in the reaction, stone processing, waste stone handling, fuel oil handling,
and limestone handling systems. The sulfur recovery system is indicated
as a subcontract item. The sulfur recovery system cost is discussed in
Appendix 0. Costs are based on end-of-1974 dollars.
Table 1-3 lists the total installed costs of the six process
systems in the demonstration plant. Table 1-4 breaks down the cost of
the reaction system into the contributions of the major equipment items
(on a total installed-cost basis). It is evident from the tables (1-3
and 1-4) that the reaction system and the sulfur recovery system represent
the dominant costs in the Initial study demonstration plant, 77 percent
of the total plant cost. The major items in the reaction system
cost appear to be gas compression equipment at 30 percent of the reaction
system cost.
EVALUATION OF INITIAL DESIGN STUDY COSTS
The $10 million total cost for the Initial study demonstration
plant results mainly from the conservative nature of the design and from
the high cost of the sulfur recovery system. A number of cost reductions
can be made without reducing the demonstration plant performance or useful-
ness as a means of CAFB demonstration.
201
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Table 1-1
EQUIPMENT LIST FOR INITIAL DESIGN STUDY
Item number
Acct. A - Towers
A-l
Acct. B - Boilers
B-l
Acct. F - Heaters
F-l
F-2 A-D
Service
Acct.
C-l
NJ
O
S3
G-2
C - General Equipment
Gl-1
Cl-2
Cl-3
Cl-4
Cl-5
CU-6
Cl-7
G2-1
C2-2
r.2-3
G2-4
G2-5
G2-6
G2-7 A&B
G2-8 A&B
G2-9 A&B
G2-10
G2-11
C2-14
G-3 A6B
G-4
G-5
G-6
C-7
G-8 A-D
G-9
G-10
G-ll
C-12
G-13 A&B
CarbonatLon off-gas scrubber
Wa<:'i> bent boiler
Start-up air heater
Hot fuel-gas burners
Fresli limestone unloading system
Filter receiver air lock
Pressure system air blower
Baghouse
Vacuum system air blower
Bnghouse rotary air lock
Rallcar unloader
Filter/receiver
Waste stone transport and loading system
Waste stone dryer rotary air lock
Recycle flue gas rotary air lock
Filter/receiver
Vacuum system compressor
Filter/receiver rotary air lock
Pressure system compressor
Silo filter receivers
Silo rotary air locks
Eliminated
Eliminated
Eliminated
Stack cyclone rotary air lock
Limestone weight feeders
Recycle flue gas baghouse
Gaslfler air screen
Regenerator air screen
PTG filter
Gaslfler cyclones
Regenerator primary cyclone
Regenerator secondary cyclone
Waste stone drier
Slurry centrifuge
Slurry agitators
Item number
Service
G-14 A&B
G-15 A&B
G-16
G-16-1
G-16-2
G-16-3
Acct. L - Reactors
L-l
L-2
Acct. M - Drums
M-l
M-2 A&B
M-3
H-4
M-5 A&B
M-6
M-7 A&B
Acct. P - Pumps
P-l A&B
P-2 A&B
P-3 A&B
P-4 A&B
& Drivers
Acct.
Q-l
Acct.
R-l
R-2
R-3
R-4
R-5
R-6
R-7
Acct.
T-l
T-2
T-3
T-4
T-5
T-6
Q - Storage Tank
R - Compressors & Drivers
T - Heat Exchangers
Duplex oil filters
Slurry agitators
Lime charge system
Quicklime charge drum
Vibrating bin bottom
Rotary air lock
Gasiflpr
Regenerator
Stone quench drum
Stone surge drums
Waste heat boiler steam drum
PTG water separator
Slurry carbonatlon drums
Liquor surge drum
Slurry carbonatlon drums
Centrifuge feed pump and spare
Centrifuge liquor pump and spare
Scrubber circulation pump and spare
Oil transfer pump and spare8
High-sulfur oil storage tank
Gasifler air blower3
Recycle flue gas blower
PTG blower3
Flue gas booster fan
SO2 booster fan
Regenerator air blower0 a
Carbonatlon flue gas blower
Regenerator off-gas air exchanger
Regenerator recycle gas cooler
PTG cooler
Scrubber circulation cooler
Oil storage tank suction heater
Oil transfer line heater
Motor drive
-------
Table 1-2
COST BREAKDOWN FOR THE INITIAL STUDY DEMONSTRATION PLANT
Description
Towers
Boilers, etc.
Process Furnaces
General Equipment
Reactors
Drums (tanks)
Storage Tanks
Pumps
Compressors
Heat Exchangers
Fuel-Gas Pipes & Valves
Solids Transfer Systems
TOTAL EQUIPMENT IN PLACE
Piping
Structures
Electrical
Buildings
Civil
Instruments
Insulation & Plant
TOTAL PROCESS MATERIALS
TOTAL DIRECT COSTS
DISTRIBUTED COSTS (18.73% of
SUBTOTAL (COST OF WORK)
INDIRECT COSTS (17. 4%)
Material
($1000)
29.30
31.00
125.00
611.80
281.70
60.20
178.00
25.02
572.26
65.80
194.42
47.73
2,222.23
433.65
38.48
88.47
10.00
4.22
150.00
9.50
909.18
3,131.41
direct costs)
Labor
($1000)
1.10
0.80
1.30
60.81
2.70
23.39
2.10
37.10
1.60
62.80
26.81
220.51
365.65
25.90
93.68
1.50
7.80
47.12
49.50
920.18
1,140.69
CONTINGENCY AND ESCALATION (18.9%)
FEE (1.18% of cost of work)
SULFUR RECOVERY SYSTEM (SUBCONTRACT)
TOTAL COST (END OF 1974)
Total
($1000)
30.40
' 31.80
126.30
672.61
281.70
62.90
201.39
27.12
609.36
67.40
257.22
74.54
2,442.74
618.74
64.38
182.15
11.50
12.02
197.12
59.00
1,829.36
4,272.10
800.00
5,072.10
880.00
957.90
60.00
3,000.00
9,970.00
203
-------
Table 1-3
SUMMARY OF MAJOR SYSTEM COSTS IN
THE INITIAL STUDY DEMONSTRATION PLANT
System Total installed cost (1000$)
Reaction A,511.00
Stone Processing 1,102.00
(slurry carbonation)
Sulfur Recovery 3,000.00
Waste Stone Handling 358.00
Fuel Oil Handling 825.00
Limestone Handling 174.00
TOTAL PLANT 9,970.00
204
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PAGE NOT
AVAILABLE
DIGITALLY
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PAGE NOT
AVAILABLE
DIGITALLY
-------
Table 1-4
SUMMARY OF MAJOR COMPONENT COSTS IN
THE INITIAL STUDY REACTION SYSTEM
Component Total installed cost (1000$)
Gasifier 580.00
Gasifier Cyclones 469.00
Gasifier Air Blower 508.50
Regenerator 186.00
Regenerator Cyclones 190.00
Regenerator Blower 85.50
Recycle Flue Gas Blower 763.00
Fuel-Gas Pipes 392.00
Fuel-Gas Valves 306.00
Burners 192.00
Stone Circulation 198.00
Other Items 641.00
TOTAL REACTION SYSTEM 4,511.00
209
-------
The costs for the three support systems seem to be based on
reasonable specifications and are of a conventional nature. The stone
processing system and the sulfur recovery system represent more than $4
million. The cost associated with the stone processing system and sulfur
recovery system depend most critically on the choice of the specific
processes to be applied. Many processing options have been identified
which reduce this cost drastically and are discussed in Appendix K. The
choice of slurry carbonation for stone processing and the Allied Chemical
process for sulfur recovery yield costs which are not feasible for this
application.
The initial design study indicates a number of options which can
be applied to reduce greatly the cost of the reaction system. These are in
the areas of reduced gas handling pressure drops, utilization of the
simplified solids transport process developed in the Esso pilot plant
operation, elimination of valves in the fuel-gas lines, minimization of
fuel-gas line length, and reduced excess blower capacity. Cost projections
were made to assess the potential of modifying the initial design study
basis. An example of a set of modifications is summarized in Table 1-5, and
indicates a potential to reduce the reaction system cost to about $3 million.
These and other modifications are discussed in more detail in Appendix J.
CONCLUSIONS
The initial design study has been used to guide the selection of
a CAFB demonstration plant design basis. Specifically, it has indicated
that
The commercial sulfur recovery process applied in the study
results in a significant cost penalty for this application.
Conventional technology utilized by the petroleum industry
for solids circulation is considered excessively expensive
in this application. Techniques developed in the pilot plant
operation result In a more economical demonstration plant.
The critical cost items in the reaction system are the gas
compression components and the fuel gas handling components,
210
-------
Table 1-5
DEMONSTRATION PLANT MODIFIED COSTS PROJECTION
(Basis - Steam for Temperature Control and Burnout)
Major equipment items
Gasifier Vessel
Gasifier Cyclones
Gasifier Air Blower
Regenerator Vessel
Regenerator Cyclones
Regenerator Blower
Recycle Flue Gas Blower
Initial study
installed
costs, $
580,000
469,000
508,000
186,000
190,000
85,000
763,000
Projection of
modified
costs, $
566,000
461,000
330,000
186,000
190,000
55,000
34,000
Modification to initial study design basis
3.05 m (10 ft) reduction in elevation;
modified internals
9.75 m (32 ft) reduction in elevation;
pneumatic fines return
55 kPa (8 psi) reduction in air circuit
pressure drop
Reduced elevation
Reduced elevation
55 kPa (8 psi) reduction in pressure drop
Elimination of flue gas recycle for purge
Fuel-Gas Pipes and Valves 698,000
Burners 192,000
Stone Circulation 198,000
Other Items 641,000
TOTAL 4,511,000
250,000
192,000
123,000
458,000
2,845,000
and transport gas; 55 kPa (8 psi)
reduction in pressure drop
Elimination of valves; 24.4 m (80ft) reduc-
tion in 1070 mm (42 in) lines; 27.4 m
(90 ft) reduction in 760 mm (30 in) lines
No modification
Utilize Esso pulsed transport system
Reduced baghouse capacity; reduced flue gas
booster fan capacity; eliminate SO- gas
heat exchanger; increase waste heat
boiler; eliminate S0?-gas blower.
-------
and these may be greatly reduced in cost by proper design.
A stone processing system should be utilized in the
demonstration plant which permits the elimination of a
sulfur recovery system (Appendix K).
212
-------
APPENDIX J
EVALUATION AND SELECTION OF REACTION
SYSTEM PROCESS AND DESIGN OPTIONS
J
-------
APPENDIX J
EVALUATION AND SELECTION OF REACTION
SYSTEM PROCESS AND DESIGN OPTIONS
Numerous process and design options are available for the
chemically active fluidized bed (CAFB) demonstration plant reaction
system. These have been identified and the demonstration plant design
basis selected from them (summarized in Appendix D). The procedure
and rationale involved in the selection of the options of major impor-
tance for the reaction system are summarized to orovide further insisht
into the nature of the demonstration plant.
BASIS OF SELECTION
The selection of the design basis for the demonstration plant
from the identified process and design options rests on a number of
general factors:
Maximum demonstration plant operability based on pilot
plant results
Flexibility to gain maximum information from the plant
Applicability of demonstration plant to commercial plant
considerations
Environmental performance and goal
Economics of demonstration plant and economic targets for
commercial plant.
These factors have been utilized in the design basis selection, in
addition to the requirement that the demonstration plant function as a
commercial power plant upon completion of the demonstration plant
program. It is clear that the demonstration plant will not represent
the ultimate CAFB plant in terms of optimum design or economics, but
213
-------
it is the goal of the demonstration plant to identify this ultimate
plant and its potential.
GENERAL PROCESS OPTIONS FOR THE REACTION SYSTEM
Table J-l lists some general process options under the cate-
gories of the reaction system concept, the general technology basis
applied to carry out the reaction system concept, and the commercial
market to which the demonstration plant should apply. The selection of
the general process configuration for the reaction system is indicated in
Table J-l, along with the rationale involved in the selection. The general
conclusion drawn from these considerations is that the demonstration
plant should be designed to utilize the same general concept being
investigated in the pilot plant (regenerative, limestone sorbent, atmos-
pheric pressure, hot fuel gas, and so forth), should utilize much of the
technology applied and developed in the pilot plant operation, based
largely on the results of the initial design study (Appendix I), and
should be concerned with the problems faced with a boiler retrofit.
REACTION SYSTEM OPERATING OPTIONS
Table J-2 lists the operating options considered and selected
for the demonstration plant reaction system. These are listed under the
categories of basic operating variables, temperature control options,
turndown and start-up options, burnout options, and flexibility and
reliability options. Those selected correspond to the design basis
summarized in Appendix D. They provide system operability and flexi-
bility while satisfying economic feasibility targets.
EQUIPMENT OPTIONS FOR THE REACTION SYSTEM
Options related to the design of equipment in the reaction
system are summarized in Table J-3. The options identified and selected
for the major equipment components in the reaction system are presented
214
-------
Table J-l
GENERAL PROCESS OPTIONS FOR THE REACTION SYSTEM
General
consideration
Options
Selection for
demonstration plant
Basis for
selection
Process
Concept
N>
M
Ol
Technology
Basis
Commercial
Market
General concept investigated by
pilot plant (regenerative process
at atmospheric pressure producing
a hot fuel gas. Air-blown regener-
ator; limestone sorbent, single-
stage gasifier; etc.)
Advanced concepts (once-through
process; fuel-gas cooling; increased
pressure; alternate regeneration
processes; alternate sorbents,
staged gasifier)
Technology applied and developed by
pilot plant (dealing mainly with
solids transport systems)
Conventional technology utilized
in petroleum industry
New plant considerations (optimized
total power plants utility or
industrial boiler)
Boiler retrofit considerations
(boiler modifications, boiler derate,
available space, gasifier location,
boiler type)
General concept
investigated by pilot
plant
Maximum
operability
Pilot plant
technology
Economics (see
Appendix I)
Utility boiler
retrofit
Market
information
-------
Table J-2
REACTION SYSTEM OPERATING OPTIONS
General
consideration
Options
Selection for
demonstration plant
Basis for
selection
Basic Gasifler operating options
Operating « Limestone size (fine particles
Variables Or -3000 microns)
Air/fuel ratio (minimum value
or large demonstrated value)
Fluldization velocity (low Listed in Appendix Maximum opera-
velocity or high velocity; D - Design Basis bility, flexi-
entrained bed reactor) bility, environ-
Bed depth (shallow, demonstrated, mental consid-
bed depth or deep bed capability) erations
Sulfur removal efficiency (meet
existing standards or demonstrate
full potential)
Others - bed temperature, lime-
stone makeup rate, etc.
Regenerator operating options
Fluidization velocity
Bed depth
Bed temperature
Temperature Gasifier
Control Stack-gas recycle
Steam or water injection Stack-gas recycle
Heat transfer surface with provision for Maximum opera-
steam and water bility and
Regenerator flexibility
Stone circulation rate Stone circulation
Heat transfer surface rate
-------
Table J-2 (continued)
General
consideration
Options
Selection for
demonstration plant
Basis for
selection
Gasifier
Turndown
Gasifier
Start-no
N>
Fuel line
Burnout
Multiple units (2)
High design velocity (^3.7 ra/sec)
Control of air/fuel ratio to
maintain operable fluidization
velocity
High start-up gas temperature
(^800°C/1472°F) and initial bed
load of limestone
Lower start-up gas temperature
and initial bed load of lime
Isolation of individual fuel-
gas lines with valves and
steam/air burnout
Backburning out fuel-gas
lines during operation
Burnout only during shutdown
of plant
Control of air/fuel
ratio
High start-up tempera-
ture with limestone as
initial bed load
Backburning of fuel-
gas lines
Economic consid-
erations and
operability
Operability
and economics
Operability and
economics
-------
Table J-3
EQUIPMENT DESIGN OPTIONS FOR THE REACTION SYSTEM
Component
Options
Selected for
demonstration plant
Basis for
selection
GASIFIER
REGENERATOR
00
SOLIDS TRANSPORT
Gasifier-to-
Regenerator
Circulation
Regenerator-to-
Stone Processing
Cyclone Fines
Recycle
Number of modules-single or multiple
Type of unit-fluldized or entrained
Freeboard height-minimized or
conservative
Vessel elevation
Distributor design-alloy steel
or refractory
Number of modules
Other options identical to gasifier
Regenerator integrated with gasifier
or separate units
Technology-Esso pulsed transfer,
other conventional dense-phase or
dilute-phase transport systems
Esso pulsed transfer
Gravity flow standleg
Other dilute.or dense-phase
transport systems
Gravity flow standleg sturtevant
pulsed flow system
Single module
Fluidized
Conservative
Minimized
Refractory
Single
Same options selected
as with gasifier
Separate units
Esso pulsed transport
Esso pulsed transport
Sturtevant pulsed
flow system
Economic
consideration
Operability
Economics and
operability
Economics (mini-
mum fuel-gas line
length)
Economics and
operability
Commercial
applicability
Same basis as
with gasifier
Demonstrated
configuration
Economics and
operability
Operability and
economics
Economics (permits
minimized fuel-gas
line length)
-------
Table J-3 (continued)
Component
Options
Selected for
demonstration plant
Basis for
selection
10
M
VO
PARTICULATE
REMOVAL
Gasifier
Cyclones
Regenerator
Cyclones
FUEL-GAS
HANDLING AND
COMBUSTION
Burners
Fuel-Gas
Lines
AIR-AND STACK-
GAS COMPRESSION
FOR GASIFIER
Number of units
Elevation
Number of units-single high
efficiency cyclone; cyclones
in series
Cyclone drainage-recycle fines to
regenerator or remove to stone
processing system
Number
Number-single with manifold or
one per burner
Design velocity
Length
Number-individual compressors for
air and stack gas or single unit
Design for pressure losses
Single cyclone per
fuel- gas line
Minimized
Operability
Economic (mini-
mized fuel-gas
line length)
Two cyclones in series Operability
Primary cyclone returns
course material to re-
generator; secondary
cyclone removes fines
to stone processing
Four
One line per burner
46 m/sec (150 ft/sec)
Minimized
Single unit for air and
stack gas
Minimize control and
line losses and margin
for bed surging
Operability
Commercial
applicability
Operability
Economics and
operability
Economics
Economics
Economics
-------
in the table. The economic conclusions drawn from the initial design
study (Appendix I) have played a large part in the selection of the
design options; at the same time care has been taken to fulfill goals
related to plant operability, flexibility, environmental performance, and
commercial applicability.
220
-------
APPENDIX K
SPENT STONE PROCESSING OPTIONS
K
-------
APPENDIX K
SPENT STONE PROCESSING OPTIONS
BACKGROUND
General
In all desulfurlzatlon processes using a solid for sulfur removal,
a stream of by-product sorbent is produced because of gradual loss of
chemical reactivity and attrition. The main variable defining the quantity
is the choice of system: sorbent regeneration or once-through. Thus, an
environmentally acceptable method of disposing of this sorbent by-product
was needed in order to complete the process design. Alternatively, a
market for the material had to be identified and developed.
To gain perspective on what processing would be required, several
preliminary studies were made:
A brief survey of what others were doing in the field of solid
industrial waste disposal
A review of solubility data for calcium compounds
A review of existing regulations for the proposed site of
the demonstration plant
Preparation of a logic diagram to outline spent stone
processing alternatives
Preparation of a preliminary material balance to estimate
probable quantities of material to be processed
e Check of the size of possible markets .
These steps are described in the following sections.
Industry Practices
A hierarchy of disposal methods for various industrial solid
wastes exists. In order of increasing environmental acceptability, they
are dumping, ponding, landfill, and stockpile. Dumping means the site
221
-------
is not expected to be usable for any other purpose for an indefinite
period. Ponding is a temporary measure used for slurries. Depending on
the settling characteristics of the wastes, the solids would eventually
be transferred to a dumping site or a landfill area. Landfill implies
structural strength and stability so that the site can be used for
construction or some other purpose after a reasonable period. To stock-
pile means to produce some material for which a future market is foreseen.
One industrial waste which presented a disposal problem similar
to the current one was fly ash from coal. Much work has been done in
recent years to utilize this material, both here and abroad, as reported
1 2
in recent symposia. ' Although domestically 80 percent of the fly ash
was still being dumped, productive use was increasing:
* Fill material
- Sanitary landfill
- Mine subsidence control
Agriculture
- Land reclamation
Soil conditioning
- Plant nutrients
Construction
Road bases
- Aggregate
Concrete products
- Dams
Mineral source
- Alumina
- Iron oxide
- Vanadium
Miscellaneous
- Ftreproofing
- Cenospheres
- Antiskid surface for winter roads.
222
-------
There had been some attempts to utilize calcium sulfate (CaSO.) as road-
building material, but the scope of such efforts was narrow. The Federal
Highway Administration reportedly was pursuing research in this area.
During the last six years, however, no new tonnage uses for fly ash had
been conceived. Further, utilization of fly ash by previously devised
methods had increased only modestly, and landfill remained the most
common disposal method.
From this review came the idea of using by-product stone as a
substitute for road-building materials and, possibly, in combination with
fly ash, as new ceramic materials. It appeared likely, however, that
landfill would be a temporary disposal method until a new market actually
emerged. Some consideration of the probable environmental impact of
leaching was therefore in order.
Solubility Data
Readily available data on the solubility of calcium compounds
was compiled and is presented in Table K-l in order of decreasing solu-
bility. It appeared that carbonates, phosphates, sulfites, silicates, and
sulfates were potential end products of calcium that might have minimal
environmental impact. Both the carbonate and the sulfite had acid forms
(bicarbonate and bisulfite) which were the most soluble of the compounds
checked.
Tricalcium phosphate (Ca.JPO,]-) was eliminated from further consid-
eration because an inexpensive source of the phosphate anion was lacking.
Phosphoric acid (H,PO.) was priced at $165/metric ton ($150/ton).
3
Flint and Wells reported solubility studies in which silicon dioxide
(SiO.) was dissolved in lime water using concentrations of calcium oxide (CaO)
up to 1.2 g/dm3. The composition of the solid phase in equilibrium with
the solution after one month was determined by the difference in the solution
composition itself, which was found to range from CaO*SiO. to 3CaO*SiO?.
Other data showed that calcium oxide was leached from anhydrous calcium sili-
cates in amounts decreasing in the order 3CaO*Si02, @-2CaO*S10., y-ZCaO'SiO.,
3CaO>2S10., and CaO-SiO.. Saturation was reached within an hour for the last
three compounds at a calcium oxide level of about one-tenth that of calcium
223
-------
Table K-l
SOLUBILITY OF CALCIUM COMPOUNDS IN WATER
Formula nat
Ca(HC03)2
Ca(HS03)2
Ca(HS)2 6H20
Form or
ural mineral Solubility, ppm
166,000
sol.
70,000-250,000
CaS04 1/2 H20 a (stable form) 3,000-8,000
8 (unstable form) 6,600
CaS04 III Soluble
anhydrite 7,000
II Insoluble anhydrite 2,100-3,000
Ca(OH)2
CaO + 2Si02
2CaO + Si02
1,300-1,400
236a
212b
CaO Si02 feeudowollastonite 95
212C
CaS Oldhamite 120-210d
CaSO. 2H20
Ca3(P04)2
CaC03 Calcite
43
25
13-14
Aragonite 12-15
Temperature, °C(°F)
20 (68)
20 (68)
20 (68)
30 (86)
20 (68)
20 (68)
0 (32)
30 (86)
30 (86)
17 (63)
15 (59)
18 (64)
25 (77)
25 (77)
Hydrolysis studies; 164 ppm free Si09 also in solution at equilibrium with solid
containing 0.103 moles CaO/mole S
Hydrolysis studies; 83 ppm free CaO also in solution at equilibrium with solid
containing 1.13 moles CaO/mole Si02.
°Hydrolysis studies; solid phase 1.07 moles CaO/mole
Decomposes .
224
-------
oxide alone. The tricalcium silicate reached saturation after eight hours
at a solute level like that of pure calcium oxide. While these levels of
solubility were encouragingly low, the actual leaching loss from such
materials in the environment was expected to be even lower because the
above data were obtained with a large excess of water. When hydrolysis studies
were done with pastes of silicates, the mixtures hardened; and analysis of
samples that had cured for one to two years showed only a fraction
of the equilibrium hydrolysis.
Among the complex hydrated silicates known are:
Ettringite 3 CaO-Al (ys CaS(y32 H20
3 CaO-Al0-CaS0- 12-13 H0
Tobennorite CaO'SiO.'H.O
Hillebrandite 2 CaO-SiCy^O
Okenite CaO-2 SiO -2 H.O.
A * *
Work reported by Minnick at the First Fly Ash Utilization Symposium in
1967 concluded that toberraorite and possibly ettringite might be chiefly
responsible for the cementing action observed in lime fly ash blends.
This supported the view that fly ash blends with spent stone might prove
to have minimal environmental impact.
Overall, processes which led to carbonate, silicate, and sulfate
forms of calcium were considered as potential candidates. Sulfite end
forms were not considered desirable in view of the experience reported
by others in wet lime scrubbing.
Existing Regulations
To include all the essential elements in a disposal process for
the demonstration plant, it was necessary to know what existing regulations
required. The following information was excerpted from the Rules and
Regulations of the Department of Health of the State of Rhode Island and
5-9
Providence Plantations .
Refuse disposal
- Refuse includes solid industrial waste.
- Refuse shall be disposed of only by sanitary landfill. . .or
other means approved by the Director of Health.
225
-------
- Refuse shall be compacted and covered daily, or more often
if necessary, with a layer of cover material at least six
inches deep after compaction. Final cover shall be at
least two feet deep.
- Cover material shall, when compacted, provide a tight seal,
and shall not crack excessively when dry.
- Refuse or leachings therefrom shall not cause or contribute
to pollution of any source of public water supply or of
any of the waters of the state.
- The deposit area shall be controlled by supervision, fencing,
signs, or equally effective means.
Particulate emissions
- Maximum particulate emissions for sources other than
combustion of liquid or gaseous fuels must be no more than
the applicable tabular value.
- Particulate matter shall be handled, transported, and
stored to prevent it from becoming airborne.
- No contaminant shall be emitted which, either alone or in
connection with other emissions, may be injurious to any
life, cause damage to property or inconvenience to property
owners, or cause a disagreeable or unnatural odor, or
obscure visibility, or interfere in any way with the
enjoyment of life and property.
Sulfur content of fuels
- After October 1, 1971, unless the Director declares a
shortage of low-sulfur fuel exists, no one shall store,
sell, or use fuel containing more than 1 wt % sulfur on a
dry basis.
- Such high-sulfur fuel may be approved for use by the
Director when combined with an approved stack-gas cleaning
process, provided the total sulfur emissions from the
stack do not exceed 1.1 Ibs S0? equivalent/1 million Btu
gross heat input.
226
-------
Maximum
emission rate
g/sec Ibs/hr
Fuel sources
Sulfur as S02 71.2 565
Nitrogen oxides as NO. 19.4 154
Particulates 6.6 52
Other process sources
Particulates 0.75 6
The Federal regulation on sulfur emissions from liquid fossil fuels is
more stringent 0.344 kg S02/MJ (0.8 Ibs S02/million Btu ). It applies to
steam generators with heat input greater than 264 GJ/hr (250 million Btu/hr ).
Particulate emissions from any single source must also be less than No. 1
Ringelmann.
Material Balance Information
As a reference point in the process development work, a prelim-
inary material balance was calculated as shown in Table K-2 for a 50 MW
plant using limestone on both a regenerative and a once-through basis.
These figures assume that the basic process consists of capturing sulfur
from the oil fuel as calcium sulfide in a gasifier and driving it
off as sulfur dioxide (S0«) in a regenerator or oxidizing it to calcium
sulfate in an oxidizer. Sulfur dioxide can be processed to sulfur or
sulfuric acid (H-SO,). Gypsum production is based on converting the
by-product stone with the aid of additional purchased sulfuric acid.
Market Size
From the 1970 Minerals Yearbook the demand for various
construction materials in millions of short tons/year in the New England
area was that shown in Table K-3.
