EPA-650/2-75-027-a
March 1975
Environmental Protection Technology Series
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EPA-650/2-75-027a
FLUIDIZED BED COMBUSTION
PROCESS EVALUATION
(PHASE I - RESIDUAL OIL GASIFICATION/DESULFURIZATION
DEMONSTRATION AT ATMOSPHERIC PRESSURE)
VOLUME I - SUMMARY
by
D.L. Keairns, R.A. Newby, E.J. Vidt, E.P. O'Neill,
C.H. Peterson, C.C. Sun, C.D. Buscagha. andD.H. Archor
Westmyhouso Research Lubornlones
Bculah 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 tields. These series are:
1. ENVIRONMENTAL HEALTH EFFECTS RESEARCH
2. ENVIRONMENTAL PROTECTION TECHNOLOGY
3. ECOLOGICAL RESEARCH
4. ENVIRONMENTAL MONITORING
5. SOC1OECONOMIC 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-a
11
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ABSTRACT
Westinghouse has carried out plant designs, cost estimates,
process evaluation, and experimental work on an atmospheric-pressure
fluidized bed residual oil gasification process for power generation.
This volume is a summary of the work which is the first phase of a
demonstration plant program. The procesr., 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 environmental emission standards for sulfur oxides, ni-
trogen 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 MWg 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 MWg 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
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,
X
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-
tlon/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 Mo. CPA 70-0 and as nart of a four-volume reoort, Volume iv
"Fluidized Bed "il Gasification/Desulfurination," in December 197°. under
2
contract *fi-02-0217. The nrevious work on residual oil gasification included:
• Concentual design and cost estimate of atmospheric-pressure
fluidized bed residual oil Rasification/desulfurization
system for utilitv 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 determine
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.
This volume (Volume I) contains a summary of the work performed
and program recommendations. Volume II contains appendices which provide
supplemental information and detailed back-up to support the summary and
recommendations.
vii
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TABLE OF CONTENTS
Page
I. SUMMARY AND RECOMMENDATIONS !
Conclusions 2
Market 3
Technology *
Economics *
Demonstration Plant Design and Cost -*
Development Requirements 7
Recommendations 7
II. INTRODUCTION 9
III. PROCESS FUNDAMENTALS 13
IV. ASSESSMENT 18
Market 18
Perspective 18
Market Factors 20
Technology 23
Environmental Impact 23
Present State of Development 25
Utility Applicability 27
Summary 28
Economics 29
Capital Costs 30
Operating Costs 38
V. PRELIMINARY DESIGN AND COST ESTIMATE 42
Background 42
Preliminary Design Scope 42
General Process Options 44
Initial Design Study 46
Selection of the Preliminary Design Basis 50
ix
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Table of Contents(Continued)
Preliminary Design 50
Design Basis 50
Energy and Material Balances 54
Demonstration Plant Design 55
Reaction System 56
Stone Processing System 64
Support Systems 69
Plant Turndown 69
Control of Coke Deposition 70
Plant Layout 70
Boiler Modifications 79
Demonstration Plant Performance 81
Energy Efficiency 82
Environmental Performance 83
Demonstration Plant Process Performance 85
Gasification 86
Regeneration 87
Sulfur Dioxide Absorber 87
Solids Transport 87
Cost Estimate 87
Demonstration Plant Cost Breakdown 88
Assessment of Plant Cost 92
Processing Options 94
Operating Cost 95
VI. DETAILED DESIGN, CONSTRUCTION, AND OPERATION 97
Scope of Work 97
Selection of Architect-Engineer 97
Detailed Engineering, Procurement, and Construction 97
Schedule 99
Cost Estimate 99
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Table of Contents (Continued)
Page
Shakedown and Commissioning Operation 104
Experimental Test Program 105
VII. COMMERCIAL PLANT DESIGN AND COST ESTIMATE 107
Basis 107
Design 108
Cost 109
Performance 109
Options 111
VIII. DEVELOPMENT REQUIREMENTS 112
IX. REFERENCES 115
LIST OF ABBREVIATIONS 117
xi
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LIST OF FIGURES
Page
1. Atmospheric Fluidized Bed Oil Gasification Power Plant 10
2. Sulfur Removal System Process Concepts 14
3. Power Plant Adder for Sulfur Removal Systems -*4
4. Oil Gasification Processing Concepts 45
5. Schematic Flow Diagram and Material Balance Base Case 57
6. Schematic Flow Diagram Material Balance Turndown Case and
Low Air/Fuel Ratio Case 59
7. Reaction System Oil Gasification-Desulfurization CAFB 61
8. L-l Gasifier and L-2 Regenerator 65
9. L-3 Absorber 67
10. Equipment Arrangement Section AA 71
11. Equipment Arrangement Section BB and CC 73
12. Equipment Arrangement Plan Upper Level 75
13. Equipment Arrangement Plan Lower Level 77
14. Detailed Design and Construction Schedule 100
15. Critical Path Diagram 1°1
xiii
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LIST OF TABLES
Page
1. Phases of CAFB Experimental Development H
2. Summary of Gasifier Conditions and Performance 15
3. Summary of Regenerator Conditions and Performance 17
4. Summary of Existing Oil- or Gas-Fired Boilers 19
5. Feedstock Selection 22
6. Environmental Impact Comparison 24
7. Present State of Development 26
8. Comparative Capital Costs for Sulfur Removal Systems 34
9. Comparative Operating Costs for Sulfur Removal Systems 35
10. Preliminary Design and Cost Estimate 43
11. Summary of Major System Costs in the Initial Study
Demonstration Plant 49
12. Summary of Environmental Performance of Chemically Active
Fluidized Bed (CAFB) Pilot Plant 84
13. Cost Summary (50 MW Pilot Plant) 89
14. Plant Cost Breakdown by Plant System 90
15. CAFB Demonstration Plant-Reaction System Process Equipment
Cost 91
16. CAFB Demonstration Plant-Stone Processing System (Dry
Sulfation) Equipment Cost 93
17. Activity Schedule 102
18. Plant Cost Summary (200 MW Commercial Plant) 110
xv
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ACKNOWLEPfi^FNTP
The results, conclusions, and recommendations presented in this
volume represent the combined. t:ork and thought of many persons at
Westinghouse and the Office of Research and Development (ORD), and
personnel at New England Electric System (NEES). Other ^RD contractors
have freely shared with us their ideas and results of their research and
development effort. Westlnghouse 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 Fewer 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. P. 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. Fangebrauck, Chief of the Demonstration
Projects Branch. Mr. S. K. Batra and Mr. p. Gendreau of >Tew England Power
Service Comnanv contributed to the demonstration plant design, provided
data on the Manchester Street Station, and participated in process
evaluation and selection. Dr. J. T
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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 efford. Many persons at the Westlnghouse
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.
xviii
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I. SUMMARY AND RECOMMENDATIONS
The atmospheric-pressure fluidized bed residual oil
gasification/desulfurization process (CAFB) produces a clean, low
heating-value fuel gas for firing in a conventional boiler. The inte-
grated process has previously operated successfully in a 750 kw pilot
plant unit, and the process has demonstrated the ability to meet environ-
mental emission standards for sulfur oxides, nitrogen oxides, and parti-
culates.
Work carried out under this contract was directed toward the
completion of a preliminary design and cost estimate for a 50 MW demon-
stration plant and a 200 MW plant design and cost estimate. Several pro-
cess and design options are evaluated. Process flow diagrams, energy and
material balances, equipment specifications, vessel drawings, equipment
arrangement drawings, site plan, an electrical one-line drawing, and uti-
lity requirements are presented for the recommended process concept.
Plant performance, environmental impact, and functional operating condi-
tions are presented and development requirements identified. Capital and
operating costs are presented for the 50 MW demonstration plant and for
commercial plants with capacities from 50 to 500 MW.
Work was also carried out to develop process design data and
criteria. Experimental tests were carried out to identify candidate lime-
stone sorbents (14 limestones were tested); to obtain data on sulfur re-
moval, sorbent regeneration, and spent stone processing (five process
concepts were tested); and to assess the environmental impact of the
spent sorbent.
Esso Research Centre, Abingdon, England (Esso) refers to the process as
the chemically active fluidized bed (CAFB) process.
Watts are used to express electrical equivalent and Btu to express ther-
mal equivalent throughout this report.
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Conclusions and recommendations are based on the results of the
work performed under this contract and on the projected costs and perfor-
mance of alternative systems for sulfur oxide emission control which are
available from outside sources. The systems considered to be alternatives
to the CAFB process are stack-gas desulfurization via wet scrubbing and
hydrodesulfurization (HDS) of residual fuel oil, including vacuum resid.
Unfortunately, comprehensive cost and performance data on these systems
are limited. The stack-gas desulfurization costs and performance are
based on the recent TVA report by McGlamery and Torstrick. The HDS
costs and performance are based on the Cities Service/Hydrocarbon
Research H-Oil process. The conclusions and recommendations for develop-
ment of the CAFB process are dependent on the validity of cost and per-
formance data available on these alternative processes. The Westinghouse
assessment of the TVA cost estimates for stack scrubbers and the SWEC
estimates for CAFB installations is that SWEC has incorporated a substan-
tial element of conservatism in terms of overdesign (25 percent excess
capacity is common in the SWEC design) and in estimating. (Piping, founda-
tions, structural supports, etc. are estimated with very liberal cost
factors for labor productivity, escalation,and contingency amounts.)
This conservatism is not included in the TVA estimates. For the actual
conditions to be found in real installations, the SWEC costs are believed
to be more representative than the TVA costs by a factor of approximately
30 percent.
CONCLUSIONS
The primary conclusions are:
• There will be sufficient quantities of high-sulfur re-
sidual fuel oils,and there are existing power plants
such that a market exists for retrofitting electrical
utility power plants to utilize high-sulfur fuel oils.
• The process offers the potential to utilize low-grade
petroleum or synthetic fuel fractions with high metals
content which cannot be economically utilized in
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alternative processes for producing clean power with
conventional boilers.
• On the basis of the development work which has been
carried out, a 50 MW plant can be built.
• The economics of the CAFB process are sufficiently at-
tractive that the process merits development for uti-
lizing low-grade, high metals content, residual fuel
oils.
• A pressurized fluidized bed residual oil gasification
process for firing a combined cycle power plant is
lower in cost than the atmospheric-pressure retrofit
system. The total power plant cost for the pressurized
power plant system will be significantly lower than for a
conventional power plant.
Conclusions from the assessments of the market, technology,
process economics, and the demonstration plant design and cost estimate
are given below.
Market
• Atmospheric-pressure operation of the CAFB process is
most applicable to the electrical utility industry as
a boiler retrofit for oil- or gas-fired boilers. The
atmospheric-pressure process is not generally attrac-
tive for new boilers or retrofit of coal-fired boilers
because of the current trend toward coal. The largest
market is for boiler sizes ranging from 50 to 400 MW.
• Vacuum bottoms or other low-grade high metals content
fuels are the most likely to be available for the CAFB
process.
• Limestone availability for CAFB may be more restricted
than for slurry scrubbing processes due to more strin-
gent requirements on sorbent physical properties with
CAFB.
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• Potential markets for CAFB by-product/waste stone are
uncertain.
Technology
• Environmentally, CAFB appears superior to lime and
limestone slurry flue-gas scrubbing because of the
reduced impact of nitrogen oxide, solid waste, and
process water requirements.
• Higher fuel efficiencies are realized with CAFB than
with HDS.
• The ability of CAFB to utilize low-grade petroleum or
synthetic fuel fractions with high metals content
provides the potential for a fuel source which may
not be feasible with HDS or stack-gas cleaning pro-
cesses.
• CAFB is in a much earlier state of development than
is lime/limestone slurry scrubbing or HDS, although
major development work is still required in slurry
scrubbing processes to demonstrate reliability.
• CAFB has the potential to offer high reliability and
lower space requirements than lime/limestone slurry
scrubbing.
• It has not been established whether there is, in
general, sufficient physical space at a majority of
boiler plant sites for either CAFB or slurry scrubbing
retrofit.
• The utilization of a clean fuel, such as from HDS, is
the most convenient technological option available to
the utility.
Economics
• The capital cost of CAFB is 25 to 50 percent greater
than that of limestone scrubbing and is comparable to
regenerative stack-gas cleaning costs.
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• The fuel adder for CAFB is competitive with limestone
scrubbing if low-grade fuels available to CAFB are 10
to 20C/10 Btu cheaper than fuels suitable for firing
in a plant with limestone scrubbing.
• HDS of vacuum bottoms with low metals content is not
competitive, even though operating at a higher load
factor, unless a unit larger than 25,000 bbl/d is
built to supply more than three 200 MW boilers, or a
unit larger than 35,000 bbl/d is built to supply two
500 MW boilers. Thus, HDS requires more immediate
commitment of capital than would a single CAFB unit.
Demonstration Plant Design and Cost
• The design basis for the gasifier-regenerator system
is based on the data obtained in the 750 kw pilot
plant unit.
• Experimental tests were carried out to evaluate five
spent stone processing options. These data indicate
- Sulfur dioxide and spent limestone from the
regenerator can be combined in a dry sulfation
process to form calcium sulfate which is suit-
able for disposal. More than 90 percent sulfation
was obtained in laboratory tests.
- High temperature (1400°C) sintering of the
spent sorbent from the regenerator is a feasi-
ble method for processing the spent limestone
for disposal.
- Slurry recarbonation of spent limestone flue
gas will produce a material which is environ-
mentally suitable for disposal and may be
marketable.
Treatment of the spent limestone with dilute
sulfuric acid will produce material containing
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primarily calcium sulfate (94% sulfation)
which is environmentally suitable for dis-
posal.
- oxidation of the calcium sulfide to cal-
cium sulfate followed by high-temperature
recarbonation of the calcium oxide is not
practical for a regenerative process and
will only be feasible in a once-through
sorbent process if the residual calcium
sulfide is not released.
• The concept chosen for the CAFB demonstration plant
uses dry sulfation of the waste sorbent to eliminate
i
the sulfur recovery step. This appears to be the most
economical processing concept other than once-through
operation with direct disposal of the utilized sorbent.
• Screening tests were carried out on 14 candidate lime-
stones for the demonstration plant. Limestone 1359
and aragonite were selected as the candidate sulfur
sorbents.
• High excess capacity factors have been applied to the
demonstration plant design to assure flexibility of
operation.
• The demonstration plant cost is about $7,400,000, in-
cluding the fuel oil system. This cost does not in-
clude start-up cost, interest during construction, and
engineering fee.
• The most important cost component in the plant is the
hot fuel-gas piping system. Options to its present
design should be considered. Other costly components
are the gasifier cyclones and fines recycle system.
• Because of cost considerations the four fuel-gas bur-
ners should be placed in a single boiler face rather
than tangentially in the four corners.
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• A new induced draft (I.D.) fan should not be installed
until demonstration plant operation indicates that it
is required to meet full boiler capacity.
Development Requirements
A number of aspects of the demonstration plant require further
development:
• The utilization of low-grade petroleum or synthetic
fuel fractions with high metals content
• The pilot-scale demonstration of the dry sulfation
spent stone processing system
• The assessment of the environmental aspects of stone
disposal and utilization options
• The demonstration of the gasifier-to-regenerator
solids transport system and sorbent pulverizers on
commercial-scale equipment
• The review of safety requirements
• The development of alternative and advanced subsystems
for the process—for example, utilization of impure
limestone, alternative spent stone processing, once-
through operation
• Plant layout and fuel-gas piping concepts.
RECOMMENDATIONS
On the basis of the market, design, and economic studies, and
assessments carried out, Westinghouse recommends that:
• Detailed design and construction of the 50 MW demon-
stration plant utilizing regular-grade, high-sulfur
residual oil not be carried out,because conventional
technology exists which can utilize these fuels and
which is cost competitive with CAFB on the basis of
available costs.
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• Development of the atmospheric-pressure fluidized bed
residual oil gasification/desulfurization process be
directed toward the utilization of low-grade, high
metals content fuels, such as vacuum bottoms, which
are not compatible with conventional boiler/stack-gas
cleaning systems.