227
-------
Table K-2
ESTIMATE OF BY-PRODUCT STONE PRODUCTION
Operating mode
1
Units
I Regenerative | Once-through
IS)
Is*
00
Net power output
Heat rate
Heat to power section
Gasification thermal efficiency
Heat fired
Heating value of fuel
Fuel fired
Sulfur content
Sulfur input
Limestone input
Sulfur removal in gasifier
Production rates
Gasifier stone, net
Regenerator stone, net
Oxidizer stone, net
Sulfur plant @ 95% recovery
Sulfur
Sulfuric acid, 100%
Gypsum plant @ 100% recovery
Hemihydrate
Sulfuric acid required
MW
MJ/kWh (Btu/kWh)
GJ/hr (106 Btu/hr)
%
GJ/hr (106 Btu/hr)
UJ/kg (Btu/lb)
kg/hr (Ibs/hr)
Wt. %
kg moles/hr (Ib moles/hr)
kg moles/hr (Ib moles/hr)
kg/hr @ 95% (Ibs/hr) 95%
metric tons/day (tons/day)
kg/hr (Ibs/hr)
metric tons/day (tons/day)
kg/hr (Ibs/hr)
metric tons/day (tons/day)
kg/hr (Ibs/hr)
metric tons/day
metric tons/day (LT/day)
metric tons/day (LT/day)
kg/hr (Ibs/hr)
metric tons/day (tons/day)
metric tons/day (LT/day)
50
10.98 (10,400)
548 (520)
95.0
577 (547)
41.9 (18,000)
13,800 (30,400)
3.0
12.9 (28.4)
12.9 (28.4)
1357 (2,992)
32.6 (35.9)
95.0
1,122 (2,474)
26.9 (29.7)
791 (1,743)
18.9 (20.8)
8.96 (8.82)
27.3 (26.9)
3,070 (4,570)
49.7 (54.8)
2.96 (2.91)
50
10.98 (10,400)
548 (520)
95.0
577 (547)
41.9 (18,000)
13,800 (30,400)
3.0
12.9 (28.4)
38.7 (85.2)
4070 (8,977)
97.7 (107.7)
95.0
2,570 (5,660)
61.6 (67.9)
3,350 (7,387)
80.4 (88.6)
5,810 (12,811)
139.2 (153.7)
62.2 (61.2)
-------
Table K-3
DEMAND FOR CONSTRUCTION MATERIAL
State
Rhode Island
Connecticut
Massachusetts
Vermont
New Hampshire
Maine
TOTAL
Sand &
Gravel
2.4
6.8
17.9
4.0
6.5
13.0
3oT?
Crushed
stone
NAa
8.4
8.0
1.5
NA
1.0
IsT?
Lime
NA
0.064
NA
0.064
Cement
0.19
0.83
1.40
1.08
0.17
0.21
3788
Gypsum
Imported
NA
Imported
a,
Not available
Thus, even on a once-through basis, the production of by-product stone
amounts to only 0.06 percent of the sand and gravel demand in the New
England states, and 0.17 percent of the crushed stone market. With
suitable values of crushing strength and bulk density plus chemical
inertness, the by-product stone was potentially a candidate for some
portion of these markets.
While a breakdown of the gypsum demand was not available, the
price structure appeared to be:
Crude gypsum - $2.50-5.00/0.9072 Mg (short ton)
Calcined gypsum - $14-19/0.9072 Mg (short ton)
Plaster - $25-60/0.9072 Mg (short ton)
It appeared that the by-product stone might be utilized in one of the
above ways.
Logic Diagram
A logic diagram (Figure K-l) was then prepared showing alterna-
tive solids processing schemes with the ultimate disposal being
Landfill
Stockpiling sulfur or calcium sulfide
229
-------
Dug. 724B633
to
W
O
Figure K -1 - Logic diagram for disposal of spent limestone
-------
Marketing sulfur, sulfuric acid, or some other as yet
undefined product.
Two attractive alternatives were:
Partial oxidation followed by landfill
Stockpiling calcium sulfide and blending with fly ash.
The former, which has been previously studied, would preserve most of the
sulfur in an available form while protecting the environment by coating
the calcium sulfide with a calcium sulfate-impervious shell. The second
would take advantage of the pozzolanic activity of typical fly ash to
tie up the sulfur in complex cement-type structures.
PROCESS ALTERNATIVES
The block diagrams presented below show 15 alternatives for
disposing of spent stone from the fluidized bed oil gasification process.
Five other methods are covered in the descriptions following, although no
flow sheets are shown. These diagrams were devised to include the essential
elements in each process so that priorities for further development work
could be assigned on a qualitative basis. Figure K-2 shows the relationship
of these processes, grouping them into regenerative and nonregenerative
(once-through) types. For convenience, the processes are designated by a
letter code.
First letter - R, regenerative; N, nonregenerative
Second letter - D, dry process; W, wet process
Third letter - S, sulfur recovery unit included;
N, no sulfur recovery unit required.
The processes are described in the sequence shown in Figure K-2,beginning
with nonregenerative. Experimental data have been obtained on selected
spent stone processing systems to support the assessment of the selected
options. Reference to the experimental support data is indicated by
noting the appropriate appendix in parentheses.
Option NDN-1 Direct Disposal
This is the simplest approach to disposal. No flow sheet is
shown, although there would be a cooling and possibly a grinding step. The
231
-------
Dug. 725BI83
NJ
LO
tsJ
Direct
Disposal
NDN-1
CaO
CaS
*
Air
*
Dry
Sul fat ion
RDN-1
ci.
I
Dry
NDN-2
CaO
CdSOi,
Oxidation Plus
Carbonat ion
MDN-3
Silica
IION-lt
1
(CaO)n. S,02
CaS
I
Deodburnmg
NDN-S
CoO
1
CaCO,
CaSOjJ
Water
Quench
*
Wet
Carbonat ion
NWS-I
CaC(
Wet
Sulfation
NWS-2
t *
r
3 CaS°«,
Direct
Disposal
RDS-1
1
CoO
CaSO,,
I
Dry
Oxidation
ROS-2
Silica
Sintering
ROS-ll
Oxidation Plus
Carbonat ion
RDS-3
(CaO)n, S,02
CaSOi,
I
Dcadburn i ng
ROS-5
\
CoO
Water
Quench
' 1 1
Wet
Carbonat ion
RWS-I
Wet Lin
Sulfation Scrub!
RWS-2 RWN-
1
Stem
Rccarbonation
NDS-6
CaCO,
S02
I
Steam
Recorbonati
ROS-6
1
-*- \
* CaCO,
img *
1
on
Steam
Rcgenerat
NDS-7
.
Sul
Recove
)0
fur
ry U
on
11 1
Su 1 fur
Recovery Uni t
Steam
Regeneration
ROS-7
C
1
lO
S
CaCO,
CoSOi,
Sulfur
or
Sul func Acid
Sul Func Acid
or
Sulfur
..Gas to Stack
CaCO,
CaSO,
Figure K-2-Oil gasification: spent stone disposal processes
-------
calcium oxide/calcium sulfide product Is preferably sold to some user who
needs lime and can tolerate the sulfide content. This is not thought of
as a general solution because it depends on local market conditions.
Option NDN-2 Dry Oxidation (Flow sheet not shown) (Appendix 0 and Reference 11)
Adding an air oxidation step could convert all of the calcium
sulfide to calcium sulfate so that the product could be marketed as an
impure lime. Since calcium sulfate is used as a retardant in cement manu-
facture, it is possible that this component would not be detrimental. The
oxidation may require grinding the limestone below 150 microns (100 mesh).
This could be done internally by a jet attriter or externally by a
micronizer-type mill. The flow sheet would be represented by the first
half of NDN-3.
Option NDN-3 Oxidation Plus Carbonation (Figure K-3) (Appendix 0)
In the event the calcium oxide/calcium sulfate from the oxidation
step is not marketable, further treatment would be required before it
could be disposed of as landfill. This flow sheet shows carbonation at
about 760°C (1400°F) with carbon dioxide recovered from calcination of
the fresh limestone. A waste heat boiler may be required in the carbonation
process to remove the heat of reaction of carbon dioxide and calcium oxide.
There is some reason to believe that successful operation will
require an excess of carbon dioxide. If this is true, carbon dioxide would
be recovered from flue gas, although there is an advantage in external
calcination of limestone in that this carbon dioxide does not become a
diluent in the product gas from the gasifier.
Option NDN-A Silica Sintering (Figure K-4)
The presence of free calcium oxide suggests the possibility of
processing the spent stone to cement-type substances by combining it with
silica. The silicon dioxide source could be sand, fly ash, or material such
as slag. The spent stone from the gasifier would be processed like cement.
233
-------
Dwg. 725B150
N>
U>
Clean Fuel Gas
CaO
CaS04
Carbonator
760°C
Waste Heat
Boiler
Cyclone
--
-------
N>
CO
Ul
(\Aake-up
Clean Fuel Gas
Flue Gas
to Stack
*-Market
Dry Solids
to Disposal
Figure K - 4-Oil gasification, nonregenerative mode-Option NDN-4: silica sintering
-------
Feed materials would be ground to less than 150 microns (100 mesh), blended
and sintered at temperatures perhaps as high as 1500°C (2732°F). Optimistically,
calcium sulfide would replace calcium oxide in complex silicate linkages.
An alternative would be to process a portion of the spent stone to
calcium oxide-silicon dioxide ([CaO] 'SiO^-type compounds and then blend
this back to coat all the calcium sulfide so that, on hydration with water,
the calcium sulfide becomes embedded in a cement matrix.
The spent limestone could be ground in a micron!zer using a
circuit of fuel gas. After the circuit has been established, only makeup
for losses would be required.
A further variation is based on work done by the Coal Research
Bureau, School of Mines, West Virginia University under a contract with
EPA in 1971. This showed that the natural pozzolanlc activity of
limestone-modified fly ash could be utilized to make calcium silicate i
bricks. The composition was adjusted by adding silica sand and
calcium oxide. Bricks were pressed and autoclaved for eight hours at
1310 kPa (190 psig) and 185°C (365°F).
The variation here is to make any necessary adjustment in
composition, hydrate the mixture, and let it develop whatever degree of
cement linkages it will. Basically, this is thought of as an ambient-to-
moderate temperature process to create enough linkages to permit disposal
as landfill.
Option NDN-5 Deadburning (No flow sheet shown) (Appendix 0)
It was known that materials like lime can be processed at high
temperatures to reduce chemical activity. This phenomenon was studied briefly
for applicability to the spent calcium oxide for both once-through and
regenerative operation. The stone was sintered at temperatures between
1200 and 1700°C (2192 and 3092°F) to render the stone inactive. A nonregen-
erative system without sulfur recovery would require inactivating the free
calcium oxide under conditions such that the sulfur would be retained.
Primary concerns for this approach are the retention of the sulfur and
the environmental activity of the processed stone.
236
-------
Option NWS-1 Wet Carbonation (Figure K-5) (Appendix M)
Spent stone from the gasifier is quenched with water and subjected
to a three-stage carbonation with carbon dioxide from external calcination
of limestone. The treated stone is dewatered, filtered, and dried. The
filtrate is used as quench water in combination with water from the off-gas
scrubber. Gases from the first-stage carbonation and the stone quench are
scrubbed in an off-gas scrubber to remove particulates. Makeup water can
be added here or in the solids separation steps. The scrubbed gas contains
hydrogen sulfide and is sent to a sulfur recovery unit such as a Glaus
or a Stretford plant.
The product is sent to disposal as calcium carbonate plus inerts
and some calcium sulfate.
Option NWS-2 Wet Sulfation (Figure K-6) (Appendix N)
An alternative to carbonation is to treat the quenched stone
with dilute sulfuric acid to form calcium sulfate from both calcium oxide and
calcium sulfide, releasing hydrogen sulfide which is expected to remain
in solution, although some odor will be apparent. The acidification step,
at least, will have to be conducted in a closed vessel. A three-stage counter-
current decantation followed by filtration will produce calcium sulfate which
may be
Marketed
Disposed of as landfill
Stockpiled for future use as a sulfur source.
Sour water is stripped of hydrogen sulfide and recycled as quench water or
as part of the wash water, depending on the residual acid content. Ideally,
the process would use somewhat less than stoichiometric sulfuric acid so
that the final product would be slightly aklaline. Hydrogen sulfide would
be converted to sulfur or sulfuric acid either on site or off site. The
net sulfuric acid requirement for the process could also be derived from
effluents from other industrial processes.
237
-------
Dwg. 725BI8?
oo
Clean Fuel Gas
Fuel r
Limestone
UUbt
1
Gasi fier
870°C
Stone
Quench
Ki I
760°(
n
1
Fresh Water
N.,H,S
"*~ f Off-Gas * *
t
1st Stage 2nd Stage 3rd Sta<
1 II *
N2> CO,
'
Dust Waste Heat
"* Removal * Boiler Dewaterer
Gas to
(* Stack
Recovery ., en
Unit H2S(V
on
je
f. IVicr ri Drv Stone
to Disposal
Figure K-5 -Oil gasification, nonregenerative mode - Option NWS-1: wet carbonation
-------
Dwg. 725B15I
10
w
vO
1
1 1 1
Limestone ^ rT^jfjPr
870°C
CaO
Pa^
\sQj
< |
Stone ~"!
Quench
1
i
Sulf uric. Acid ». Acidification
I
uust
-" Removal
us i
2 Sulfur
J Stripper » Kecovery
1 . llnif
i unit
* ~fc
H2S
H90
*
\A/ac i \A/^ch lA/ach «
1st Stage * 2nd Stage * 3rd Stage
41 4 f 41
Filter «
L_
Gas to
Stack
^ Sulfur to
Outside
Ar-iH Dlinl
AGIO riant
Fresh
Water
CaSO
4 Wet Solids
to Disposal
Figure K -6- Oil gasification, nonregenerative mode-Option NW S-2: wet sulfation
-------
Option NDS-6 Steam Recarbonation (Figure K-7)
Another chemical reaction is available for converting the spent
stone to an environmentally acceptable form. The stone from the gasifier
is first pulverized in a micronizer-type mill using compressed flue gas
as a grinding medium. Some carbonation may occur in this step. The stone
is transferred to a reactor operating at about 705°C (1301°F) and 1520 kPa
(15 atm) where it is contacted with flue gas and steam to produce calcium
carbonate and release hydrogen sulfide.
It is possible to regard this as a route to regeneration different
from that which corresponds to the R-series of options. One advantage would
be that the limestone would not be exposed to the higher temperature levels
(1070°C [1958°F]) and presumably would be less subject to deactivation
(dead-burning).
Option NDS-7 Steam Regeneration (No flow sheet shown)
A simplication of Option NDS-6 is to produce calcium oxide rather
than calcium sulfide by treating the gasifier stone with steam alone. The
hydrogen sulfide released would be processed to sulfur. It appears that
temperatures of 1200 to 1300°C (2192 to 2372°F) would be required to produce
an equilibrium mixture containing even three percent hydrogen sulfide.
If the calcium oxide had to be discarded, this would present
problems comparable to those of Option NDN-2 or NDN-1. It is also possible
that the calcium oxide could be recycled, which would correspond to a
method of regeneration alternative to air oxidation.
Option RDN-1 Dry Sulfation (Figure K-8) (Appendix L)
Sulfur dioxide released during regeneration is cooled to about
870°C (1598°F) in a waste heat boiler and recontacted with the portion of
regenerated stone that would have been discarded. In the presence of air,
calcium sulfate reforms. The treated stone is cooled by incoming air and
may be marketed as a gypsum or discarded. Since the reason for discarding
240
-------
Dwg. 72SB 188
Clean Fuel Gas
Gas to
*" Stack
Sulfur or
Sulfur! c
Acid
Dry Sol ids
to Disposal
Fines to
Disposal
FigureK-7-Oij gasification, nonregenerative mode - Option NDS-6: steam recarbonation
-------
ro
js
M
Clean Fuel Gas
Limestone
Air
Dust
Removal
Gasifier
870°C
CaO
CaS
Regenerator
1070°C
Absorber
870°C
I
CaO
(CaS04,CaS)
SO,
Waste Heat
Boiler
CaSO,
(CaO, CaS)
Figure K-8-oil gasification, regenerative mode-Option R DN-l: dry sulfation
Dry Solids
to Disposal
Fines to
Disposal
-------
stone in these regenerative processes is to maintain activity and particle
size distribution, theoretically the makeup rate could be less than raole-
for-mole on sulfur. Option RDN-1 would require an increase to one-for-one.
Option RDS-1 Direct Disposal (Figure K-9)
This is the simplest version of the regenerative processes as far
as stone processing is concerned. A water quench prior to disposal may
release some hydrogen sulfide, but this is expected to be small enough to
permit venting to a stack. It is possible that a market can be found for
the slaked lime product.
Option RDS-2 Dry Oxidation (Figure K-10) (Reference 13)
Spent stone from the regenerator is oxidized in a micronizer-
type mill to calcium sulfate/calcium oxide. This can be offered as a
lime substitute.
Option RDS-3 Oxidation Plus Carbonation (Figure K-ll) (Appendix 0)
In the event that a market cannot be found for the product from
RDS-2, a further process step could be added in which carbon dioxide from
flue gas is used to recarbonate the stone. The data, however, do not
confirm this.
Option RDS-4 Silica Sintering (Figure K-12)
This is similar to Option NDN-4 except that the product contains
much less sulfur since sulfur dioxide is released in the regeneration step
for conversion to acid or sulfur.
Option RDS-5 Dead-burning (Figure K-13) (Appendix 0)
As in Option NDN-5, the concept is to subject the stone to be
discarded to temperatures of about 1500°C (2732°F) to inactivate the
calcium oxide. Because most of the sulfur has been released in the
regeneration, there is less likelihood of having significant amounts of
243
-------
Dug. 725BI91
Clean Fuel Gas
Gas to
Stack
Su If uric
Acid or
Sulfur
Wet Stone
to Disposal
Fines to
Disposal
Figure K - 9-oil gasification, regenerative mode - Option RDS-1: direct disposal
-------
Dwg 72 SB I 53
N>
*
Ol
Clean Fuel Gas
-
Waste Heat
Boiler
Sulfu
Recove
Unit
r
ry
/
i
^ <
c
<
CaO
(CaS04)
Gas to
*" Stack
Sulfuric Acid
or
Sulfur
Dry Solids
"*" to Market
Figure K -10-Oil gasification, regenerative mode-Option R D S-2: dry oxidation
Fines to
Disposal
-------
Clean Fuel Gas
Limestone
Air
Flue Gas
Gasifier
8?0°C
CaO
CaS
Regenerator
1070°C
Dust
Removal
H , SO
2 2
-------
Dug. 725BI89
_ Clean Fuel Gas
Gas to
'Slack
Market
Dry Solids
1 to
Disposal
Figure K - I2~0il gasification, regenerative mode - Option RDS-4: silica sintering
-------
ro
*
oo
Clean Fuel Gas
Dust
Removal
bdbiner
870°C
1 CaO
1 CaS
Air _
1070°C
Fuel
CaO
(CaS04,CaS)
Cyclone
Dead-Burner
1500°C
I
Flue Gas
CaO
(CaS04)
- 775B152
Waste Heat
Boiler
auiiur
Recovery
Gas to
Stack
Sulfuric Acid
or
Sulfur
_^ Dry Solids
to Disposal
Fines to
Disposal
Figure K-13-Oil gasification, regenerative mode-Option RDS-5: dead-burning
-------
sulfur dioxide in the off-gas from the deadburner. In any case, the gas
can be recycled as shown to the flue gas cooler and back to the gasifier,
or perhaps to the stack.
Option RWS-1 Wet Carbonation (Figure K-14) (Appendix M)
This represents one primary option investigated for the
demonstration plant. A modification may be made in doing the carbonation
with three parallel streams of carbon dioxide-rich flue gas rather than
countercurrent as shown. The final gases are scrubbed to remove parti-
culates before being recycled- to the gasifier to reabsorb the small
amount of hydrogen sulfide released by carbonation.
Option RWS-2 Wet Sulfation (Figure K-15) (Appendix N)
This is a regenerative version of Option NWS-2. If hydrogen
sulfide is released, the amount may be small enough to permit venting to
a stack, recycling to the gasifiers, or sending to a sulfur recovery unit.
Option RWN-1 Lime Scrubbing (Figure K-16)
This is another way of designing the spent stone disposal process
to be internally sufficient in raw materials. Quenched regenerated stone
is used to produce a lime slurry. Calcium sulfate can possibly be
separated off. The lime slurry is used to scrub sulfur dioxide from the
regenerator gases after they have been cooled in a waste heat boiler.
The product is a mixture of calcium sulfite (CaSO.), calcium bisulfite
(Ca[HSOJ2), and calcium sulfate (CaSO.). After drying, the material is
disposed of as landfill.
Option RDS-6 Steam Recarbonation (Figure K-17)
This is similar to Option NDS-6. One difference is that the
recarbonation can be carried out at 101.3 kPa (1 atm) by raising the temperature
to 750°C (1382°F). The objective is conversion of the calcium oxide rather than
recovery of hydrogen sulfide. The concentration and amount of hydrogen sulfide
249
-------
Dwg 7258149
Ul
O
Clean Fuel Gas
CaO
(CaS04,CaS)
Flue Gas
Flue Gas
Cooler
Bag
Filter
so2
NJOJ
Waste Heat
Boiler
515°C
Sulfur
Recovery
Stone
Gas to
' Stack
. Sulfuric Acid
or
Sulfur
Carbonation
3rd
j
Dewa
I
Filt
Stage
t
terpr -
er
Carbonation
2nd Stage
1 t
I »
guencn
I
_+ Carbonation
1st Stage
1
CaCC-3
«-
(Ca(OH)2,CaS04.(
" r
_», Off Gas
Scrubber
:aS)
Fresh
Water
^ Dry Stone
to Disposal
Fines to
Disposal
Figure K - 14-Oil gasification, regenerative mode-Option RW S-l: wet carbonation
-------
Dug. 7256)86
_Clean Fuel Gas
*i
Gasi f ier
870°C
t-o
In
1
'2S
1 0
2
Wash
1st Stage
N2. S02
(0,)
(CaO)
Gas to
Stack
Sul fur or
Sul furic
Acid
Fresh
Water
Flue Gas
Flue Gas
Cooler
Bag
Filter
1
Wet Solids
to Disposal
Fines to
Disposal
Figure K - 15 Oil gasification, regenerative mode - Option RWS-2: wet sulfation
-------
725BI&*
Clean Fuel Gas
Gas to
Stack
Dry Solids
to Disposal
Fines to
Disposal
Figure K -16-Oil gasification, regenerative mode - Option RWN-1: lime scrubbing
-------
in the off-gas is expected to be low enough to permit venting directly to
a stack. A second difference is that sulfur is released in the regeneration
as sulfur dioxide, which is probably best processed to sulfuric acid rather
than sulfur.
Option RDS-7 Steam Regeneration (No flow sheet shown)
A simplification of Option RDS-6 is to regenerate calcium oxide
from calcium sulfide rather than calcium carbonate by treatment with steam
alone. There would still be a sulfur recovery plant because of the air
regenerator. It is expected that the.residual sulfide in the regenerator
store would be low enough that the gases from the steam treatment could
be vented to a stack. Disposal of the calcium oxide would present the
same problems as Option RDS-1 or RDS-2.
INITIAL SELECTION OF CANDIDATE PROCESSES
The above set of 20 alternatives was next subjected to a quali-
tative review to identify a smaller number of processes for more detailed
study. In general, the objective of these evaluations was to
Minimize the air/water pollution potential
Minimize the investment and operating costs
In the absence of a market, minimize the production rate of
the stone
Maximize the possibility of alternative utilization of the
stone, even though this might occur sometime in the future
Favor dry processes over wet ones
Favor processes for which the technology is the farthest
advanced.
Discussions with utility companies confirmed the expectation
that they were Interested in as simple a process as possible. The generation
of electricity itself was beset with enough problems; they were very
reluctant to add to them those of a chemical plant.
A direct disposal process would be simple and should have low
capital investment. RDS-1, a regenerative process shown on Figure K-9,
253
-------
Dwg. 72SB185
to
Ui
JLIean Fuel Gas
Gas to
Stack
Sulfuric Acid
or Sulfur
Dry Sol ids
to Disposal
Fines to
Disposal
Figure K -17-oil gasification, regenerative mode - Option RDS-6: steam recarbonation
-------
was selected in preference to NDN-1 (figure not shown) because the material
to be discarded contained a relatively small amount of sulfur as calcium
sulfate rather than a large amount as calcium sulfide. There were
potentially more ways of disposing of the former than the latter.
Dry oxidation can convert the sulfide to the sulfate, but there
would be substantial amounts of calcium oxide remaining. Initial labor-
atory data confirmed the difficulty of obtaining complete oxidation of
the sulfide because of the rapid formation of a shell of sulfate. The
effluent stone could be ground to expose unoxidized calcium sulfide, but
the data suggested that the particle size would have to be quite small
(perhaps less than 74 urn) in order to achieve complete conversion of
the sulfide. NDN-2 (no flow sheet shown) was favored over RDS-2 (Figure K-10)
because the latter would require a sulfur recovery plant. The high temper-
ature of regeneration (1070°C/1958°F) would also tend to inactivate the
remaining calcium oxide and presumably make dry carbonation less successful.
Of the four options in this group, NDN-3 (Figure K-3) was thus selected
for further consideration. The feasibility was demonstrated by tests on
sulfided limestone which showed that 70 percent recarbonation of -74 ym
stone was obtained at 720 to 7608C (1328 to 1400°F) and 101.3 kPa (1 atm) by
contacting with 16.5 percent carbon dioxide plus 3 percent water in
nitrogen for five hours. Data are presented in Appendix 0.
Dry sulfation, as shown in RDN-1 (Figure K-8), showed promise as
a means of avoiding the necessity of sulfur recovery facilities by recon-
tacting the by-product stone with the regenerator off-gas in the presence
of air to tie up the sulfur as calcium sulfate. The feasibility was
demonstrated with BCR 1359 limestone in which 85 percent conversion of
calcium oxide to calcium sulfate was obtained by contacting CAFB regenerator
stone with 1/2 mole % of sulfur dioxide and 4 mole % oxygen in nitrogen
at 870°C (1598°F). The stone was less than -74 ym. With 500 pm stone,
the amount of sulfation has not been more than about 40 percent. Support
data are presented in Appendix L.
Silica sintering, in which the effluent stone would be blended
with fly ash or other siliceous material and sintered at a suitable
255
-------
temperature, showed promise in view of all the work that had been done by
others on fly-ash utilization. NDN-4 (Figure K-4) was favored over RDS-4
(Figure K-12) because the latter would require both a sintering section
and a sulfur recovery section.
Deadburnlng, as represented by NDN-5 (no flow sheet) and RDS-5
(Figure K-13), was Intended to utilize the knowledge that exposure to
high temperature had the effect of making the calcium oxide less reactive
chemically. While it is conceivable that suitable process conditions
could be found both to inactivate the lime and to retain the sulfur as
calcium sulfate, the regenerative version was Judged more immediately
practicable. Feasibility was demonstrated by heating CAFB regenerator
stone for two hours at 1450°C (2642°F). Contacting the cooled product
with water in the ratio of five parts water and one part stone produced
a temperature rise of only 0.8eC (32.5°F)/gram of stone versus 4.3°C (40°F)
for uncalcined stone. Dsta, are presented in Appendix 0.
Steam recarbonation as in NDS-6 (Figure K-7) would release the
sulfur as hydrogen sulfide, which is more amenable to sulfur recovery than
is sulfur dioxide. It would require elevated pressure, however, to achieve
a reasonable hydrogen sulfide concentration. The regenerative version can
be conducted at 101.3 kPa (1 atm) by raising the temperature to 750°C (1382°F)
A simplification is obtained by omitting the recarbonation and
merely regenerating calcium oxide. Both NDS-7 and RDS-7 (no flow sheets)
would be market dependent since disposal of the resulting calcium oxide
in a landfill operation is not believed acceptable. There are the further
questions of slow rates of reaction and incomplete conversion, especially
for the nonregenerative version. Further consideration of these four
options was therefore deferred.
With the addition of a water quench, three wet processes were
formulated. Lime scrubbing, RWN-1 (Figure K-16), would recontact the
sulfur dioxide from the regenerator with a slurry made from the spent
stone. It was not considered further because it was expected that
direct stack-gas scrubbing with lime would prove less costly. Wet carbon-
ation resulted in one of the more insoluble compounds of calcium. Laboratory
256
-------
data (Appendix M) also supported the view that essentially complete con-
version of the calcium oxide and residual calcium sulfide to calcium car-
bonate could be obtained. With three hours' contact time at 85°C(185°F),
90 percent recarbonation of CAFB regenerator stone was obtained for 300
to 3000 pro particles using 14 percent carbon dioxide in nitrogen at
two times stoichiometric and a water/solid ratio of 40:1. An alternative
process, wet sulfation, visualized treating the spent stone with sulfuric
acid to make gypsum (Appendix N). Treating CAFB gasifier stone with
100 percent excess of 1:1 sulfuric acid resulted in a weight gain equi-
valent to a 72 percent conversion of the calcium oxide/ calcium sulfide
to calcium sulfate. RWS-1 (Figure K-14) was favored over NWS-1(Figure K-5)
because of the smaller lime throughput, hence smaller recarbonation
facilities required, and because the CAFB pilot plant work was far more
complete for the regenerative mode. NWS-1 was favored over the regenerative
version for the initial selection because the stone would not have been exposed
exposed to 1070°C (1958°F), only 870°C (1598°F), and because the sulfur
would be released as hydrogen sulfide rather than as sulfur dioxide.
The initial seven processes were then examined in more detail,
still qualitatively, as shown in Table K-4. Wet carbonation, RWS-1, had
the largest number of process steps (17) versus 10 to 12 for the others.
Dead-burning required the highest operating temperature:1500°C (2732°F)
versus 1070°C (1958°F) for the other regenerative processes and 870°C
(1598°F) for the nonregenerative candidates. Only one, direct disposal,
was completely market dependent. Dead-burning was not, but on the other
hand was not considered to produce a marketable product. The other five
processes produced potentially marketable products. Both wet sulfation
and silica sintering would require additional raw materials besides the
limestone. Overall, it was decided to eliminate silica sintering from
further consideration as a prime candidate because of the length of the
development program required. Work would continue in this direction,
however, because of the long-term promise of this approach.
Using the material balance shown in Table K-2, a brief look at
the gross realization from the wet sulfation process was taken as shown
in Table K-5 as three times stoichiometric and, instead of a regenerator,
257
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an oxidizer is used to convert calcium sulfide to calcium sulfate. The
balance of the calcium oxide is converted to gypsum by treating with purchased
sulfuric acid. In the regenerative version, sulfur dioxide driven off
in the regenerator is recovered as acid in a sulfuric acid plant and used
to convert the calcium oxide in the by-product stone to gypsum; makeup
sulfuric acid is purchased.
Because of the large excess of lime, the once-through option
shows a net cost of $2/0.9072 Mg (ton) of gypsum or $312/day added cost over
merely disposing of oxidizer stone as landfill. A gross realization is possi-
ble in the regenerative version because most of the acid can be generated
from the sulfur recovered. Using a service factor of 0.9, the annual gain is
$172,800. With 16 percent capital charges, 50 percent taxes, and even a
20-year payout time, however, this gain will support an investment of only
$655,000. Out of this must come facilities for a sulfuric acid plant, a
gypsum plant, and gypsum handling. Gypsum in this calculation was valued
at $14/0.9072 Mg (ton) of hemihydrate. There are, of course, other costs
associated with operation of a landfill. Gypsum may be worth $19/0.9072 Mg
(ton), and perhaps the product can be upgraded to plaster uses.
Regarding market potential, the gypsum production would be
16,330 Mg (18,000 short tons) of hemihydrate per year, using a 90 percent
service factor. The Minerals Yearbook shows that domestic production has
been constant at about 9,073-1000 Mg (10,000,000 short tons)/year over the
period 1967 to 1970. Thus, it would take 110 to 250 MW plants to match the
demand for gypsum. This shows that the market is large enough for at least
some locations to use this option for stone utilization.
Laboratory data showed that long residence times in the presence
of excess acid were likely to be required for conversions significantly
better than 70 percent. This was interpreted to mean that a significant
effort would be required to develop optimum conditions. It is known that
in the phosphoric acid industry production of by-product gypsum has been
hampered by filtration problems and the difficulty of washing essentially all
the phosphoric acid out the gypsum. Progress has been made, as in the
Kellogg-Lopker process, wherein by use of carefully controlled con-
258
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Table K-5
REALIZATION FROM GYPSUM PRODUCTION
End use
Landfill
$/day
Gypsum
$/day
Once-through
Transportation @ $4/0.9072 Mg (ton)
Sulfuric acid @ $36/1.016 Mg (LT)
Value of product
Gross cost
Incremental cost, $/day
$/0.9072 Mg (ton) gypsum
354
NAa
354
0
354
615
2203
2818
2152
666
312
2.03C
Regenerative
Transportation @ $4/0.9072 Mg (ton)
Sulfuric acid @ $36/1.016 Mg (LT)
Value of product
Realization (cost)
Incremental gain, $/day
$/0.9072 Mg (ton) gypsum
83
NA
83
0
(83)
219
105
324
767
443
526
9.60
NA - not applicable.
Dollars per 0.9072 Mg (ton).
261
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dltions some phosphate rocks do not even require grinding before being
fed to the process.
It was concluded that wet sulfation was only marginally attractive
at this stage of development, and further consideration of it was deferred.
ECONOMIC COMPARISON OF CANDIDATE PROCESSES
Having narrowed down the number of processes to be examined more
closely, a preliminary heat and material balance was made for one case,
wet carbonation, and detailed equipment lists and a cost estimate were
obtained from SWEC. Cost estimates were then made for the other four
processes on a differential basis. The flow sheet of Figure K-14 gives
the essential details of RWS-1. Table K-6 shows the breakdown of the
Investment required for each case.
Observations on Equipment Costs
The base equipment cost for RWS-1 was estimated by SWEC as
$2,222,000. Of this, about 27 percent is common for all five processes.
This would include the fuel oil handling equipment, the limestone unloading
system, the by-product stone handling system, and the gasifier air blower.
Seven items account for another 44 percent - the gasifier, the regenerator,
the two sets of cyclones, the recycle gas blower, the fuel-gas valves,
and the large fuel-gas lines. On a systems basis, the breakdown is as
follows:
Gasifier, ex. common items 27.5 %
Regenerator 11.3
Recycle gas and PTG* systems 15.4
Carbonation system 17.2
Common & miscellaneous 28.6
100.0 %
Thus, only 17 percent of the base equipment cost is attributed to the
method of processing the by-product stone.
Three of the four alternatives show a lower equipment cost;
*purge transport gas
262
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the fourth is about the same.
The lowest cost process is, as might be expected, direct
disposal: RDS-1. At this point, a commercial use for the effluent stone
had not been identified, so the ultimate disposition was thought of as
dumping or landfill. Work to this date by Consol showed that calcium
sulfide does become inactive, that is, resistant to carbon dioxide plus
steam regeneration even at 760°C (1400°F). On the basis that the small
amount of calcium sulfide that might remain on the stone will be inactive,
a simple water quench to cool and hydrate the calcium prior to disposal
is used. It is of interest that even with this simplest approach the
saving is only about 16 percent ($372,000). On an installed basis
including indlrects, and so on, the absolute savings would be large,
perhaps three times the above figure, although the percent of saving
would probably be about the same.
The four alternatives to RWS-1 share a common advantage over
wet carbonation: a saving of $380,000 (17%) for deleting the carbonation
system. This is offset by the additional facilities required in each case.
The RWS-1 cost is low because a particulate removal system for
the sulfur dioxide stream and the sulfur plants was not in the estimate.
Dead-burning (RDS-5) also offers significant savings: $229,000
or about 10 percent. The major questions are whether the stone remains
inert over a long period of time and how inert is it when made on a
tonnage basis.
Oxidation plus carbonation, NDN-3, is a stand-off with wet
carbonation. However, the data thus far show only 81 percent oxidation
of calcium sulfate and 67 percent carbonation of the calcium oxide. A
significant cost is the oxldlzer air blower needed to operate the
pulverizer ($201,000). One advantage over RWS-1 results from the higher
limestone rate: less recycle flue gas for heat balance on the gaslfier
is required.
Dry sulfation appears slightly less expensive than RWS-1. It
has an advantage over NDN-3 in that the sulfur dioxide available nearly
matches the calcium oxide, as opposed to the roughly 200 percent excess
for the once-through process.
265
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Selection of By-Product Stone Process
Despite the cost advantage of direct disposal, it was felt that
a more general solution would be a process yielding a product that
could be disposed of in applications such as landfill. Both wet carbon-
ation and dead-burning were encumbered by the need for a sulfur recovery
plant. This would add $3,000,000 investment to the approximately
$7,000,000 cost for the wet carbonation system. The latter figure was
built up from the $2,222,000 bare equipment cost by the addition of other
direct costs, indirect costs, contingency, escalation, and fee.
The investment for dry sulfation was somewhat lower than that for
dry oxidation plus carbonation. Since the technology was more fully
developed for dry sulfation, it emerged as the prime candidate for a spent
stone process.
STATUS OF TECHNOLOGY OF PROCESS SELECTED
Although the initial laboratory work has established the technical
feasibility of dry sulfation, further tests are needed to provide the
balance of definitive design information. The particle size range is one
prime variable. Smaller particles would be utilized to a higher degree,
but would require more grinding energy and would be more difficult to
recover from the effluent gas streams. They would also be more difficult
to fluidize. Gas residence time, percent excess air in the absorber, fluid-
ization requirements, solids transport requirements, particle collection
efficiencies are areas that need additional supporting data. Another
possible problem is partial sintering due to eutectic formation between calcium
sulfate and calcium sulfide and possibly aided by limestone impurities.
After the demonstration plant has been operated successfully, it
may be possible to discharge the absorber off-gas directly to the stack.
It is for the present being recycled to the gaslfier and guarantees meeting
the sulfur dioxide emission regulations.
266
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OPTIONS FOR THE DRY SULFATION PROCESS
Size Reduction
The evidence thus far is that the spent stone from the regenerator
must be reduced in size to below 74 ym to achieve maximum utilization
of the calcium oxide. The design includes size reduction in an external
jet pulverizer using a stream of compressed air as the grinding energy
source. The manufacturer's literature claims a higher energy efficiency
than in grinding in a fluidized bed. This is explained by observing that
the available mechanical energy of the jet is distributed over more
particles in the fluidized bed, so fewer of them accelerate to a velocity
high enough for fracture by interparticle collisions. It is possible
that the jet energy can be more efficiently utilized in the fluidized
bed, but the jet mill
Separates the process functions of bed fluidization and
size reduction and
Automatically removes the product by elutriation when it has
reached the desired size.
In the fluidized bed, some of the product could be subject to reentrainment
in the jet and be overground. Inquiries made of manufacturers of con-
ventional grinding equipment revealed that the stone would most likely
have to be cooled to essentially ambient temperature first. Jet pulver-
ization has been used at 482°C (900°F) and higher.
Grinding Medium
The design uses compressed ambient air as the grinding medium to
avoid other problems. It is possible that recycle gas could be used for
grinding if the particulate load in it was sufficiently low. One advantage
would be heat conservation; a second is avoidance of heat release due to
oxidation in the pulverizer. An alternative is to use the regenerator
off-gas, after cleaning and cooling, as the grinding medium so that some
absorption and oxidation could occur during the grinding. With favorable
results, the absorber vessel could be eliminated in favor of two or more
pulverizers in parallel*, contingent on process temperature requirements.
267
-------
To avoid the problem of how to introduce the dusty air from the
pulverizer cyclone into the absorber, the oxidizing air is provided as a
separate stream. Costs probably could be reduced if this second air
stream could be eliminated or at least reduced.
The residual nitrogen from these air streams enters the product
fuel gas as a diluent although it permits some reduction in the recycle
gas rate needed for heat balance on the gasifier.
Heat Conservation
One design criterion is to minimize the number of process steps
and especially heat exchange steps. A significant variable Is the amount
of calcium sulfide and carbon left on the effluent stone from the regen-
erator. Oxidation of calcium sulfide to calcium sulfate releases more
heat than absorption of sulfur dioxide on calcium oxide and oxidation of
calcium sulfite to calcium sulfate. Any carbon left simply adds to the
cooling duty required in subsequent processing.
The air rate used for grinding in the pulverizer is sufficient
to absorb the heat of oxidation of calcium sulfide and cool the stone to
500°C (932°F). This is based on 0.5 wt % sulfur as calcium sulfide on
the regenerator stone, which is equivalent to 1.0 mole % of the
calcium. It was assumed that all the carbon was burned off in passing
through the regenerator.
If the heat duty proves to be greater, there is a trade off
between the amount of air used and its supply pressure.
The air supplied to the absorber is five to ten times theoretical.
This helps the oxidation of sulfur dioxide to sulfur and also permits
elimination of a cooling coil for heat balance on the absorber. For larger
plants, it may prove economical to reduce the air rate and put in a
second waste heat boiler.
268
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REFERENCES
1. Faber, J.H., W.E. Eckard, J.D. Spencer. Ash Utilization Proceedings:
Third International Ash Utilization Symposium. March 13-14, 1973.
Bureau of Mine Information Circular 864.
2. Faber, J.H., J.F. Capp, J.D. Spencer. Proceedings - Fly Ash Utilization
Symposium, March 14-16, 1967. Bureau of Mines Information circular
8348.
3. Flint, E.P., and L.S. Wells, J. Res. NBS 12:751, 1934. RP 687; and
in R.H. Bogue The Chemistry of Portland Cement, 2nd ed. New York,
Reinhold Publishing Corporation, 1955. Ch. 22, p. 28.
4. Minnick, L.J. in Faber, J.H., J.P. Capp, J.D. Spencer, Fly Ash
Utilization Symposium, March 14-16, 1967. Bureau of Mines Information
Circular 8348, pp. 287-315.
5. Rhode Island Clean Air Act, General Laws of 1956, Title 23, Chapters 25,
25.1 as amended to 1971.
6. State of Rhode Island and Providence Plantations, Dept. of Health.
Rules and Regulations Establishing Minimum Standards for Permissible
Types of Refuse Disposal Facilities, January 1969.
7. . Rules and Regulations for the Prevention; Control, and
Abatement and Limitation of Air Pollution. Regulation 1-13. As of
January 1973.
8. . Solid Waste Law. Ch. 192, P.L. of 1968.
9. . Coated Resource Management Council. Title 45 General
Laws, Chapter 279. P.L. of 1971.
10. U.S. Bureau of Mines Minerals Yearbook, Vol. II, 1970. Table 2, p. 2.
11. Keairns, D.L., D.H. Archer, R.A. Newby, E.P. 0"Neill, E.J. Vidt.
Evaluation of the Fluidized Bed Combustion Process. Volume IV.
Environmental Protection Agency. Westinghouse Research Laboratories.
Pittsburgh, Pennsylvania. EPA-650/2-73-048d. November 1973.
12. Coal Research Bureau, West Virginia University, Morgantown, W. Va.
Pilot Scale-up of Processes to Demonstrate Utilization of Pulverized
Cool Flyash Modifier by the Addition of Limestone-Dolomite Sulfur
Dioxide Removal Additive. Proposed for EPA Oct. 1971. NDISPB 213 639.
13. Staff Bureau of Mines, Bulletin 650. Mineral Facts and Problems, 1970
edition Chapters on sulfur, gypsum, calcium, phosphorus, clay, and sand
and gravel.
269
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APPENDIX L
DRY SULFATION EXPERIMENTAL PROGRAM
L
-------
APPENDIX L
DRY SULFATION EXPERIMENTAL PROGRAM
All fuel desulfurization processes generate as a by-product
spent sulfur sorbent and sulfur in some formfor example, sulfur (S) or
sulfuric acid (H.SO,). The oil gasification process produces a gas stream
containing up to 10 percent sulfur dioxide (SO.) and spent calcium oxide
(CaO) at a temperature of 1070°C (1958°F). Although the process uses a
regeneration cycle, the overall makeup rate of calcium required to
sustain adequate desulfurization yields an approximately 1:1 molar ratio
of calcium to sulfur in the by-product streams. Independent treatment
of these by-products requires additional processing facilities to convert
sulfur dioxide treatment for the hot waste calcium oxide which contains
residual sulfide and has retained sufficient activity as hydrolysable
lime to initiate combustion on contact with air and organic waste matter.
A potential solution to these problems lies in combining the waste
materials under suitable conditions to form calcium sulfate. The immediate
advantage of the process is that a dry, chemically stable solid is
formed as the sole waste material released to the environment.
The initial process concept is outlined in Figure L-l.
TECHNICAL ASSESSMENT
There is a large body of literature on the use of lime to capture
sulfur dioxide during or after fossil fuel combustion, so it proved
possible to outline closely the expected conditions under which the process
could be successfully applied to the CAFB by-product streams.
271
-------
Dwg, 62U8A33
0.5%
To Gasifier
CaO+S(L+V20,
2 t
(H20)
CaSO,
Sulfided
Sorbent
CaS/CaO
From Gasifier
Regenerated
Sorbent CaO
To Gasifier
S02<10%
Regenerator
1070°C
(1958°F)
Air
Dry Sulfator
870° C
Air
Stone By- Product
(Ca0)0.1(CaS04'o.9
Spent Sorbent
CaO
Figure L-l- The dry sulfation process
272
-------
The immediate problems which required consideration were:
The utilization criterion or extent of conversion achieved
before the process is judged feasible from the point of
view of total sulfur emissions from the plant and residual
activity of unsulfated lime
The particle size of lime required to achieve a given degree
of sulfation or utilization of the lime.
This utilization criterion was arbitrarily set at >90 percent
conversion of calcium oxide to calcium sulfate (CaSO,), at >90 percent
fixation of sulfur dioxide in calcium sulfate. This would result in an
overall sulfur dioxide removal efficiency greater than 81 percent at
90 percent sulfur retention in the gasifier. The efficiency could be 90
percent of total sulfur if the by-product sulfur dioxide which is not
converted to calcium sulfate is recycled to the gasifier.
Particle Size
Decreasing particle size generally increases the calcium utili-
zation. Approximately 40 percent utilization of the calcium oxide is
the highest value observed in fluidlzed-bed capture of sulfur dioxide
by calcined lime using about 500-micron particles. It is clear that
much smaller particle sizes would be essential to obtain the high
calcium oxide utilization required to match the production of sulfur
dioxide regenerator off-gas and the calcium/sulfur one-molar makeup
rate of stone.
This point may be illustrated by showing the rate of sulfation,
as a function of stone utilization, for 45-micron limestone particles and
the corresponding contact time in a fluidlzed bed required to achieve
90 percent sulfation of the calcium oxide in the bed at those rates, as
in Figure L-2.
Borgwardt, however, has demonstrated that utilization in
sulfation increases markedly as particle size decreases plugging of
the gas-transport pores in the solid becoming less of a barrier to
reaction as the diffusion thickness of the product decreases.
273
-------
Curve 678108-A
1
o
c
o
Sio'3
L_
o>
Q_
E
ZJ
O
*o
I ,-*
Required
Utilization, and
Rate Constants
u
ro
OJ
IS
or
to
"o
s
Q^
H
O
o
CO
to
10
IT = Contact Time .
(Seconds) for 90%
| S02 Retention
T = 0.5
T = 1.5
T=?-S
T=3.0
1 1
20 40 60 80 100
Stone Utilization
Figure L- 2- Rate criteria for sulfation of waste in
TG experiments at 0.5%
274
-------
A preliminary experiment on the sulfation of calcined lime-
stone was carried out on the thermogravimetric (TG) apparatus.
Limestone 1359, ground to smaller than 37 ym particle diameter,
was calcined in 14 percent carbon dioxide/nitrogen by heating to 871°C
(1600°F) in TG 295; when calcination was complete, the solid was sulfated
at 871°C (1600°F) in 0.5 percent sulfur dioxide, 4 percent oxygen in
nitrogen at atmospheric pressure, shown in Figure L-3.
The rate of reaction was rapid up to about 62 percent utiliza-
tion (16 minutes), and it slowed down considerably after this point. The
utilization reached 92.4 percent in 130 minutes. The rate of reaction
at the 90 percent utilization level was 1.749 x 10~ sec" (i.e., 1.749
_3
x 10 % of the total calcium oxide was converting to calcium sulfate
per second).
Using this rate data, the bed height necessary to retain 90
percent of the sulfur dioxide was calculated. It was assumed that first-
order kinetics holds up to the 10-percent sulfur dioxide levels expected
2
from the regeneration process. Work at TVA has demonstrated first-order
kinetics with respect to sulfur dioxide in the concentration range 2 to 10
percent. Using the equation
R = 100 [l-e~kT],
from the model shown in Figure L-4 the residence time required for the gas
in the bed at 90 percent retention is 0.75 seconds. (R is the percent
sulfur retained in the bed, k is the first-order rate constant, and T Is the
gas residence time.) This requires a bed-height about 1.2 times the
superficial gas velocity (in seconds) through the bed and is an encouraging
result. A literature review of fluldlzed-bed results yielded two sets
of data which bear on the problem.
The first experiments, carried out by PER on 44 um limestone
1359 at high-gas superficial velocity in a coal bed, at calcium/sulfur mole
ratio of 2.6, gave sulfur dioxide retention of 77 percent, with a bed
utilization of 30 percent. The real gas residence time must have been of
the order of 0.05 seconds. The relevance of this result to the dry sulfa-
tion process can be assessed by comparing models of sulfur retention for
275
-------
Curve 6?8l07-A
100
90
80
70
60
5 50
^^)
^^3
'§ 40
C 30
C
O)
CD
Q_
ro
O
20
10
I I
TG 295 < 34|im CaO from Limestone 1359
871°C (1600°F)
Atmospheric Pressure
I I I I
T
1
10 20 30 40 50 60 70
Time/Minutes
90 100 120 130 140
Figure L-3-The sulfation of fine particles of calcined limestone
-------
Dwg.
CaO+S02+l/202-*CaS04
S02, (Ce)
H
so2, (C0)
Air
V
p = Moles Ca/ ml
(CaO)01(CaS04)09
"~ U e
-kr
He , k1 p
T= -TT k= -7T-
o
o
to
o>
Q_
O
CO
0)
eg
CO
ro
O
Rate Curve
from TG
Experiment
atC1 = [S021
Utilization 0.9
V
Figure L-4-First-order model for fluidized bed sulfation of CaO
-------
In-bed generation of sulfur dioxide
Below-bed generation of sulfur dioxide.
In the latter case all the sulfur dioxide passes through the entire bed,
and sulfur dioxide retention is improved over the former process. The
Westinghouse model of first-order reaction in the two cases shows that if
77 percent retention is noted in the former case, under the conditions of
the PER experiment 98.8 percent sulfur dioxide retention will occur in
the latter case. Since the density of the bed (Ca in moles/ml) is not
known for the PER experiment, the group kpr has to be inferred as 4.5
from the retention figure noted. To achieve 90 percent sulfur dioxide
retention, the group kpt must not fall below 2.3. According to the model,
since T can increase by a factor of * 30 (making residence time = 1.5
second), k can fall by a factor of * 50. In the TG experiment, the rate
at 90 percent utilization was 1/50 of the rate at 30 percent utilization,
indicating the feasibility of 90 percent sulfur removal at 90 percent uti-
lization.
4
Argonne tested the reactivity of fine particles of limestone
1359 (25 um) in the fluidized bed combustion of coal. While laboratory
tests showed 85 percent utilization of the calcium oxide in the elutriated
solids, they did not observe such good results in the fluidized bed. They
concluded that fine particles had insufficient residence time in the bed.
There was, however, a significant correlation between the extent of
calcination and extent of sulfation of particles collected from the
fluid bed, primary cyclones, and secondary cyclones. The indication was
that 80 to 90 percent sulfation occurred with fully calcined particles.
The by-product for disposal is not freshly calcined material
but lime which has been subjected to recycling between gasifier
conditions (870°C/1600°F) and regenerator conditions (1070°C/1958°F). It
was anticipated that material from the plant would sulfate at a lower
rate because some degree of sintering must occur under process conditions.
Two more experiments were carried out using regenerator bed
material from run 7 on the CAFB plant at the Esso Research Centre,
Abingdon, England (Esso). In six hours of sulfation in 0.5 percent
sulfur dioxide, 4 percent oxygen in nitrogen at 870°C (1600°F), 74 percent
278
-------
of the residual calcium oxide was sulfated. Calculation using the
first-order fluidized bed model shows that 90 percent removal would be
achieved at a gas residence time of 1.5 seconds, at the reaction rate
noted for 74 percent utilization.
In a second, repeated, sulfation run the effect of water vapor
on the slow sulfation rate was tested. The rate increased by a factor
of 10 when * 3 percent water vapor was introduced into the sulfating
gas stream.
These results were sufficiently encouraging to proceed with
tests using a concentration of sulfur dioxide which more precisely
simulates the process conditions. It was also necessary to seek an
optimum temperature for the process, since the 871°C (1600°F) figure
was chosen to maximize the rates at 0.5 percent sulfur dioxide in the
sulfating gas.
TG Tests Using Five Percent Sulfur Dioxide as the Sulfating Gas
The results obtained using a simulated inlet concentration
confirmed that first-order predictions hold from 0.5 to 5 percent
sulfur dioxide; over 90 percent of the calcium oxide was sulfated in
several experiments.
The rate of sulfation at 5.0 percent sulfur dioxide, 10 percent
oxygen in nitrogen was accelerated
By increasing the temperature from 871 to 910°C (1600 to
1670°F)
By saturating the air component of the sulfating mixture
with water at 20°C (68°F).
The rate of reaction was decelerated
e By decreasing the temperature to 825°C (1517°F)
By increasing the particle size from < 74 urn diameter to
[74-149] urn diameter.
The course of reaction for four runs is shown in Figure L-5.
The projected sulfur dioxide concentrations exiting from a
fluidized bed dry sulfation reactor were calculated, and the results for
279
-------
Curve 657023-A
to
CD
O
60
o 40
CAFB Run No 7 (Limestone 1359), Ground as Noted
Sulfated in 5% S02, 82% N2, 10% 02, 3% H20
CaO + S02 + 1/2 02 - CaS04
TG317 v 920°C(1688°F)>74< 148pm
TG316 o 8250C(1517°F)<74um
TG314 A 871°C(1600°F)<74Mm
TG315 n 910°C (1670°F)<74pm
40
60 80
Time/Minutes
Figure L-5-The suifation of spent sorbent from the CAFB regenerator bed
-------
a bed which is 90 percent sulfated are shown in Table L-l. The projec-
tions are probably underestimates, since they do not allow for the
presence in the bed of material which is less heavily sulfated. Because
there is not a sharp decrease in rate at 90 percent sulfation, reaction may
proceed past 90 percent sulfation; and there will be a distribution of
particles in the bed with sulfation levels of between * 40 and 96 percent.
The rate-constant used to calculate the residence time in the bed is less
than the true average rate constant.
Table L-l
PROJECTED SO RETENTION IN FLUIDIZED BEDS OF CAFB SPENT SORBENT
£ (90% Sulfated)
Experiment
316
314
315
317
T/°C (°F)
825 (1517)
871 (1600)
910 (1670)
920 (1688)
Particle size
(ym)
< 74
< 74
< 74
> 74 < 149
% S02 Retained in bed
nominal gas residence time
2 sec | 4 sec
71.3 91.8
75.5 94.0
97.5 99.9
77.0 94.7
The runs carried out are summarized in Table L-2. The rate data
for TG 314 are plotted in Figure L-6. These rate data were later used to
project the performance of the 25.4 mm (1.0 in) fixed-bed reactor.
TESTS OF THE DRY SULFATION PROCESS IN A 25.4 mm (1 in) DIAMETER REACTOR
The thermogravimetric studies have demonstrated the technical
feasibility of the dry sulfation process and identified the optimum
conditions to be: particle size smaller than 74 ym, temperature range
870 to 910°C (1600 to 1670°F), and gas composition of 5 percent sulfur dioxide,
10 percent oxygen, and 3 percent water vapor in nitrogen. The purpose
of this work was to study the dry sulfation reaction in a 25.4 mm (1.0 in)
diameter fixed or fluidized bed and verify that optimum conditions and
design bases for the dry sulfator of the demonstration plant could be
projected from the kinetic model based on the thermogravimetric analyses.
281
-------
Table L-2
TG EXPERIMENTS PROBING CONDITIONS FOR THE DRY SULFATION PROCESS
TG No. Solid
295 Limestone
1359
296 "
297 "
312 "
313 "
314
315
316
317
318 "
319 "
Size/iim Pretreatment T/°C (°F) %
< 37 Calcined in 871 (1600)
14% C09/N9 at
871°C (IBOO'F)
< 74 CAFB Run #7 "
regenerator
ii ii ii
ii ii ii
ii ii ii
M ii ii
910 (16/0)
825 (1517)
< 149 " 925 (1697)
> 74
< 74 " 910 (1670)
ii ii ii
% CaO
S00 % 00 % CaSO/./time (min)
0.5 4 92.4/130
0.5 4 74/370
0.5 4
5.0 10 90/118
5.0 10 > 97/
5.0 10 91.4/77
5.0 10 95.2/43
5.0 10 96/115
5.0 10
5.0 10
0.5 4
Comment
Rate increased
by introduction
of H20
Accelerated
rate due to H_0
Residual CaO
did not recar-
bonate
-------
5 \
ISJ
00
J_
O)
c
in
-------
EXPERIMENTAL SYSTEM
A 25.4 mm (1 in) quartz reactor was designed and constructed
to investigate the dry sulfation process in a bench-scale fluid bed
reactor under the optimum conditions identified in the thermogravimetric
tests. Because of the hygroscopic nature of the fine bed material and
the small diameter of the quartz bed, problems associated with bed
agglomeration, slugging, and channeling became so severe that it was
virtually impossible to achieve any good quality fluidization, even
with the assistance of a bed vibrator. Fixed bed reactions, therefore,
were conducted in this unit for most of the dry sulfation tests.
Figure L-7 shows a schematic diagram of the 25.4 mm (1 in) reactor
test system. By monitoring the sulfur dioxide concentration of the exit gas,
the sulfation reaction in the bed could be followed when the sulfur dioxide
concentration entering the bed was known. A Dynascience sulfur dioxide
monitor with SS 330 sensor measuring up to 10 percent of sulfur dioxide
was used. At the upstream 10 percent sulfur dioxide in nitrogen was
diluted with an equal volume of air (either dried or saturated with water
vapor at 25°C [77°F]) to form 5 percent sulfur dioxide and 10 percent oxygen
in nitrogen, which was preheated by passing through the heated quartz
beads in the bottom section of the quartz tube. The exhaust was filtered,
partially bled out, and the remainder passed to the sulfur dioxide
sensor to be monitored for its sulfur dioxide rejection by the bed.
EXPERIMENTAL RESULTS
Twenty bench-scale dry sulfation tests using the 25.4 mm (1 in) quartz
reactor, designed to study the effects of temperature, particle size, and
gas velocity are summarized in Table L-3. The degree of sulfation was
determined by integration of the sulfur dioxide absorption profiles, as
well as by analyses of the final bed materials by wet chemical method,
gravimetric method, and X-ray diffraction. For runs DS1 to DS8 where
fluidization was attempted, the gas-flow rate was chosen between the
calculated minimum and terminal fluidization velocities so as to provide
maximum fluidization visually at room temperature without suffering too
284
-------
Air
Flow
Meter
Particulate
Thermocouple #1 Filter
CAFB Regenerator
Material
Fritted Quartz
Distribution Plate
Quartz
Beads
Drierite
b
Open-end
Manometer
5%S02,10%02 Thermocouple
in N2 #2
(wor W/oH20)
H20 Bubbler
at Room Temp.