• Fuel availability for the process be investigated to
assess market need and market potential for the pro-
cess. Fuels to be investigated should include but not
be limited to vacuum bottoms, Venezuelan bitumens, re-
sidual oil from shale, tar sands, and liquefaction of
coal.
• Development work be directed toward the following
critical areas:
- Process demonstration with low-grade, high-
metals content petroleum fuels in the Esso
CAFB pilot plant
- Demonstration of dry process(es) for disposal/
utilization of sulfur dioxide and spent stone
in pilot-scale tests
- Demonstration of the pulsed flow solids trans-
port system between the gasifier and regenera-
tor in a commercial-scale facility
• Detailed design and construction of a 50 MW demonstra-
tion plant utilizing low-grade, high metals content
fuels be carried out if completion of the critical ex-
perimental tests identified and review of fuel projec-
tions demonstrate process operability and fuel
availability.
• A program be initiated to assess the application of a
pressurized fluidized bed residual oil gasification/
combined cycle power plant system to utilize regular
high-sulfur residual oil and low-grade, high metals
content fuels.
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IT. TNTPOmJCTTON
Achievement of the national goal of fuel resource independence
rests largely on the success of massive development programs in coal and
nuclear fuel utilization. It is clear that netroleum derivatives will
continue to supply a large portion of the Tinited States' electric utility
fuel demand during the vears before national energy independence is
realized. Two options presently eyist for the environmentally sound
utilization of atmospheric residual fuel oils for electric power generation:
conversion to low-sulfur fuel oils (hydrodesu]furization) at the refinery
or stack-gas cleaning (limestone slurry scruhbing, line slurry scrubbing,
and others) as a utility boiler retrofit or new boiler feature. In most
cases the utilization of refinerv vacuum bottoms as a fuel oil is not
economically feasible with these two options because of their high
metals content. The goal of the fluidlzed bed residual oil gasification
process is to provide a third alternative which utilizes heavy fuel oils
(atmospheric residual or vacuum bottoms) to generate clean power in terms
of gaseous emissions and liquid and solid wastes.
The Environmental Protection Agency (EPA) through their Office
of Research and Development (ORP) is funding the development of an
atmospheric-pressure fluid!zed bed oil gasification process for firing
in a conventional power plant. A simplified illustration of the process
concept is shown in Figure 1. The Esso Petroleum Company, Ltd., Abingdon,
England (Esso) invented the chemically active fluidized bed (CAFB) pro-
cess and is carrying out pilot-plant studies. Residual fuel oil is added
to a simple, open, fluidized bed of limestone with sufficient air—to
25 percent stoichiometric—to maintain the bed at about 870 to 970°C
(1600 to 1778°F) and to react the oil to produce a hot fuel gas. A sum-
mary of the Esso development work is shown in Table 1. With EPA funding
Esso has operated a 750 kw CAFB pilot plant more than 2600 hours and has
provided design and operating information required for the development
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Dwg.2967A73
Fluid Bed
Gasifier
\
Limestone
Conventional Boiler and
Steam-Turbine Generator
Clean Fuel Gas '
Particulates
Air Preheat
rmynn
80%
Spent Stone (CaO, CaS,
CaS04) to regenerator and/
or processing for disposal
or utilization
20%
Stack
Induced
Draft
Fan
Forced
Draft
Fan
Oil
Gasifier Fan
Figure 1-Atmospheric fluidized bed oil gasification power plant
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program. EPA contracted with Westinghouse in 1969 to carry out commercial
plant conceptual designs and an evaluation of the CAFB process.
Table 1
PHASES OF CAFB EXPERIMENTAL DEVELOPMENT
Phase
Pre-EPA
EPA Phase I
| Accomplishments
Feasibility in batch units
o Comparison of fuels and
| Time period
1966-1970
June 1970-March 1972
EPA Phase II
EPA Phase III
limestones in batch units
e Intensive process
variables study in batch
units
0 Design, construction, and
commissioning of continuous
pilot plant
• Comparison of fuels and
limestones in batch units
e Study of process variables,
design alternatives, and
process problems in con-
tinuous pilot plant
e Support for demonstration
plant design
e Investigation into trace
element retention
July 1972-Dec. 1973
July 1973-Jan. 1975
In November 1972, New England Power Service Company of New
England Electric System (NEES) agreed to participate in the program,
and boiler No. 12 at the Manchester Street Station, Narragansett Electric
2
Company, Providence, R.I. was selected as the demonstration plant site.
A three-phase demonstration plant program has been conceived:
• Phase I. Preliminary design and cost estimate of demonstra-
tion plant installed on an existing boiler, followed by an
assessment by all parties and a decision to stop or proceed
11
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e Phase II. Detailed design and construction of demonstration
residual oil gasification process
o Phase III. Developmental operation of the gasification
process and Integrated power plant.
Westinghouse, as prime contractor to EPA, subcontracted with
Stone & Webster Engineering Corporation (SWEC) fo'r the engineering ser-
vices for Phase I. Combustion Engineering, Inc. (CE), was engaged in a
boiler retrofit and burner evaluation study.
Objectives for the work performed under this contract included
e Assessment of the market for the process
• Selection of complete, Integrated fluidized bed oil
gasificatlon/desulfurization/power plant concept
• Preparation of a preliminary design and cost estimate for
the 50 KW demonstration plant
• Preparation of a conceptual design and cost estimate for a
200 MW plant
• Specification of a work scope required to carry out the
demonstration plant program
o Assessment of alternative technology in light of the prior
tasks
The results of the work carried out to meet these objectives are reported
in this two-volume report.
12
-------
III. PROCESS FUNDAMENTALS
The basic fluidized bed residual oil gasification or CAFB pro-
cess includes four processing steps:
• Oil gasification/desulfurization
• Lime regeneration
• Waste stone processing
• Sulfur recovery (optional).
A number of variations of these processing steps have been con-
1 2
sidered in process evaluations ' —for example, a once-through limestone
mode of operation where the sorbent is highly utilized in the gasifier up
to 90 percent and the sorbent regeneration step is eliminated. These
concepts are illustrated in Figure 2. The regenerative mode has been
selected as the demonstration plant basis and for commercial assessment
of the CAFB process because it has been successfully tested on a pilot-
plant scale (^ 1 MM), and it permits operating with the minumum quantity
of sorbent. A number of variations involving waste stone processing and
sulfur recovery are still under active investigation. A dry sulfation
process which contacts the waste stone with the regenerator sulfur di-
oxide product-gas and eliminates the sulfur recovery step has been se-
lected for the demonstration plant preliminary design and commercial
assessments.
The gasification/desulfurization step has been previously dis-
4 5
cussed in some detail. ' A summary of the gasifier conditions, reactions,
and performance is shown in Table 2. Residual oil (atmospheric or vacuum
bottoms) is injected into a bed of lime (870°C/1600°F) along with 1/2 to
1-1/2 times the stoichiometric amount of makeup limestone and about 17 to
25 percent of the air required for complete combustion of the oil. The
oil cracks, carbon is deposited on the lime particles, the fresh lime-
stone calcines, and hydrogen sulfide (and other sulfur
13
-------
Fresh
Limestone *
(or Dolomite)
Residual Fuel Oil
or Synthetic »
Fuel Fraction
Fluidized Bed
Gasification/
Desulfurization
System
bpent Stone
Spent Stone
Soent Stonp .—
_
Regenerated
Stone
Was!
Stor
(Ca/S
Spent Stone ,-
PfinanaraifrflH —
IxcljcncrdlcO
Stone
Ca/S>l
Spent
Proce
Stone C
ssing
Regen
t? »
le
<1)
Regent
We
Sti
(Ca/J
pratjnn
Spent Sto
Processir
aratinn
iste
3ne
)~1)
;a/s>i
Sulfur Rich
Gas
ne
ig
Sulfur Rich
Gas
1 1
Stone
Processing
Dwg 6256A37
Utilization
Disposal/Utilization
*" (including sulfur
recovery options)
•^Sulfur/
Sulfuric Acid
^Disposal/Utilization
(including sulfur
recovery)
Solids Disposal/
"Utilization
Figure 2-Sulfur removal system process concepts
-------
Table 2
SUMMARY OF GASIFIER CONDITIONS AND PERFORMANCE
Conditions
• Temoerature - 870 to 970°C (1600 to 1778°F)
o Pressure - atmospheric
e Air/fuel ratio - 17-25% stoichiometric
a Fresh limestone rate - 0.5-1.5 times stoichiometric
9 Bed sulfur content - 4-6 wt ? sulfur as CaS
o Bed carbon content - 0.1-0.3 wt % carbon
G Fluidization velocity - 1.2-1.8 m/sec (4-6 ft/sec)
e Bed denth - 0.6-1.2 m (2-4 ft)
Reactions (two-region fluidized bed)
Oil thermal cracking •*• Carbon + H2 + hydrocarbons + H-S, CS-, COS
Oxidizing region
c + [02) -> ro2> en
CaS + 3/2 02 -»• CaO + S02
CaS + 202 •* CaSO^
Oil + 02 -»- C02, H20
Reducing Region
CaCO, -»• CaO + C0«
H2S + CaO f CaS + HZ
CaSO, + hydrocarbons -»• CaS
/t
Performance - demonstrated in batch tests and pilot-plant runs
e QO1? sulfur removal or greater
e Removes 100% of fuel vanadium, 75? nickel, 20% sodium
o
e Produces fuel gas having hot heating value 7.450 kJ/m
(200 Btu/scf or greater)
e Combustion of the fuel gas possible with low excess air: acts
as two-stage combustor which yields reduction In nitrogen
oxides (from 270 opm to 155 opm on oilot olant)
o High reliability demonstrated
e Limestone makeup rates down to 0.5 of stoichiometric are
oossible-nominallv operates at stoichiometrlc makeup in the
regenerative system.
• Carbon deposition (tars) and lime accumulation may occur in
the fuel-gas lines and cvclones.
15
-------
compounds) are released. The deposited carbon, and some fuel oil, is
combusted while the hydrogen sulfide reacts with the lime to form
4 5
calcium sulfide. A number of other side reactions also occur ' and
are discussed in Appendix D. The gaslfier temperature is controlled
by the air/fuel ratio and either stack-gas recycle, water, or steam
injection. Control not only of sulfur dioxide but also of nitrogen
oxides and fuel oil metals is realized.
Tie details of the regeneration chemistry have been previously
4 5
described ' and are summarized in Table 3 along with the regenerator per-
formance. The regenerator chemistry is also discussed in Appendix D. The
calcium sulfide produced in the gasifier is circulated to the regenerator
and is contacted by air at 1050 to 1100°C (1922 to 2012°F) to generate cal-
cium oxide, some calcium sulfate, and a sulfur dioxide gas of about 8- to -
10 volume percent sulfur dioxide. Deposited carbon is also combusted.
The regenerated sorbent is recycled to the gasifier. The rate at which
the sorbent is circulated between the gasifier and regenerator controls
the regenerator temperature. The sulfur dioxide is at a concentration
suitable for sulfur recovery or sulfuric acid production.
A number of orocesses arc being considered for converting the
spent solids from the regenerator into an environmentally acceptable form
for disposal or by-product utilization. The regenerator material is about
92 wt % calcium oxide, about " wt % calcium sulfide, about 3 wt %
calcium sulfate, and the remainder inert compounds. The technical
SUDDort work to develop these processes is contained in the appendices.
Examples of options which might be utilized, depending on market
conditions and power plant location, are dry sulfation of the waste with
the regenerator sulfur dioxide (Appendix L), dead-burning of the waste
solids (Appendix 0), dry oxidation of the calcium sulfide followed by
recarbonation of the waste or silica sintering (Appendix G), all of which
produce a dry solid product. The options of wet carhonation of the stone
(Appendix M) cr wet sulfation (Appendix N) have also been considered but
are considered less attractive. The waste solids may also be valuable
without further processing for their calcium oxide or their vanadium
content, which may be about 1 wt « in some cases.
16
-------
Table 3
SUMMARY OF REGENERATOR CONDITIONS AND PERFORMANCE
Conditions
e Temperature - 1050'-1100°C (1922-2012° F)
e Pressure - atmospheric
• Fluidization velocity - 1.2-1.8 m/sec (4-6 ft /sec)
o Bed depth - 0.61-1.22 m (2-4 ft)
e Bed sulfur content - 0.5 - 1 wt % sulfur as CaS
e Bed carbon content - 0% carbon
• About 0.1 mole % oxygens in product gas
Reactions
C + 02 -»• CO,
CaS + .1/2 02 -»• CaO + S0?
CaS + 202 •»•
1/4 CaS + 1/4 CaSO^ •»• CaC + S02
4 5
Performance '
o SO. content of product gas - 8-10 mole percent
e CaSO, production in regenerator - amount produced depends
on stone carbon content, operating temperature, excess
oxygen level
17
-------
IV. ASSESSMENT
MARKET
Perspective
A preliminary review of the total market for oil gasification
plants indicates that the majority of plants range between 50 to 600 MW
for both new and retrofit applications and can be used in either the uti-
lity or industrial markets. The advantages of economy of scale may become
limited by the physical size of the units, thus requiring a modular design
for the large-capacity plants. Current trends toward energy conservation
have led to reduced income for utilities as well as delays in power gene-
ration expansion across the country. This consequence, the projected uti-
lization of coal, the potential for large (800 to 1400 MW) nuclear-powered
stations, and the uncertain petroleum availability for power generation
could minimize the application of oil gasification for new installation.
If current delays in power generation expansion continue, however, there
could be a market for a pressurized fluidized bed residual oil
gasification/combined cycle power plant concept.
Coal-fired units have been eliminated from this analysis. It is
believed that the current shift to coal and resources conservation ethics
will preclude any significant future conversion of coal units to oil.
The retrofit market for existing oil- or gas-fired units is, therefore,
considered to be primary. A summary of existing oil- or gas-fired boilers
of commercial size is shown in Table 4. On the basis of this survey, it
would appear that a 200 MW unit would be a good choice for modular design.
The total potential application is approximately 200 in number. This
represents a much smaller market segment when compared to the number of
potential 0 to 100 MW applications. On the basis of these considerations,
a 200 MW unit has been selected for commercial comparison.
18
-------
Table 4
SUMMARY OF EXISTING OIL- OR GAS-FIRED BOILERS2
FPC Region b
1
2
3
4
5
6
7
8
c
Total
Unit capacity (MW)
0-100 | 101-200
167 33
nosu.Ly
54 17
197 50
Mostly c<
53 18
471 d 118
| 201-400
13
13
35
jal and hydro -
8
69
| 400-1000
7
5
29
6
47 6
a Considers boilers whose primary fuel is oil or gas regardless of the
unit design fuel.
The Federal Power Commission (FPC) regions are illustrated in the
following outline map of the United States:
cIf all of the boilers which could be fired with coal were eliminated
from consideration in the tables, Region 1 would be effectively
reduced to no boiler candidates, the number of candidates in Region 3
would be reduced by one-third, and Regions 5 and 8 would be essen-
tially unaffected.
d'Essentially all the boilers smaller than 50 MW are 20 to 40 years old.
eOf the units between 400 and 1000 MW, 11 are greater than 600 MW and
2 are greater than 800 MW.
19
-------
The specific application of the oil gasification process will need to
consider other factors in addition to those mentioned. The process eco-
nomics are Influenced by the power station's load factor, utility costs,
space availability, proximity to refineries, and so on. The distribution
of fuel is an important factor for small plants and for plants which
would utilize low-grade fuels such as vacuum bottoms. The availability
of raw materials may also influence the decision to utilize this process.
The dlsposability of the by-product material may be either an asset or a
liability, depending on the specific application. Thus, each potential
application must be fully evaluated for its individual merits.
The major market areas are located on the East, West, and Gulf
Coasts. The market is almost equally distributed in these areas.
Market Factors
Environmental Standards
Elimination or relaxation of current emission standards would
permit the direct use of high-sulfur fuels and reduce the trend for in-
creased use of low-sulfur fuels. Thus, the emphasis on developing vari-
ous pollution abatement apparatus is also diminished. While a reduction
in localized air pollution requirements may be considered for brief
periods, long-range national programs are directed toward overall improve-
ment of air quality standards.
Environmental Legislation
The continually increasing interest in utilizing existing coal
resources may prompt legislation restricting the use of alternative fuels
for specific applications. The value of crude oil feedstocks and their
derivatives may result in legislation prohibiting their use for power
generation. It is considered more likely, however, that oil feedstocks
will be limited by supply. Thus, power utilities would be forced to use
coal or refinery vacuum bottoms for fossil units. Vacuum bottoms, while
oil derivatives, represent the end product of the fuel distillation pro-
cess.
20
-------
Feedstock Availability
The availability of high-sulfur fuel feedstocks 'is critical to
the marketability of the residual oil gasification process. An assess-
ment of various feedstocks (Table 5 and Appendix A) results in the con-
clusion that vacuum tower bottoms or other high metal content fuels are
the most likely to be economically attractive for the CAFB process. The
utilization of these low-grade petroleum residues is, unfortunately, un-
clear. Several processing alternatives have been proposed, including
blending with low-sulfur and/or low-metals residual oils.
Commercial processes are being developed to desulfurize fuel
oil from many crude oils. The cost of desulfurizing high-sulfur, high
metal content residual oil fractions, however, is currently too high to
permit the production of a fuel economical for power generation.
What quantity of the low-grade fuel would be available for uti-
lity use with the CAFB is not clear; however, quantities could be avail-
able to meet approximately 25 percent of the present oil-fired power
generation capacity (Appendix A).
The ability of the CAFB process to handle very heavy high metal
content resids is important. Large quantities of fuel liquids could be
produced from such sources as the Alberta tar sands, the Orinoco tar
belt, oil shale deposits, and coal deposits. The high boiling residues
(+565°C/1050°F) of these "syncrudes" would be suitable for industrial
and/or utility application by utilizing the CAFB process. These fuels
could significantly increase the available feedstocks.
Feedstock Availability - Limestone
Limestone is generally available throughout the country. '
While all Vine's may be suitable to some extent, the process may require
some modifications ir order to handle varied oarticulate loading levels
because of different limestone attrition rates. Thus, process economics
are Influenced by the rate of use of limestones based on their nerformance.
Results of tests of various limestones from all sectors of the country
21
-------
Table ">
FEEDSTOCK SELECTION
Feedstock
Comment
Hieh-Sulfur Crude
High-Sulfur Residual
High-Sulfur "acuum Bottoms
Unlike]v to be available for power genera-
tion use as oart of a resources conserva-
tion program because of the valuable
bv-nroducts which it contains.
Foreign
Currently 23% of our residual imports have
high-sulfur content (Appendix A). While
more widely available from European refin-
eries because of their operational technique
(up to 40% of every barrel of crude processed
becomes resid), this report has conserva-
tively precluded their use, assuming a
limited import policy.
Domestic
While high availability is forecast, this
report precludes their use, assuming
restrictive legislation requiring them to
be processed to vacuum bottoms, also part
of the resources conservation program
(Appendix A).
Foreign
Unlikely because of current refinery opera-
tion mode and transportation economics;
added high costs and equipment modifications
would make its importation impractical.
Pomestic
Most likely alternative based on availabil-
itv, see Appendix A.
22
-------
are provided in Appendix Q. Tn general, availability is no nroblem;
limestones, however, must be selectively identified in each area for a
specific duplication.
Competing Technology
Competing technology or the advent of advanced technological
development resulting in a more practical and economical method of
desulfurlzlng fuel would shift the problem of pollution control from the
power utility to the fuel supplier. Some of these aspects are reviewed
in the technologv assessment below.
TECHNOLOGY
Technological aspects of the PAFB process are assessed and
compared with two competing technologies for fuel utilization and
pollution control — lime/limestone slurrv flue-gas scrubbing and
hydrodesulfurization of residual fuel oils. The environmental impact,
the present state of development, and the applicability of the tech-
nologies to the electrical utility industry are reviewed.
Environmental Impact
The environmental impact of the CAFB process and lime/limestone
slurry flue-gas scrubbing are compared in Table 6. Details of CAfTJ
environmental performance are presented in Appendix R. The basis of
comparison is a 200 MW oil-fired power plant (retrofit). The CAWB process
appears to be superior environmentally to limestone and lime slurry scrubbers.
Sulfur removal capabilities for the processes are comparable, but the
CAFB process provides a considerable reduction in nitrogen oxide emissions
and consumes an order of magnitude less process water than do slurry
scrubbing processes. Similar nitrogen oxide control could be achieved by
the addition of special combustion equipment on the boiler. Auxiliary
power requirements are comparable at three to four percent of plant capacity.
The limestone usage is nearly Identical for the processes — about 1 mole
of calcium per mole of sulfur removed from the fuel — but the CAFB process
23
-------
Table 6
ENVIRONMENTAL IMPACT COMPARISON
200 MW conventional oil-fired
power plant with limestone or
lime slurry flue-gas scrubbing
200 MW conventional oil-fired
power plant with CAFB oil
gasification and dry sulfation
of solid product
Plant Emissions
S02 kg/CJ (Ib/mm Btu)
NO^ kg/CJ (Ib/mm Btu)8
Total Participates kg/GJ (Ib/mm Btu)
CaO
Solids Mg/day (tons/day)
chemical form
Liquids Mg (tons/day)
chemical form
0.15 (0.35)
0.34 (0.80)
0.043 (0.10)
None
None
145 (160)
CaSO-, CaSO^, CaCO.. Inerta (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 CaSO., CaO, Inerts
None
Hone
10
Resource Usage
Limestone Mg (tons/day)
Process water, liters/day (gal/day)
Auxiliary power (kv)
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 104-4 x 10* (5-10) per year
per 3 m (10 ft) depth of pond
Yes
No
11B (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 fluldized bed combustion.
Waste 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, H.S. Rosenberg, "Lime/Limestone Sludge Disposal - Trends in the Utility
Industry"; both presented at the EPA Flue Gas Desulfurlzatlon Symposium, Atlanta, Georgia, Nov. 1974.
cProcess water and auxiliary power for limestone and lime slurry scrubbing baaed on G.G. McGlamery and R.L.
Torstrlck, "Cost Comparisons of Flue Gas Desulfurizatlon Systems," presented at the EPA Flue Gas Desulfurlzatlon
Symposium, Atlanta, Georgia, Nov. 1974.
Based on steam for stack-gas reheat.
-------
could ootentially reduce this consumption 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. CAFB also permits the utilization of high-metals vacuum bottoms
as a fuel. Additives exist which mav oermit a boiler to utilize these
high metal fuels directly, with reduced boiler deterioration.
Hydrodesulfurization (HDS) requires the production of hydrogen
and oxygen, which contributes to a substantial energy inefficiency.
Substantial amounts of process water and cooling water are also used.
HDS produces elemental sulfur as a by-product, but its market value Is
uncertain. Nitrogen oxide emissions would be reduced to some extent in
the power plant using a hydrodesulfurlzed fuel oil without the addition
of nitrogen oxide control equipment to the boiler burners. The ability
of HDS to utilize vacuum bottoms economically is currently in doubt
because of the effect of high metal content of the oil on the catalyst
. 8
performance.
Present State of Development
The status of the development of the three alternative tech-
nologies is considered in Table 7. CAFB is in the pilot-plant stage
with extensive testing with atmospheric residual oil already completed.
Vacuum bottoms utilization has been investigated on batch-scale equipment.
CAFB stone processing techniques have been investigated only on laboratory-
scale aoparatus. Lime and limestone slurry scrubbing represents the most
o
advanced of the flue-gas scrubbing technologies.
Regenerative flue-gas cleaning techniques are In stages of development
similar to CAFB. Although lime/limestone scrubbing is near commercial
status, it has received only limited utility acceptance. HDS represents
commercial technology, although the utilization of vacuum bottoms is
economically bevond the current limits of the technology.
Development areas exist for all of the competing technologies,
but CAFB, by the nature of its development, has the widest range. None
of the CA^B development areas are believed to be critical to technological
25
-------
Table 7
PRESENT STATE OF DEVELOPMENT
CAFE
Status
Lime /limes tone
scrubbing
HDS
• Pilot plant (atm. • Near commercialization • Commercial
Major
Development
Areas
IS)
residual oil utilization)
• Batch units (vacuum
bottoms utilization)
• Laboratory units
(stone processing)
• Stone processing/
sulfur recovery
• Vacuum bottoms Demon-
stration
• Peducet1 limestone
requirements
• Control of coVe
deposition
• Stone circulation
• retailed design and
operation of
demonstration plant
Pilot plant (sludge
processing)
• Control of erosion,
corrosion, plugging,
scaling
• Sludge disposal
• Petrofit structural
problems
• Other mechanical
problems such as
bypass dampers
Improved operation
vith vacuum bottoms
-------
success. Lime/limestone slurry scrubbing processes have been making
progress in reliability, but further development and demonstration is
required. ' At the present tine, the ability of HDS to utilize
vacuum bottoms appears uneconomical, although nrocessing alternatives do
8,12
exist.
Utility Applicability
The degree to which any of the three alternative technologies
meet the requirements of the utilitv industry depends on considerations
such as:
• Process reliability and maintenance requirements
• Process control ability
• Process complexity
• Retrofit boiler performance and life
• Process safety
• Process space requirements
• Raw materials availability.
It is likely that the utility industry would favor the firing of a
clean fuel (such as that from a HDS process) over retrofitting either
CAFB or lime/limestone slurry scrubbing processes based on these
considerations. HDS does not require the utility to adapt to new tech-
nology. CAFB and lime or limestone slurry flue-gas scrubbing appear
comparable in these aspects, with CAFB possibly providing greater
reliability, lower space requirements, and improved boiler life.
High reliability has been demonstrated by CAFB in the pilot
plant, except for the complications associated with coke-lime deposits in
the cyclone Inlets. This appears to be a minor problem in the light of
Improved operating methods to reduce the coke deposits now being investigated
in the pilot unit and the effect of larger scale equipment on the time
between shutdowns. Large land usape is required for disposal ponds with
lime or limestone slurry flue-gas scrubbing, but CAFB requires only
limited storage of dry waste/by-product material. Some plants may not
provide the physical space required for a retrofit by either process. CAFB
27
-------
removes much of the fuel oil metals (vanadium, nickel, and sodium) associated
with boiler tube corrosion. The availability of suitable limestones may
restrict the application of both CAFB and wet scrubbing processes (Appendix
Q).
The ultimate choice of technologies must be made on the basis
of these technological issues combined with complex and unstable non-
technical Issues such as:
• Process cost
• Fuel cost, availability, reliability of supply
• National policy and pollution standards.
Geographical location will also play a role in the relative merits of these
technologies. If a reliable and cheap source of vacuum bottoms is
available for CAFB, the utility may be placed in a more attractive market
situation.
The utilization of a pressurized fluidized bed residual oil
gasification orocess with current combined cycle power generation technology
2
was previously evaluated. Application of CAFB at higher pressures offers
the potential for higher plant efficiencies (up to 50 percent power plant
efficiency) than can be obtained with conventional power plants; the ability
to utilize more economically low-grade, high-metals content residual fuel
oils in new power plants, and the ability to reduce construction time over
conventional plants for new installations. These factors are important
enough in the overall energy and environmental effort to warrant further
consideration of the pressurized concept.
Summary
Environmentally, CAFB appears superior to lime and limestone
slurry flue-gas scrubbing because of the reduced impact of nitrogen
oxide, solid waste, and process water requirements. Higher fuel
efficiencies are realized with CAFB than with HDS. The ability of CAFB
to utilize vacuum bottoms provides the potential for a fuel source which
may not be feasible with HDS or stack-gas cleaning processes.
28
-------
CAFB is in a pilot-plant stage of development while lime/
limestone scrubbing is near commercial status and HDS represents commercial
technology.
The applicability of the three technologies to the electrical
utility industry is a comnlex consideration. Technologically, HDS
provides many advantages to the utility operator. CAFB provides a
number of potential advantages over lime/limestone scrubbing — possibly
improved reliability, lower snace requirements, and iirproved boiler life.
When unstable factors such as process cost, fuel cost, fuel availability,
and national policy are considered, many of the technological Issues may
be relatlvelv unimportant. The greatest potential advantage of CAFB
operating at atmospheric pressure or pressurized for combined cycle power
appears to be its ability to utilize vacuum bottoms and, thus, to provide
a fuel source to the utility which might be otherwise unavailable.
ECONOMICS
An assessment of costs for several alternative means of sulfur
oxide pollution control has teen prepared. The most likely alternatives
to fuel desulfurlzatlon via CAFB processing, such as vacuum resid HDS
and stack-gas desulfurization via wet scrubbing have been reviewed.
Capital and operating costs have been assembled for the Cities Service/
Hydrocarbon Research H-Oil process, a limestone slurry process for
stack-gas purification under development by Bechtel and others, and a
catalytic oxidation (Cat-Ox) process for stack-gas purification under
development by Monsanto.
H-011 was chosen to represent oil HDS for two reasons:
• The fast percolating bed of catalyst used in commercial
anolications of the H-Oil process provides the possibility
for continuous catalvst replacement in the desulfurization
of high metals content vacuum resid.
• Recent data are available from a reliable source (Hydro-
13
carbon Processing Magazine's 1974 Refining Handbook issue )
on costs for a commercial-scale vacuum resid processing
installation.
29
-------
The limestone slurry nrocess was chosen to represent the lowest
cost system for sulfur dioxide removal from stack gas, and the Cat-Ox
process by Monsanto was chosen to represent a high-cost system for sulfur
dioxide removal from stack gas. It should be noted that the high cost of
the Hat-Ox nrocess is partly due to the avoidance of sludge or waste stone
disposal problems. Both limestone slurry and Cat-Ox process costs were
obtained from a recent TVA/EPA paper, "Cost Comparisons of Flue Gas
3
Desulfurization Systems" by McGamery and Torstrick.' A substantial effort
was made to place all published costs on comparable bases, as will be
discussed in some detail below.
Capital Costs
SWEC prepared capital cost estimates for two CAFB plants, a
50 MW demonstration plan (see Appendix E) and a 200 MW demonstration
plant (see Appendix H). The tabulation in Table « presents these two
estimates by SWEC as costs I (50 MW) and II (200 MW). The SWEC estimates
have been modified to a slightly different form in Table 9, by showing the
new burner costs broken out from the process equipment costs and Identified
as allowances. Other amounts have been added to the SWEC estimate for
contingency, fee, start-up costs, and interest during construction. The
percentages used for these added amounts were derived from the McGlamery
and Torstrick paper mentioned above. Also allowances for new induced
draft (I.D.) fan costs were added to the SWF.C estimates. Thus, bare cost
estimates by SWEC of $7,360,100 and $21,583,500 for single-train demon-
stration plants of 50 and 200 MW, respectively, have become total capital
costs of $9,497,000 and fi28,l'55,noO in Table S.
A careful review of the SWEf cost estimates discussed above
permits the following conclusions to be reached'
• The 200 MW estimate, as well as the 50 MW estimates in Table
by SWEC, are botfr based on a 20 percent air/fuel (A/F) ratio.
The 200 MW 20 percent air/fuel ratio case (Appendix !!)
prepared by SWEC assumes 16 gas burners.
• The 200 MW estimate Includes new limestone storage and ash
storage silos. Pirce the 50 MIC design used existing coal
30
-------
and coal-ash storage facilities for fresh and spent stone,
no costs for storage silos were Included in that estimate. The
200 MW design includes costs for nearly three weeks' storage
of fresh and spent stone.
• The 200 MW estimate includes compressor costs that were
derived from the 50 Mt\T estimate by direct multiplication
bv four.
• The 200 MW instrumentation materials cost is 2.7 times the
cost of instrumentation materials in the 50 MW estimate.