For
Dilution
Bleed out /
Exhaust
Dyna-
science
S02
Meter
Re-
corder
Figure L-7-Schematic diagram of the 25 mm quartz reactor test system for dry sulfation
-------
much elutriation of fines. It must be mentioned that the quality of
fluidization achieved was poor in all cases and worse with finer parti-
cle size (smaller than 74 microns). It improved significantly when the bed
was vibrated with an electronic vibrator attached at the top of the quartz
tube. Fixed-bed reactions were carried out in runs 9 through 20, with
the exception of DS14. The nominal residence time was determined by the
bed height divided by the linear gas velocity. The reaction was usually
terminated when the sulfur dioxide absorption ceased or reached a nearly
steady rate. In most cases where there was poor or no fluidization the
bed material agglomerated at the end of the reaction, probably because
of local melting. This was especially true for the finer (smaller than
74 micron) particles.
Sulfur dioxide absorption profiles of three different combin-
ations of bed weights and gas rates are shown in Figures L-8, L-9, and L-10.
The shaded areas indicate the reaction time required for 100 percent
sulfation of the bed material, at 100 percent sulfur dioxide absorption
level for each combination of sulfur dioxide flow, and the amount of calcium
oxide in the bed. In Figure L-8 fluid bed tests of the same amount of
sulfur dioxide flow and bed material, and, therefore, the same gas-solid
contact time, were compared. The following general trends were noted:
The sulfation reaction was faster for the finer particle
size, smaller than 74 microns
The sulfation reaction was accelerated by the addition of
1.5 percent water vapor for the same particle size.
Repeated dry sulfation runs with two different gas velocities
were plotted in Figures L-9 and L-10 separately. Although the reproduci-
billty was not good under any chosen set of controllable parameters , the
absorption profiles followed a general shape. Greater than 90 percent
sulfation of the bed material (CAFB reg., < 74y) was obtained in this
reactor geometry at approximately 30 percent sulfur dioxide removal of
an inlet sulfur dioxide concentration of 5 percent.
As it was extremely difficult to fluidize the fine powder
(< 74p) of CAFB regenerator material in the quartz unit, it was mixed with a
286
-------
PAGE NOT
AVAILABLE
DIGITALLY
-------
is)
CO
\o
90 120
Time, min.
180
Figure L-8-S02 absorption profiles of the following dry sulfation runs in the 25 mm quartz
reactor with 10 gm CAFB regenerator material and 500 cc/min of 5% S02, 10%
in N2 with or without H20 saturation at 25°C:
DS2 < 297n, 1.5%H2Oadded, 910°C (1670°F)
DS3 < 297M, 1.5% H20 added, 870°C (1600°F)
DS4 < 74M, L5%H20added, 870°C (1600°F)
DS5 < 297u, No H20 added, 870°C (1600°F)
DS6 < 74jj, NO H20 added, 870°C (1600°F)
The shaded area indicates 100% sulfation of the CaO.
-------
Curve 658lf90-A
VO
o
100% Sulfation at
100% Absorption
DSW, 62% Sulfation
DS17, 95% Sulfation
DS18, 888, Sulfation
DS20, 80% Sulfation
180
Time, min
Figure L-9-S02 absorption profiles for dry sulfation runs at 870°C (1600°F), of
7 gm, < 74n CAFB material with 250 cc/min of 5% S02, 10% 02, 1.5%
H20 in N2
-------
100
c
o
to
\o
CM
O
i2 60
o
8
fe 40
20
0
60
120
Curve 65#f88-A
T
100% Sufaltion at
100% S02 Absorption
180 240
Time, min
300
360
420
Figure L-10-S02 absorption profiles for dry sulfation runs at 870°C (1600°F), of 7 gm,
<74p CAFB material with 150 cc/min of S02, 10%02, 1.5% H20 in N2.
-------
larger size of calcium sulfate (< lOOOy Drierite) in various proportions.
This was done in DS7, DS8, and DS14. As a result it was possible to
fluidize the mixtures, and the final beds were not hardened. In all three
runs the quality of the fluldization was poor, with high particle
carry-over. In Figure L-ll the percent of sulfur dioxide absorption was
plotted as a function of the total volume of 5 percent sulfur dioxide
feed per gram of calcium oxide in the bed, assuming no particle elutriation.
It can be seen that the sulfur dioxide absorption decreased rapidly.
This is attributed to the higher gas velocity, short contact time, and
significant calcium oxide elutriation loss.
The effects of temperature, particle size,and gas velocity were
shown in Figures L-12 and L-13, where the degree of calcium oxide sulfation
was plotted as a function of sulfur dioxide/calcium oxide ratio in the
feed stream. The dotted straight line indicates the theoretical maximum
sulfation. In these figures the relative positions of the curves, in
other words, the relative deviations from the theoretical maximum line,
were significant. Observation of these figures revealed the following:
At any SO./CaO feed, a higher degree of sulfation was
obtained at higher temperatures in the range of 770° to
870°C (1418 to 1600°F).
Higher degree of sulfation was achieved with smaller
particle size.
For particle sizes smaller than 74 microns, a higher percentage
of sulfation was obtained with lower gas velocity and longer
gas residence time.
The effect of the gas-flow rate was reversed for particle
sizes smaller than 297 microns.
EXPERIMENTAL VERSUS PREDICTED
Calculations were also made to predict the sulfur dioxide
absorption profiles of the fixed-bed reaction in the 25.4 mm (1 in) reactor
from the kinetic data obtained by the TG studies under similar operating
conditions.
292
-------
VO
LO
2gmCAFB,<74M
6gm Drierite,<1000u
DS14
4gmof CAFB,<74u
4gmDrierite,<1000M
100% Sulfation
at 100% S02 Absorption
5 10
Total 5% S02 Feed per Gram of CaO in the Bed, liters/gm
15
Figure L-11-Comparison of dry sulfation runs at 870°C (1600°F) using mixtures of fine
CAFB powder « 74u) with larger Drierite « lOOOu) with gas flow of 1440
cc/minof5%S02J&, 10%02, 1.5%H2OinN2
-------
100
VO
o
"fa
"5
CO
5
§ 50
CD
Q-
tq
o
/^Theoretical
Maximum
DS10,870°C
(1600°F)
DS15, 820° C
(1508°F)
0.2
0.4
0.6 0.8 1.0
Molar Ratio, SO^CaO
DS16, 770° C
(1418°F)
1.2
1.4
1.6
Figure L-12-Percent of CaO utlization as a function of the total S02 feed, showing the effect
of bed temperature
-------
100
IS)
SO
Ui
co
oo
250
O)
0>
Q_
k_
eg
O
0.2
/^Theoretical
/ Maximum
DS13,<74M, 150cc/min
DSll,<297u,
250cc/min
DS12,<297M, 150cc/min
0.4 0.6 0.8 1.0
Molar Ratio, SO^CaO
1.2
1.4
Figure L-13-Percent of CaO utilization as a function of the total S02 feed, showing
the effects of particle size and gas flow rate
-------
Two approaches were investigated. In the first model a rate
law was used to model the TG data over the entire range of utilization,
and an explicit model was then derived for the bed conversion as a
function of time. In the second model the TG data for rate as a function
of utilization were assumed constant in each of ten ranges of calcium
oxide utilization, and a numerical method was used to calculate the
retention of sulfur dioxide in the bed as a function of time.
Explicit Model Using an Averaged Overall Rate
The explicit model was devised by mass balancing of sulfur dioxide
and calcium oxide on a differential volume of the bed and integrating it
throughout the total bed height. The mathematics show the following
relationships:
S 1
100 At
L-l
At
L-2
At . BZ .
e + e -1
At .
6 -1 - L-3
At . BZ n
e + e - 1
where s = the percentage of sulfur dioxide absorption by the bed
pky« ,
Z = bed height expressed in moles of CaO
L = total bed height, mole
t = reaction time, min
P = total pressure, atm
296
-------
k = rate constant of the reaction, CaO + S09 + 1/2 C< » CaSO,,
-1 -1
as deduced from T6 studies fc-mole , min
y = S0« concentration expressed in moles of SO- per mole of N?
y - initial S0_ concentration
G = moles of M_ passing through the bed per minute
e = bed porosity
R = gas constant, atm-£-deg -mole
T = reaction temperature, °K
a = fraction of CaO sulfated.
The computation was complicated because the first-order* plot
from the TG data, -In (1-a) versus t, was not a simple straight line but
several segments with a range of slopes. Thus, a range of k values were
taken into consideration in calculating S from Equation L-l. Figures L-14
and L-15 shew the predicted sulfur dioxide absorption profiles in
Figures L-8 and L-9, respectively. Comparing Figures L-7 and L-14, as
well as Figures L-9 and L-15, revealed that the first part of the reaction
agreed rather well with the predicted values, but the second half, the
lower portion of the reversed "S" shaped curve, did not. The bed was
capable of removing a much higher percentage of sulfur dioxide from the
gas stream than oredicted from the kinetic model at greater than 50
percent calcium oxide utilization. This seemed to indicate that the
sulfation of calcium oxide proceeded initially as the explicit model
predicted but was later influenced by some complex mechanism (perhaps
involving the reaction product calcium sulfate) which resulted in
much greater sulfur dioxide removal. The formation of white fume when
the bed was partially sulfated (t 50 percent) seemed to promote this
hypothesis further, although the mechanism was not understood.
To check the validity of this exercise, the process was reversed
to predict the TG results by setting the bed height, Z, to zero in
Equation L-3 and calculating the fraction of calcium oxide conversion,
First-order in these instances refers to solid concentration.
297
-------
N>
\o
CO
140
120
Curve &58486-A
c
o
o
cs>
CVJ
O
CO
60
o
c 40
o>
20 ^«
0
0
60
120
180
Time, min
i i
100% Sulfation at 100%
S02 Absorption
240
360
Figure L-14-Calculated S02 absorption profile for dry sulfation reaction at 870°C
(1600°F) of 7 gm, < 74p CAFB material with 250 cc/min 5% S02, 10%
02, 1.5%H2OinN2
-------
to
VO
VO
7
100% Sulfation at
100% S09 Absorption/
Time, min
Figure L-15-Calculated SC>2 absorption profile for dry sulfation reaction at 870°C (1600°F),
of 7 gm, < 74|a CAFB material with 150 cc/min 5% S02, 10% 02, 1.5% H?0 in
N2
-------
a, at the first layer of the fixed bed. Figure L-16 shows the predicted
a versus t curve in the shaded region and compares it with the actual
data from TG 314 under the same operating conditions.
Numerical Prediction of the Sulfur Dioxide Retention in a 25.4 mm (1 in)
Diameter Batch Reactor
The reactor was modeled as follows. The total bed was
considered as a series of n segments of height Ah, where n*Ah = total
bed height, and n*At = total gas residence time in the bed. In each
segment reaction was assumed to proceed homogeneously over a finite
interval of sulfation, taken as 10 percent. For each interval of 10
percent sulfation, the average rate was read from the graph in Figure L-6,
so there are ten rate constants, each of which applies to any segment at
some point during sulfation of the bed.
The sulfur-dioxide-containing gas is considered as entering
the base of the bed. Reaction takes place uniformly in the first segment,
and the fraction of sulfur dioxide entering the second segment is computed.
This fraction remains constant until 10 percent sulfation of the first level
is achieved, when a lower rate constant begins to apply in this segment.
The computation is repeated through each layer of the bed until the fraction
of the sulfur dioxide which escapes from the bed has been computed. This
fraction changes every time one of the segments sulfates through a 10 percent
change.
The computation was programmed on a Wang 2200 B calculator,
and the program, shown in Table L-4, was run for several numbers of bed
segments using the TG rate data of 314 for the 10 rate constants. Apart
from smoothing the curve of sulfur dioxide retention with time, the
number of bed levels considered did not significantly alter the result.
The result for a 40-layer bed is shown in Figure L-17 and compared with
the experimental measurements in the 25.4 mm (1 in) unit. The parameters
input to the model were chosen as the experimental conditions in DS4
in the 25.4 mm (1 in) unit.
300
-------
Curve 658487-A
100 -
o Experimental
Shaded - Calculated
Time, min
Figure L-16-Comparison of the calculated dry sulfation reaction curve
and the experimental TG data
301
-------
Table L-4
PROGRAM FOR NUMERICAL PREDICTION
OF S02 RETENTION
o
ro
10 SELECT PRINT 205 (110)
20 V=40
30 DATA 4400,340,210,0
40 DIM Z(11),E(11)
SO FOR I = 1 TO 4:REA[) Z(I)
00 E(I)=EXP(-Z(I)*.01293*.2177/V)
70 NEXT I
80 PRINT "FIXED BED S02 ABSORPTION"
90 DIM T(40),U(12),F(40),J(40),Q(40)
100 K=218.075
110 FOR B = 1 TO 40: Q(B)=.3333:NEXT B
120 FOR I = 1TO 4:U(I)=.3333*I:NEXT I
130 A=0
140 P=2000
150 P=100*P
160 P=2*P
170 M-l
180 FOR B=l TO 40:S=Q(B):S=S/.3333
190 IF S]ll THEN 440
200 F(B)=E(S)
210 IF F(B)=1. THEN 260
220 H=U(S)
230 T(B) -(H-J(B)J*K/(ri*(l-F(BJ))
240 M=M*F(B)
250 IF P]T(B) THEN 420
260 NEXT B
270 M-l
280 FOR B = 1 TO 40:A-H-(M*F(B))
290 M-M*F(B)
300 J(B)=J(B)+P*A/K
310 IF Q(B)-J(B)L.0001 THEN 400
320 NEXT B
330 X=X+P/60
340 IF XLU+2 THEN 380
350 PRIUTUSING 360,M.X
360% *.###*
370 W=X
380 IF J(40)[.9THEN 150
390 STOP
400 Q(B)=Q(B)+.3333
410 GOTO 320
420 P=T(C)
430 GOTO 260
440 S=ll
450 GOTO 200
-------
CONCLUSIONS AND RECOMMENDATIONS
In a fixed bed of smaller than 74 micron CAFE regenerator
material, more than 90 percent sulfation was obtained at
30 percent sulfur dioxide absorption, at 870°C (1600°F) with
5 percent sulfur dioxide, 10 percent oxygen, 1.5 percent water
vapor in nitrogen, at three seconds nominal gas residence time.
A greater sulfation rate was observed with smaller particle
size and lower gas velocity.
A greater sulfation rate was observed at higher temperature
in the range of 770 to 870°C (1418 to 1598°F).
The addition of 1.5 percent water vapor accelerated the
sulfation reaction.
Agglomeration of the bed material was a problem when good
fluidizatlon was not achieved.
The reaction mechanism, including the formation of white fume
when the bed was approximately 50 percent sulfated, was
imperfectly understood.
The fluidization behavior in the 25.4 mm (1 in) unit of fine
particles of CAFB regenerator material was not satisfactory. The
hygroscopic nature, electrostatic forces, light density,
fine particle size, and wall effects were among the sources
for the channeling, slugging, and elutriation problems observed.
Further dry sulfation studies should be conducted in the 100 mm
(3.9 in) m fluid bed laboratory test unit.
Correlation of the results obtained from the TG data and the
25.4 mm (1 in) unit was reasonably good. It can be concluded
that the design numbers for the dry sulfator which were based
on the TG data were valid.
303
-------
g
2 i.o
E .8
o
i= 7
^^
I '6
S .5
8
o
CO
.3
2
Curve 678106-A
i i i i i i i i i i i i i i i i i
DS4 870°C <74u 5% S00 500cc/min
£.
I I I I I
40 Layer
Model
i I l i i i
I i i i i l
100 130 140 160 180 200 220 240
Time/Minutes
Figure L- 17-Fixed bed experimental results
-------
REFERENCES
1. Borgwardt, R.H., and R.S. Harvey. Environ. Sci. Technol. 6, (4),
350 (1972).
2. Hatfield, J.D., Y.K. Kim, R.C. Mullins, and G.H. McClellan.
Investigation of the Reactivities of Limestone to Remove Sulfur
Dioxide from Flue Gas. Prepared for the Air Pollution Control Office
by TVA, (1971).
3. Robison, E.B., A.H. Bagnulo, J.W. Bishop, S. Ehrlich. Interim Report
on Characterization and Control of Gaseous Emissions from Coal-Fired
Fluidized Bed Boilers. Pope, Evans and Robblns to NAPCA, October 1970.
4. Jonke, A.A. et al. Reduction of Atmospheric Pollution by the Applica-
tion of Fluidized Bed Combustion. Annual Report, Argonne National
Laboratories. Report ANL/ES-CEN-1002. 1969/1970.
5. Li, K. Private Communication. August 1974.
305
-------
APPENDIX M
SLURRY RECARBONATION EXPERIMENTAL PROGRAM
M
-------
ABSTRACT
The technical feasibility of slurry recarbonation for processing
the spent stone from oil gasification has been demonstrated. Under opti-
mum conditions 85°C (185°F), two hours, and twice stoichiometric carbon
dioxide (CO.) 96 to 97 percent recarbonation of the regenerator
spent stone from the oil gasification process is obtained. Slurry
settling characteristics, particle size distribution, and chemical
composition of the recarbonated product have been determined.
307
-------
APPENDIX M
SLURRY RECARBONATION EXPERIMENTAL PROGRAM
INTRODUCTION
There are many options for processing the spent stone from the
oil gasification process before disposalsuch as direct disposal, dry
sulfation, calcium sulfide (CaS) oxidation, dead-burning, silica sintering,
acid sulfation, and slurry recarbonation, as described in Appendix K.
Stone and Webster Engineering Corporation (SWEC), as subcontractor, has
completed an initial design for the demonstration plant, incorporating a
slurry recarbonation system which is considered to be of minimal techni-
cal risk.
For all practical purposes, dry carbon dioxide (CO.) will not
react with calcium oxide (CaO) at ordinary atmospheric temperatures; but
in the presence of water vapor or liquid water recarbonation occurs readily.
The hydration reaction of calcium oxide is strongly exothermic. The
reaction of water with waste stone from the oil gasification process not
only causes release of heat to the environment but also liberates
sufficient malodorous and toxic hydrogen sulfide (H.S) to be considered a
nuisance or health hazard. It seems, therefore, that the most reliable
way to forestall the weathering effects of ambient water vapor, rain, and
atmospheric carbon dioxide on the waste stone is to subject the stone
to thorough treatment with water and carbon dioxide prior to disposing
of it. This concept forms the basis for the slurry recarbona-
tion process as a method of pretreating spent stone from the oil gasifica-
tion process. The slurry recarbonation process, using flue gas, converts
the calcium sulfide present in the waste stone to hydrogen sulfide
gas, which is recycled to the gaslfier or could be sent to the
sulfur recovery plant. The recarbonated product is expected to
309
-------
be environmentally suitable for disposal. Furthermore, the sensible heat
of the hot waste stone and the heat from the reaction can also be uti-
lized in this process.
A liquid-phase Chance reaction which uses carbon dioxide to
convert calcium sulfide to calcium carbonate in aqueous solution was
studied by Consolidation Coal Company. At the Westinghouse Research
Laboratories a bench-scale experimental program has been set up to inves-
tigate slurry recarbonation of the CAFB spent stone. This section sum-
marizes the experimental results, including the optimum conditions for
wet carbonation reactions, and identifies the design basis for the slurry
recarbonation process as a method of pretreating spent stone from fluidized
bed oil gasification.
CONCLUSIONS AND RECOMMENDATIONS
The technical feasibility of slurry recarbonation for process-
ing spent stone from the oil gasification process has been demonstrated.
From these studies the following conclusions and recommendations are made:
Under optimum conditions85°C (185°F), two hours, and twice
stoichiometric carbon dioxide96 to 97 percent recarbona-
tion of chemically active fluidized bed (CAFB) regenerator
bed material is obtained.
Greater than 85 percent of the calcium sulfide present in
the CAFB waste stone is collected as gaseous hydrogen sul-
fide. The recarbonated product contains less than 0.2 per-
cent of residual sulfide.
The slurry half-settles in 4 minutes and completely settles
in 45 minutes.
The recarbonated solid is a white powder with the following
particle size distribution:
- 3 percent larger than 74 microns (u)
- 3 percent of 44 microns to 74 microns
- 94 percent of smaller than 44 microns
The average particle size is approximately 10 microns.
The marketability of the recarbonated producta fine, white
powder averaging 10 microns in sizeshould be investigated.
310
-------
EXPERIMENTAL DATA
Apparatus
Figure M-l shows a schematic diagram of the experimental set-up.
The liquid-phase carbonation reactions
CaO + H20^=
Ca(OH)2 + C02^=± CaC03 + H20
CaS + H20 + C02 ;F±- CaC03 + H2S
were studied in batch tests in a glass laboratory apparatus at atmospheric
pressure. The reaction vessel was a 500 ml, three-neck, round-bottomed
flask equipped with a thermometer, a gas inlet with a frit bubbler, and a
gas outlet which was connected to a train of acidified iodine bubblers.
The reaction mixture was heated in a hot water bath and a magnetic stir
used to mix the slurry.
Procedures
The spent stone from the regenerator of Esso Research Centre,
2
Abingdon, England (Esso) batch test No. 7 was used for these studies. The
starting material as received had a particle size range of from 300 to 3,000
microns and was ground to various mesh ranges for different runs . Fourteen
percent carbon dioxide in nitrogen was used to simulate the carbon dioxide
level in flue gas. In a typical run, 200 ml of deionized water was heated
in the reaction flask to the desired temperature. Then a calculated flow
of 14 percent carbon dioxide in nitrogen was admitted, and 5 grams of CAFB
No. 7 stone of a particular particle size range was added and magnetically
stirred. The exhaust gas was passed through two iodine solution traps and
the total hydrogen sulfide emission analyzed immediately after each slurry
recarbonation run. The parameters explored in these experiments included
temperature, particle size, water/stone ratio, amount of carbon dioxide,
and reaction time.
311
-------
H2S / C02 / N2
U)
M
ro
Flow
Meter
compressed
14%C02
in
N2
Figure M-1-Schematic diagram for the slurry recarbonation system
-------
Results
At the end of each reaction run, the mixture was filtered on a
Buchner funnel with S&S 589 medium-rate filter paper and moderate vacuum.
The precipitate was then washed with deionized water and dried at 500°C
(932°F) for one hour to convert any calcium hydroxide (CafOH]-) to calcium
oxide. The dried product was then calcined at 1000°C (1832°F) for one
hour, and the molar percent of recarbonation was determined gravimetrically as
. fc. (W500eC " W10QO°C)/44
% recarbonation = r:
1000°C/56
in which the starting material was assumed to be 100 percent calcium oxide.
Table M-l summarizes all the slurry recarbonation runs, their reaction
variables, and molar percentages of recarbonation as determined by the
gravimetric method on calcination of the reaction product. It can be
noted that the carbonation reaction is shifted further to the right with
smaller particle size, higher temperature, higher carbon dioxide flow,
and longer reaction time when the water/solid ratio is sufficient to
allow thorough mixing of the slurry. If the particle size of the starting
material (300 to 3000 microns), the reaction temperature (85°C [185°F]), and
the liquid/solid ratio (40/1) were held constant, the maximum degree of
recarbonation (91 to 92 percent) was achieved in a minimum reaction time
of two hours and twice the stoichiometrically required carbon dioxide.
Since the CAFB No. 7 regenerative material contains approximately 94 percent
calcium oxide and 2 percent calcium sulfide, however, the true percen-
tage of recarbonation of the total carbonatable material is the value
indicated above divided by 0.96; that is, under optimal experimental
conditions 96 to 97 percent of the carbonatable components of the CAFB
regenerative material is recarbonated.
Chemical Analyses
In order to understand the slurry recarbonation process better,
the following chemical analyses were carried out for runs 16 through 24:
a S~ in the exit gas by bubbling the exit gas through a
313
-------
Table M-l
SLURRY RECARBONATION TEST RESULTS
Test
Starting
material
& parti-
cle size
00
Reaction
time
(hours)
H20/
solid
Reaction
temperature,
°C (°F)
Total C02
used (times
theoretical)
Percent of
recarbonation
1
2
3
4
5
6
7
8
9
10
12
13
14
15
16
17
18
19
20
21
22
23
24
CAFB No. 7 3
reg. bed
300-3000
n 2
11 o
3
II Q
4
2
3
< 430 3
< 430 3
300-3000 3
500-1000 3
300-3000 3
3
n 2
i. 2
2
2
0.5
" 1
n 2
1
1
5:1
5:1
36:1
40:1
30:1
50:1
40:1
40:1
40:1
40:1
40:1
40:1
40:1
40:1
40:1
40:1
40:1
40:1
40:1
40:1
40:1
40:1
40:1
75 (167)
75 (167)
85 (185)
85 (185)
85 (185)
85 (185)
85 (185)
85 (185)
85 (185)
85 (185)
85 (185)
85 (185)
85 (185)
85 (185)
85 (185)
85 (185)
85 (185)
85 (185)
85 (185)
85 (185)
85 (185)
85 (185)
85 (185)
much excess
much excess
much excess
much excess
much excess
much excess
much excess
20 x C02
20 x C02
4 x C02
10 x C02
4 x C02
4 x C02
2 x C02
1.5 x C0_
1.5 x C0_
1.5 x C02
1 x C02
1 x C02
1 x C02
2 x C02
2 x C02
3 x CO.
37
39
85
92
88
92
90
91
91
90
88
92
90
89
85
90
85
88
68
79
91
79
84
314
-------
series of acidified iodine traps for hydrogen sulflde
and back titrating the unreacted iodine with sodium thiosulfate
using starch indicator
S~ in the filtrate by iodometric method and back titrating
with sodium thiosulfate
« S~ in the solid product same as above
I I
e Ca in the filtrate titrating with ethylenediaminetetra-
acetic acid solution buffered to pH = 10 using Eriochrome Black
T indicator
11
Ca in the solid same as above.
Table M-2 summarizes the results. The following observations
can be made:
At optimum conditions for the slurry recarbonationminimum
of twice stolchiometric carbon dioxide, 2 hours, and 85°C
(185°F) ~ 83 to 87 percent of the total sulfide was collected
in the gaseous exit, 2 to 3 percent in the supernatant, and
10 to 15 percent in the solid (runs 22 and 25).
Less than 0.2 percent of S~ was found in the recarbonated
solid product
The presence of carbon dioxide increased the release of
hydrogen sulfide as a gaseous product. The effect of
carbon dioxide was demonstrated by comparing runs 21, 23, and 24.
Insufficient carbon dioxide and/or too short a reaction
time resulted in higher dissolved sulfide concentration and
less hydrogen sulfide released from the gas exit. Run 19 was
an example with low S~ concentration in the gaseous exhaust
but high S~ in the filtrate, even though the percent of
recarbonation achieved with the set of reaction conditions
(1 x CO,, 2 hours, 85°C [195°F]) was marginally satisfactory.
I |
Higher dissolved Ca concentrations were found in the filtrates
of those runs with less complete recarbonation (runs 20, 21, and
23). This was expected because of the higher dissolved
calcium hydroxide.
315
-------
_ Table M-2
AND S= DETERMINATION FOR THE SLURRY RECARBONATION REACTIONS
Run
Reaction
conditions
Z
Recar-
bonatlon
S~. lodometrlc method
Gas
Wt Z ln_
total s"
In the
starting
material
Filtrate
Mg/ml
Wt Z ln_
total S
In the
starting
material
Solid
Wt Z In
solid
product
Wt Z ln_
total S
in the
starting
material
Ca*+. EOT A Method
Filtrate
Mg/ml
Solid
Wt Z In
carbonated
solid
Wt Z In
calcined
solid
6
8
15
16
OJ
C* 17
18
19
20
21
22
23
24
25
85'C (185'F). 4 hr
excess (X>2
85°C (185°F). 3 hr
20 x C02
85'C (185°F), 3 hr
2 x C02
85'C (185'F), 3 hr
1.5 x CO,
2
same as No. 16
same as No. 16
85'C (185*F), 2 hr
1 x 2
85°C (185'F), 1/2 hr
1 x C02
85°C (185'F), 1 hr
1 x COj
85'c (185'F), 2 hr
2 x C02
85'C (185'F), 1 hr
2 x C02
8S°C (185'F), 1 hr
3 x C02
85'C (185'F), 2 hr
92
91
89
85
90
85
88
68
79
91
79
84
92
<0.01
0.024
0.016
0.272
,
89 0
85 0.083
36 0.166 51
10 0.354 78
9 0.288 73
83 0.011 2
22 0.248 57
77 0.013 3
87 0.0096 3
<0.05 0.107
0.116
0.218
0.280
0.120
0.170
0.092 13 0.190
0.15 12 0.755
0.19 18 0.643
0.17 15 0.147
0.22 21 0.413
0.22 20 0.176
0.094 10 0.144
38.4
39
41 68
41 69
40
46
43
40
42
40
40
2 x CO,
-------
series of acidified iodine traps for hydrogen sulfide
and back titrating the unreacted iodine with sodium thiosulfate
using starch indicator
S~ in the filtrate by iodometric method and back titrating
with sodium thiosulfate
S~ in the solid product same as above
j
o Ca in the filtrate titrating with ethylenediaminetetra-
acetic acid solution buffered to pH = 10 using Eriochrome Black
T indicator
j^j
Ca in the solid same as above.