New estimates using the low air/fuel ratio (17 percent) data
presented in SPEC'S flow diagram and material balance, Figure 3,
were made in-house bv Westinghouse on the basis of the facts presented
above and on the SWEC cost detail summary. The new estimates proportioned
costs for the low air/fuel case, 50 MW unit equipment, from the 20 percent
air/fuel case, 50 MW equipment, by taking the ratio or equipment capacity
to the 0.7 power, and multiplying costs hy the resultant number. For
example, the gasifier air and recycle gas (stream 2) in the low air/fuel
case is 80 percent of the same stream in the base case. That capacity
ratio, 0.8, to the 0.7 power is O.R55. The gasifier blower cost for the
base case is thus multiplied by 0.855 to arrive at a cost for the low
air/fuel case. New costs for the 50 MW and ?00 MW units were derived in
this way. At the same time, the 200 MW costs were also adjusted to reflect
smaller stone storage silos (one week instead of three weeks' storage),
to factor compressor costs from 50 MW to 200 MW using the 0.7 power
instead of the direct multiplier SWEC used, and to factor instrumen-
tation materials from a SO MW to a 200 MF single-train olant using a
1.7 multinlier rather than the ?.7 multiplier SWEC used. The smaller
multiplier is more annrooriate for reflecting the same number and
comolexitv of instruments in the large and small olants alike, but longer
tubing and wire runs in the larger plant because of the nhysical size of
the increased system. Thus, bare cost estimates of S6,£48,000 and
$17,252,000 for single-train, low air/fuel ratio olants of 50 and 200 MW,
respectively, were obtained by Westinghouse estimates, and these become
$8,528,000 and $22,305,000 total capital cost, as summarized in Table 8.
31
-------
Curve 678952-B
to
ro
o>
o
l_
O>
CAFB(20%A/F)
Cat-Ox Scrubbing
o-CAFB(17%A/F)
Limestone Scrubbing
45,000bbl/day
•—o*.
H OS with H?prod.
on Vac. Resid
H h
_L
100 200 300 400 500 600 700
Plant Capacity, Megawatts
900
1000
1100
FigureB- Power cost adder for sulfur removal systems
-------
To arrive at costs for CAFB units retrofitted to a 500 MW
two-boiler, oil-burning Installation, each 200 MW unit cost (one 20 percent
air/fuel ratio, one 17 percent air/fuel ratio) was Increased to 250 MW
capacity using the 0.7 power, and then the cost was doubled to reoresent
two complete trains of 250 MW each. The 750 MW CAFB gasifier would be
approximately 13.7 m (45 ft) in diameter, approaching the size of the largest
size cat cracker. The McGlamery and Torstrick caper also uses two 250 MW
scrubbing systems for their 500 MW cost studies. It can be concluded
that, both from a size standpoint and an equipment reliability standnoint,
it is advisable to assume that a 500 MW capacity system would be retro-
fitted with two CAFB units as has been assumed in this economic assessment.
The 500 MW CATO total capital costs, as summarized in columns V and VI of
Table 8, are $65,253,000 and $51,085,000 for the 20 percent air/fuel
and 17 percent air/fuel cases, respectively.
The first process cost to compare with CAFE costs was the
vacuum resid HDS H-Oil process. Tost and utility data available for the
H-Oil system at 17,000 bbl/d permitted the evaluation of this process at
two capacities bracketing the base point capacity (9000 bbl/d and 45,000
bbl/d). The 9000 bbl/d capacity of the small HDS unit approximates that
required to fuel a 225 MW power plant, and the 45,000 bbl/d capacity of
the large HDS unit is as large an installation as is practical to handle
the resid from two or more two-stage crude oil stills of large capacity
(160,000 bbl/d each) at a single site.
Based on H-Oil's $710 per dally barrel direct costs plus the
hydrogen production systen costs from a private communication with Stanford
Research Institute, 1972, and the use of a cost exponent of 0.65 derived
for HDS systems by Guthrie of Fluor Coporation in a 1970 article in Chemical
Engineering. the total direct costs of the grass roots HDS units were
$18,200,000 for the QOOO bbl/d installation and $50,400,000 for the 45,000
bbl/d installation (S2020/daily bbl and $1120/daily bbl, respectively).
The total capital costs for those two units, as summarized in columns
VII and VIII of Table 8, are $29,523,000 and $83,35«,000. Due to the
relatively established nature of the H-Oil process, half of the start-up
33
-------
COMPAMHVE CAPITAL cosn fa IULFIB HMOVU svsnms
Procat EQUV in Puce
ProcniMlirtlUtor
IOUI Orrccft
OiUrituUUn
Suttow
Indirai Coin
IDll Birt Cnl
OMinatnry
loui Proem Invnl 1
Nn 1 0 Fin'
Burner fW
loui Imml
Start up Coin
Internl during Cam!
loui CJpitil Call
,/..
CAFB
SOMWurd!
IZOlAfFnll
• 208,800
2803200
• 9 187000
899000
• 0.032000
1 250 000
• 7287 (DO
5H. on
25" 003
• 1.099 Oil
50 mi
7U030
• I 187000
699000
095000
19497000
190
11
CAFB
ZOOMWunil
l20>A/r ltrilneu.1
1 7099000
9400000
116949030
7900000
• 19049.000
2799 OU
•71 in on
1695000
871.000
121 733 030
200 000
TO 000
tn 771 000
1942.000
1942000
128.199.000
141
III
CAFB
SOMHrunrl
imA/Fnt 1
• Z. 146. ODO
2.929.000
• 4. 679. ODD
767.000
• 9.412.090
1 128,090
• 6970000
467,000
237,000
• 7274090
78. ODD
• 7.392.000
so. on
5B.OOO
• 8.53.00)
171
TV
CAfB
ZOOMVfuna
tt.iOr.gi. etc.)
• 5.015.000
8,162.000
• 11.177.000
1.9)1.000
• 19,168,009
1.796.000
• 14.964. 030
1118,000
699.0X
118,911,030
288,000
•H. 221. 000
1.938. ODD
1 938.000
• 21105.000
112
V
CAFB
gooiiwiinii
ISlWFol 1
• 16. 496. 000
22. in. no
• 38.685000
9 149. ODD
• 44.S30.000
9.271.000
149.831 OB
1.868.090
1.934.030
• 55.603.030
650,030
• 96.231030
4. 500. OX
4.509.000
• 69. 251. OB
111
in
CAFB
9OMWuifl
1171 WF at 1
1 12.6Z5.030
I7.5i2.030
ISO. IS7.000
4.SM.OD9
• 14,148,000
4.111.000
138.161.010
1,019.030
1. 909.000
• 41.389.00
090.000
• 44.039.030
3.523.030
1.921.030
151.09.000
102
VII
Brill rail KDSiinll
•tthrbpnaidlan
llOOObU/diinrI)
imply lor ZSMWI
1 7761.000*
10.419.000*
III. ZOO. 000
2790.000
• 20950000
2.48D.030
• 21430030
1820.000
910.000
126.160.000
ZOO.aXIICJt.1
• 26.360.000
1054.030*
2.109.030
• 29.923.000
111
VIII
Gnu roan KDSurdl
olihH,indudlon
145.009 u/rj Drill
1 umlr lor 1125 MW)
• 21.492.000'
28, «, 000*
190,400,030
7.600.030
• 98,000.00)
6867.000
164 867.030
5040.030
2.920.000
• 72 427 000
2.0000)910.1
«74,427.0D
2.977.0001
9.9X030
• 83, 351 OX
74
IX
UojHMt
sinjbWrt)
amm*
1 1.00.003*
i.*i mitt
1 4.696000'
6119000'
1 11,011 030
l.6MOOOb
1 12 679,000
1900 000°
• 14,179000
1 101 000
550 CO)
• 15826000
709030 (Ml
1 10 976 000
1091000
1 122000
• 19901030
98
XIV
CJI-0>
OIOCBI
gOOMWunu
1 10 is> oni *
139MOOO'
124199,000
1680.000
• 28,035 000
1.119.000
• 11 154000
2416001
1 718 OJH
1)5003000
) ifauwiM
• 30716m)
3 674000
1 9io an
• 43349000
87
'proportnnM Iron Mil oVKhlo> CUB ZOtWI ZOO MW unit
"MI. IOCATB;S«C*
C«JIMnct
(S«C illMnu
'mil ngroal di4rgt tut U HuncM UJor ol HOS OtvHoownt
-------
COMPUAIM OnuOING COSTS FOB SUflB KMOVAL SYSTCJ4S.
Ul
1 II III
CAIB ure an
so M* i»n 200 MIT MM so M* ml
I20» XII rill 1201 MF 1 tuln to. 1 117* «/F M 1
UreaUntaUUpl • 50000 1 100000 t 50.000
UMf t SitWnloOpem 14?. BO 149 BO 149 BO
Slum Nq.Nq.Hq.
Wer Nfg.Nq.Mg.
Pwer 106500 D6100 140.000
MjlMimnu 207 100 661 HO ID.ODO
IMrCull 77000 27000 DODO
CCIUI Clurgn 1419100 4,195100 1Z70.KD
PUMOHid. 111,100 '12 900 UO.nO
UMrCMM. 15000 15000 15(00
IOIM 17HS.no 16,407700 ll.940.no
HjSO, a SMIui Crdlt
Fud tor Prouii Hen
Mel
MUli/kWi 6.73 411 5.54
• 110*8111 62,5 411 55.4
IV V VI VII
CAFB Ofl CAB 5,,,, ran WS ml
ZOOM* tint SOONWunft SOOMWud dui H, oradudtaT
117*4/1 llrtnl l!»*/Ftal imVFHLI lfggg BUd ^m
IM ittrjgt.ct , a(0|> ^ jjj mn
1 200000 1 500.003 I 500.000 • 156.400
149.BO 1I9.BO 149.80 149.BO
Nq. Kg. Mg. B.SOO
Meg. Nq. Mjg. 12 BO
StAOD 2.065.0DO 1.409000 311. OO
527 100 1.50.400 1,207.500 721X000
27.000 D.OOO 0.000 45.600
1.B3.4DD 9722.709 7.611.700 4.W.WO
252.700 757.109 556.900 S7.»
15.000 15.009 15.000 15.000
15.05S.OOO 114.714.701 111,467.900 16114.100
en. 7001
•.19.600
17.771, 000
161 472 121 416
16.1 C2 a a Q»
VIII
imi rail HOS inn
«Uh Hj ontkidloi
145.000 ttOla unit!
iiiaiyhr 11BMW)
1 I.7B.OOO
149.530
192. B)
6420)
1559.200
2.016.000
41109
12 420 »
18.401
15.000
(19. W. BO
(1.431.3001
a,m.2oo
I26.144.7t0
2.96
296
IX
UBntm
50 MW will
• 50.000
210. BO
H.500
2.000
71700
201.800
45.600
6U.TO
114900
21.000
• 1.445.600
411
411
Umtm
1 200000
210200
138,000
(.000
111.000
610500
45600
LI* ODD
265500
21000
•1,70) BO
2.61
265
XI
Umnunt
•cruttlng
SOOMWutt
500,000
Z10ZO
345000
20,000
787.400
1. 191 400
45600
1699.100
519900
21000
• 7.140000
2.10
210
XII
CH-Oi
proeni
SOMWunil
n no
61 100
12500
2500
«J 400
130 600
41000
928000
W400
630D
II 36). 100
147 U»
417 100
II 711000
497
497
cl%
pncnl
iOOKWimil
69.100
61 100
50100
toon
361 an
unto
41000
2105690
194.700
6300
14.149.100
I1S&.4QOI
1 671 100
•i 6ii no
402
40.2
XIV
CJI-O.
prom
UOMWiniH
172800
61 100
125100
75000
•04400
974201
40.000
6451001
428.000
6300
19206100
1471 0001
4 177 909
117 911 COO
169
16,9
•um
151 SOI 14/lonSttn«70i/l£OJII)S™. H/Wti, HIOOUHf. U91/|rClp. Ouign. (U5/llirltrGil4r|ll MTIaiif HJO. OUtdSyaWBulorWMil Mlrl «8V,rUnm. O/p CU8.
HOSnidlcUOIndlimdBnl IMar* tl/Uvtiaa FlnlOiM. it TOld OtHCoiti, utwOMtf. «»»(
-------
expense percentage used for CAFB was used to arrive at start-up costs for
the HDS units. Contingency, fee, and interest during construction percentages
remained the same for HDS and CATO estimates.
Both limestone and Cat-Ox processes for stack-gas cleanup were
assessed using the previously mentioned x*ork of McGlamery and Torstrlck for
TVA/EPA. The costs given in their paper for scrubbers applied to 200 MW
and 500 MV oil-fired units burning ?..5 percent sulfur oil were used as a
basis for comparison with CAFB. The mid-'72 start-of-engineering cost
basis was escalated to early-175 start-of-engineering by using a 1.25
cost multiplier derived from Chemical Engineering's cost index. Also,
the new unit ?00 MW costs given in the TVA study were factored to
retrofit cost bases by applying cost adders to the new unit costs derived
from 500 MW studies of new and retrofit cost differences for oil-fired
systems. These cost adders take into account that a new unit is designed
for a 30-year life versus a 25-year life for a retrofit design and, more
significantlv, that a new unit does not have the costs associated with
installing scrubbers in an existing plant with limited space and many
existing interferences that must be moved to accommodate the retrofit
scrubbers.
In order to compare capital cost estimates fairly to CAFB and
HDS costs, the scrubbing process costs used were the total direct costs,
and the same percentages were applied to those total directs as applied for
the CAFB and HDS units. Thus, distributables (in other words, construction
management and overhead costs) were applied at about 15 percent of total
directs, as SWEC had done, instead of the 11 percent to 13 percent used in
the TVA study. Also, indirect costs (engineering, procurement services,
and so on) were applied at about 13.5 percent of total directs as SI-TEC
had done instead of the 9 to 11 percent used in the TVA study.
All other costs, such as contingency, fee, and so on, were also uniformly
applied on the same basis as in the TVA study. The 200 MW and 500 MW costs
in S/kw were plotted in semilog paper as a function of capacity In mega-
watts and were extrapolated to give $/kw for 50 MW units. These extra-
polated costs were then summarized in Table fl, along with costs derived
directly from the TVA study, described above. Limestone scrubbing total
36
-------
capital costs for 50 MW, 200 MW, and sno MW installations, as summarized
in coluraes IX, X, and XI of Table 8, are $4,537,000, $17,725,000, and
$24,829,000, respectively. Cat-Ox scrubbing total capital costs for 50 MW,
200 MW, and 500 MW installations, as summarized in folumes XII, XIII, and
XIV of Table 8, are !?6,?2R,non, $19,501,000, and $4^,349,000, respectively.
All of the costs discussed above are subject to additional factors that
could, during the construction and starf-ur of these facilities, be
responsible for as much as doubling the total capital charges summarized
above. As pointed out in the TVA study, premium labor charges to cut the
construction schedule or to obtain adequate manning, plus special R&D
costs for first of a kind units, olus cost escalations at higher rates than
those discussed above, unusual equipment delivery delays, anv provisions for
spare equipment to Increase reliability, and so on could — and usually do —
make actual installations much more expensive than those In Table 8.
Those added costs can be applied uniformly to all costs tabulated, however,
and should not affect the comparative economic process assessment.
Pressurized oil gasification for firing in a combined cycle
power plant is estimated to be lower in cost than the atmospheric-pressure
retrofit system cost because of:
• More compact vessels and fuel-gas piping at ]520 bars
(15 atm) pressure
• Shorter, less complex fuel-gas piping system
• Higher plant efficiency, which permits lower limestone
and oil rate per MW of plant capacity
• The regeneration process which generates hydrogen sulfide,
permitting elemental sulfur recovery
• The plant layout which is designed for low investment
and maintenance.
The gasification process (including oil feeding, limestone handling,
regeneration system, waste stone processing system, waste stone handling
system, sulfur recovery process, and gasification system) for a 260 MW
combined cycle plant is estimated to cost about $60 kw (direct cost)
based on end-of-l?7A costs. A low-cost combined cycle plant (consisting
37
-------
of two Westinghouse 501B gas turbines and a single 0.7 m (28.5 In) steam
turbine nrovides a total oil gasification combined-cycle power plant
capital cost of about $250/kw compared to about S600/kw for a conventional
oil-fired oower plant with stack-gas cleaning. This large cost difference
Is credited to the low cost of the combined cycle olant — maximum shop
fabrication and short construction time.