Table M-2 summarizes the results. The following observations
can be made:
At optimum conditions for the slurry recarbonationminimum
of twice stoichiometric carbon dioxide, 2 hours, and 85°C
(185°F) 83 to 87 percent of the total sulfide was collected
in the gaseous exit, 2 to 3 percent in the supernatant, and
10 to 15 percent in the solid (runs 22 and 25).
Less than 0.2 percent of S~ was found in the recarbonated
solid product
The presence of carbon dioxide increased the release of
hydrogen sulfide as a gaseous product. The effect of
carbon dioxide was demonstrated by comparing runs 21, 23, and 24.
Insufficient carbon dioxide and/or too short a reaction
time resulted in higher dissolved sulfide concentration and
less hydrogen sulfide released from the gas exit. Run 19 was
an example with low S~ concentration in the gaseous exhaust
but high S~ in the filtrate, even though the percent of
recarbonation achieved with the set of reaction conditions
(1 x CO., 2 hours, 858C [195°F]) was marginally satisfactory.
11
Higher dissolved Ca concentrations were found in the filtrates
of those runs with less complete recarbonation (runs 20, 21, and
23). This was expected because of the higher dissolved
calcium hydroxide.
315
-------
_ Table M-2
AND S= DETERMINATION FOR THE SLURRY RECARBONATION REACTIONS
Run
Reaction
conditions
Z
Recar-
bonatlon
S", lodometrlc method
Gas
Wt Z ln_
total s"
In the
starting
material
Filtrate
MB/n,l
Wt Z ln_
total S
In the
starting
material
Solid
Wt Z In
solid
product
Wt Z ln_
total S
In the
starting
material
OB**. EOT A Method
Filtrate
fte/ml
Solid
Wt Z In
carbonated
solid
Wt Z in
calcined
solid
u>
6
8
15
16
17
18
19
20
21
22
23
24
25
85'C (185'F).
excess O>2
85'C (185'F),
20 x C02
85'C (185'F),
2 x 2
85'C (185'F),
1.5 x C02
same as Ho. 16
same as Mo. 16
85'C (185'F),
1 x COj
85'C (185°F),
85'C (185'F),
1 x C02
85'c (185°F),
2 x C02
85'C (185'F),
2 x C02
85'C (185°F),
3 x C02
85"C (185°F),
2 x C02
4
3
3
3
2
hr
hr
hr
hr
hr
1/2 hr
1
2
1
1
2
hr
hr
hr
hr
hr
92
91
89
85
90
85
88
68
79
91
79
84
92
<0.01
0
0
89 0
85 0
36 0
10 0
9 0
83 0
22 0
77 0
87 0
0.024
.016
.272
.083
.166 51
.354 78
.288 73
.011 2
.248 57
.013 3
.0096 3
<0.05 0
0
0
0
0
0
0.092 13 0
0.15 12 0
0.19 18 0
0.17 15 0
0.22 21 0
0.22 20 0
0.094 10 0
.107
.116
.218
.280
.120
.170
.190
.755
.643
.147
.413
.176
.144
38.4
39
41 68
41 69
40
46
43
40
42
40
40
-------
Percentages of Ca in the. solid products before and after
calcination were in good agreement with the expected 40 percent
for calcium carbonate and 71.4 percent for calcium oxide.
The repeatability of the slurry recarbonation experiments was
demonstrated in runs 22 and 25, both of which were carried out
at the optimum conditions twice stoichiometric carbon dioxide,
two hours, and 85°C (185°F). Recarbonation of 91 and 92 percent,
respectively, was achieved. Results from the sulfide and
calcium analyses were also in good agreement.
The starting material (CAFB No. 7 regenerator bed stone), as
well as the carbonated solid product and filtrate from run 6, were analyzed
by the analytical chemistry staff for S~, SOT, and cationic impurities,
using emission spectroscopy and wet chemical methods. Table M-3 presents
the results.
The filtrate from the slurry recarbonation run is slightly
basic with pH = 7.9, probably due to some dissolved calcium hydroxide.
Its buffering capacity is shown in Figure M-2 where 10 ml of a typical
recarbonation filtrate (run 6) were neutralized with 0.1 normal
hydrochloric acid.
Slurry Settling Characteristics
The recarbonation reaction slurry from run 25 was allowed to
cool to room temperature before it was transferred into a 250 ml graduated
2
cylinder (h = 239 mm, A = 10.47 cm ). It was then diluted to 250 ml
with deionized water and thoroughly mixed for five minutes, including
inversion of the cylinder by hand. The sedimentation height at the
bottom of the menicus was then visually taken at 0.5-, 1.0-, 1.5-minute
intervals. Figure M-3 shows a slurry settling curve, from which it is
noted that the time required for the slurry to settle completely was 45
minutes, and the time required to settle to half the volume (125 ml)
was 4.0 minutes.
317
-------
Table M-3
CHEMICAL ANALYSES OF Ca"1"4", S=, AND CATIONIC IMPURITIES FOR A
TYPICAL SLURRY RECARBONATION TEST
Species
CAFB No. 7 regenerator
bed material
(wt %)
Run No. 6
solid product
(wt %)
Run No. 6 filtrate
(ug/ml)
Al
Ag
B
Ca
Cr
Cu
Fe
Mg
Mn
Mo
Na
Ni
Si
V
S=
S04
pH
3.4
0.02
0
69.0
0.2
0.2
0.9
1.4
0.9
0
0.8
0.1
1.4
0.9
1.0
2.6
__
0.81
0.004
0
40.0
0
0.1
0.4
0.8
0.8
0
0.4
0.2
0.05
0.2
< 0.05
1.4
__
10.5
0.2
2.0
107.0
0
5.0
1.0
213.0
0.5
10.5
53.0
3.0
3.0
64.0
10.0
240.0
7.9
318
-------
8.0
7.0
6.0
5.0
PH
4.0
3.0
2.0
1.0
i I i I i i I i i
I I
I I I I
0 .05 .10 .15 .20 .25 .30 .35 .40
ml of 0. 1 HC I
.45 .50
Figure M-2 -Neutralization curve of 10 m I slurry recarbonation
run 6 filtrate with 0.1NHC I
319
-------
Curve 6565*»7-B
t, = 4 Min for Settling to Half Volume
t = 45 Min for Complete Settling
50
100
Settling Time, min
150
200
Figure M-3 -Slurry settling curve for run 25
-------
Particle Size Distribution
To further the studies of the slurry recarbonation process for
treating spent stone before disposal, the particle size and distribution
of the recarbonated product were investigated. Two methods were employed
during this study!
Mechanical Sieving
Standard dry sieving, with sieve sizes extending to 37 microns
or more commonly 44 microns, was not satisfactory in this case. The
agglomerates of the recarbonated powder, even though they adhered
loosely together, would not pass through the smallest sieve sizes,
thereby distorting the gradation and making it appear coarser than it
actually was. A wet sieving procedure was found more suitable. A
stream of water was applied at moderate pressure to wash down the putty,
and the agglomerates then readily broke up into smaller particles and
passed through the smaller sieve sizes. The residues on variously
sized screens were then dried at 500°C (9328F) and the weights recorded.
The percentage of residue on each sieve was calculated on the basis
of the total weight. Table M-4 summarizes the particle size distribution
of the recarbonated product from run 25 as determined by the wet sieving
method.
Table M-4
PARTICLE SIZE DISTRIBUTION OF THE RECARBONATED
PRODUCT FROM SLURRY RECARBONATION RUN 25
Sieve Size
+ 200 mesh
+ 325 mesh
- 200 mesh
- 325 mesh
| Particle Size
> 74y
44 to 74y
< 44y
I Distribution
2.6%
3.2%
94.2%
321
-------
Microscopy
Ordinary microscopy, scanning electron microscopy, and photo-
microscopy are often employed to determine particle size, shape and form,
and size distribution. Figure M-4a is a typical optical microphotograph,
at 200 X, of the recarbonated material showing some agglomerates of
smaller particles with dark-field background. The average particle size
is seen to be approximately 10 microns. Figure M-4b shows a typical
electron scanning microphotograph of the same sample, at 2000 X, revealing
the fiber-like microstructure of the recarbonated product.
322
-------
(a)
(b)
Figure M-4a^\ typical optical microphotograph of the
slurry recarbonated product from run 22
(200X)
Figure M-45-A typical electron scanning microphoto-
graph of the slurry recarbonated product
from run 22
(2000X)
323
BM-59739
-------
REFERENCES
1. Esso Research Centre, Abingdon, England. Final Report to U.S.
Environmental Protection Agency, February 1974, Contract No. 68-02-03-00.
324
-------
APPENDIX N
ACID SULFATION EXPERIMENTAL PROGRAM
N
-------
ABSTRACT
The technical feasibility of acid sulfation for processing
spent stone from the oil gasification process has been demonstrated and
the design bases identified. The reaction rate is slow. The sulfation
obtained was 94 percent with 0.95 stoichiometric sulfuric acid in 7.2
hours at 95°C (203°F) with the feed particles ground to smaller than
1000 microns. Slurry characteristics and the particle size distribution
of the sulfated stone have been investigated and the chemical composi-
tions determined. The economic practicality of the acid sulfation
process has not yet been established. A precipitate has been produced
which may have market potential.
325
-------
APPENDIX N
ACID SULFATION EXPERIMENTAL PROGRAM
INTRODUCTION
One method for processing spent stone before disposal is
treatment by acid sulfation. The spent stone from the oil gasification
process reacts with dilute sulfuric acid (H-SO.) to form calcium sulfate
(CaSO.) which is environmentally suitable for disposal. The sulfur con-
tent is collected as gaseous hydrogen sulfide (H^S) and is passed on for
further processing. The major debit for this process is likely to be
the cost of sulfuric acid.
The Westinghouse Research Laboratories have set up a laboratory-
scale experimental program to investigate the liquid-phase sulfation
reactions of calcium oxide (CaO) and calcium sulfide (CaS) with dilute
sulfuric acid. This report summarizes the experimental results, including
optimum conditions for the acid sulfation reactions, and identifies the
design basis for the acid sulfation process as a method of pretreating
the spent stone from fluidized bed oil gasification.
CONCLUSIONS AND RECOMMENDATIONS
The reaction of the waste stone from the oil gasification
process with dilute sulfuric acid was slow. With 0.95 stoichiometric
sulfuric acid in 7.2 hours at 95°C (203°F) with the feed particles crushed
to smaller than 1000 microns (y), 94 percent sulfation was obtained.
When the spent stone was used without additional grinding
(300 to 3000y), 94 percent of sulfation was accomplished only when the
reaction parameters exceeded the following sets of minimal conditions:
90°C (194°F), 3 x H2S04, 4 hours
90°C (194°F), 2 x H2S04, 6 hours
95°C (203°F), 2 x HS0> 5 hours
327
-------
The excess sulfuric acid not only is uneconomical but also creates prob-
lems in handling.
More than 90 percent of the sulfide in the waste stone was
collected as gaseous hydrogen sulfide. Little dissolved sulfide was
found in the filtrate. The sulfated stone contained less than 0.02 percent
sulfide.
When the sulfated product was precipitated in 70 percent
methanol, more than 99 percent of the particles were smaller than 44
microns, with an average size of less than 10 microns. When the precip-
itate was allowed to digest in aqueous solution for five days, however,
the precipitate grew into needle-shaped crystallines of 50 to 500 microns
in length. This product may have market potential.
Slurry settling characteristics were determined for the two
types of precipitates described above. Results showed t-,- ^ 580 minutes
and tf 2? 800 minutes for the former, and t. ,« = 5.5 minutes and t, = 15
minutes for the latter.
Although the technical feasibility of the acid sulfation process
has been demonstrated, there are economical penalties to be paid unless
Sulfur dioxide from the regenerator off-gas can be converted
to sulfuric acid which, in turn, is used to sulfate the
spent stone
A marketable end product can be produced.
EXPERIMENTAL DATA
Apparatus
Figure N-l shows a schematic diagram of the experimental set-up.
The liquid-phase sulfation reactions
CaO + H2S0
CaS + H2S0
were studied in batch tests in a glass laboratory apparatus at atmospheric
pressure. The reaction vessel was a 500 ml, three-neck, round-bottomed
flask equipped with a thermometer, a gas inlet for passing a moderate
328
-------
H2S/N2
H2S04
Spent Stone
Magnetic Stir -
Heating Plate
Figure N-1-Schematic diagram for the acid sulfation system for
spent stone processing
329
-------
flow of inert gasin this case nitrogen (N.) over the reaction,
and a gas outlet which was connected to a train of acidified iodine
bubblers. The reaction mixture was heated in a water bath, and a
magnetic stir was used to mix the slurry.
i
Procedures
The spent stone from the regenerator of Esso Research Centre,
Abingdon, England (Esso) batch test No. 7 was used for these studies.
The starting material as received had a particle size range of from 300
to 3,000 microns and was ground to various mesh ranges for different
runs. In a typical acid sulfation run, a calculated amount of concentra-
ted sulfuric acid was diluted to 130 ml in a reaction flask and heated
to the desired reaction temperature in a heated water bath. Then
5 gm of the CAFB starting material of a specific particle size range
was added with magnetic stirring. A moderate flow of nitrogen was
flushed over the reaction mixture. The exhaust gas, which contains
\
small amounts of hydrogen sulfide in the nitrogen stream, was bubbled
through acidified iodine solutions from which the total hydrogen sulfide
was analyzed Immediately after each acid sulfation run. Parameters
explored in these experiments Included reaction temperature, particle
size of the waste stone, amount of sulfuric acid used, and reaction
time.
Due to the relatively high solubility of calcium sulfate in
water 0.21 gm/ 100 ml at 30°C (86°F) and 0.16 gm/100 ml at 110°C (230°F)
the degree of sulfation could not be determined by precipitating calcium
sulfate from the aqueous reaction mixture directly. It was found that the
solubility of the calcium sulfate was reduced significantly in a 70 percent
methanol (CH.OH) medium. An alternative method is described in the following.
At the termination of the reaction, the mixture was allowed to
cool before methanol was added. The precipitate was allowed to settle in
70 percent methanol overnight. The precipitate was then filtered, dried
at 110°C (230°F) for 2-1/2 hours, and weighed. The HO°C (230°F) dried
solid was again heated to 500°C (932°F) for one hour for the decomposition
330
-------
of calcium hydroxide (Ca[OH].) and hydrated calcium sulfate, if they
were present.
Results
The molar percentage of sulfation was calculated by the
following formula:
v i* ^ (W500°C " Winitial)/80 .__
% sulfation = - :j - x 100%
Winitial/5g
where 58 was the average molecular weight of the CAFB No. 7 regenerator
waste stone containing 96 percent calcium oxide, 2 percent calcium
sulfide, and 2 percent calcium sulfate; and 80 was the difference between
the molecular weights of calcium sulfate and calcium oxide.
Table N-l summarizes all the acid sulfation runs, their reaction
variables, the molar percent of sulfation as determined gravimetrically
by the above formula, and the molar ratio of water (H.O) to calcium sulfate
of the sulfated products. The sulfation reaction was accelerated with
smaller particle size, larger amounts of sulfuric acid, and higher
temperature .
The molar ratio of water to calcium sulfate was calculated from
the weight loss on heating the solid from 110°C (230°F) to 500°C (932°F)
by the following formula:
molar ratio of H.O/CaSO
.. - * /,,*
2 4 W° x * sulfation/136
No trend had been found to explain the variation of the molar ratios.
Thermal gravimetric work indicated that the water content was due to
calcium hydroxide, which was probably the Intermediate product of the
liquid-phase sulfation of calcium oxide.
From Table N-l it can be noted that more than 90 percent
sulfation of the waste stone (equivalent to less than 94 percent sulfa-
tion of the sulfatable portion of the stone) can be achieved only when
the reaction parameters exceed the following sets of minimal conditions:
331
-------
Table N-l
A SUMMARY OF THE ACID SULFATION EXPERIMENTS
Run
No.a
Temp t
°C
(°F)
Time,
hr
Amount of
H2S04,
times of
stoichiometric
required
CaS04
precipitated
in
% of b
sulfation
Molar
ratio
H?0/CaSO,
Comment
AS1
AS2
AS3
AS4
AS5
AS6
AS 7
ASS
AS9
AS10
ASH
AS12
AS13
AS14
AS15
AS16
85 (185)
90 (194)
90 (194)
90 (194)
90 (194)
90 (194)
90 (194)
90 (194)
90 (194)
90 (194)
90 (194)
95 (203)
95 (203)
95 (203)
95 (203)
95 (203)
6.0
6.5
4.0
4.0
6.0
6.0
8.0
4.0
2.0
3.0
8.0
6.5
6.5
2.0
4.0
8.0
5
3
3
2
2
1.33
1.33
2
3
3
1.33
1.33
0.95
3
2
1.33
H2S04/H20
70% CH3OH
30% H2S04/H20
same as above
same
same
same
same
same
same
same
same
same
same
same
same
same
56
90
92
82
93
66
83
83
74
70
79
58
85
86
86
1.3
0.15
0.63
1.12
discarded
0.35
0.57
0.25
0.16
1.25
1.50
0.39
0.55
0.49
-------
Table N-l (Continued)
Run
No. a
Temp,
°C
Time,
hr
Amount of
times
stoichiometric
required
CaSC>4
precipitated
in
% of
sulfation"
Molar
ratio
H,0/CaSO,C
Comment
AS17 95 (203) 7.2 0.95 same 90 0.66
AS18 95 (203) 7.2 0.95 70% CH3OH 59 0.67
CO
AS19
AS20
AS21
95 (203)
5.0
95 (203) 5.0
95 (203) 5.0
30%
Slurry used 0.51
for determin-
ation of its
settling char-
acteristics ;
percentage of
sulfation not
determined
0.95
same as AS18
same
90
68
0.16
0.5
No CH OH added; slurry
settling for 5 days and
white needle-shaped
crystalline formed in
water solution
Same reaction conditions
as AS19
aCAFB reg. bed stone used (300-3,OOOy) in all runs except those specified otherwise.
Percent of sulfation =
°Molar ratio of lUO/CaSO, = rr
dCAFB No. 7, 500-lOOOy.
100% .
AW
(110-500°C)/
18
«, x % sulfation/136
x 100% .
500°C
-------
1. 90°C (194°F), 3 x H2S04, four hours, 3,000p (A53)
2. 90°C (194°F), 2 x H2S04> 6 hours, 3,000y (A55)
3. 95°C (194°F), 2 x H^O^, 5 hours, 3,000u (A520)
4. 95°C (194°F), 0.95 x H^O^,, 7.2 hours, l.OOOp (AS17)
Under the first three sets of sulfating conditions, the process suffered
from excess sulfuric acid, which not only was uneconomical but also
created problems in handling. When 0.95 times stoichiometric sulfuric
acid was used with the purpose of providing just enough sulfating agent
to result in a near neutral solution (the fourth set of conditions), 90
percent sulfation could be accomplished in 7.2 hours, but only if the spent
stone was first crushed to smaller than 1,000 microns (comparing runs 17 and 18)
Although the technical feasibility has been demonstrated in these
studies, the economic practicability may only be justified with a marketable
gypsum product.
Chemical Analyses
The following chemical analyses were carried out in the
Westlnghouse laboratory:
S in the exit gas bubbling the exit gas through a
series of acidified iodine traps for hydrogen sulfide and
back titrating the unreacted iodine with sodium thiosulfate,
using a starch indicator
S~ in the filtrate by iodometric method and sodium
thiosulfate back titration
S in the solid product same as above
SO, in the solid product gravimetric determination as
barium sulfate (BaSO,).
I [ ^
Ca in the filtrate titration with ethylenediaminetetra-
acetic acid solution buffered to pH = 10 and using Eriochrome
Black T indicator
11
Ca in the solid same as above.
334
-------
It was found that more than 90 percent of the sulfide in the
CAFB regenerator material was collected as gaseous hydrogen sulfide.
Very little sulfide was found in the filtrate, and the sulfated solid
product contained less than 0.02 percent sulfide. The weight percentages
of calcium agreed reasonably with the theoretical value, 29.4 percent for
calcium in calcium sulfate.
The starting material (CAFB No. 7 regenerator bed stone), as
well as the solid products and filtrates from runs AS5 and AS19, were
analyzed by the analytical chemistry staff using emission spectroscopy
and wet chemical methods. Table N-2 presents the results.
Rate of Filtration and pH Change
Additional work was carried out to determine the rate of
filtration and the pH of the resulting filtrate and washing solutions.
At the termination of the reaction, the slurry mixture was
allowed to digest in 70 percent methanol overnight. It was then filtered
on S&S filter paper of medium texture (white ribbon) using a Buchner
funnel with moderate vacuum. The time required for filtering the mixture,
when the thick slurry was all contained within the volume of the Buchner
funnel (130 ml), to form a solid cake was eight minutes. In each washing,
130 ml of water at 51°C (123.8°F) was used. The wash water was collected
separately and the pH measured. Table N-3 summarizes the time required
for and the pH of each washing.
Particle Size Distribution
Two methods were used to study the particle size distribution of
the acid sulfated product.
Mechanical Sieving
Standard dry sieving with sieve sizes extending to 37 microns
or more commonly 44 microns was not satisfactory in this case. The
agglomerates of the recarbonated powder, even though they adhered
335
-------
Table N-2
CHEMICAL ANALYSES OF Ca"*"1", S=, SO^ AND CATIONIC IMPURITIES FOR
ACID SULFATION REACTANT, PRODUCTS, AND FILTRATES
Species
CAFB No. 7 regenerator
bed material
(wt %)
ASS
solid
(wt %)
ASS
filtrate
(yg/ml)
AS 19
solid
(wt %)
AS19
filtrate
(yg/ml)
Al
Ca
Cu
Fe
Mg
Mn
Na
Ni
Si
V
S=
so:
3.4
69.0
0.2
0.9
1.4
0.9
0.7
0.1
1.4
0.9
1.0
2.6
0.06
27
0.002
0.03
0.001
0.0006
0.1
0.006
0.7
0.0006
< 0.02
63
40
8
8
250
160
40
80
80
660
660
< 5
Excess
0.005
27
0.0005
0.02
0.0005
< 0.0003
0.03
< 0.003
0.05
0.0005
< 0.02
63
13
660
2
13
13
2
8
2
21
26
< 11
Excess
336
-------
Table N-3
RATE OF FILTRATION AND pH OF THE WASHING SOLUTIONS
FOR THE ACID SULFATION REACTION MIXTURE ASS
Filtration with
moderate vacuum
Time
pH
Reaction mixture (130 ml)
1st wash with 130 ml H.O
at 51°C (123.8°F) L
2nd wash with 130 ml H.O
at 51°C (123.8°F)
3rd wash with 130 ml HO
at 51°C (123.8°F)
4th wash with 130 ml H.O
at 51°C (123.8°F)
5th wash with 130 ml H.O
at 51°C (123.8°F)
8.0 min
1.5 min
1.0 min
50 sec
35 sec
20 sec
< 2.0
2.0
3.0
4.0
6.0
7.0
337
-------
loosely together, would not pass through the smallest sieve sizes,
thereby distorting the gradation and making it appear coarser than it
actually was. A wet sieving procedure in which a stream of water, applied
at moderate pressure and used to wash down the putty, was found to be
more suitable. The agglomerates then readily broke up into smaller
particles and passed through the smaller sieve sizes. The residues on
variously sized screens were then dried at 500°C (932°F) and their weights
recorded. The solid product (dried at 110°C [230°F]) from ASS was wet
sieved. It was found that more than 99 percent of all particles
were smaller than 44 microns.
Microscopy
Ordinary microscope, scanning electron microscopes, and photo-
microscopes are often employed for determining particle sizes, the shape
and form of the particles, and particle size distribution. Figure N-2a is a
typical optical microphotograph at 200X of the wet sulfated material (ASS)
showing some agglomerates of smaller particles with dark-field background.
The average particle size is seen to be smaller than 10 microns. Figure N-2b
shows a typical electron scanning microphotograph of the same sample at 2,OOOX,
revealing the fine microstructure of the acid-sulfated product.
Slurry Settling Characteristics
Acid sulfation runs AS19 and AS20 were carried out with identical
conditions. At the end of the five-hour reaction, the mixture of AS20 was
allowed to digest in 70 percent methanol for complete precipitation of
calcium sulfate, thus making possible the gravimetric determination of
the degree of sulfation (90 percent), while the AS19 slurry mixture was
transferred to a 500 ml graduated cylinder with enough water added to
make the total volume SOO ml. After thoroughly mixing (about S minutes shaking),
the sedimentation heights were recorded at 0.5-, 1.0-, 1.5-, and 2-minute
intervals. Figure N-3 shows the settling curve, from which it is estimated
that t. , ^ 580 minutes and t, * 800 minutes for half-volume and complete
338
-------
(a)
Figure N-2a-An optical microphotograph of the acid-sulfated
product from run AS5 (200X).
(b)
Figure N-2b-An electron scanning microphotograph of the acid-
sulfated product from run AS5 (500X).
339
RM-59943
-------
500
400
300
ET
'to
200
"eo
o>
e
100
Curve 67039-A
I
j = 580 min
^ s tf =800 min
1
1
I
I
200 400
600 800
Time, min
1000 1200 1400
Figure N-3-Slurry settling curve for acid run AS19 immediately following the
sulfation reaction
-------
settling, respectively. Since the critical data on Figure N-3 had been
missed, it was decided to reshake the slurry after four days and observe
the settling behavior once again. Figure N-4 shows the second settling
curve. The rate of this settling (t. .« B 5.5 minutes and t, = 15 minutes)
was much faster than that found on Figure N-3. This can be explained by
the growth of calcium sulfate particle size during the four days of
settling in the dilute sulfuric acid solution.
The precipitate was filtered, washed, and dried after a total
of five days in aqueous solution. The final product was white, needle-
shaped crystalline material of a much larger particle size than the sulfated
products (very fine powder) from the previous runs. Figure N-5 compares
some photomicrographs from ASS and AS19. The digestion of AS19 precipi-
tate slurry in aqueous solution during the five-day settling experiments
was most likely responsible for the difference in crystalline form and
size. The latter may be marketable.
341
-------
Curve 657Q1+0-A
500
400
300
'£ 200
c
o>
E
CO
100
~5.5 min
t = 15 min
I
50
100
Time, min
150
Figure N-4- Slurry settling curve of AS19 after 4-day settling and reshaking to form the
new slurry
-------
(a)
(b)
Figure N-5-Comparison of photomicrographs of the acid sulfation
products from (a) AS5 and (b) AS19 (200X).
343
RM-59942
-------
REFERENCES
1. Consolidated Coal Company. Monthly Progress Report No. 3. EPA
Contract No. 68-02-1333. September 1973.
2. Esso Research Centre, Abingdon, England. Final Report to U. S.
Environmental Protection Agency, February 1974, Contract No. 68-02-
03-00.
344
-------
APPENDIX 0
DRY RECABBONATION AND SINTERING EXPERIMENTAL PROGRAM
0
-------
APPENDIX 0
DRY RECARBONATION AND SINTERING EXPERIMENTAL PROGRAM
INTRODUCTION
Spent sorbent from the chemically active fluidized bed (CAFB)
process may contain considerable quantities of residual calcium oxide
(CaO). The residue from the regenerative process may contain 95 percent
of the calcium as calcium oxide; in once-through processes the residual
calcium oxide decreases as stone utilization in sulfur removal increases.
Although the stone has been rejected from the regenerative
cycle in order to allow input of fresh reactive limestone, the activity may
be too great to permit disposal of the stone without neutralization of
the residual activity of the calcium oxide. Stone and Webster Engineering
Corporation (SWEC) indicated the nature of this problem by adding some
water to a small pile of the regenerator bed material from run 7. The
sample self-heated sufficiently to combust the paper carton holding the
stone. Apart from the possibility of burning organic material in contact
with the waste and damaging plant and animal tissue, the material undergoes
substantial swelling on absorption of and reaction with water, rendering
it unsuitable for landfill.
Two processes to overcome this problem were investigated. In the
first process, high-temperature recarbonation of the sorbent, using flue
gas, was considered. In the second process dead-burning the material by
sintering it at high temperature was investigated.
SUMMARY OF HIGH-TEMPERATURE RECARBONATION
To dispose of spent sorbent as calcium carbonate (CaCO.), calcium
oxide recarbonation at high temperatures (700 to 800°C/1292 to 1472°F) was
studied. The results may be summarized as follows:
345
-------
Regenerator bed material cannot be substantially recarbon-
ated, because exposure to the relatively high temperature
in the regenerator (1080°C/1976°F) probably sinters the
calcium oxide.
Once-through material can be recarbonated after oxidation
of the calcium sulfide fraction.
The optimum process conditions projected from the thermo-
gravimetric (T6) data are (assuming atmospheric pressure
and 14 percent carbon dioxide and 3 percent water in
nitrogen as recarbonation gas):
- 725°C (1337°F)
- Fine particles (about 74 ym)
- Approximately twice stoichiometrlc carbon dioxide
- Fluidized bed gas residence time approximately 6 seconds
(based on superficial gas velocity).