Operating Costs
The TVA/EPA study was used as the basis for operating cost
comparisons. It was felt that this study, using power cost at 10 mills/
kWh steam at 70C/454 kg (1000 Ib), potable water at 8C/3785 liters (1000 gal),
14.9 percent per year for capital based charges, and so on was a better
basis for comparison of plants in the 200-to-500 kw range than the operating
costs suggested by NEES for the Manchester Street Station and used by
SWEC (Appendices E and H). Table 9 summarizes the annual cost so
calculated. Power plant systems are calculated for 7000 hr/yr and HDS
systems for 330 days/yr (7920 hrs) to arrive at mills/kWh. The TVA
study assumes that all newly installed pollution control systems will be
operated 7000 hrs/yr for the first two years, so that this is a valid
basis for comoarison. Limestone is assumed to be available to CAFB and
limestone-scrubbing systems at $4/Mg (ton), the cost assumed by
TVA. This may be too low, but not enough is yet known about stone require-
ments to state this with any certainty. The limestone scrubbing units are
charged with 25-year ponds for waste slurry containment. It has been
assumed that the dry sulfated CAFB waste stone will be removed at no cost
to the process.
Labor and laboratory costs have been assumed to remain constant
from 50 to 500 MW size units. This may not be strictly accurate, but
these costs are such a small part of the total that such an assumption
does not significantly affect the results of the comparison. Utility
costs have been assumed to be a direct function of plant capacity, a'
simplifying assumption that also will not affect comparative results.
Plant overhead is taken as 20 percent of labor and utility costs, and labor
38
-------
overhead is taken as 10 percent of direct labor and supervision at an
average hourly rate of $8.
Operating costs in mills/kWh are plotted against capacity in MW
in Figure 3. It is evident from this chart thaf
• The present TAFB design is not cost competitive with the
limestone scrubbing system design and cost projection and,
therefore, must be salable on the basis of
- Utilization of fuels which cannot be economically used in
alternative processes
- No sludge for disposal
- Lower water requirements
- Reliability
• HDS is not competitive, even though operating at a higher
load factor, unless a larger than 25,000 bbl/d unit is built
to supply more than three 200 MW boilers, or larger than a
35,000 bbl/d unit is built to supply two 500 MW boilers.
Thus, HDS requires more immediate commitment of capital
than would a single CAFB unit.
• Above about 100 MW, CAFB is lower in cost than a stack-
scrubbing retrofit system producing sulfuric acid from the
recovered sulfur values.
Catalvst costs for HDS should be noted. These were based on the
processing of a West Texas sour vacuum resid having the following
properties:
• Boiling range ^65°C (1050°F)+
• Sulfur content 5.1 wt %
• API gravity 12.7°
• Conradson carbon number 20
• (V+Ni) content 130 ppm
A Kuwait atmospheric resid having anproximately half of the vanadium
and nickel content of the West Texas resid also had half of the catalyst
costs (6c/bbl versus 12c/bbl). Very high heavy metal content resids
cannot be economically processed in HDS units. This can be shown if one
assumes that catalvst consumption and, thus, catalvst cost are roughly
39
-------
proportional to heavy metal content. The TOS catalyst cost for an
Amuay vacuum resid tested in the batch unit at Esso would then be over
four times the catalyst cost for Vest Texas resid, or about 50c/bbl based
on 536 pom of vanadium and nickel. Furthermore, the Venezuelan
Bachaquero vacuum resid would require about SI of catalyst per bbl of
resid, based on 1040 ppm of vanadium and nickle. The CATO is expected
to utilize these resids without cost nenalty, but the increased catalyst
costs noted above would add 0.71 and 1.65 mills, respectively, to the HDS
system costs. These additions to HDS costs would be sufficient to make
the largest practical size vacuum resid HDS unit uncompetitive, even
though base load operated, compared with a 200 MW 17 percent air/fuel
ratio CAFB unit.
The ability of CAFB to handle very heavy high metal content
resids must be confirmed by additional tests on the Esso pilot plant.
Such an ability could be important, not only in the utilization of
petroleum-derived resids but also in association with the production
of large quantities of fuel liquids from such sources as the Alberta
tar sands, the Orinoco tar belt, the Colorado/Green River oil shale
deposits, and highly-volatile, high-sulfur Illinois and western Kentucky
coal deposits. Such processing facilities could produce wide boiling range
materials generally characterized as "syncrudes1' and suitable as feedstocks
for existing or new two-stage refinery distillation systems. The fraction
of these svncrudes boiling above 565°C (1050°F) would be produced as vacuum
bottoms and would also be suitable for 100 percent conversion into clean,
noncorroslve fuel gas by CAFB processing. Trace metal contents in such
syncrude-derlved vacuum bottoms mav make conventional coking and combustion
of the coke so derived either exnenslvr. or ohysically difficult. For
example, ash fusion characteristics of the coke produced from Atabasca
tar sands created severe problems in the boilers used to burn that coke.
Such problems would be eliminated bv the use of CAFB technology. To
quote from a recent technical paper bv Dr. Kett of Exxon's Florham Park
P&n Center, "In the future, progressively more and more of the crude barrel
will need to be converted into lower boiling products ..." and the resi-
40
-------
dues from conversion of such lower boiling products, whether from heavier
crudes or from syncrudes,can be effectively turned into fuel gas useful
in an environmentally acceptable manner by existing power generating
boilers.
41
-------
V. PRELIMINARY DESIGN AND COST ESTIMATE
BACKGROUND
The first phase of the demonstration nlant program is to
preoare a preliminary design and cost estimate. This phase of the
program has included four major tasks. The tasks, participants, and
schedule are summarized in Table 10 . NEES is the cooperating utility
2
for the demonstration plant program. NEES personnel and the staff at
the Narragansett Electric Company, Manchester Street Station provided
plant design data and information on plant load requirements, assisted
in the selection of the architect-engineer, and participated in the
selection of process options and the evaluation of the preliminary
design work.
The design effort was based on the experimental data from
Esso on its batch units and 1 MW continuous pilot plant, the spent
stone processing data from Westlnghouse (Appendices K-0)t the conceptual
1 2
design and assessment by Westinghouse, ' and the operating and environ-
mental constraints established by NEES and EPA. EPA has provided
general guidance and funding for the Phase I effort. The recommendations
developed from this preliminary design and cost estimate will be reviewed
by the project participants.
PRELIMINARY DESIGN SCOPE
Three stens were carried out to select the design basis for
the CA^B demonstration plant:
• An identification of the process requirements, operating
parameters, and design and process options (Appendices
J and K)
42
-------
Table 10
PRELIMINARY DESIGN AND COST ESTIMATE
Task
I
Participants*
Schedule
Ld
1.
3.
Selection of the architect-
engineer and contract
negotiation
Completion of initial
design study and cost
estimate
4.
Evaluation of the initial
design study, selection"of
the design basis recommend-
ed for the demonstration
plant, contract negotiation
for preliminary design
Preparation of the prelim-
inary design and cost
estimate
Westinghouse
NEES
Esac
EPA
Westinghouse
SWEC
Esso representative on-site
consultant
Allied Chemical Corporation
NEES
Esso
EPA
Westinghouse
NEES
Esso
EPA
Westinghouse
SWEC
Combustion Engineering
NEES
Esso
EPA
January 18, 1973-August 1973
Sentember 6, 197?-March 19, 1974
March 1974-July 1974
August 6, 1974-December 9, 1974
Westinghouse served as prime contractor to EPA. NEES is the cooperating utility. Esso has developed
the process and is carrying out pilot-scale tests under contract to EPA. SWEC was a subcontractor
to Westinghouse to perform, engineering services. Allied Chemical Corporation provided design and
economic data on sulfur recovery. Combustion Engineering provided an evaluation of boiler
performance.
-------
• An initial design study to identify the critical cost
and performance factors (Appendix I)
• An assessment of the process options and the initial
design study to select the design basis for a demonstra-
tion plant which would meet the goals of the program
(Appendices J and K).
General Process Potions
The general process consists of six processing systems:
• The reaction svstem (gasification and regeneration)
• The stone processing svstem
• The sulfur recoverv system
• The limestone handling system
• The fuel oil handling system
• The waste/by-product stone handling system.
The general function of the reaction system is to generate a low
heating-value fuel gas for utilization bv the boiler. The stone
processing system converts the spent sorbent into a form suitable
for disposal or by-product utilization. The sulfur recoverv system
produces elemental sulfur from the sulfur captured in the gasifier.
The remaining three systems store and carry the raw materials
(residual oil or other liquid fuel and limestone) and the waste or by-
product materials which are utilized in the process. The three major
processing systems can be arranged in a number of ways to generate
alternative process schemes. Figure 4 illustrates the genera] process
options. The general regenerative process requires the utilization of
stone processing system and a sulfur recovery system arranged in parallel.
A number of processes can be utilized for both of these systems. The
general once-through process scheme eliminates the regeneration function
from the reaction system and reauires a stone processing system which
produces an environmentally suitable solid product utilizing all of the
sulfur removed from the fuel oil. Again, many options exist for this
44
-------
U1
Limestone
| Hot Fuel
Gas to Boiler
Recycled
Gases
*
Sulfur/
Sulfuric Acid
Limeston
Oil
Air
Hot Fuel
eas to Boiler
Recycled
Sulfided
Lime
Air
Calcium Sulfate disposal/utilization
Limestone
Oil
Air
Hot Fuel
Gas to Boiler
Recycled
Gases
Stone product disposal/utilization
GENERAL REGENERATIVE PROCESS DEMONSTRATION PLANT PROCESS
Figure 4-Oil gasification processing concepts
ONCE-THROUGH PROCESS
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stone orocessing system. A third general processing option is shown
in the figure which combines the regenerative and once-through
features to eliminate the sulfur recovery system and generate a high
sulfur-content solid product.
Numerous nrocess and design options have been identified
for the reaction system (Appendix J). These deal with areas such as
the technology basis, the market (new boiler versus retrofit,
utility versus industrial boiler), advanced concepts yet untested,
equipment and operating options. A variety of stone processing
options are discussed in Appendices K and 0. These include dry sulfa-
tion (Appendix L), slurry carbonation (Appendix M), acid sulfation
(Appendix N), dead-burning, and dry oxidation followed by drv
recarbonation (Appendix 0). Alternative processes for sulfur
recovery have also been investigated (Appendix p). Considerations
such as sulfur recovery versus sulfuric acid production, and
existing sulfur recovery processes versus developing technology,
have been made.
Initial Design Study
An initial design of the 50 MW CA^B demonstration plant was
carried out bv SWEC under contract to Westinghouse Research Labora-
tories. The basis for the initial design study was the utilization
of conventional catalytic cracking technology developed in the
petroleum industry for the solids transport systems, the specification
of equipment to provide extensive system flexibility and reliability,
and the inclusion of sulfur recovery from the snent stone. The study
identified the critical design parameters and cost items in the plant.
These data provided a basis for evaluation and identification of
alternative design and operating parameters to reduce the t»lant cost
while maintaining system flexibility and onerability.
The critical factors assumed in the design basis for the
initial design study are summarized for each of the processing systems.
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Reaction System
*
The initial design basis included items relating specifically
to process flexibility and reliability:
• Once-through ooeration canabilitv in addition to regener-
ative operation
• 2.44 m (8 ft) bed depth capability
• About 50 nercent excess capacity for stack-gas recycle
• Extremely high nressure drons assumed in air and stack-
gas circuits; °fi.5 kPa (14.n PS!) for control, line losses,
bed losses, and distributor losses and eauinment design
• Low gasifier nlenum start-up temperature and lime storage
svstem for start-up material
• Valves in the fuel-gas lines to isolate the fuel lines for
burnout during operation, fuel-gas lines based on IP.3 m/s
(60 ft/sec) velocity
• A solids transport system to circulate solids between the
gasifier and regenerator based on known catalytic cracker
dense phase pneumatic transport technologv. Required tall
standlegs 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 supnort plate on gasifier
• A stone ciuench drum used to cool waste solids before they
enter the stone processing system*
Sulfur Recoverv System
A commercial process to recover elemental sulfur from the
regenerator product gas (R volume percent S09) was specified. The
process (A]lied Chemical) is described in Appendix p.
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Stone Processing System
The stone processing svstem selected for the initial design
study was a slurry carbonation nrocess which removed carbon dioxide ^.--^i
from the plant stack gas to contact the regenerator waste stone and
produce a dry calcium carbonate product. The hydrogen sulfide (H.3)
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.
Waste Stone Handling System
The design basis for the waste stone handling system is a
four-times-stoichiometric limestone feed rate.
Table 11 lists the total installed costs of the six process
systems in the initial study demonstration plant in end-of-1974 dollars.
It is evident from the table that the reaction and sulfur recovery
systems represent the dominant costs In the initial study demonstration
plant, or 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. Detailed costs are presented in Appendix I.
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 will result in a more economical demon-
o
stration plant.
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Table 31
SUMMARY OF MAJOR SYSTEM COSTS IN
THE INITIAL STUDY DEMONSTRATION PLANT
System Total Installed cost (1QQQ $)
Reaction 4,511.00
Stone Processing j 1,102.00
(slurry carbonation) !
i
Sulfur Recovery 3,000.00
Waste Stone Handling 358.00
Fuel Oil Handling 825.00
Limestone Handling 174.00
TOTAL PLANT 9,970.00
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» The critical cost items in the reaction system are the gas -
compression components and the fuel-gas handling components*
and these may be greatly reduced in cost by proper design.
• A stone processing system should be utilized in the demon-
stration plant which permits the elimination of a sulfur
recovery system (Appendix K).
Selection of the Preliminary Design Basis
The selection of the design basis for the demonstration plant
from the identified process and design options and the results of the
initial design study rests on a number of general factors:
• Maximum demonstration plant operability based on pilot-plant
results
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o Cost of removal of No. 10 and No. 11 boilers to provide
space is not included, and cost of refractory-lined
1067 mm (42 in) piping for hot fuel gas is based on four
burners located together in near side of No. 12 boiler.
Tangential firing will require at least 50 percent more
1067 mm (4? in) pining.
e Boiler house, overhead coal bunker, ash silos, railroad
siding, rail car spotting equipment, and oil barge
unloading facilities are all existing and usable.
c Steam, boiler feedwater, service cooling water, electrical
power, plant air, Instrument air, and low-sulfur fuel
oil are all available by connection to existing supplies
at no cost.
• Continuous firing of present normal fuel at a 5 percent
minimum of boiler capacity; 50 MW power requires 191,000 kg/hr
(422,000 Ib/hr) steam. The hot fuel gas is estimated to
raise 195,000 kg/hr (430,000 Ib/hr). The boiler capacity
is 204,000 kg/hr (450,000 Ib/hr).
• Fuel oil feed weight-off has sufficiently fast response to
boiler fuel rejection signals so there is 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.
• The design parameters and effectiveness of absorber
reaction at 94 percent minimum efficiency (not more than
6 percent of sulfur dioxide and of calcium oxide unreacted)
will be tested and proved in the pilot plant.
9 Complete reaction of residual calcium sulfide in regenerated
stone occurs in the absorber (so that waste stone is
unreactive toward natural waters) and will be tested and
proved in the pilot plant.
• Suitable limestone supply is 483 km (300 miles) by rail.
Suitable waste solids disposal is 48 km (30 miles) by truck.
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a Horizontal pulsed flow for controlled stone circulation will
be tested and proved in a full-scale cold flow model.
» Air jet pulverizing of regenerated stone for feed to absorber
will be tested and proved in manufacturer's test equipment.
o The 50:1 scale-up of processes proved in the Esso pilot
plant is feasible and appropriate.
A detailed summary of the design basis and assumotions is given in
Appendix D. Imoortant factors in the design soeclfication are listed.
• Plant size - l>f> VW
• Fuel oil rate - 12,700 kg/hr (28,000 Ib/hr)
• Plant sulfur removal efficiency - 90%
• Plant turndown ratio - 4:1
• Fuel oil properties - Table D-l
• Limestone properties - Table D-l; a final selection of
limestone for the olant has not vet been made. Laboratorv
tests and pilot-plant tests are being carried out in order
to make the selection (Appendix Q).
• Gasification p'rocess concent selected - regenerative
process with dry sulfation of spent stone
• Reaction section conditions
Oasifier conditions/design
Operating temperature - 880°C (16160*1)
Pressure - boiler oressure nlus pressure losses through
cyclones, gas ducts, and burners 17.Q kPa (2.6 osi)
Air/fuel ratio - °OZ of stoichiometrlc maximum (includes
oxygen in flue Has recycle)
Superficial Huidizatlon velocity - 1.83 m/s (6 ft/sec)
Bed deoth (maximum) - 1.22 m (4 ft)
Sulfur removal efficiencv - to yield 90 percent olant
efficiency
Limestone makeun rate - 0.5 to 1.5 times stoichiometric
Number of gasifier modules - 1
Gaslfler temperature control - flue gas recycle with
orovision for steam and water In-fection
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Burnout - provision for continuous steam/air injection
into cyclone inlets/or burnout with flue gas during
shutdown
Turndown ratio - 4/1
Turndown method - control of air/fuel ratio with minimum
velocity reduction
Number of fuel-gas ducts - 4 (one oer burner)
Elutriation rate of sorbent from gasifier - 3,175 kg/hr
(7,000 Ib/hr)
Regenerator conditions/design
Operating temperature - 1080°C (1976°F)
Pressure - 0 to 2.5 Pa (0 to 10 in. H20) less than the
gasifier pressure
Air rate - 0.1% oxvgen in product gas
Superficial velocity - 1.83 m/s (6 ft/sec)
Bed depth - 1.22 m (4 ft)
Number of modules - 1
Temperature control - by stone circulation rate
emergency inert gas auench for temperature overshoot
Solids circulation system
Esso dense phase pulsed flow technique
Gasifier cyclones and fines circulation system
Four cvclones on gasifier set at minimum elevation
Cvclone fines returned to regenerator bv commercial
plug flow technique
Air circuits/flue gas circuit pressure drops
Minimize losses by proper design of blower control, gas
lines, and so on
Single blower for air and stack gas to gasifier
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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 instal-
lation to reflect the uncertainties involved in the scale-uo 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.
• Dry sulfation system
Pulverizer
Spent stone reduced in size to about 75 micron average
diameter while hot in commercial jet pulverizer (Jet-
0-Mizer, Fluid Energy Processing and Equipment Company,
HatfieId, Pa.)
Absorber
Bed temperature - 870°C (1598°F)
Fluidization velocity - 0.122 m/s (0.40 ft/sec)
SO- absorption efficiency - 90 percent
Energy and Material Balances
Common assumptions have been made for generating energy and
material balances having sufficient accuracy for a preliminary design
task. These include:
e No heat losses
e Steady state with no accumulations
o Air to blowers is at 25°C (77°F) and bone dry
e Purge and transport gas streams have been neglected
• The pulsed nature of solids transport systems have been
neglected by assuming steady flow behavior
When possible, data and correlations generated by the Esso
experimental programs have been utilized to carry out balances — for
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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 operating and environmental consideration.
Material and energy balances have been carried out for three cases:
• The base case condition at full capacity
• A case with plant turndown to 25 percent of capacity
o A case with a 17 percent of stoichiometric air/fuel
ratio as compared with the base value of 20 percent of
stoichiometric.
The balances are summarized in Figures 5 and 6.
Demonstration Plant Design
The Manchester Street Station, boiler No. 12, is a 46 MW,
balanced draft, CE unit producing 204,120 kg (450,000 Ib) steam/hr
at 510°C (950°F) and 83 bars (1200 psig). The boiler is designed
for both coal- and oil-firing. The plant site is described in
Appendix C.
The demonstration plant is a facility which will demonstrate
system operability, environmental performance, boiler performance, and
process retrofit capability. It will also provide a basis for establishing
commercial plant economics. 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
55
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three systems store and carry the raw materials (residual oil and
limestone) and the waste or by-product materials (sulfated lime) which
are utilized in the process. A detailed flow diagram of the reaction
system and the dry sulfatlon system is shown in Figure 7. Numbered
streams in the figure correspond to the numbered streams in the material
and energy balance figures.
Reaction System
The reaction system consists of
a A gasifier
• A regenerator
e Solids transport components
e Particulate removal components
a Air and stack-gas compression components
o Fuel-gas handling and combustion components
a Heat exchange components
a A purge and transport gas (PTG) system.
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 (HgS) 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 hot fuel gas flows to four boiler burners through
four separate refractory-lined pipes where combustion is completed
for steam generation. The sulfided lime (i> 6 wt % calcium sulfide
[CaS]) is transported pneumatically to the regenerator vessel at
31,400 kg/hr (69,419 Ib/hr). The calcium sulfide in the stone and the
carbon (^ 0.3 wt % CHi/2) are reacted with preheated air at a temperature
of 1075°C (1970°F) to generate calcium oxide (CaO), calcium sulfate,
sulfur dioxide (SO.) (ft mole %), and water vapor and carbon dioxide.
Regenerated sorbent is recirculated to the gasifier at 32,400 kg/hr
(?!,«.5 Ib/hr). The sulfur dioxide gas at 3,980 kg/hr (8,778 Ib/hr)
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and a regenerated stone purge stream at 297 kg/hr (655 Ib/hr) are
removed from the reaction system for nrocesslng In the stone processing
system (dry sulfation process).
The gasifier and regenerator vessels are conventional refractory-
lined fluidized bed vessels with carbon steel shells (shown in Figure 8).
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 3.7 m (12 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 152 mm
(6 in) above the grid through three independent oil rings surrounding
the vessel, which distribute oil to about 60 horizontal injection tubes,
2 2
with one injection point per 0.47 m (5 ft ) of bed cross section. A
single vertical baffle wall is placed in the bed and the regenerc tad
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 152 mm (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) operation temperature. No flow
baffles are utilized in this regenerator. In addition to the circu-
lating stone inlet and outlet, the regenerator has a waste stone drain
line located in the wall at about 0.92 m (3 ft) bed depth.
Control
The major control nolnts in the reaction section of the
demonstration plant are the gasifier temoerature and bed deoth, and the
regenerator temoerature and bed rtenth. The gasifier temnerature is
controlled at about *7i°r (1600°F) by the rate at which stack gas is
recycled to the unit. The demonstration olant can also utilize steam
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or water injection as a method for gasifier temperature control. The
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 gasifler-to-regenerator transport leg is controlled
by the regenerator temperature measurement. The regenerator bed depth
is controlled by the regenerator-to-gasifler 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 are 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.
Stone Processing System
Dry sulfation has been selected as a means of orocessing the
waste stone from the reaction system in the demonstration plant. The
dry sulfation stone processing svstem consists of the following
components:
e An absorber
0 A stone pulverizer
e Gas compressor
o Farticulate removal components
0 Heat exchangers.
The absorber, a fluidized bed reactor, is shown In Figure9.
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The dry sulfation process contacts the waste stone from the
reaction system and the boiler cyclone with the sulfur dioxide oroduct-
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 fluidized bed absorber. In order to enhance the rate of absorption and
the degree of sulfation of the stone, the stone is pulverized while hot
before being contacted with the sulfur dioxide. The dry sulfation process
design specifications are based on the kinetic study presented in Appendix
L. The sulfated by-product stone from the absorber is cooled and trans-
ported to the waste stone handling system.
Support Systems
The limestone, fuel-oil, and waste stone handling systems
function to suooort the two major processing systems in the demonstration
plant which have been described. These three support systems are conven-
tional and, therefore, 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 (29,000 Ib/hr). The waste/by-product stone
handling system transports and stores waste/by-product stone at about
1360 kg/hr (3,000 Ib/hr).
Plant Turndown
The demonstration plant is designed to operate stably at
specified levels of performance down to a partial load of 25 nercent of
full plant canacity. The stability of the reaction system is limited by
the minimum operable fluidization velocity in the regenerator and
gasifier. The minimum cmerable fluidization velocity of these units
is one-half of the full capacity design velocity. Thus, turndown of the
reaction system to 50 percent of full capacity can be carried out by a
simole linear reduction by "50 percent in the capacity of all process
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streams. Further reduction in the plant cutout must be carried out by
increasing the gasifier air-to-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.
Control of Coke Deposition
The deposition of coke on the walls of the fuel-gas pipes and
in the gasifier cyclones may be limited by proper design of the system
and may be controlled by periodic burnout of the system. Smoothing all
bends and transitions in the fuel-gas wipes reduces the formation of
coke-lime deposits at these points. Based on pilot-plant performance
(burnout required about every 200 hours of operation) the demonstration
plant may require burnout about every 2000 hours of continuous operation.
Demonstrated burnout procedures exist and are of the order of one percent
of the time between burnouts. The probable plant operating schedule will
be such that the gasification process may be burned out during scheduled
plant shutdowns.
Detailed descriptions of the demonstration plant are given in
Appendices B and E.
Plant Layout
Preliminary demonstration plant layout drawings are shown in
Figures 10 through 13. The drawings, plan views and elevations, show all
the major plant components, including the hot fuel-gas piping. Vessel L-l
is the gasifier; vessel L-2, the regenerator; vessel L-3, the absorber for
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stone processing; vessels G-8 (four of them), the gasifier cyclones; and
vessels G-9 and 6-19, the regenerator cyclones (in series). Vessel G-20
is the spent stone pulverizer and vessel G-17 is the pulverizer cyclone.
The absorber cyclone is vessel G-18. The arrangement of the equipment
has not been determined with regard to plant cost.
Boiler Modifications
An evaluation of the interfaces between the demonstration plant
boiler and the fluidized bed residual oil gasification process was carried
out. CE provided certain engineering services designed to establish
parameters for firing the high-temperature, low heating-value gas in the
CE steam generator (unit No. 12) at the Providence, Rhode Island plant
of NEES (see Appendix C). Westinghouse provided CE with a design basis
taken from conservative CAFB performance estimates (20 percent air/fuel
ratio; high stack-gas recycle requirement).
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 be-
cause of l.D. fan limitations. To achieve a 204,120 kg (450.000 Ib)
per hour steaming rate new l.D. fans must be supplied. New burners
will be required to fire the low heating-value gas. CE concluded that
the following modifications should be made to the boiler (Appendix F
presents further details of the evaluation):
• Tangential firing should be used. The base design using
all four burners in one boiler face may lead to materials
problems. Tangential firing provides advantages over
single-face firing. Two of the four fuel-gas pipes would
have to be extended about 15.2 m (50 ft) each for tangen-
tial firing capability.
• The unit side-walls must be modified for tangential
firing.
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• A feedwater line must be moved in order to insert air
ducts, and so forth. Other plant/boiler components
may require modification for access to the boiler.
• New air ducts for tangential firing must be added.
• A flame safeguard system must be provided for both the
low heating-value gas and No. 6 oil (dual fuel boiler).
• New supports designed for 704°C (1300°F) gas tempera-
ture must be added at the lower back of the economizer.
The normal gas temperature entering the economizer with
oil-firing is about 593°C (1100°F). The economizer
will be steaming with the low heating-value firing.
• In order to generate 204,000 kg/hr (450,000 Ib/hr) of
steam (the maximum boiler capacity) a new I.D. fan
would be required (based on the gasification process
base design conditions). Because of increased flue gas
volume leaving the boiler (higher flue gas temperature)
the existing I.D. fan would allow only about
182,000 kg steam/hr (400,000 Ib steam/hr) to be gene-
rated (again based on the gasification process base
design condition).
• Refractory plugs will be added to the existing horizon-
tal burners to maintain present firing capability.
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 per-
spective is of interest, it may not be necessary or desirable to adopt
this point of view for the demonstration plant. The study was also influ-
enced by the basis which CE used for the gasification process.
The evaluation CE made was strongly influenced by a less than
optimal understanding of the gasification process, its flexibility, and
its unique conditions. Certainly, some modifications would be required to
adapt the existing boiler to its role in a low heating-value gas
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demonstration plant. The 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 main-
tain maximum boiler efficiency, but this is not a requirement of the
demonstration plant. It appears that a substantial savings could be made
by using the four-burners-in-one-wall approach to boiler firing, and with
fewer potential problems than with longer, more convoluted fuel-gas pipes.
If this approach caused some loss in boiler efficiency it would not be sig-
nificant to the overall program. In addition, 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 should not be required for the demonstration
plant:
• The design 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.
Further review of the boiler performance and the recommended
modifications is required before a decision can be made regarding material
and equipment requirements.
DEMONSTRATION PLANT PERFORMANCE
The overall performance of the demonstration plant may be sepa-
rated into three performance categories:
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• Energy efficiency
• Environmental performance
• Technical performance of the components.
Energy Efficiency
The efficiency of the demonstration plant is dependent on three
plant aspects:
• The fuel oil energy losses in the gasification process
(thermal losses and carbon losses)
• The auxiliary power and steam required by the plant
• The efficiency with which the boiler utilizes the low
heating-value fuel gas (stack losses).
The energy losses in the gasification process consist only of
the energy required to heat and calcine fresh limestone feed and the neg-
ligible heat losses from the process equipment. The coke deposited on the
gasifier bed is circulated to the regenerator, where it is combusted to
raise steam as the regenerator sulfur dioxide gas passes through a waste
heat boiler. The gas stream from the dry sulfation absorber is recycled
to the gasifier. The losses from limestone calcination and heating amount
to about one-half of a percent of the fuel oil energy input to the plant.
Total losses from the gasification process would be expected to be about
one percent.
The auxiliary power and steam requirements for the demonstra-
tion plant (based on the base design specifications) are summarized
below (from Appendix E):
Steam - kg/hr (Ib/hr): 1030 kPa (150 psij»), 1R5°C (366°F)
Produced 1*10 (4,000)
Used (fuel oil feed system) 1410 (3,115)
Exported 401 (8fi5)
Power - kw 2944
Steam is produced at a net rate of 401 kg/hr (885 Ib/hr) for export from the
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hot regenerator-sulfur dioxide gas. The steam used by the plant is for
the fuel oil system. Thus, there is no net steam consumption for the
demonstration plant except for normal fuel oil system heating require-
ments. The auxiliary cower requirement represents about 6 oercent of
the total oower plant capacity for the base case conditions, Most of
this power is associated with the compression of gasifier air and
recycled stack gas. At an air/fuel ratio of 17 percent of stoichiometric
the plant auxiliary power would be reduced to about 2,200 kw or about
4.5 percent of the power plant output. These power requirements are
based on low efficiency blowers for gasifier air and stack-gas recycle
(50 to 60 percent adiabatic efficiency). Since the boiler existing F.D.
fan now handles only 80 percent of the normal combustion air quantity, some
power savings would be realized over the normal F.D. fan power reouirement.
CE (Appendix F) has estimated that (for the base conditions)
the boiler efficiency with the low heating-value gas will be about 82.5
percent, as compared with about 87 percent with oil firing. This is due
to higher stack-gas exit temperatures with the low heating-value gas.
With a 17 percent of stoichiometric air/fuel ratio in the gasifier, the
volume of gas passing through the boiler will be reduced, leading to
higher boiler efficiency. The demonstration plant boiler will be dual-
fired with No. 6 fuel oil to permit operation of the plant during periods
of gasification process shutdown. While the normal steam superheat
temperature of 510°C (950°F) can be achieved with the low heating-value
gas, a steam temperature of only 4R2°C (900017) can be achieved with No. 6
fuel oil. Maintenance of some damaged boiler baffles will increase the
No. 6 fuel oil superheat temperature.
Environmental Performance
One of the major goals of the CAFB demonstration plant program
is to demonstrate the ability of the process to utilize residual fuel
oils and vacuum bottoms in an environmentally attractive way. The
demonstration plant has been designed on the basis of the pilot-plant
results summarized in Table 12. The environmental impact of the process
is discussed in Appendix G.