SUMMARY OF STONE SINTERING
To dispose of spent sorbent as calcium oxide, the possibility
of deactivating the lime by subjecting It to a high-temperature sintering
process was briefly investigated. The results showed that sintering is
a feasible method for disposal-processing of material from a regenerative
CAFB plant. The process would produce a dry, inactive, nonpowdered solid;
it would be used in conjunction with a process for recovery of sulfur
dioxide from the regenerator off-gas. Activity tests during hydration
showed that sintering for two hours at 1400°C (2552°F) almost totally
prevents the hydration. Further tests are required to establish the extent
of emissions of trace elements from the stone during sintering and the
rate at which ions are leached from the solid during subsequent weathering
of the stone.
CALCIUM OXIDE RECARBONATION
Recarbonation experiments in the TG apparatus were carried out
on three materials (See Table 0-1):
346
-------
Table 0-1
TG RUNS FOR SPENT SORBENT DISPOSAL
TG No. Solid Pretreatment
Recarbonation of residual CaO
Reaction Temperature, Result
°C (°F)
248 CAFB 7 Heated to 500°C Air oxidation and 0-1100 Incomplete recarbonation
regenerator bed (932° F) in N.
249 CAFB 7 regener- Taken to 800 °C
ator material (1472°F) in NZ
256 Denbighshire Calcined in NZ
limestone
recarbonation (32-2012) over 12 hours. Sulfide
oxided corresponded to 2%
wt. of stone.
(1) Air oxidation 800-1080 (1) Sulfide oxidized corres-
(1472-1976) ponded to 0.9% wt of
stone
(11) Recarbonation 700 (1292) (11) 11% of CaO recarbonated
in 16.4 hours.
Recarbonation in 650-780
13% C02 (1202-1436)
4 Powdered CaCO. Calcined in Argon Recarbonation in 750 (1382) 91% CaCO, in 200 minutes
3 at lOK/mln to
780°C (1436°F)
100% CO, J
2
283 CAFB 7 regener- Air oxidized at Recarbonation in Scan 800-700- 20% CaCO, formed in 200
ator bed material 750°C (1382°F)
ground < 74 urn and at 800° C
(1472°F)
17% CO./N. 800 (1472- minutes. ESTIMATED MAXIMUM
1 i 1292-1472) RATE AT 750°C (1382°F)
285 Limestone 1359 Calcined in 17% Recarbonation in 750 (1382) 70% CaCO. formed
800-2000 pm CO /N2 at 850°C 17% CO,/N, J
(1562°F)
286 Limestone 1359
< 74 vm
L I.
" 750 (1382) 80% CaC03 (285 minutes)
287 " Calcined in 17% 750 (1382) 34.5% CaCO in 91 minutes
CO./N2 at 871 °C, J
(1660° F) sintered
70 minutes at 1100°C
(2012'F)
288 " Calcined in 14% Recarbonation 14.5% 750 (1382) %CaO recarbonated: 85%
CO /N at 871°C CO /3% HO in 70 minutes
(IBOOoF) L i
289 " Calcined at 871°C Recarbonation: 14.5% 750 (1382) Sulfidation to 31.4 % CaS
(1600°F): sulfided CO,/ 3% H.O after sul- % CaS oxidized - 50.2%
0.5% H2S in fuel fiae oxidation % residual CaO
gas; oxidized,
BOO°C (1472°F)
air, recarbonated - 77.21%
Wt% CaS 11.90
CaSO, 22.50
CaCO, 55.8
CaO J 9.22
imp .58
290
Limestone 1359
420-500 urn
Calcined, sulfided,
oxidized (3X) re-
carbonated
Recarbonation: 14.5% 750 (1382)
O>2/3% HO after multi-
pie sulfide oxidation
Sulfidation to 30.3% CaS
% CaS oxidized 80.6%
% residual CaO recarbonated
67%
347
-------
Regenerator bed material from run 7 on the Esso Research
Centre, Abingdon, England (Esso) pilot plant
Limestone and dolomite
Sulfided/oxidized limestone.
Regenerator Bed Material
The recarbonation of residual calcium oxide in CAFB regenerator
bed material from run 7 was attempted. Runs 248 and 249 showed that
about 10 percent of the residual calcium was recarbonated in a crushed
sample at 700°C (1292°F) in a 10 percent carbon dioxide/air mixture over
a 16-hour exposure. Reference to an earlier run on calcined powdered
calcium carbonate reagent (TG4) showed that 91 percent recarbonation could
be achieved in 200 minutes in pure carbon dioxide at 750°C (1382°F).
A sample of run 7 material was ground to less than 74 urn diameter
particle size, the air oxidized to convert residual sulfide to sulfate, and
the recarbonated in 17 percent carbon dioxide in nitrogen. The temperature
was raised through the interval (700 to 800°C/1292 to 1472°F) to find the
maximum reaction rate. After 200 minutes, 20 percent calcium carbonate
had formed, and the maximum recarbonation rate was estimated to occur at
750°C (1382°F), in agreement with other studies. It was concluded that
recarbonation of the regenerator bed material at high temperature is too
slow to permit a feasible process.
Limestone and Dolomite
The effects of stone particle size, stone sintering, and water
vapor on the recarbonation reaction were tested. Limestone 1359, < 74 ym,
was calcined by heating it to 870°C (1598°F) in 14 percent carbon dioxide
in nitrogen and then recarbonated at 750°C (1382°F) in the same gas, at
atmospheric pressure, in TG 286. The recarbonation was initially extremely
rapid, reaching 60 percent of theoretical within four minutes, as shown
in Figure 0-1. The rate slackened appreciably, however, and settled into
a slow phase. Here the fraction reacted depended linearly on the
348
-------
OJ
.7
n i« 3
O
H^
**. 4
s
=> ?
«J
t_j
_i
S-2
.1
0
0
LIMESTONE 1350
74 MICRON
CALCINED AT 871 C
RECARRONATED AT 750 C
16.5?? C02/N2
(WITH
KO
TIME/MINUTES
TG 28.
(SINTERED)
FIGURE n-I- RECARBOHATION OF CALCIUM OXIDE
-------
logarithm of the elapsed time. Projection of the data in the range 10 to
50 minutes up to 300 minutes accurately predicted the observed extent of
reaction (easily within 5 percent of the actual value). These data indicate
that 80 percent recarbonation is achieved with a residence time of the order
of two hours in excess carbon dioxide, but 90 percent conversion would
require about two days (44 hours).
Two additional runs were carried out on limestone 1359 (the
starting material for run 7), at 750°C (1382°F) and the conversion noted
was 70 percent calcium carbonate (< 420 ym particles) and 82.5 percent
calcium carbonate (< 74 ym particles).
Since the recarbonation of regenerator bed material (which has
been recycled between fluidlzed beds at 870 and 1080°C/1598 and 1976°F)
proceeded much more slowly than in the case of limestone calcined at 871°C
(1600°F), the effect of sintering was briefly investigated. In T6 287,
limestone 1359, < 74 ym was calcined in 17 percent carbon dioxide in
nitrogen and then heated in nitrogen to 1100°C (2012°F) and held at that
temperature for 70 minutes. The solid was then cooled to 750°C (1382°F)
and recarbonated in 17 percent carbon dioxide in nitrogen. As Figure 0-1
shows, conversion to carbonate was slow and incomplete (34.5 percent calcium
carbonate in 91 minutes), indicating that the high-temperature exposure had
produced a stone of similar activity to that obtained from the CAFB regenerator.
The effect of adding water vapor to the recarbonation gas was
tested in TG 288 after a preliminary test had shown that water vapor caused
a marked acceleration in reaction rate. The course of reaction as shown
in Figure 0-1 Indicated that, while the rate was much faster initially, the
Increased extent of recarbonation was about 6 percent once the slow
diffusion-controlled reaction rate regime had been established.
Oxidized Sulfided Limestone
In order to determine the Interaction between oxidation of calcium
sulfide to calcium sulfate and recarbonation of the residual calcium oxide,
two experiments were performed.
In the first experiment, limestone 1359 was sulfided to 31 percent
calcium sulfide, and the stone was then oxidized in air at 800°C (1472°F).
350
-------
Fifty percent conversion of the sulfide to sulfate was achieved. The stone
was then cooled, ground to less than 74 ym and recarbonated at 750°C
(1382°F), with 14.1 percent carbon dioxide in nitrogen saturated with
water at 20°C (68°F). Reaction was continued to 77 percent conversion
of the residual calcium oxide. It is interesting to note that the calcium
sulfate loading on the stone had only a minor effect on the rate of
recarbonation. The reaction rate during recarbonation is shown in Figure 0-2
and was used to outline the required dimensions for a fluidized bed
recarbonator.
In the second trial of the two stages, an attempt was made to
increase the extent of conversion of the calcium sulfide by successive
stages of cooling, grinding, reheating, and oxidizing. After the third
oxidation, when no further reaction was apparent, the stone was recar-
bonated, and 67 percent of the residual calcium oxide was converted to
calcium carbonate.
It can be concluded that a combined sulfide oxidation/oxide
recarbonation process would yield a finely ground stone composed
approximately of:
Component % by weight
CaS 12
22
CaC03 56
CaO 10
Further development of this process would require leaching tests
on larger samples than are produced in TG experiments. The purpose of the
leaching experiments would be to determine the rate of release of sulfur
from the residual sulfide and the overall mechanical stability of the
spent sorbent.
Reactor Design
The data from TG 289 were plotted in the form rate of recarbon-
ation versus extent of recarbonation. As Figure 0-2 shows, the rate was a
351
-------
10'
o
o
5
c
to-1
DC
"S
CD
*-«
oc
Curve 678263-A
T
TG289
Limestone 1359 «74mn)
Pretreated
(CaO).,67CaS, CaS04
Recarbonation
14%C02, 3%H20/N2
Atmospheric Pressure 760°C (1400°F)
1
l
1
1
I
0.10 0.20 0.30 0.40 0.50 0.60 0.70
Calcium Oxide Fraction Recarbonated
Figure 0-2-Rate of recarbonation of residual CaO in spent
sorbent
352
-------
logarithmic function of the extent of the reaction( y> between 30 and 70
percent recarbonation:
. _ [5.8736 - 12.593 y]
K e .
Using this rate, the retention of carbon dioxide in a fluidized bed of
calcium oxide particles smaller than 74 ym (delivered at a rate of 1.24 x
10 kg/68.1 Ib moles of calcium per hour) was calculated, assuming first-order
kinetics for gas absorption by the solid. From the carbon dioxide
retention, the required feed rate of carbon dioxide to the bed was
calculated, as shown in Figure 0-3. Up to 50 percent of the calcium
oxide can be recarbonated with virtually total retention of the carbon
dioxide in the bed. The extent of recarbonation for a 1 times stoichio-
metric feed to the bed is 62.5 percent. To achieve 80 percent recarbon-
ation, 8 times stoichiometric is required.
The effect of lowering the bed height is shown in Figure 0-4.
In the 60 to 70 percent recarbonation range, halving the bed height reduces
the extent of recarbonation by about 5 percent at a given carbon dioxide
feed rate.
The bed area required at 62.5 percent conversion is 4.57 m (15 ft)
in diameter, assuming the recarbonation gas is delivered at 128 kPa
2
(18.7 Ib/in ); the solids residence time is 16.87 hours.
STONE SINTERING
Since lime is normally produced from limestone as a chemically
active material for hydration to calcium hydroxide, the calcination process
is usually carefully controlled to ensure maximum stone activity. It is
known that calcination at temperatures above 1000°C (1832°F) may lead to
reduced activity of the product lime. Several studies have shown that
the loss in chemical activity is paralleled by a consolidation of the
pore structure in the lime and a growth in the crystallite size which
reduces the surface area of the solid. These changes are harnessed in
the high-temperature calcination and sintering of dolomite to produce
353
-------
Curve 678262-A
9.0
8.0
7.0
c
§ 6.0
& 5.0
O
*S
! 4.0
3.0
2.0
1.0
CaO1 1111
<74jjm, Fed to Bed at 5.73 moles sec"1
760° C
14.1%C02, H20, N2
0.5ft. sec"1 (Superficial Velocity)
7.25ft. (Expanded Bed Height)
at 50 % Voidage
First-Order Absorption of C02
Atmospheric Pressure Reaction
Rates fromTG-289
i
i
2 x Stoichiometric
for Extent of /
Recarbonation/
/
Stoichiometric
for Extent of
Recarbonation -
1
1
1
.1 .2 .3 .4 .5
CaO Fraction Recarbonated
.6
.7
Figure 0-3-C02 requirement for CaO recarbonation
354
-------
O)
o
"o
10
Curve 678261-A
10
10
0
\\ -^- 4x Stoichiometnc-
a=0.7
= 0.65 -
1 xstoichiometric
= 0.6
Fed to Bed at 5. 73 Moles sec
760° C, Atmospheric Pressure
15.24cm- sec"1 (Superficial Velocity)
-14. 1% C02 in N2 (H20)
First-Order Absorption of C02
Reaction Rate fromTG 289
a = CaO Fraction Recarbonated
i _ I _ i _ i
""1
I
0 50 100 150 200 250
Expanded Bed Height/cm
300
Figure 0-4-Effect of bed height on C02 requirement
for recarbonation of CaO
355
-------
dead-burned dolomite which is used as an inert refractory lining for
high-temperature process vessels. The idea that the CAFB spent sorbent
(which is already partly deactivated by exposure in the regenerator to a
temperature of 1080°C/1976°F), could be dead-burned has an immediate
attraction. Because sintering involves an almost homogeneous reaction
throughout the solid particle, it does not depend on gas/solid contact
to achieve some target for conversion of the solid. It thus differs
fundamentally from the dry sulfation or oxidation/recarbonation processes.
The process is primarily intended to apply to spent sorbent
from a regenerative process with a low-sulfur loading (< 1 wt % sulfur).
While the sulfur content would certainly be liberated as sulfur dioxide,
the quantity involved would not be large and could be recycled into the
gasifier. In a once-through process, the procedure would generate a high
concentration of sulfur dioxide, depending on the gas throughput required
to heat the stone to the sintering temperature.
Experimental Work
Two phases of experimental tests on spent sorbent sintering were
carried out. In the first phase tests designed to measure the activity
of the stone during hydration were investigated. In the second phase a
set of CAFB regenerator bed material samples were sintered at temper-
atures from 1200 to 1550°C (2192 to 2732°F) and tested for residual
activity during hydration.
Preliminary experiments are listed in Table 0-2. The procedure
was to add a known weight of the calcium oxide sample to 20 ml of deion-
ized water in a test tube held in a thermos flask containing water at room
temperature. The temperature difference between the slaking lime/water
mixture, agitated with a magnetic strirrer, and a reference thermocouple
held in another thermos filled with water, was monitored using Cr/Al ther-
mocouples connected with a sensitive amplifying voltmeter (Flukemeter) and
output to a chart recorder. The temperature difference slowly rose to a
maximum over a 15-minute period. Table 0-2 shows the maximum temperature
rise associated with hydration of CAFB run 7 regenerator bed material before
356
-------
Table 0-2
EXOTHERMIC HYDRATION OF WASTE LIME IN 20 ml WATER
Experiment
No.
Sample
Pretreatment
Temperature
recorded
Time
to
AT
max
AT
max
Min.
AT(C)per
gm mat'l
CAL 1 CAFB run 7
regenerator
bed material
CAL 2 CAFB run 7
regenerator
bed material
CAL 3 CAFB run 7
regenerator
bed material
CAL 4 CAFB run 7
regenerator
bed material
CAL 5 Raw Denbighshire
limestone
CAL 7 KOH pellets
Thermometer 12.7 15.5 4.12
Thermocouple 8.5 13.0 4.3
Thermocouple 13.89 12.8 4.6
Sintered 2 hrs Thermocouple 2.56
at 1450°C
(2642°F)
Thermocouple None
60.
0.84
Thermocouple 14.6 2. 5.82
357
-------
and after sintering in air at 1450°C (2642°F). Sintering apparently
decreased activity by a factor of five. As a measure of the heat liberated,
two experiments were carried out in which KOH pellets were dissolved in
20 ml deionized water. These result^ indicated a heat capacity for the
system water + thermocouple + stirrer of 167 J/K (^ 40 cal/K) or < 10
percent of the calcium oxide in the sample sintered to 1450°C (2642°F)
available for hydratlon.
2
Temperature rises measured by Murray in lime-slaking with seven
parts of water by weight to one of quicklime show that rises of ^ 3 °C
(37.4°F) can be expected for material calcined for 67 minutes at 1372°C
(25008F).
The second stage of preliminary experimentation consisted of
attempts to establish a base-line against which small temperature rises
could be measured.
An ice bath was chosen as the most convenient starting point, and
the temperature difference between the ice bath and 10 ml of water held in
a test tube in the ice bath was recorded for several overnight runs. The
10 ml of water was continuously agitated using a magnetic stirrer. It
was found that readings taken every hour over a ten-hour period were
remarkably constant, indicating a temperature difference of 0.196 + .006°C
(0.353 + 0.011°F). While this should give sufficient sensitivity to
measure a very slight degree of reaction, the experiments also have to be
conducted at higher temperatures, since lime hydrates more extensively the
higher the temperature of the reaction.
Experimentation was continued with a change in procedure:
the addition of water to a small quantity (t< 1.5 gm) of CAFB material in
a test tube in a water bath. The thermocouple was positioned in the
solid at approximately the center of the end hemisphere of the test tube.
The quantity of water added was varied from 0.2 ml up. It was found that
temperature rises of 182°C (363°F) could be achieved on addition of
*> 1.6 ml water. On passing through the 100°C (212°F) mark, there was
always a halt in the temperature rise, so the water addition was reduced
to 0.2 ml in order to keep the temperature maximum below 100°C (212°F).
358
-------
A set of CAFB samples from the regenerator bed were then sintered
in a platinum boat in a muffle furnace under flowing nitrogen. Temperature
was attained within 2.5 hours, held for one hour, and cooled to about
600°C (1112°F) over a four-hour period. The samples were then transferred
to desiccators. Testing the samples by addition of 0.2 ml of water gave
the results shown in Table 0-3. There is no significance in the
difference recorded for the samples in the range 1400 to 1550°C (2552 to
2822°F).
Table 0-3
TEMPERATURE RISE ON HYDRATION OF SPENT SORBENT
CAFB RUN 7 REGENERATOR BED MATERIAL
Sintering temperature, °C (°F) Temperature rise on hydration
No treatment 77.5 (171)
1200 (2192) 28.5 ( 83)
1400 (2552) 0.38 (100)
1500 (2732) 0.36 (0.97)
1550 (2822) 0.24 (0.75)
For the purposes of hydration, it can be assumed that sintering
at 1400eC (2552°F) is sufficiently high a temperature to render the spent
stone inert. The next phase of technical feasibility studies requires
leaching tests to determine the environmental impact of trace-element
emissions from the stone and to determine the long-term stability of the
sintered material. Further work is proceeding in these areas.
359
-------
REFERENCES
1. Drehmel, D.C. Test to Evaluate Reactivity of Boiler-calcined Limestone
Used in Air Pollution Control. Ceramic Bulletin, 50, 666. 1971.
2. Murray, J.A., H.C. Fischer, and D.W. Sabean. The Effect of Lime and
Temperature of Burning on the Properties of Quicklime Prepared from
Calcite. Proc A.S.T.M. 1263, 1951.
360
-------
APPENDIX P
SULFUR RECOVERY
SYSTEM DESIGN AND EVALUATION
-------
APPENDIX P
SULFUR RECOVERY
SYSTEM DESIGN AND EVALUATION
Sulfur can be recovered from the sulfur dioxide (SO.)-rich
off-gas from the regenerator in a number of forms: for example, sulfur,
sulfuric acid, liquid sulfur dioxide, gypsum. Sulfur was selected as
the preferred product on the basis of an assessment of the alternatives
presented in Appendix S. The economics of sulfuric acid production from
the off-gas was projected to compare with the sulfur production costs.
PROCESS OPTIONS
Several processes are available for converting sulfur dioxide
1 Q
to sulfur. Allied Chemical has developed a process for catalytic
reduction of sulfur dioxide to elemental sulfur using natural gas as the
reductant. This process has been and will be used to reduce sulfur
dioxide to sulfur from smelter plant off-gas and power plant stack-gas
cleaning process off-gases. The process consists of three main steps:
gas purification, sulfur dioxide reduction, and sulfur recovery. The gas
purification essentially removes excess water vapor as well as gaseous
and solid impurities. In the reduction step, nearly half of the
sulfur dioxide is converted to elemental sulfur by reaction with methane,
along with simultaneous formation of hydrogen sulfide (H.S). In the
sulfur recovery step, sulfur formed in the reduction system is
condensed out of the gas, and the remaining sulfur dioxide and hydrogen
sulfide are reacted in a multistage Claus conversion system to produce
additional sulfur.
The citrate system (Bureau of Mines) is another system for
removing sulfur dioxide and producing sulfur. Prolonged pilot tests have
shown that the Bureau of Mine's buffered sulfur dloxide-hvdroeen sulfide
process is capable of removing 95 to 99 percent of the sulfur dioxide
361
-------
remaining sulfur values not converted to elemental sulfur in the off-gas
treatment step. These gases are recycled to the front of the adsorber
where the sulfur values are adsorbed on the char. The carbon dioxide,
water vapor, and nitrogen are nonreactive materials at 170°C (350°F) and
therefore pass through the adsorber bed.
While the conversion of the concentrated sulfur dioxide to
sulfur offers the potential for a simplified process, it has not been
demonstrated on a commercial scale. The Bureau of Mines(Morgantown)
and Foster Wheeler are investigating the use of coal to convert sulfur
dioxide to sulfur. Their work is also in the development stage.
DEMONSTRATION PLANT
Technical and economic information was obtained on the recovery
of sulfur from the sulfur dioxide-rich off-gas produced in the regenerative
18
mode of operation. Several processes were considered; the Allied
Chemical direct reduction technology, however, is the only compatible
process which has been commercially demonstrated. Allied Chemical
Corporation provided the following information on its sulfur dioxide
reduction technology applied to the SO MW fluidized bed oil gasification/
desulfurization demonstration plant:
Preliminary flow diagrams with utility and raw material
requirements and effluent streams
Budget cost estimate capital and operating for the
sulfur recovery plant. Boundary conditions: sulfur-
containing gas received from waste heat boiler, Incinerated
tail gas, sulfur shipment facilities
Space requirements for gas cleaning, sulfur recovery, incin-
eration, sulfur storage, sulfur transport
Operating constraints, for example start-up, shutdown, load
follow.
Westlnghouse specified regenerator performance and a sulfur balance
for their technical and economic study which is summarized in Table P-l.
The report on the sulfur recovery plant is attached. Allied Chemical's
362
-------
TABLE P-l
DEMONSTRATION PLANT
(Projected Regenerator Performance Data)
Item
Component
SO,
co2
H20
°2
N2
Comment
Concentration (% by vol.)
8 (6-10%
4
2
< 0.1
85.9-86
range)
Sulfur Dioxide Output
Sulfur
Exit Gas Temperature
Exit Gas Particulate Loading
Gas Pressure
180-725 kg/hr (400-1600 lb/hr)a
(4/1 turndown)
90-360 kg/hr (200-800 Ib/hr)
1070°C (1960°F)
3
.05-.09 grams/m (0.02-0.04 gr/scf)
might be achieved; 2.3-4.6 grams/m3
(1 to 2 gr/scf) may be emitted
depending on the particulate removal
following the regenerator.
Controlled pressure close to atmospheric
(estimated to be 130 to 250 mm H.O or
5 to 10 inches H0) *
uaily sulfur to sulfur recovery plant may range from 8.7 Mg/day (9.6 tons/day)
to zero; typical sulfur balance around the plant for regenerative operation
is: 380 kg sulfur/hr (840 Ib sulfur/hr) input from oil 12,700 kg/hr (28,000
Ib/hr) 9.1 kg sulfur/hr (20 Ib sulfur/hr) leaves with spent stone, 350 kg
sulfur/hr (780 Ib sulfur/hr) goes with the regenerator off-gas and 18 kg
sulfur/hr (40 Ib sulfur/hr) leaves in the stack gas.
363
-------
estimate for the direct cost is $2 million. Westinghouse estimates the
total installed cost, including indirect costs, will be approximately
$3 million or $60/kw for the 50 MW demonstration plant. The cost of the
sulfur recovery system can be projected based on a 0.5 exponent with
sulfur dioxide concentration and capacity. Allied Chemical personnel
considered this a reasonable basis for cost projection up to the equivalent
of ^ 600 MW plant size.
SULFUR ACID PLANT
Costs were projected for a sulfuric acid plant in order to
compare them with conversion to sulfur. These projections, based on data
published in June of 1969 by Lurgi Corporation in Chemical Engineering
Magazine, indicated that a "battery limits" investment of about $1 million
in 1974 would be required for the production of nearly 9072 Mg (10,000 tons)
per year of 98% sulfuric acid from sulfur dioxide-containing gas. When
off sites investment for the supply of process water, cooling water, and
electric power are added to the "battery limits" investment, and costs for
piping the dilute sulfur dioxide stream to the acid plant are considered,
the total installed cost would approach $1.5 million for a sulfuric acid
plant. This acid plant cost, when reviewed briefly by Allied Chemical
personnel, was considered to be a minimum cost. The size exponent for
capacity of sulfuric acid plants was considered by Lurgi to be 0.65.
364
-------
REFERENCES
1. Bischaff, W.F., FW - BF Dry Adsorption System for Flue Gas Clean-Up.
U.S. Environmental Protection Agency Flue Gas Desulfurization Sym-
posium, May 1973.
2. Steiner, P., H. Tuntgen, and K. Knoblanch. Process for Removal and
Reduction of Sulfur Dioxides from Polluted Gas Streams. Presented at
the 16th National Meeting of the American Chemical Society.
3. Rosenbaum, J.B., et al. Sulfur Dioxide Emission Control by Hydrogen
Sulfide Reaction in Aqueous Solution. The Citrate System. Bureau of
Mines Report No. 7774, 1973.
4. Wright, J.P. Reduction of Stack Gas S0« to Elemental Sulfur. The
Journal of World Sulfur, May/June 1972.
5. Hunter, W.D. and J.P. Wright. S0« Converted to Sulfur in Stack-gas
Cleanup Route. Chemical Engineering, Oct. 2, 1972.
6. Haas, L.A. Sulfur Dioxide: Its Chemistry as Related to Methods for
Removing It from Waste Gases. Bureau of Mines Report No. 8608, 1973.
7. Hunter, W. D. and A. W. Michener. New Elemental Sulphur Recovery
System Establishes Ability to Handle Roaster Gases. Engineering and
Mining Journal. 1972.
8. Hunter, W. D. Application of SO, Reduction in Stack Gas Desulfuriza-
tion Systems. Allied Chemical. (Flue Gas Desulfurization Symposium.
New Orleans. May 14-17, 1973.) p. 15.
365
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ALLIED CHEMICAL SO REDUCTION
TECHNOLOGY
APPLICATION IN WESTINGHOUSE
ELECTRIC CORPORATION
OIL GASIFICATION DEMONSTRATION PROGRAM
ALLIED CHEMICAL CORPORATION
INDUSTRIAL CHEMICALS DIVISION
SULFUR POLLUTION CONTROL SERVICES
P. O. BOX 1139R
MORRISTOWN, NEW JERSEY 07960
January 15, 1974
367
-------
INTRODUCTION
Presented herein is technical and economic information
pertaining to Allied Chemical Corporation technology for
direct reduction of S0~ gas to elemental sulfur. The basis for
this presentation is tfie letter from Westinghouse Electric
Corporation's Dr. D. L. Keairns to Mr. W. D. Hunter, Jr.,
dated November 19, 1973.
This information is intended for use by Westinghouse
in connection with preparation of preliminary design and
cost estimates (Phase IA) of a fluidized bed oil gasification
system for retrofit on the 50 MW unit #12 at the Manchester
Street Plant of New England Electric System.
The SO? reduction system in this application will recover
sulfur dioxide, in the form of elemental sulfur, from the
off-gas produced in the calcium sulfide regenerator in the
oil gasification system.
368
-------
PROCESS DESCRIPTION
In the application of Allied Chemical's SO- reduction
technology to various emission sources, one requirement is
that the entering gas stream must be free of all particulate
matter and sulfuric acid mist. Sulfuric acid mist formation
generally occurs when a process gas stream contains both
SO, and water vapor. Some SO, is formed under certain con-
ditions when oxygen is present in the S02 containing gases.
The conventional method of eliminating particulates and
acid mist is by wet scrubbing. This adds to the cost of
the overall plant and creates another problem in the dis-
posal (neutralization) of a weak sulfuric acid. A favorable
feature of the Westinghouse oil gasification process is
that the off-gases from the CaS regenerator contain very
little oxygen, and consequently, the formation of SO3 by
the oxidation of S02 at high temperatures is not expected.
This offers the possibility of incorporating bag filters to
achieve efficient particulate removal in place of a wet
purification system once the gases are cooled to the
proper level. It is assumed, therefore, that the gases
entering the SO, reduction system are clean and at a tem-
perature of 300*F.