83
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New federal Source Performance standards for fossil fuel-fired
steam generators of greater than 263.75 GJ/hr (250 million Btu/hr) heat
inputs are 0.34 kg/GJ (0.8 lb/106 Btu) for sulfur dioxide and 0.43 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/ln Btu).
Table 12
SUMMARY OF ENVIRONMENTAL PERFORMANCE OF
CHEMICALLY ACTIVE FLTIini7ET> BED (CAFB) PILOT PLANT
Sulfur Removal Potential
Sulfur Product
NO Control Potential
x
Limestone Makeup Rate
Removal of Metals
Fuel Oils Utilized
Processing Steps "Demonstrated
Up to 95%
Sulfur fixed as solid calcium sulfate
Normal 270 pom reduced to 155 ppm
0.5 to 1.5 times stoichiometric
inn% fuel vanadium, 757 rickel,
40% sodium
Atmospheric residual (vacuum bottoms
on batch unit tests)
Gasification and regeneration
Since the particulate emission from the oil gasification plant
consists mainly of calcium oxide, the environmental inpact of calcium
oxide particulate emission is also of interest. The stacV exhaust from
the oil gasification demonstration Plant is eauivalent 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.16
The rate of dry solids production from the demonstration plant
dry sulfation system is about 1360 kg/hr (3,000 Ih/hr). Several alterna-
84
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tives are being investigated for the disposal or utilization of this
material (Appendix S): land disposal, ocean dumping, deep mine injection;
or marketing for bv-product recovery as gvosum, elemental sulfur, or
sulfuric acid. Uses as construction material (concrete, aggregate, etc.)
are also being explored.
The CAFB demonstration plant compares favorably environmentally
with limestone and lime slurry scrubbers. Sulfur removal capabilities for
the processes also 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.
e 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.
0 The CAFB process 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.
Demonstration Plant Process Performance
The technology which is being utilized In the demonstration
plant preliminary design comes from three sources:
• Conventional chemical process plant practice
o Esso pilot-plant developments
o Laboratory bench-scale experimentation.
The process technical performance and reliability depends largely on
the source of the technology. Many aspects of the demonstration plant
design are based on improved plant reliability.
The provisions for oil feed, air feed, flue gas recycle,
limestone feed, and waste stone handling are all by standard equipment
in commercial use, and no abnormal problems are expected. Overcapacities
of 15 percent on oil, 25 percent on gas blowers and baghouse, and 50
85
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percent on limestone and waste stone are considered adequate to provide
for variations from process design requirements. The maximum oil feed
rate of 115 percent normal provides for 14,500 kg/hr (32,000 Ib/hr) fuel,
or 620 GJ/hr (588 MM Btu/hr) higher heating value to the system, or at
2 percent gasification heat loss, 615 GJ/hr (583 MM Btu/hr) to the boiler.
In addition to larger than normal flue gas flow, the requirement of
diverting 20 percent of combustion air to the gasifler with 80 percent
through the air preheater to the boiler results in an abnormally high
temperature from preheater to stack. Modifications to the preheater are
under consideration but are not included in the cost estimate of this
study.
The reliability and full capability of the boiler is protected
by supplemental firing with the normal low-sulfur fuel oil.
Gasification
The pilot-plant work provides a sound basis for anticipation of
90 percent or better removal of sulfur at a level of 2.6 percent in the
fuel. The depth of the bed of fluldized lime in the demonstration unit
is more than double that In the pilot plant. This should provide a
safety factor adequate to compensate for the degree of grossness in
air and fuel distribution caused by practical mechanical designs and
to allow for a modest recycle of sulfur dioxide from the absorber.
A possible 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. Pilot-plant runs have demonstrated
simple and short duration burnout procedures. The utilization of vacuum
bottoms has not been demonstrated in the Esso pilot-plant program; only
limited batch runs have been carried out.
86
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Regeneration
The pilot-plant work provides a sound basis for anticipation of
sulfur rejection at 8 percent sulfur dioxide in the off-gas and loading of
sulfur on the stone sufficiently low to provide pickup in the gasifier.
Sulfur Dioxide Absorber
No pilot-plant data on the absorber reaction has been generated,
although tests in large Esso batch units are scheduled. The basis for
the preliminary design has come from laboratory studies which Westinghouse
has carried out (Appendix L).
Solids Transport
The circulation of solids between the gaslfler and regenerator
by a dense-nhase pulsed transnort scheme developed by Esso has proved to
be a reliable circulation method, well suited to the CAFB process.
Full-scale cold modeling of the demonstration plant circulation system
will be carried out.
The Sturtevant systems for fines transport have been proved
in all-metal systems at temperatures up to 390°C (750°F), but there is
no experience with slug flow in refractory-lined nine. In case of
inoperability, substitution of Incolloy pipe Is a possible modification.
COST ESTIMATE
SWEC has completed a orellmlnarv cost estimate for the demon-
stration plant. The plant cost is given on the basis of end-of-1974
dollars and includes the fuel oil system, limestone handling system,
waste stone handling system, reaction system, stone processing system
(dry sulfation), and boiler modifications. The total plant cost does
not include fee, contingency, start-up cost, and interest during construc-
tion. The total olant cost of $7,350,000 resulting from the SWEC
study represents the cost for a demonstration plant with high excess
capacity factors for many items of equipment and no optimization where
high-cost plant components (such as the fuel-gas piping system) are
87
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involved. A breakdown of the plant cost is presented, and the contribu-
tion of the various equipment items and systems is discussed. The
costs of some processing options and design modifications are also
presented and discussed.
Demonstration Plant Cost Breakdown
The following tables summarize the demonstration plant capital
investment. Table 13 presents the total SWEC estimated plant cost by
general equipment classification. The high piping cost indicated in
the table (65 percent of total process equipment) is due to the fuel-gas
piping system.
Table 14 provides a cost breakdown according to the major plant
systems — reaction, stone processing, limestone handling, waste stone
handling, and fuel oil:
• The reaction system in the table includes boiler modifications
(fuel-gas burners and wind box) and plant buildings.
• The fuel-gas piping system (including supports and refractory)
has been shifted from the process materials category as it
appears in Table 13 to a process equipment item since it does
represent unusual process piping.
• The limestone handling system includes all handling equipment
from receiving UP to the point of injection of limestone into
the gasifier.
• The stone processing system accepts regenerator waste
stone and the cooled and cleaned regenerator sulfur-dioxide
gas and generates the waste/by-product stone.
The reaction system process equipment cost is about 65 percent of the
total plant equipment cost, and the stone processing system is about
17 percent of the total plant equipment cost. The reaction system and
stone processing system contribute about 77 percent of the total plant
direct cost.
Table 15 summarizes the costs of the equipment items in the re-
action system. The code identification for each item refers to the SWEC
88
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Table 13
COST SUMMARY
(50 Mi.' Plant)
Item
PROCESS EQUIPMENT
A. Towers
B. Boilers, S1 heaters
F. Process furnaces
C. 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. Insulation1"
N. Paintingb
Refractory lining
TOTAL PROCESS MATERIALS
TOTAL DIRECT COST (DM +DL)
DISTRIBUTABLE ACCOUNTS
VI. Insurance (excluding all risk)
V2. Federal and scaces taxes
X. Temporary construction facilities
Y. Field office (Including Insurance t
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
I'laterlal 'S)| Labor (S) I Total
32,000 -00 32
151,200 5,100 156
775,000 45,900 820
504,000 27,800 531
500 200
355,500 35,500 391
3,100 300 3
429.500 30.100 459
—
58,700 1.800 60
2,309.500 147.300 2,456
1,048,000 545,000 1,593
169,800 48,800 218
158,800 221,400 380
24,500 25,400 49
65,800 110,100 175
129,200 54,600 183
54,000 — 54
87,900 — 87
60.000 — 60
1,798,000 1,005,300 2,803
4,107,500 1,152,600 5,260
ind taxes) 850
850
6,110
1,250
1,250
7,360
(S)
,600
,300
,900
,800
700
,000
,400
,600
,500
,800
,000
,600
,200
,900
,900
,800
,000
,900
,000
,300
,100
,000
,000
,100
,000
,000
,100
Not included
15X
"From original SWEC table.
Subcontract.
89
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Table 14
PLANT COST BREAKDOWN BY PLANT SYSTEM
Demonstration olant system
a
Reaction
system
Stone processing
system
Limestone
handling
Waste stone
handling
T'uel oilc
System
Tocal
Process Equipment
vo
o
Material 1,830,655 504, 76O
Labor 211,600 22,120
Total Process 2,042,255 526,880
Equipment
N t
Process Materials ""
Material 1,003,085
Labor 521,000
Total Process 1,524,085
Materials
TOTAL DIRECT COST 4,OQ3,220
88,000
19,900
107,900
36,500
31,300
67,800
175,700
120,000
42,300
162,300
22,600
63,300
85,000
248,200
276,900
1,200
278,100
?14,400
239,700
454,100
732,200
2,820,315
297,120
3,117,435
1.276.S85
855, 30n
2,131,885
5,249,320
Includes major boiler modifications, plant buildings, and fuel-gas pining system (as a process equipment
item).
Includes limestone weight feeders and limestone star valve.
Demonstration plant is dual-fueled with residual oil and low-sulfur oil.
-------
Table IS
CAFB DEMONSTRATION PLANT
REACTION SYSTEM PROCESS EOIITPMF.NT COST
Item
I Materlal(S) I LnborfSl I Totnim |
Fxccna capacity
flaqlfler fl,-])
Regenerator i'L-7)
Gaslfler Tvclones
(C-RA-D)
Casifier Finns Recvcle
Svstem (C'l-A-n),
(G-22 A-D), (i- -•»!),
(G-23 A-D), (r.-23 F.-H)
Regenerator Cyclones
and Trickle Valves
(C-9), (G-9A). (G-10),
(G-10A)
Gaslfler Air/Flue Gas
Blower and Screen
(»-8>, (G-5)
Regenerator Air Blower
and Screen (D-6), (0-6)
Fuel-Cas Plnlnp 'iv'tom
ReRencrntor SO -r:as
System- ?
Waste Meat Roller (R-l),
Stenm Hrum (M-3), Hent
F.xchanRer (T-l), B«q
Filter, Wrapper, nnd
DlachnrRe VnJvp fn-1<>),
Flue Cns Rervcle Svstem-
Stack Cyclone Lock Hopper,
Valves and Connecting
PlplnR (R-2P), «:-V) A.B),
(G-37), Recvcle Flue Cae
Ba^house, Screw Conveyor
and Airlock (G-4).
(G-4 A.B), BaRhouse Lock
Hnpner, Valves and Con-
necting Plplnp. (r,-?o)i
(G-11 A-B), (C-33), Flue
f.as Booster Fan (R-4)
Purge f< Transport Gas
Svstem- PTC Blower
Suction Cooler (T-7),
PTC Blower fn-1), PTG
Cooler (T-3), PTG Water
Separator (M-4), PTG
Filter (R-7)
Boiler Modifications-
Fuel-Gas Burners (F-2A-D)
and Wlndbox (F-24-D)
Start-up Air Heater (F-l)
Portable Chunck Trap Vessel
(G-36)
722,000
41.000
211,000
188,100
47,600
208,750
39,230
••'1.415
58.300
86,100
fi 1,460
82,000
56,500
5,000
13,400
23"). 000
000 41,900
1.800 214,800
5,800 191,90n
900 48,500
20,300 229,0.r>0
1,800 41,030
150,000 671,415
1,500 59,800
Rased on 707 air/fuel ratio:
nlani. mny operate at 17Z air/
fuel ratio
10% volumetric flow excess
About 40-minute storage capacity
In lock hoppers - much larger
than required
15X volumetric flow excess
?OT excess power
30Z excess volume: 50Z excess
power
About 207 excess pipe diameter
7,000 03,100 About 257 volumetric flow excess;
757 power excess
1,400 64,860 About 257 excess volumetric flow,
3SZ excess power
3,400 85,400
1,100 57,600
300 5.300
TOTAL PROCESS EQUIPMENT 1,830,655 211,600 2,042,255
91
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identifications appearing in the process flow diagram (Figure 7). The
fuel-gas piping system equipment cost is about 33 percent of the total
reaction system equipment cost at $671,000. This piping cost includes
the shop-fabricated carbon steel pipe sections (1.17 m [46 in] and 1.22 m
[48 in] diameters at a total length of 74.7 m [245 ft]), elbows, tees, re-
ductions, flanges, welds, testing, bolt-up cost, hangers and supports, a
15 percent design allowance, refractory, and additional structures re-
quired for pipe system support. The table also lists the degree of over-
capacity SWEC has specified for each equipment item.
Table 16 provides a similar breakdown for the equipment in the
stone processing system. The major cost item in the stone processing
system is the absorber, about 48 percent of the total equipment cost.
More detailed cost breakdowns are included in Appendix E.
Assessment of Plant Cost
The estimated overcapacity (beyond the normal level
of safety factor) built into many of the items listed in Tables 15 and
16 adds approximately $500,000 to the total plant cost. The single item
having the largest overcapacity cost effect is the fuel-gas piping.
The original design basis of 45.7 m/s (150 ft/sec) gas velocity
(Appendix D) has been reduced by overcapacity to less than 30.1 m/s
(100 ft/sec), resulting in about a 21 percent increase in piping diameter.
The layout of the fuel-gas piping system and the gasifier location has
not been optimized for the preliminary design. Potentially, large cost
reductions could he realized by optimization. But, since the SWEC
design basis places the four burners in a single face of the boiler, a
plant with tangential burners at the four corners of the boiler would
result in a more comnlex and costly fuel-pas oioinp system (possibly
as rcuch as ^300,000 additional equipment cost for ?'0.1 m [100 ft]
additional length). The need for tangential burners has not been
definite]v established, although CE has recommended It.
The gssdfier is not an extremely imoortant cost item in the
total plant. Operating the p.asifier at 1.22 m/s (4 ft/sec) velocity
92
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VO
Table 16
CAFE DEMONSTRATION PLANT
STONE PROCESSING SYSTEM (DRY SULFATION) EQUIPMENT COST
Item
Absorber (L-3)
Absorber Cyclone and
Trickle Valve (G-18),
(G-lftA)
Jet Pulverizer (G-20)
Pulverizer Cyclone and
Trickle Valve (G-17),
(G-17A)
Absorber Air Blower
and Screen (P-9)
Absorber SO, Blower
(R-10)
Pulverizer Air Blower
and Screen (R-ll) , (G-24)
Solid Waste Cooler/
Conveyor (G-27)
Absorber Heater (F-3)
! Material ($)
241,000
24,ooo
34,200
7,600
33,130
67,000
37,730
47,500
12,600
1 Labor f$)
13,500
300
560
300
1,200
1,700
1,800
2,200
560
1 Total(S)
754,500
24 , 300
34,760
7,°00
34,330
68,700
39,530
49 , 700
13,160
| Excess capacity
35% excess
45% excess
27% excess
15% excess
oower
15% excess
power
5% excess
power
50% excess
volumetric
volumetric
volumetric
flow; 652
flow: 15%
flow
flow
flow
excess
excess
flow: 37? excess
flow
TOTAL PROCESS EQUIPMENT
504,760
22,120 526.880
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rather than at the 1.83 m/s (6 ft/sec) design velocity which has been
applied would result in a total plant cost increase of only about 5120,000.
Pilot-plant studies Indicate that a n.61 m (2 ft) bed depth may he
sufficient for the achievement of 90 percent sulfur removal, resulting
in a 0.61 m (2 ft) decrease in bed depth from the base design, rcn the
basis of pilot-plant results, this decrease would substantially reduce
the bed and pasifier distributor pressure drop and could reduce the fines
elutriation rate (with a 1.22 m/s F4 ft/sec1 velocity) by an order of
magnitude. Elimination of the gasifier cyclones and fines recycle system,
if possible, would result in a total Plant cost reduction of about $900,000
and an additional reduction in fuel-gas piping system complexity.
The preliminary design of the demonstration plant does not
represent a prototype of a commercial retrofit system. Optimization
of the plant (piping system layout, gasifier location and operating
conditions, and so on) and elimination of excess capacity would substan-
tially reduce the demonstration r-lant cost. The piping system represents
the largest cost and the most sensitive cost item in the plant.