The SO, is forced through the reduction system by the
main blower (1). The reducing agent (natural gas) is intro-
duced at the blower discharge. The reaction between S02
and the reducing gas is exothermic, but because the SO2
concentration in the feed gases is relatively low, some
oxygen (air) is admitted at the blower inlet to maintain
a thermal balance. The feed gas mixture is then passed
through the feed gas preheater (2) where its temperature
is raised above the dew point of the sulfur formed in the
primary reactor system (3) .
The principal function of the catalytic reduction sys-
tem is to achieve maximum utilization of the reductant
while producing both sulfur and H2S, so the H2S/S02 ratio
in the gas stream leaving the system is essentially that
required for the subsequent Glaus reaction. Although the
chemistry of the primary reaction system is extremely com-
plex and includes reactions involving 11 different elements
and compounds, it may be summarized in the following equations:
369
-------
2S02 > C02 + 2 H20 + S
Because of the exothermic nature of these reactions,
the gases leave the reduction reactor system at a substan-
tially higher temperature. A portion of this heat is used
to preheat the incoming gases in feed gas heater (3) as
shown in the flow diagram Figure 1.
The elemental sulfur that is formed in the primary
reactor system is condensed in a horizontal shell-and-tube
steaming condenser (4). This represents over 40% of the
total recovered sulfur. The process gas stream then enters
the first stage (5) of a two-stage Claus reactor system
where the following exothermic reaction occurs:
2 H2S + S02 V 3/2 S2 + 2 H20
After the first stage of Claus conversion, the gas is cooled
and additional sulfur condensed by passage through a vertical
steaming condenser (6). Further conversion of H2S and S02
to sulfur takes place in the second stage Claus reactor
(7). This sulfur is condensed in a third steaming unit (8).
A coalescer (9)/ containing a mesh pad, then removes en-
trained liquid from the gas stream. Molten sulfur from the
three condensers and the coalescer is collected in a sulfur
holding pit (10) from which it is pumped to storage. Res-
idual H,S in the gas from the process is oxidized to S02
in the presence of excess air in an incinerator (11) before
being exhausted to the atmosphere.
INVESTMENT COST
The estimated total direct cost given in Table II is
order of magnitude quality, based on December 1973 prices,
and is for a battery limits unit. A process building,
sulfur storage tank and truck loading station are included.
Indirect costs are not included, i.e., engineering,
construction supervision, temporary construction facilities,
purchasing, license fees, etc.
370
-------
Offsites facilities are not included; i.e., steam generation
and boiler feedwater facilities, electrical substation, admini-
stration and service facilities, fire protection, site
development (piling, roads, railroads, fencing), sewers,
process control room, laboratory, pipe bridges, yard lighting,
etc.
OPERATING COSTS
The operating costs relating to natural gas are based on
incineration of the gases from the SO? reduction system at
1250°F to assure complete oxidation of all sulfur compounds
to SO2- When the percentage of inerts in the feed gas is
high, as from Westinghouse regenerator, the fuel cost associated
with incineration represents a substantial part of the total
natural gas cost. Some reduction in this fuel cost may be
realized by incinerating at 900°F, if trace quantities of some
sulfur compounds other than S02 (less than 500 ppm) can be
tolerated. It is believed that substitution of fuel oil for
natural gas in the incinerator may be preferred, and this is
an acceptable alternative for the incinerator fuel requirement
shown in Table I as natural gas.
OPERATING MANPOWER
The operating manpower requirement, shown in Table II,
is 1 1/2 men per shift. Actually, one operator controls
the process. Only a fraction of an additional man's time is
needed for inspection of operating equipment in the field. It
is assumed that adequate manpower performing other duties
in adjacent areas will be available to serve this purpose.
SPACE REQUIREMENTS
For battery limits S02 Reduction unit, including start-up
heater, sulfur pit, sulfur storage tank and truck loading
station:
80' x 90' = 7200 Sq. Ft.
Max. Ht. = 45 Ft.
Building size
60* x 60' x 45' (within above area)
Roadways are not included in the above space requirements.
371
-------
PRESSURE DROP
Pressure drop through the unit is estimated at approximately
5 psig. Incinerator outlet gas pressure has been assumed to
be plus 0.5 psig.
TURNDOWN
The unit is capable of turndown to 1/4 of rated capacity.
START-UP, SHUTDOWN
Estimated on stream factor: 90%
Estimated time to bring unit on stream following a
cold shutdown: 80 hours
Estimated time to bring unit on stream following a hot
"stand-by" shutdown: Immediate
Estimated time to take unit off stream: Immediate
372
-------
TABLE I
PROCESS AND OPERATING DATA
FEED GAS TO REDUCTION UNIT
Equivalent LT Sulfur/DAy in SO2
Composition % by volume, wet basis
SO, 8.0
CO, 4.0
H20 2.0
°2 °-1
Nj (balance) 85.9
Temperature °F
Gas Volume SCFM (60°F, 1 atm)
SULFUR PRODUCTION @ 95% recovery LT Sulfur/Day
TAIL GAS
% S02 by volume, wet basis
Volume SCFM (60°F. 1 atm)
Temperature, exit incinerator °F
UTILITIES - IMPORTED
Natural Gas (as 100% CH ) M SCF/Day (60°F, 1 atm)
Process Reductant
Incinerator Fuel @ 1250°F
@ 900°F
For "Hot Stand-by" during plant shutdowns
Boiler Feed Water (130 psig) #/hr
Cooling Water (90°F) g/hr
Auxiliary Steam (50 psig) #/hr (when unit shutdown
only)
Instrument Air (100 psig) SCFM
Electricity
8.58
Motors HP
Other (lighting, etc.)
KW
UTILITIES - EXPORTED
Steam (15 psig) #/hr
Cooling Water #/hr
Condensate #/hr
Boiler blow down #/hr
300
1970
8.15
0.31
3440
1250
133
113
73
35
1975
1500
550 (max.)
100
65
20
1600 (max.)
1500
500 (max.)
180
373
-------
TABLE II
INVESTMENT AND OPERATING COSTS
INVESTMENT COSTS - Total Direct
OPERATING COSTS (Basis 330 day operation)
Operating Labor 7 shift men plus 1 utility man
on days, $6/hr. Total 8 men
Supervision 1 operating supv. @ $19M
1 technical supv. @ $19M
Tests and Inspection - Included in above
Operating Supplies 10% of (Oper. & supv. labor)
Normal Maintenance 7% of Total Direct Cost
50% labor/50% materials
$2,100,000
Annual $
99,800
38,000
0
14,000
147,000
Utilities
Electricity @ 9 mills/KWH 4880
Natural Gas as 100% CH, @ 80C/MSCF
Reductant 35,120
"Hot Stand-by" Allowance (35 days/yr.) 980
Incineration Fuel as 100%
CH4 @ 80C/MSCF 29,840*
Auxiliary Steam @ 65C/Mlbs. (during shutdown only) 300
Instrument Air 270
Boiler Feed Water @ 25C/M Ibs.
(less condensate return @ 12 1/2$/ M Ibs.) 3,420
Cooling Water - No net consumption 0
Total 74,810*
Overheads @ 63.5% of all labor 134,200
TOTAL OPERATING COSTS 507,810*
NOTE: Excludes depreciation, taxes, insurance, start-up
costs, any equipment revisions and debugging
BY-PRODUCT CREDITS
Sulfur @ $25/long ton 67,200
L.P. Steam @ 30C/M Ibs. 3,800
NOTE: *Incineration Fuel Cost and Total Operating Costs will
be reduced by $10,570/yr. with 900°F Incinerator operation.
374
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ALLIED_CHEMICALS09 DEDUCTION-TECHNOLOGY
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-------
APPENDIX Q
LIMESTONE SELECTION
Q
-------
APPENDIX Q
LIMESTONE SELECTION
INTRODUCTION
The fluidized bed oil gasification process requires a supply
of limestone suitable for desulfurizatlon. Selection of the best material
for use at any particular plant site requires definition of the criteria
to be met by candidate stones in the light of current Esso Research Centre,
Abingdon, England (Esso) experience, and location and testing of available
stones. Testing the stones is subject to the development and refinement
of the selection criteria.
The stone selection process described here focused on finding
limestone suitable for use in the Providence, Rhode Island area but should
be typical of the stone selection process for almost any area. Most of
the fossil-fuel power plants in the eastern United States are located
near limestone-or dolomite-producing counties; the Carolinas and the
New England states, however, are exceptions. The search for a suitable
limestone is correspondingly more difficult in these areas.
INITIAL CRITERIA
As a result of batch reactor and pilot-plant experience with the
stones listed in Table Q-l, Esso laid out the initial criteria for stone
selection.
Table Q-l
LIMESTONES TESTED BY ESSO
Limestones |
Denbighshire
Limestone 1359
Limestone 1691
Limestone 1690
| Location |
Wales, U.K.
Virginia
New York
New York
Wt % CaO
55.2
54.1
45.0 -
26.7
49.1
377
-------
The primary emphasis was on purity: the calcined stone should have a calcium
oxide content of more than 50 percent, as shown in Table Q-2, in accordance
with the Esso finding that "the purer the stone is the better it performs"
(in the pilot plant). In addition, it was thought that very fine-grained
limestones should be ruled out, in view of a possible tendency to
generate excessive fines in the fluidized beds.
Table Q-2
LIMESTONE SPECIFICATION21
Component
CaO
MgO
sio2
M00_
2 3
Concentration
> 50
< 3
< 10
< 2
(wt %)
aJanuary, 1973.
LOCATION OF CANDIDATE STONES
The original schedule called for selection of five stones for
shipment to Esso at Abingdon so that batch unit tests could be performed.
The best stone was then to be tested in a continuous run on the Esso
pilot plant.
Geographical Location
The New England Electric System (NEES) plant in Providence,
Rhode Island is not typical of power plants in the eastern part of the
United States in that it is relatively remote from limestone- and
dolomite-producing areas. Quarries which produce limestone in this area,
Rhode Island, Connecticut, Massachusetts (with the exception of the Pfizer
Quarry in Adams), and Maine do not yield high-calcium limestone.
The possible sources of limestone may be grouped roughly as
follows:
378
-------
Sources in adjacent New England states
Sources in the Middle Atlantic states (New York, New Jersey,
Pennsylvania)
Long-distance suppliesfor example, Virginia, the Bahamas,
Canada.
Many sources were used to locate likely suppliers of limestone.*
Suppliers suggested by NEES
Suppliers listed in the Minerals Yearbook, U.S. Department
of the Interior
The New York office files of the Unites States Bureau of
Mines, Albany N.Y.J limestone quarries in the states of
Vermont, Massachusetts, Rhode Island, Connecticut
An EPA limestone inventory and suppliers list
Suppliers mentioned in reports by contractors on both the
Fluidized Bed Combustion and Coal Gasification Programs.
The stones which were eventually obtained are listed in Table Q-3,
which includes some stones not yet tested.
The location of the limestone quarries in relation to Providence
is shown in Figure Q-l and Q-2.
Testing
Two forms of screening were carried out on the stone samples
obtained. The reactivity of the limestones in a stream of fuel gas
containing 0.5 percent hydrogen sulflde (H.S) was measured in a thermo-
gravlmetric (TG) apparatus at 871°C (160CPF), and at atmospheric pressure.
At a later stage in the program, measurements of attrition losses during
calcination of the stones in a fluidized bed unit were carried out. Additionally,
samples of the raw stone were polished, etched, and photographed by
optical microscopy to ascertain the grain structure of the material.
* To these should be added, for future reference, the geological surveys
of any of the states in the vicinity of a plant site and the list of
commercial lime producers published in 1970 by the National Lime
Association, Washington, D.C.
379
-------
Dwg. 1671B61
Figure Q-1-Location of limestone supplies in relation to Providence, Rhode Island
380
-------
1672B62
Providence,
Rhode Island
Scale of Miles
0 100 200
Figure Q-2-Location of limestone deposits in eastern United States
381
-------
Table Q-3
CANDIDATE MATERIALS - COMPOSITION
Stone
Limestone 1359
Ocean Industries (aragonlte)
Limestone 1691
(Seneca Falls, N.Y.)
Conklln limestone
Lee limestone
Rockland limestone
Farber white limestone
Pfizer calclte
Pfizer dolomite
Vermont dolomite
(Denbighshire limestone)
De Vault dolomite
Bellefonte limestone
Origin
Virginia
Bahamas
New York
Rhode Island
Massachusetts
Maine
New Jersey
Massachusetts
Connecticut
Vermont
Wales, U.K.
Pennsylvania
Pennsylvania
% Ca %
% Ignition
MR wt. loss
38-39.5 0.04 43.4
38.9 0.04
33.1 0.2 38.6
24.6
27.9 5.8 42.5
29.5
^54
36.3 3.1
39.8
21.9
20.9 9
21.0 12
42.9
47.2
.7
.9
38.2 0.23
382
-------
Reactivity to Hydrogen Sulfide
The sulfidation of calcined limestones has been discussed in a
2
previous contract report. A description of the apparatus and procedure
is given in Appendix R.
The sulfidation of calcined limestone at the temperature of
oil gasification proceeds initially at the maximum rate permitted by mass
transfer from the gas phase to the solid surface. As the stone is sulfided,
the rate of reaction falls off in a manner which reflects the amount of
calcium available for reaction. Because the rate of reaction in the range
of interest can be closely simulated by a first-order equation, the
reactivity of a particular stone can be characterized by the time taken
to reach a particular extent of reaction at a fixed gas concentration.
Since the highest extent of sulfidation noted in the chemically active
fluidized bed (CAFB) process has been less than 30 percent, the time
taken to reach 30 percent sulfidation of 1200 to 1000 micron particles of
stone in 0.5 percent hydrogen sulfide in a fuel-gas mixture at atmospheric
pressure was chosen as an index of the stone reactivity. The results
obtained for a particular stone were repeatable, as shown in Figure Q-3.
Little variation in the results obtained was noted for the different
samples testeddolomites containing 21 percent calcium react as
rapidly as do limestones in this range. Further work to develop a more
meaningful reactivity test which can be used to project performance in
a fluidized bed is reported in Appendix R.
Attrition
The attrition of sorbent limestones or dolomites leads to loss
of fines from the fluidized bed reactors. The resistance of the stone
to attrition, therefore, is an important property of the sorbent in a
process such as fluidized bed oil gasification in which there is a high
recirculation rate of stone between gasifier and regenerator. The
maximum deleterious effect is noted where a significant percentage of
the stone breaks down into fines on calcination.
383
-------
oo
in
t 25
o
Q£
LU
C£
UJ
O
iI
X
O
=3
ii
CJ
20
15
in
t
DENBIGHSHIRE LIMESTONE
1200-1000 MICRON
CALCINED
SULFIDATION AT 87I°C
0.5% H2S/FUEL GAS
TG 255 +
TG 251 .
TG 252 *
TIME/MINUTES
FIGURE Q-3 REPEATABILITY OF STONE SULFIDATION KINETICS
-------
Some limestones dust to a fine powder so rapidly during cal-
cination that it is impossible to maintain a fluidized bed of the stone.
There is, however, an intermediate case in a once-through system without
regeneration where relatively high attrition rates may be helpful in
increasing the sorbent reactivity and utilization in desulfurization.
Fine sulfided particles from such a process would be easy to oxidize to
calcium sulfate for disposal. As will be discussed in more detail later.
Esso has noted that the attrition rate depends on the process occurring
in the fluidized bed. Loss rates from the Esso pilot plant and batch
reactors have been measured during calcination, kerosene combustion, and
oil gasification. The results have shown that losses under one set of
operating conditions do not correlate directly with losses under another
3
set of conditions.
Fundamental research on the relation between the properties of
carbonate rocks and their mechanical strength in different applications
such as when used as sulfur dioxide (SO.) sorbentsis in progress at
the Illinois Geological Survey, and some preliminary results have become
available. In this work attrition by decrepitation was tested by the
routine method used in the glass industry, and although the results were
in general agreement with data on the relative attrition in some batch
fluidized tests, no general correlation with the mineral, chemical, and
4
petrographic properties of the samples was apparent.
The approach adopted here was to calcine samples of carbonate
rock in a 50 mm (2 in) fluidized bed unit under the conditions listed In
Table Q-4.
The 50 mm (2 in) diameter reactor was run at approximately 1.4
atmospheres for the attrition studies. The reaction vessel, constructed of
Inconel 600 alloy, has a capacity of about 50 gm (0.11 Ib) of dolomite in
the heated zone. Although provision exists for metered introduction
of solids, batch reaction was used here. Fluidizing gas passes through
a preheating coll wound externally on the reactor, between the furnace
and reactor, so that it arrives at the distributor plate at temperature.
The distributor plate has 97 equally spaced 1.0 mm (0.040 in) diameter
385
-------
Pfizer (Adams) Calcite 50X
. vMJ^PIUPUl11 IT-:.
Rockland, Rockport Limestone 50X
Figure Q-4-Grain structure of candidate stones for the NEES plant
386
RM-62187
-------
Table Q-4
CONDITIONS FOR 51 mm (2 in) UNIT ATTRITION INDEX TESTS
Bed charge * 105 gm (0.23 Ib)
Particle size 18/40 (1,000 - 420 ym) (.039 - .0165 in)
Temperature of calcination 750°C (1382°F)
Time at temperature 'v. 100 minutes
Time to heat to temperature ^ 3 hours
Gas flow 26 1. min'1 N_ [or 2.6 x 10~2 m3. min"1 N ]
-1 -1
Superficial velocity 700 mm sec (2.3 ft sec )
holes. Fine material elutriated from the bed during a run was trapped
on filter paper capable of trapping particles (larger than 4 ym diameter)
and weighed. The percent weight loss from the bed (calculated on the
basis of fully calcined stone) was used as an index of resistance to
attrition; low losses from the bed are taken as indicating high
attrition resistance.
Since a number of the stones tested in the 51 mm (2 in)
unit were also tested on Esso batch reactors or the Esso CAFB pilot plant,
an indication of the usefulness of the attrition index can be obtained
by examining the results in Table Q-5. Here the indices of attrition are
noted and compared with the Esso experience.
There appear to be three ranges of decrepitation behavior:
low, high, and intermediate.
Low Attrition
Limestone 1359 gave the lowest attrition index, 0.3. Comparison
of the loss rates observed by Esso for four stones (BCR 1359, Denbighshire,
Pfizer Adams [designated NEES 1] and BCR 1691), under conditions of kerosene
387
-------
Table Q-5
COMPARISON OF ATTRITION INDEX FROM 51 mm (2 In) UNIT
AND RESULTS NOTED BY ESSO
Stone sorbent
Attrition index
(51mm unit)
Esso CAFB
experience
Limestone 1359
Pfizer (Adams) calcite 1921
Conklin limestone
Aragonite
Ocean Industries
(Bahamas)
Bahaman aragonite
Newton N.J.
Denbighshire limestone
Limestone 1359
0.3
6.7
13.9
6.0
0.7
6.6
0.3
Low attrition rate
Very high attrition (but
not by normal mechanism)
during cycling8
Very high attrition ;
Decrepitated badly
during calcination
Dust losses during
calcination small
During calcination had twice
the loss rate of 1359; on
sulfidation/regeneration
cycling had lowest attri-
tion rate of stones tested
Had lowest attrition rate
on calcination and a higher
rate than Denbighshire on
cycling
Craig, J.W.T., G.L. Johnes, G. Moss, J.H. Taylor, D.E. Tisdall, Monthly Report
on Contract 68-02-0300. No. 11 Esso Research Centre, Abingdon, England to
Environmental Protection Agency, June 1973. P. 7.
388
-------
combustion showed that BCR 1359 gave the lowest loss rate (the other
being R(1359) - 0.1 R (1691)
- 0.32 R (1921)
0.5 R (Denbighshire).
In gasification/regenerative cycles the rates observed by
Esso were:
R(1359) - 1.4 R (Denbighshire)
= 0.3 R (BCR 1921)
= 0.42 R (BCR 1691).
Thus, BCR 1359 showed the lowest attrition loss on calcination
in the 51 nun(2 in) unit and the lowest bed loss rate during kerosene
combustion. It was slightly more attrition-prone than Denbighshire
stone in gasification/regeneration cycles.
It may be concluded that stones with an attrition index as low
as around 1.0 are prime candidates for the fluidized bed oil gasification
s -
process if they fulfill other criteria.
High Attrition
Some of the samples, for example Lee limestone, exhibited an
attrition index >10.0. This high loss rate should exclude consideration
of the stone as a candidate for the CAFB process. It should be pointed
out that the calcination procedure in the 31 mm (2 in) unit where the
stone is heated to 750°C (1382°F) over a two-hour period is mild in
relation to the thermal shock experienced by injection of cold stone
into the CAFB fluidized gasifier at 871°C (1600°F). The only direct
comparison possible with CAFB experience is the Conklin limestone, which
gave attrition indices of 14.0 and 6.0 in two tests, and dusted so badly
in the CAFB batch reactor that it was rejected for further testing on the
pilot plant.
Intermediate Attrition
Results of attrition measurements which lie in the third
interval, 1 to 10, are ambiguous. Thus, Denbighshire limestone and Pfizer
389
-------
calcite (Adams) gave similar results, 6.0 and 6.7, respectively. Den-
bighshire limestone has been, to date, the most satisfactory limestone
tested in the CAFB process; the Pfizer calcite was unsatisfactory because
of the elutriation of fines and a rapid reduction of mean particle size '
in the fluidized bed. The relative performance of these stones, however, as
Esso has noted, was highly dependent on the process in the fluidized bed;
under conditions of fuel oil combustion, the Pfizer calcite was marginally
more attrition resistant (0.83) than the Denbighshire stone; and under
conditions where kerosene was combusted, the order was reversedDen-
ibighshire stone showing 0.65 times the loss rate of the Pfizer calcite.
Esso has suggested that the deteriorating performance of the calcite
is due to the thermal shock experienced under cycling between gasifier
and regenerator.
While the results to date are not unambiguous, it is thought that
stones with a high attrition index should be rejected for further testing;
stones with a low attrition index should be favorably considered; and
stones with an intermediate attrition index should be considered for
further testing.
The measurements of attrition in the 51 nun(2 in) unit should
be extended to cover a wider range of stones and operating conditionsfor
example, sulfidation and sulfidation/oxidation cycling.
Grain Structure
The samples of the candidate limestones were mounted in plastic
resin, polished, etched, and photographed. The grains in limestone were
evident after the stone had been etched in dilute acetic acid. Some of
the dolomites required a stronger solution of 10 ml acetic acid, 10 ml
phosphoric acid, 5 ml lactic acid, and 75 ml water to develop the image
of the grains. The longest linear dimension of the grains was taken
to Indicate the grain size of the material. Figure Q-4 shows the massive
grain structure of the Pfizer (Adams) calcite and the finer structure
of Rockland limestone. Comparison of the numbers with quantitative
grain size measurements at TVA showed sufficient agreement to regard
the procedure as a guide to grain size.
390
-------
Stones Sent to Esso for Further Testing
The stones shipped to Esso for testing on their batch reactors
were:
e Pfizer calcite (Adams)
Pfizer dolomite (Canaan)
o Ocean Industries Aragonite
Conklin limestone.
The Pfizer calcite suffered high attrition losses in gasification/
regeneration cycles, and Esso tentatively modified the criteria to exclude
highly crystalline limestone from consideration. Pfizer dolomite was
shipped because it showed adequate reactivity on the TG sulfidation test,
is quarried around 100 miles from Providence, and had been reported by
Consolidation Coal Company (Consol) as the most attrition-resistant
stone they had tested in their bench-scale sulfidation/regeneration
unit. It was later noted, however, that samples of this stone crumbled
to a fine powder after sulfidation on the TG apparatus, and Consol
confirmed that the stone had poor attrition resistance in the fully
calcined state. Apparently its attrition resistance is high only
when it is used as half-calcined dolomite (CaCO /MgO). Esso was
advised not to proceed with batch testing of this stone.
The Conklin limestone was shipped because of the proximity
of the quarry to the NEES site. Despite its unpromising composition and
attrition index, it was thought necessary that its potential usefulness
be accurately known. It was then rejected on the basis of high
attrition losses in batch tests. The aragonite from the Bahamas was
shipped to Esso, gave satisfactory results on the batch reactors, and
is scheduled for further testing on the continuous pilot plant.
Tymochtee dolomite was also tested on the batch reactors. An assessment of
the results will permit a decision on the possibility of using dolomite
rather than pure limestone in the CAFB process. The entire stone
selection results are summarized in Table Q-6.
391
-------
Table Q-6
SUMMARY OF CANDIDATE STONES FOR THE NEES PLANT
Stone
% Ca
Pfizer calcite, 39.9
Adams, Mass.
Aragonite 38.1
Ocean Industries
(Bahamas)
Farber limestone 36.4
Franklin, N.J.
Rockland limestone 29.5
Rockport, Me.
Lee lime 27.9
Lee, Mass.
Conklin limestone 24.6
Lincoln, R.I.
Pfizer dolomite 21.9
E. Canaan, Conn.
Vermont dolomite 20.9
Florence, Vt.
Denbighshire 39.8
limestone
Wales, U.K.
Limestone 1359 38.7
Stephens City, Va.
Reactivity Grain size Attritl
T33% CaS (um) (% wt
8.5 Coarse 6
10.2 600
8.0
11.0 Coarse 24
600
8.2 Medium 18
140
8.5 Medium 48
40-100
Medium 11
63-250b
20.0 Coarse0
250
8.0 Coarse
600
7.0 6
on index
loss) Action basis Comment
.7 Batch tested at Severe attrition on
Esso plant
Shipped to Esso
.9 Rejected
(attrition, purity)
.22 Rejected
(attrition, purity)
.4 Rejected
(attrition, purity)
.5 Shipped to Esso Rejected because of
high attrition
Shipped to Esso TG indicated high
but not tested attrition in full-cal-
cined state; confirmed
by Consol
Rejected (marble)
.6 Standard, already
tested
Fine 0.31 Already tested Provides index for com-
4-63
parison of stone proper
at 33% CaS is the time required to react 0.1 gms sulfur/gin raw stone.
bHarvey, R.D. Paper presented at 3rd Limestone Symposium, St. Petersburg, Florida, December 1967.
CHatfield, J.D., et al. Investigations of the Reactivities of Limestone to Remove S02 from Flue Gas,
TVA Report to APCO, 1971.
-------
CONCLUSION
The results of the testing procedures have not, as yet, revealed
an outstanding candidate stone for the NEES plant. The stones which have
been tested can be ranked in order of preference:
Limestone 1359
Aragonite
Limestone 1691
subject to final reporting of the Esso tests on aragonite.
Efforts are continuing to locate and test additional candidate
materials. Table Q-7 lists suppliers who responded to a request for
samples for testing and whose samples are currently being tested or will
be tested. Probably the eastern half of Pennsylvania is the most
promising search area. It contains numerous high-calcium limestone
deposits, and the future Impact of demand for wet-scrubbing materials
(lime and limestone) in this area is lower than the national average.
Additional sources in Virginia should be examined in an endeavor to find
stones which have properties similar to those of limestone 1359. Finally,
the possibility of shipping stone from the coastal areas of Florida
should be investigated, since this may prove to be competitive with
shipping aragonite from the Bahamas.
The more general case of selecting candidate stones for any
area has benefited from the experience gained in the search for a stone
for Providence. The attrition and grain-size tests provide tools for
screening materials. The reactivity tests have been less useful to date
but may prove valuable at a later stage. Current Esso research on
increasing the extent of sulfidation achieved in desulfurization by
increasing the bed height is aimed at decreasing the calcium/sulfur molar
feed rate required for 90 percent sulfur removal in the process. Increasing
the extent of reaction will depend on sulfidation within the stone
structure; and the order of reactivity experienced may reflect the results
of TG investigations. The selection process will become easier if
forthcoming tests at Esso using Tymochtee dolomite prove successful,
since a wide variety of dolomites are mined in Pennsylvania, Ohio,
393
-------
Michigan, Indiana, and Illinois. Further selection criteria will be
developed as the waste stone processing systems are studied and as
leaching tests on material for disposal continue. Finally, the restriction
to high-calcium limestones may be relaxed if stone hardening processes,
such as those developed by Consol, can overcome the high attrition rates
experienced with impure stones.
Table 0-7
RESPONDENTS TO SAMPLING BEQUEST
Supplier
Location of stone
quarry or supplier
Nature of stone
6. & W.H. Corson, Inc.
The General Crushed Stone Co.,
M.J. Grove Lime Co.
Riverton Corp.,
Dixie Lime and Stone Co.
Premier Resources
Armco Steel
Plymouth Meeting, Pa. >97% CaCO
Easton, Pa. (Watertown
N.Y; Chattanooga,
Tenn.)
Middletown, Va.
Riverton, Va.
Sumterville, Fla. 96.5% CaCO,
Costa Mesa, Calif. 85% CaCO
Piqua, Ohio
394
-------
REFERENCES
1. Craig, J.W.T. (Esso Research Centre, Abingdon, England). Letter to
S.K. Batra (NEES). January 25, 1973.
2. Keairns, D.L., D.H. Archer, R.A. Newby, E.P.O'Neill, E.J. Vidt.
Evaluation of the Fluidized Bed Combustion Process: Vol.IV,
Fluidized Bed Oil Gasification/ Desulfurization. Environmental
Protection Agency. Westinghouse Research Laboratories. Pittsburgh,
Pennsylvania. EPA-650/2-73-048d. 96 pp. and appendices.
3. Craig, J.W.T., G.L. Johnes, G. Moss, J.H. Taylor, and D.E. Tisdall.
Chemically Active Fluid-Bed Process for Sulphur Removal during
Gasification of Heavy Fuel Oil Second Phase. Environmental Protection
Agency. Esso Research Centre. Abingdon, England. EPA-650/2-73-039.
November 1973. 204 pp.
4. Harvey, R.D. Fracture Surfaces of Carbonate Aggregates: Scanning
Electron Microscope Study. University of Illinois.(Reprinted from
Proceedings of Twentieth Annual Highway Geology Symposium. Urbana-
Champaign, April 17-19, 1969)(Illinois State Geological Survey
Reprint Series 1970V.) 20 pp.
5. Harvey, R.D., and J. C. Steinmetz. Petrography of Carbonate Rocks by
Image Analysis. (Proceedings of the 7th Forum on Geology of Industrial
Minerals, Tampa, Florida Department of National Resources, Division
of Interior Resources - Special Publication 17, April 28-30, 1971.)
9 PP.
6. Drehmel, D.C., and R.D. Harvey. Carbonate Rock Properties Required by
Desulfurization Processes. (Proceedings of the 10th Forum on Geology
of Industrial Minerals, Columbus, Ohio State University, 1974.)
7. Malhotra, R. and R.L. Major. Electric Utility Plant Flue-Gas
Desulfurization: A Potential New Market for Lime, Limestone, and
Other Carbonate Materials. Illinois State Geological Survey, Illinois
Minerals Note 57. June 1974. 19 pp.
395
-------
APPENDIX R
THERMOGRAVIMETRIC STUDIES
ON THE SORBTION OF SULFUR BY LIME
-------
APPENDIX R
THERMOGRAVIMETRIC STUDIES
ON THE SORBTION OF SULFUR BY LIME
INTRODUCTION
Desulfurization using lime during oil gasification is generally
represented by the reaction
CaO + H2S -» CaS + H20.
With this reaction to model the desulfurization process, thermodynamics
can be used to find the maximum sulfur retention in a fluidized bed of
lime. It should also be possible to define how closely a fluidized bed
of given dimensions and operating conditions approaches the thermo-
dynamic limits for sulfur retention by studying the kinetics of the
reaction as a function of the bed operating conditions.
A fundamental parameter of the fluidized bed system is the
calcium molar feed rate required to desulfurize the fuel gas generated
in the bed. As the calcium oxide (CaO) in the bed sulfldes, the rate of
sulfidation will decrease, thereby lowering the efficiency with which
the process prevents escape of sulfur (S). Unless the calcium oxide is
replenished, the rate of sulfidation will eventually fail to fix enough
sulfur to meet the process requirements (90 percent sulfur removal).
In experiments conducted by Esso Research Centre, Abingdon, England (Esso),
high calcium to sulfur mole ratios (about 15/1) have, of necessity, been
maintained in the fluidized beds while achieving high sulfur retention.
In contrast, previous thermogravimetric (TG) studies of the sulfidation
of calcium oxide do not show an appreciable decline in the rate of
reaction until about 30 percent of the available calcium oxide has reacted.2
If similar kinetics obtained in the fluidized bed, there should not be
397
-------
any decline in the efficiency of sulfur removal until the calcium molar
feed rate fell to less than 3/1. It follows that there is a distinct
difference between T6 results and fluidlzed bed results in regard to
calcium utilization. TG data predict that less calcium is needed in
the bed to desulfurize the oil than is observed in practice. In order
to probe this difference, TG experiments and calculations were carried
out on the sulfidation of dolomite.
EXPERIMENTAL WORK
The object of the experimental work was to find the factors
which would decrease the rate of sulfidation of calcium oxide. The
following factors were considered:
Stone particle size
Calcination history
Carbon deposition on the stone
Stone sintering
Exposure of the stone to alternate oxidative and
reducing-gas compositions.
The TG experiments are summarized in Table R-l.
The equipment used was a DuFont 951 thermogravimetric balance
2
which, modified as described by Ruth, was capable of operation in
flowing hydrogen sulfide (H.S) at elevated temperatures. The balance,
controlled by a DuPont 990 console, sends out a continuous electric signal
which provides a reading of the weight of the sample as a function of time.
In experiments reported here all work was carried out at atmospheric
pressure which in Pittsburgh is 97 ± 6 kPa (28.7 ± 0.2 in Hg). The fuel gas
used to simulate the reducing gasifier conditions contained 26 percent
hydrogen (H.), 10.2 percent carbon monoxide (CO), 2.54 percent methane
(CH,), 16.9 percent carbon dioxide (CO.), 44.4 percent nitrogen (N2).
Calcination Pretreatment
The majority of the TG experiments have been carried out on
limestones calcined by heating to temperature (871°C/1600°F) at a heating
398
-------
Table R-l
TG TESTS ON SULFIDATION OF LIME
TG no.
241
242
243
252
253
457
Stone
Denbighshire
limestone
1200-1000 vm
Denbighshire
limestone
1200-1000 vim
Denbighshire
limestone
1200-1000 urn
Denbighshire
limestone
1200-1000 ym
Denbighshire
limes tone
1200-1000 ym
Denbighshire
limestone
^ 3,000 ym
Pre treatment
Rapid calcination
Stone coated with
Apiezon N grease
Stone coated with
graphite
Predried at 500°C;
calcined in N?
Predried at
500° C; calcined
in N2
Calcined in N_
at 871°C
Reaction
Sulfidation
in fuel gas
0.5% H2S
Calcination
in nitrogen
Calcination
and sulfication
0.5% H2S
Sulfidation (0.5%
H S- with one oxi-
dation/reduction
cycle
Sulfidation 0.5%
H S in fuel gas
with seven oxida-
tion/reduction
cycles
Sulfidation in
0.5% H2S in
fuel gas
Temp.
871°C
(1600°F)
871°C
(1600°F)
871°C
(1600°F)
871°C
(1600°F)
871°C
(1600°F)
871°C
(1600°F)
Comment
Rate unaffected by prior
calcination rate
Calcination unaffected:
no visible carbon layer
noted
Sulfidation rate not
affected in range
(0-30% CaS)
See text
See text
Slow Sulfidation
-------
rate of 10°C (50°F) per minute in flowing nitrogen. The calcination
history of the stone has been shown to be exceedingly important for
reaction with sulfur dioxide, in that during the calcination process the
pore structure, in which product calcium sulfate (CaSO.) grows, is formed.
3
Experimental tests reported previously show that the calcination
pretreatment greatly influences the overall sulfur retention and the
calcium oxide utilization expected in fluidized bed reactors. A test
was carried out to see if the calcination of the limestone at rates
closer to those likely to obtain in the fluidized bed gasifier would
influence the rate of sulfidation of the product lime. In TG 241
Denbighshire limestone was calcined as rapidly as the Instrument heating
rate permitted, and calcination was complete within ten minutes. This
treatment proved to be without affect on the sulfidation rate of the
product.
Oscillation between Oxidizing and Reducing Conditions
Two tests were carried out on 1200 to 1000 ym samples of
Denbighshire limestone. The samples were dried at 500°C (932°F) to
prevent "popping" or decrepitation. By heating them in nitrogen at
10eC/min (18cF/min) to 800°C (1472°F) they were then calcined. Sulfi-
dation in 0.5 percent hydrogen sulfide/fuel gas was interrupted when
26 percent of the calcium had sulfided. After they had soaked in
3
nitrogen for ten minutes, air was admitted at 400 ml/minute (24.4 in /
min) for one to two seconds, resulting in a weight increase of 0.125 mg which
may be attributed to oxidation of calcium sulfide to calcium sulfate. An
eight-minute soak in nitrogen caused a decrease in weight greater than
the increase during oxidation by a factor of 1.3. This weight loss should
have been 1.33 times the earlier weight gain if calcium sulfide and calcium
sulfate combined to reject sulfur dioxide according to the equation
1/4 CaS + 3/4 CaSO^ -» CaO + S02
400
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Resulfidation of the solid was initially rapid, but within 0.5 minutes
it fell to the rate normally observed at that extent of utilization and
continued unexceptlonally to more than 80 percent sulfidation of the
calcium oxide.
In the next run interruption of sulfidation, followed by a
pulse of air, under the same conditions as previously, was repeated
seven times without deleterious effect on the rate.
The purpose of these experiments was to see if oxidation of
calcium sulfide in calcium sulfate could cause a gradual accumulation of
sulfate in the pores of the calcium oxide, thereby preventing access to
unreacted calcium sulfide. This process was expected to cause a gradual
decay in the reactivity of the solid. The observed rates of sulfidation
4
were marginally increased. Craig suggested that the cycle carried out
in the TG apparatus duplicates the entire regenerative process and that
surface calcium oxide was regenerated on passing from an air to a nitrogen
feed. This mechanism is in accord with the observation that the average
ratio of weight loss to weight gain in the oxidation/reduction cycle was
1.28 over six cycles. The postulated mechanism requires a ratio of 1.33.
If this mechanism is accepted as a working assumption, then it predicts
a loss of sulfur dioxide from the stone as it passes from the oxidizing
to the reducing zone of the gasifier. This loss of sulfur dioxide may
contribute to an overall rate of sulfur retention lower than that
expected from the absorption of hydrogen sulfide.
Carbon Deposition
Two experiments were carried out to determine the effect on the
sulfidation rate of,a coating of carbon on the outer surface of the stone.
The first experiment failed because coating the sample with Apiezon grease
failed to cause any deposit of carbon during calcination. In the second
experiment, the limestone was coated by mixing it with a graphite paste.
Although there was a definite black layer on the stone, no evidence of
a decrease in sulfidation rate was noted over the first 30 percent of
calcium oxide sulfidation.
401
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*>
o
s:
UJ
o
Ul
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a
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-------
Stone Particle Size
A large particle was selected from the Denbighshire limestone
6-16 mesh sieve fraction. The particle was roughly 3,000 pm in diameter.
On calcining, the stone was sulfided and the sulfidation rate
fell drastically after the first five percent of sulfidation, as compared
with the smaller particles (1,000 to 1,200 urn) which had been previously
tested, as shown in Figure R-l.
The data obtained in this run were entered into the first-order
model* of a fluidized-bed desulfurization process. The experimental
parameters used were from Esso batch tests using 3,000 to 1,000 um
particles of limestone. As Figure R-2 shows the fluidized bed results
are less optimistic than the TG data. They are bounded reasonably by
the TG results for the two extremes of particle size. The agreement
between TG data and fluidized bed results, however, is much poorer than in
the case of sulfur dioxide retention under oxidizing conditions.
Further work is required on this system since it is important
to devise a TG test which will accurately predict the fluidized bed
performance of a given sorbent. At the moment it is clear that routine
tests on candidate limestones for the oil gasification process should be
extended to include larger particle sizes. The range of fluidized bed
conditions for which test results are available should increase greatly
because of the extended bed-height capability coming on line at the Esso
plant and because of the 100 mm (4 in) fluidized bed unit which is coming
into operation at Westinghouse.
*The fluidized-bed model used is described and discussed in detail in
the companion two-volume report under this contract on fluidized bed
combustion.
403
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§
CO
100
.E 80
0>
s
a:
=3
CO
»*
o>
o
k_
o>
Q_
60
40
20
TG Data on
Sulfidation of Calcined
Limestone Applied to First-Order
Model of Fluidized Bed
o Esso CAFB Fluidized Bed Results
for Run 14 Using 13,000 -1,0001 nm Limestone
1
1
1
1
I
1
I
1
2.0 4.0 6.0 8.0 10.0 12.0 14.0 16.0 18.0
CaO Percent Utilization
Figure R-2-Comparison of Esso CAFB oil gasification desulfurization and the
projection of fluidized bed from thermogravimetric data
-------
REFERENCES
1. Craig, J.W.T., G.L. Johnes, G. Moss, J.H. Taylor, D.E. Tisdall.
Chemically Active Fluid Bed Process for Sulphur Removal during
Gasification of Heavy Fuel OilSecond Phase. Environmental
Protection Agency. Esso Research Centre, Abingdon, England.
EPA-650/2-73-039.
2. Ruth, L.A. Reaction of Hydrogen Sulfide with Half-Calcined Dolo-
mite. Thesis. Department of Chemical Engineering, The City College,
The City University of New York. New York 1972.
3. Keairns., D.H. Archer, R.A. Newby, E.P.O'Neill, E.J. Vidt. Evaluation
of the Fluidized Bed Combustion Process. Vols. I and IV. Environmental
Protection Agency. Westinghouse Research Laboratories. Pittsburgh,
Pennsylvania. EPA-650/2-73-048 a and d. December 1973.
4. Craig, J.W.T. Private communication. February 1974.
405
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APPENDIX S
SPENT LIMESTONE DISPOSITION
-------
APPENDIX S
SPENT LIMESTONE DISPOSITION
Implicit in the process development work has been the belief
that a satisfactory method of disposition of the effluent stone would be
available. Several alternatives were investigated:
Land disposal
Ocean dumping
Deep mine injection
Potential markets.
LAND DISPOSAL
Land dumping does not appear to be a long-range solution to the
problem of sorbent disposal. The dump area would not be usable for any
other purpose at any foreseeable time. As an interim measure, however,
while research to develop other methods continues, it may be an acceptable
route. In the vicinity of Providence, Rhode Island, three of six disposal
contractors are willing to undertake to remove approximately 36.3 Mg (40
tons) a day of dry by-product stone from the dry sulfatlon process at costs
in the range of from $2.50 to $3.81/Mg ($2.75 to 4.20/ton). Standard roll-off
type containers would be used to collect the sorbent on-site. The contain-
ers would be removed every two days. Preliminary indications are that no
problem exists with leaching of trace elements from the spent stone.
DEEP MINE INJECTION
Discussions were held with members of the Pennsylvania Department
of Environmental Resources, formerly the Department of Mines, regarding
the feasibility of injecting the waste material into abandoned mines. Even
407
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if the mine owner and/or the Department could be certain of the solubility
of the material, it would be difficult to obtain the necessary permits to
conduct deep mine injection activities.
OCEAN DUMPING
For locations in and near coastal areas, the possibility of
ultimately disposing of spent solid sorbent by ocean dumping exists. Such
disposal, however, is now covered by final regulations and final criteria
for the evaluation of permit applications established in accordance with
Title I of the Marine Protection, Research, and Sanctuaries Act of 1972,
Public Law 92-532 and Sections 403 (c) and 402 of the Federal Water
Pollution Control Act Amendments of 1972, Public Law 92-500. The basic
intent of these regulations is not to prohibit such dumping but to regulate
it strictly. There is also an International Convention on the Prevention
of Marine Pollution by Dumping of Wastes and Other Matter which was
ratified by the U.S. Senate in 1973 and which will come into force when
ratified by 15 nations. These regulations form the basis for the
operating program to enforce this convention.
As applied to industrial wastes, no material may now be dumped
into the territorial waters of the United States without an EPA permit.
It is possible to obtain the following classes of permits to cover the
disposal of spent sorbent from oil gasification processes:
General permit - for small quantities of nontoxlc waste
Special permit - for wastes containing amounts of certain
elements not in excess of prescribed limits except if such
wastes are or rapidly become harmless after dumping. Harm
includes rendering edible marine organisms unpalatable or
endangering health of humans or other life forms. The spent
sorbent would appear to come under the classification of a
material requiring special care.
Interim permit - for materials which do not meet the require-
ments of the regulations but for which a plan is under active
development to bring the dumping within the regulations.
408
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Research permit - to investigate the impact of spent sorbent
on the marine environment.
One other class, the emergency permit, is judged not applicable.
One serious restriction is the language of the act regarding
renewal of permits. There is no provision for indefinite extension of the
permits; they may be renewed upon application to and review by the
Administrator. Research permits last 18 months, special permits three
years; both may be renewed. Interim permits last no longer than one
year and cannot be renewed; instead, a new one must be obtained. There
is an implication to the general permit that it is to cover situations of
an intermittent nature. Overall, it would still seem possible to obtain
one or the other types of permits for a period of a few years while research
continues on developing alternative uses for spent sorbent.
The procedure for obtaining a permit is lengthy, however, as it
includes the following steps:
Accumulation of technical information to support the
application
Preparation of the application
Public notice via the newspapers, the Post Office, or to
any person, group, or state or federal agency requesting it
Notice to the water pollution control agency for the affected
state
Review by the Corps of Engineers, the Coast Guard, the
Departments of the Interior and of Commerce, and the Regional
Director of the NMFS-NOAA which has jurisdiction over fish
and wildlife resources of the affected state.
Holding hearings, if written requests from any person are
received to hold such hearings. Such hearings shall be held
no sooner than 30 days after receipt of the request. No
time limit is set on the length of the hearings, although
they are to be conducted expeditiously. The recommendations
of the presiding officer at the hearing shall be forwarded to
the appropriate Administrator within 30 days after adjournment
409
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of the hearings. The Administrator has 30 to 180 days to
take final action.
It is of interest that the act lists as alternatives to ocean dumping
landfill, deep-well injection, shallow-well injection, and spread of the
waste material over open ground. The latter amounts to land dumping.
Section 228.52 also provides that solid wastes of natural minerals or
materials compatible with the ocean environment may be generally approved
for ocean disposal if they meet criteria of insolubility and settling or
dispersal rates. Section 227.71 provides that the limiting concentration
of a constituent must not exceed one percent of that shown to be toxic to
certain sensitive marine organisms or otherwise detrimental to the
marine environment.
The oil to be desulfurized contains up to 400 ppm vanadium and
40 ppm nickel. It is expected that most of the vanadium and a substantial
fraction of the nickel will be picked up by the sorbent. In the regener-
ative version of the process, the makeup calcium/sulfur molar ratio is
less than 1.0. For an oil containing 3.2 percent sulfur, assuming complete
conversion to calcium sulfate, the metals will be concentrated by a factor
of 100/13.6 or about eight for a makeup ratio of approximately one. It
has been suggested that if, indeed, such metal capture is obtained, the
spent sorbent might be sold to a vanadium recovery operation. Alterna-
tively, one of the concerns will be specifying an adequate dispersal
method if ocean dumping is utilized.
To follow up this avenue of disposal, the Westinghouse Ocean
Research Laboratory has been brought into initial consultation on this
subject and has outlined a program which would essentially cover the
accumulation of technical back-up information. This would include toxicity
experiments designed or approved by EPA. The proposal is being reviewed
to assess the probability of its overall success as a viable alternative.
POTENTIAL MARKETS
In addition to exploring the above disposal techniques, the area
of by-product recovery was also investigated. Major attention was focused
410
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on locating suitable uses for gypsum, elemental sulfur, or sulfuric acid.
Sulfur Market
The removal of sulfur from fuels offers the possibility of
marketing it as sulfur, sulfuric acid, or in some other commercial form.
Considering the potential scale of the desulfurizatlon process, both
current and projected markets were investigated. The following table is
based on data from the Bureau of Mines.
Table S-l
SULFUR RESOURCES IN THE UNITED STATES
Native sulfur
Pyrite
Sulfide ores
Natural gas
Petroleum
Oil shale
Coal
Anhydrite and gypsum
Sea water
TOTALS
I Resources
548 (539)a
258 (254)
52 (51)
52 (51)
51 (50)
1,548 (1,524)
13,421 (13,209)
19,100 (18,798)
NAC
35,030 (34,426)
| Years supply
59b
29
6
6
6
168
1,453
2,069
3,855
Millions of Mg (millions of long tons).
bYear's supply at 1968 production rate of 9.23 x 106 Mg (9.085 x 106
long tons).
cNot available.
The table shows that the first five sources, which are relatively easily
exploited, can meet current needs domestically for 104 years. This is
extended for another 20 centuries with the aid of the energy in oil shale
and coal. Given a supply of energy, refractory sources such as gypsum,
and eventually even sea water, can be exploited.
411
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The next step is to compare these figures with the potential
production of sulfur from the 50 MW demonstration plant. This works out
to 8.94 Mg (8.8 LT) day or about 2500 Mg (2900 LT) year. Sulfur produc-
tion would approach one percent of the 1968 domestic rate when the
utility consumed enough fuel to generate 1700 MW. The demonstration plant
will, thus, be an insignificant factor in the sulfur or sulfuric acid
market. The minimum economic size for an acid plant is probably no less
than 25.4Mg (25 LT) acid/day.
Examination of the end-use pattern for sulfur reveals the
following conclusions:
87% consumed as acid
50% used to manufacture phosphate fertilizers
89% supplied by elemental sulfur of which 92% was Frasch sulfur
The second largest use of sulfur, acid for inorganic pigments,
was only about one-tenth the size of the fertilizer market.
Since essentially all of the Frasch sulfur comes from Texas and Louisiana,
this confirms that it is economical to ship sulfur rather than acid. This
supported the view that stone processing should favor sulfur production
rather than acid.
Because of the heavy dependence of the sulfur market on the
fertilizer market, a brief check of this aspect was made. The 1968
domestic demand was 3,152,000 Mg (3,475,000 short tons ) of phosphorus,
of which 76 percent was for fertilizer phosphorus. Since there are other
processes for extracting phosphorus from rock, these serve to limit the
price of sulfur. This means a limited product credit for by-product
sulfur over the long term.
A reasonable question is the stability of the fertilizer market.
The following table gives domestic and world reserves of phosphorus,
expressed in millions of megagrams (short tons).
412
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Table S-2
PHOSPHORUS RESERVES
Location
Marketable product,
million Mg
(millions short tons)
Phosphorus content,
million Mg/
(millions short tons)
Domestic U.S.
Known reserves
Potential reserves
TOTAL
Rest of world
TOTAL
7,162 (7,895)
49,370 (54,420)
56,532 (62,315)
939 (1,035)
5,245 (5,782)
6,184 (6,817)
13,543 (14,983)
19,777 (21,800)
The known domestic reserves will last 2200 years at the 1968 demand rate.
Taking into account the geographical distribution of phosphate rock, there
may be local shortages, but nationwide the supply is adequate for far more
than 22 centuries.
The demand for both sulfur and phosphorus is increasing, however.
There is both a population increase and a per capita Increase. Some uses
will grow with the population, which is estimated at 1.6 percent year.
Others will grow at a faster rate. The problem is what growth rate to use.
Use of even a modest rate like 4 percent/year would result in a cumulative
demand that would exhaust the known reserves of phosphate rock in 65 years,
and the total reserves in 110 years. Similar results would obtain for
sulfur, although it would last somewhat longer. The population itself
will have doubled in 44 years. It appears reasonable, therefore, to adopt
the following view:
Resources of sulfur and phosphorus are adequate for time spans
far beyond the life span of the proposed desulfurization
processes.
Except for local situations, it does not appear that a
significant economic credit will be obtained for by-product
sulfur. At best, these credits will merely offset the cost
of the process for producing such products.
413
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If the by-product sulfur is discharged to the environment, in
some suitable form, this should be done without regard for
future recovery of sulfur.
Local situations might permit economical marketing of the sulfur.
Both Du Pont and Monsanto were contacted regarding the potential of
utilizing sulfur or sulfuric acid produced from the Westinghouse process.
Both felt that conversion processes for sulfur dioxide to sulfur or
sulfuric acid would be expensive to Install and not practical on such a
small scale (i.e. 50-MW plant scale). Questions regarding the sulfur
purity levels also suggested its uncertain marketability.
This information was regarded as confirmation of the above
analysis of long-term markets. In effect, if a product is to be developed
from the spent sorbent, it is not sulfur or sulfuric acid.
Gypsum Market
Ideally, the effluent stone from the dry sulfation process, or
from an alternative process, would be 90 percent calcium sulfate. It is
possible this material would be a substitute raw material for gypsum
products. For additional information on current industry interest,
leaders in the gypsum products industry were contacted to determine if
the material produced by the oil gasification process could be used in the
manufacture of products such as wallboard and plaster. United States
Gypsum, Johns Mansvilie, and Flintkote have requested samples for testing
in order to evaluate the by-product material's compatibility with existing
materials of construction for gypsum products. Cinder Products Corporation
has requested a sample to test in the manufacture of cinder blocks.
The market potential has been discussed to some extent in
Appendix K. For further perspective, data from the Bureau of Mines showed
that domestic reserves of gypsum are estimates at 18.14 million Mg (20
million tons), or 2000 years' supply at the 1968 production rate of 9.07
million Mg (10 million short tons/year). United States demand was
14,137,000 Mg (15,600,000 tons); the balance was imported principally
from Canada and Mexico. Among the areas deficient in gypsum are the
414
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New England states. There are also enormous reserves of anhydrite and
by-product gypsum from phosphoric acid production.
The median growth rate for domestic gypsum demand is estimated at
30 percent/year, corresponding to a demand of 36.29 million Mg (40 million
short tons) in 2000. Hence, there is no threat of a shortage of gypsum.
Transportation, however, is an important factor. Overland routes are
forced by economics to be short. Where ocean transport is available,
however, gypsum can be shipped long distances: Nova Scotia to New York
City and Mexico to Seattle.
It appears the quality of the spent stone might permit utilizing
the stone as a raw material for gypsum products. To check out this
alternative, efforts are under way to prepare sample quantities of sulfated
stone to send to the above producers for evaluation.
Other Markets
Besides direct use as landfill, there are a variety of outlets
in the general construction field that offer market potential. The spent
stone will probably have 0 to 10 percent of the calcium as oxide in the
regenerative process and possibly two-thirds as oxide in the once-through
version. This can be blended with fly ash, which has a high silica content,
to make cement-type materials. These may be sintered and hydrated or simply
hydrated at ambient temperature. The object is to produce concrete,
aggregate, or other materials usable in large-scale construction as in
roads, dams, embankments, or sites preparation.
415
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REFERENCES
1. Staff Bureau of Mines, Bulletin 650. Mineral Facts and Problems,
1970 edition. Chapters on sulfur, gypsum, calcium, phosphorus, clay,
and sand and gravel.
416
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TECHNICAL REPORT DATA
(I'lt asc read tuitnictions on the rcitisc before completing)
1 BbPOHT NO.
E PA-650/2-75-027-b
3. RECIPIENT'S ACCESSION-NO.
4. TITLE AND SUBTITLE
Fluidized Bed Combustion Process Evaluation
Phase I--Residual Oil Gasification/Desulfurization
Demonstration at Atmospheric Pressure (Vol. n)
5. REPORT DATE
March 1975
6. PERFORMING ORGANIZATION CODE
7. AUTHORED. L. Keairns ,R. A. Newby,E. J. Vidt,E. P.
O'Neill, C. H. Peterson, C. C. Sun, C. D. Buscaglia, and
D H Archer
8. PERFORMING ORGANIZATION REPORT NO.
9. PERFORMING ORG \NIZATION NAME AND ADDRESS
Westinghouse Research Laboratories
Beulah Road, Churchill Borough
Pittsburgh, PA 15235
10. PROGRAM ELEMENT NO.
1AB013: ROAP 21ADB-009
11 CONTRACT/GRANT NO.
68-02-0605
12 SPONSORING AGENCY NAME AND ADDRESS
EPA, Office of Research and Development
NERC-RTP, Control Systems Laboratory
Research Triangle Park, NC 27711
13. TVPE OF REPORT AND PERIOD COVERED
Phase I; 5/73-12/74
14. SPONSORING AGENCY CODE
15. SUPPLEMENTARY NOTES
16 ABSTRACTThis volume contains appendices resulting from an evaluation of the atmos-
pheric-pressure fluidized bed residual oil gasification/desulfurization process, a
process referred to by its inventor (Esso Research Centre, Abingdon, England) as
the chemically active fluidized bed (CAFB) process. The CAFB produces a clean,
low heating value fuel gas for firing in a conventional boiler. The integrated process,
previously operated successfully in a 750 kW pilot plant unit, has demonstrated the
ability to meet environmental emission standards for sulfur oxides, nitrogen oxides,
and particulates. Work carried out under this contract was directed toward comple-
tion of a preliminary design and cost estimate for a 50 MW demonstration plant and
a 200 MW plant design and cost estimate. Several process and design options are
evaluated. Process flow diagrams, energy and material balances, equipment speci-
fic ations, vessel drawings, equipment arrangement drawings, a site plan, anelec-
trical one-line drawing, and utility requirements are presented for the recommended
process concept. Plant performance, environmental impact, and functional operating
conditions are presented and development requirements identified. Capital and oper-
ating costs are presented for the 50MW demonstration plant and for commercial
plants with capacities from 50 to 500 MW. Limestone sorbent support data is given.
7.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.lDENTIFIERS/OPEN ENDED TERMS C. COSATI Field/Group
Air Pollution
Combustion
Fluid Bed Processing
Residual Oils
Gasification
Desulfurization
Boilers
Nitrogen Oxides
Limestone
Air Pollution Control
Stationary Sources
CAFB Process
Particulates
13B 13A
21B 07B
13H, 07A
21D
8 DISTRIBUTION STATEMENT
Unlimited
19. SECURITY CLASS (ThisReport)
Unclassified
21 NO. OF PAGES
441
20 SECURITY CLASS (Thispage)
Unclassified
22. PRICE
EPA Form 2220-1 (9-73)
417
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