Processing Options
In addition to the demonstration plant concept which was
selected for the preliminary design study (regenerative process with
dry sulfation of the waste/by-product), two other process concepts could
be investigated in a demonstration plant — once-through operation and
regenerative operation with sulfur recovery. From the initial design
study (Appendix 1) and the results of the final preliminary study on the
dry sulfation processing concept the cost of a demonstration plant using
sorbent regeneration, recovery of sulfur from the regenerator sulfur di-
oxide, and slurry carbonation of the waste/by-product material is esti-
mated to be about $10.5 million; the sulfur recovery process is estimated
to be about $3 million; the reaction system and support systems are as-
sumed to cost about the same as in the SWEC study; and the slurry carbo-
nation system cost is about the same as that of the dry sulfation system.
This cost would be reduced by the same factors discussed for the final
94
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preliminary design (in other words, optimization and overcapacity reduc-
tion) .
Once-through operation could result in a low-cost demonstration
plant if two assumptions were valid—very high lime utilization (more
than 75 percent) can be achieved and the highly utilized calcium sulfide
waste can be marketed without further processing. Assuming the validity
of these propositions, the cost of the demonstration plant would be about
$5.8 million without additional optimization or excess capacity reduc-
tion. With optimization of the plant (elimination of the gasifier cy-
clones and fines recycle systems; reduction in piping system length and
complexity) and reduction in the excess capacity, the once-through demon-
stration plant could cost as little as $4.5 million.
The processing concept using regenerative operation with the
sulfur recovery system used in the initial design study did not appear
attractive from a cost aspect. Other processes being developed, however,
may render this option economically competitive. Once-through processing
is potentially attractive from a cost aspect, but the assumptions of high
sorbent utilization (Appendix R) and direct disposal of calcium sulfide
(Appendix S) require extensive investigation.
Operating Cost
SWEC has estimated the demonstration plant annual operating
costs. Unit costs used are as provided by NEES, basis 1976; internal
power 23 mills/kWh; No. 6 fuel oil; $2.25/MM Btu (HHV); capital charges,
20 percent of bare cost/year; maintenance (including construction tools,
supplies and labor), 5 percent of bare cost/year.
Annual costs are based on 7000 hr/vr stream operation.
Bare Cost
Direct Operating Costs
Power
Operating labor
Limestone su^rlv
Solid Waste disposal
$7,3*0,100.00
104,500.00
76,?50.on
40.fi80.00
95
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Maintenance
Total 1,2*4,585.00
Indirect Operating Costs
Capitol charpes 1,47?,020.00
Total Direct & Indirect Costs 2,736,605.00
This annual operating cost amounts to a fuel adder of about 76C/MM Btu
or 7.8 mills/kWh. For an optimized plant with overcapacity reduced,
the fuel adder could be substantially reduced. Operation of the
plant with a cheaper vacuum bottons fuel, if available, could result
in a fnel credit relative to the cost of No. 6 fuel oil.
96
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VI. DETAILED, DESIGN, CONSTRUCTION, AND OPERATION
A scone of work, schedule, and cost estimate are presented here
for the Phase II work, detailed design, and construction. Shake-down
and commissioning operations and an experimental test program projected
for the plant are reviewed. Implementation of the second phase of the
program, however, detailed design and construction, is contingent on
the decisions reached by the project participants after evaluation of the
Phase I work.
SCOPE OF WORK
Selection of Architect-Engineer
A list of qualified architect-engineers will be preoared.
Specifications and procurement documents based on the preliminary design
will be prepared and distributed to the candidate architect-engineers.
Proposals for the detailed design and construction, and cost estimate
for operation (Phase III), will be received and evaluated and an
architect-engineer selected.
Detailed Engineering, Procurement, and Construction
The services of the architect-engineer will include but not be
limited to the following categories of work:
• Preparation of layouts, architectural drawings, and flow
diagrams to reflect changes from the preliminary design
caused by size and specifications of equipment finally
selected
• Completion of mechanical equipment soecifications and
vessel drawings to show actual sizes and orientation of
all interconnections
97
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• Preparation of structural detail and foundation drawings
for buildings and equipment
• Provision of a piping model plus piping orthographies for
all unmodeled piping
• Provision of electric lighting, power wiring, and heat
tracing installation drawings, including motor control
center layouts
• Provision of instrument sizing and selection, instrument
installation details, and final panel layout drawings
• Provision of ten copies of all manufacturers'parts lists,
spare parts recommendations, operating instructions, and
shop drawings
• Provision of critical path scheduling to monitor on a
bimonthly basis the timely performance of the work
• Provision of a Division of Work and a Mechanical Completion
Definition to allocate responsibility between Westinghouse,
NEES, and the architect-engineer for all phases of the work
• Fabrication and/or procurement of all necessary materials,
equipment, and fabrication services to complete the plant
• Performance and completion of all construction and installa-
tion work necessary to erect the demonstration plant at
the Manchester Street Station, Narragansett Electric Company,
Providence, R.I.
• Receiving and storing of all material and equipment, and
procurement and furnishing of all manual labor, supervision,
and subcontractor services necessary to perform the work to
the best advantage.
• Arranging for the provision of all utilities and services
for the construction and testing of the facility, payment of
all applicable taxes and obtaining of all necessary permits.
• Connecting the demonstration plant into the existing
electrical, steam, water, sanitary sewage, compressed air,
natural gas, and storm sewer systems at battery limits
98
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• Performance of pressure testing of piping and vessels,
testing of mechanical functioning of the equipment, and
final calibration of the instruments as a minimum, prior
to certification of mechanical completion.
SCHEDULE
The overall program, as shown on the computerized critical path
plot figure, is projected to require three years to complete. The value
of the critical path scheduling technique is that it enables priorities
to be established on the basis of material lead times and provides a
logical sequence for interconnecting activities. When looking at Figures
14 and 15, one can see each activity described above a line between two
boxes with numbers.
The lead box is called the predecessor and the follow box is
called the successor event, as shown on the critical path listing
attached.
The information below the line is the estimated duration of
the specific activity in weeks. The most critical path is located on
the top of the page, and subsequent paths are listed by their criticality.
The durations for the various activities shown on the schedule
(Table 17) are assumptions generally valid for labor efficiencies and for
labor and material availabilities prevailing in the fall of 197A.
Westinghouse expects that individual activity durations may vary considerably
from some of those shown on the schedule. This is, however, unlikely to
affect significantly the overall time required to complete the demonstration
plant construction.
COST ESTIMATE
An estimate tabulation is presented in Table 13 to show the de-
tailed costs of the oil gasification system with dry stone sulfation in-
stalled at the Manchester Street Station, including new high-sulfur fuel
oil tankage, piping and pumps, new hot-fuel gas burners, and new limestone
99
-------
o
o
a
&
a-
a
o
8
I
7
1
1 '
1 '
1 ' !t
1
1
. SSifMStt" -i
1 nee . fa v^_>_
1 '
1 '
1 '
onjM nm. »»«.
1
1 l
I
;x.
iSSffii.il'
i
1 m* i
l
i
L
:x,
' "'5.- •'«.!, I ^
1 ' AffikfiJ1," •
iiu it oy-n.r
^—^—^ risoac MIGH itir »«i*n
1 - mo u.o/-n.t l
m™B — "^^ £.a/-to.» pCjoo- 's?iaL >UWBI"
1 1 j- sau IB o/-n r
' MBCUM ELIHRICM. CBUir
, t , ' ' "" "-1"-"' ' J flTMQSPHER]C
• nn t o/-n t
OIL GRS1F1CPTION
l "™". lS?J1^;.'3,I-™.t i j HER! TRflNSFEfl Di VIS JON
HBSji.tfJ'^ """ "I"«-5"1«"'-1™ i ' LEGEND
"S'-'n"^. . _,, ' ~ ""•' J NON-CB.T PflTM
^ L-MIJ— «»_ ^^.jj,- ~-^ FLBBl
raBOBf run. SWCH
- 1UI M O/i.t I
"" """•'
cn
1 11-M It/*
0.0 10.0
20.0 30.0 10.0 SO.O 60.0 70.0 80.0 90.0 100.0 110.0 120.0 130.0 140.C 150.0 160 0 170.0 IbO.O
nra/ii/rv Mio/nsTf jn/io/n JM/tm ocr/n/rr ww/it/io
-------
in iv "»H)*i '»" ™«omieiii
lien ilnnt - mil a o/n.»
Stiff,,
.J
0.0
10.0 ZO.O 30.0 40.0 50.0 60.0 70.0 80.0 90.0 100.0 110.0 120.0 130.0 140.0 150.0 160.0
JUL/C/TS m/n/71 wn/i tsti niG/ii/Tf jiu/ikviT JM/UTI on/n/n
170.0 180.0
Figure K -Critical path diagram
-------
FOtAS SYSTEM AT»GAS
DEPORT 22 - ACTIVITY SCHEDULE SORTED BY
Table 17
ACTIVITY SLACK
DATA DATE- 2/IS/7S
o
to
INPUT
ORDER
1*
2*
3*
4«
S«
6*
7«
8*
9«
10*
11*
12*
13*
14*
15*
16*
17
18
19
20
2l
22
23
24
25
26
27
28
29
30
31
32
33
34
35
36
37
38
39
10
1
2
3
4
5
ft
14
19
29
31
40
41
8
10
II
13
30
12
9
18.
49
54
57
42
43
58
59
27
28
37
38
39
16
17
44
45
44
50
SI
22
PRED
20005
2000S
2001
2002
2002
2004
2005
2100
2215
9001
9002
3015
3998
501
502
SOS
9001
502
501
3998
3998
3999
4000
9002
3115
9001
3635
9002
33(5
3968
9002
3215
2100
2195
3015
3135
3042
3015
30SS
21TO
SDCC
2OOI
2Q01
2Q02
2004
ZOOM
2Q05
2 100
2215
9 001
'002
30l5
3998
5CI
502
503
?999E
9(302
*999E
502
3999
3999
IOOD
»999E
3 1 | 5
3998
3635
3968
3315
3998
3998
3215
3998
2|95
3998
3|35
3042
3968
3055
3998
2294
ACTIVITY
DESCRIPTION
REC APPROVAL PHASE 2
REC APPROVAL PHASE 2
ISS POX ENGINEERING CONSULTANT
FINALIZE FLOW DIAGRAM
PHEP HEAT 6 MATERIAL BALANCE
APP FLOW DIAG I H£AT BALANCE
DESIGN MAJOR EQUIPMENT
SPECIFY PURCHASE RESALE EQUIP
PPEP PLOT PLAN
BUILD SCALE MODEL
SPECIFY / PIPING
PROCURE PIPING
PKEP INSTRUCTION ROOK
PREP START-UP MANUAL
APP START-UP MANUAL
SHIP START-UP MANUAL TO SITE
PKEP PIPING i, INSTR DIAG.
SHIP INSTRUCTION ROOKS TO SITE
APP INSTRUCTION BOOK
INSTALL BOILER MODIFICATIONS
INSTALL'' MECH EQUIPMENT
INSTALL INSULATION
INSTALL PAINTING
SPECIFY HIGH TEHP VALVES
PROCURE HIGH TEMP VALVES
SPECIFY / SUPPORT STEEL
PROCURE SUPPORT STEEL
SPECIFY / ELECTRICAL EQUIP
PROCURE ELECTRICAL EQUIP
INSTALL SUPPORT STEEL
SPECIFY / INSTRUMENTATION
PROCURE INSTRUMENTATION
SPECIFY BOILER MODIFICATIONS
PROCURE BOILER MODIFICATIONS
SPECIFY / "ISC FOUNDATIONS
ISS PO/ MISC FOUNDATIONS
IMSTALL/MISC FOUNDATIONS
SPECIFY / PIPE SUPPORTS
PROCURE PIPE SUPPORTS
DESIGN FUEL SYSTEM
DURATION EXPECTED
EST EFF START
4|0
1.0
1.0
3.0
3.0
2.0
23,0
H.O
2.0
10.0
14,0
6S.Q
16.0
6.0
2.0
1.0
8.0
4.0
3.0
12.0
12.0
6.0
6.0
I6.Q
60.0
11.0
60.0
11.0
60.0
io. n
n.o
S2.Q
20.0
62.0
11.0
1.0
B.Q
13.0
21.0
11.0
ISFEB7S
ISFEB75
1SHAR75
!2ApR7S
I2APR75
03MAY75
17MAY75
ZSOCT7S
3IJAN76
14FfB76
24APR76
HA(.iG76
I2N0V77
04MAR78
I5APR78
29ApR78
I1FEB76
ISApR7B
0«tMAR78
I2N0V77
I2NQV77
04FEB78
18MAR78
24ApR76
HAuG76
HFEB76
22MAV76
24ApR76
3|JuL76
I6JUL77
24ApR76
3|JuL76
250CT7S
13MAR76
I1AU676
20NQV76
I80FC76
11AuG76
25N0CT775
EXPECTED
FINISH
12MAR75
I2HAR7S
Q9APR75
30APR7S
30APR75
11MAY75
220CT75
28JAN76
IIFEB76
2IAPR76
| IAUG76
09NQV'7
OIMAR78
(2APR7B
26APR78
21MAY78
07APR76
IOMAV78
22MAR78
OIFEB78
OIFEB76
I5MAR78
26APR78
IIAUG76
Q50CT77
19MAY76
I3JUL77
28JUL7A
2ISEP77
21SEP77
28JUL76
27JUL77
IOMAR'6
I8MAY77
|7NOV76
ISDEC76
09FEB77
|ONOV76
06APR77
28JAN76
SLACK
•27.2
-27.2
-27.2
• 27.2
-27.2
-27.2
-27.2
-27.2
-27.2
-27.2
-27.2
-27.2
-27.2
-27.2
-27.2
-27.2
-25.2
-25.2
-21.2
-23.2
-23.2
-23.2
-23.2
-22.2
-22.2
-20.2
-20.2
• 20.2
-20.2
-20.2
-12*2
-12.2
-2.2
•2.2
1.8
1.8
1.8
3.8
.8
.B
3
S
EPA
NEES
PUR3
A
A
ENGI
A
A
A
USM
A
PUR3
1BW
IBN
ENGS
CAP
A
CAP
ENGS
FS2
FS2
FS2
FS2
A
PURS
A
PURS
A
PURS
FS2
A
PURS
A
PURS
A
PURS
FS2
A
PUR 3
-------
Table 17(Continued)
FOCAS SYSTEH ATM&AS
REPOeT 22 • ACTIVITY SCHEDULE SORTED BY
ACTIVITY SLACK
DATA DATE- 2/15/75
I
2
3
^
5
6
7
a
9
10
11
13
11
IB
16
17
19
v I_
)ER
21
20
IS
52
53
35
36
32
33
34
17
18
23
21
25
55
56
26
7
PRED
229i|
2215
2100
30|5
3075
3635
36|S
9001
3515
3522
9002
37|5
3315
3335
3312
3015
3095
33|S
2000S
SuCC
3998
*99B
3998
3075
3999
3968
3515
3522
3966
3715
3998
3335
3312
»999E
3Q9S
IQOO
3335
*999E
PROCURE
PROCURE
PROCURE
SPECIFY
PROCURE
SpEclFv
ISS pO/
SPECIFY
ISS PO/
INSTALL
SPECIFY
ISS PO/
SPECIFY
ISS PO/
INSTALL
SPECIFY
PROCURE
DESCRIPTION
FUE«- SYSTEM
PURCHASE RESALE EQUIP
MAJOR EQUIPMENT
/ INSULATION
INSULATION
/ SUPPORT STEEL INSTL
SUPPORT STEEL INSTL
/ EQUIP FOUNDATIONS
EQUIP FOUNDATIONS
Equip FOUNDATIONS
/ MECH ERECTION
MECH ERfECTION
/ ELECTRICAL INSTALL
ELECTRICAL INSTALL
ELECTRICAL
/ PAINTING
PAINTING
PREP owes/ ELECTRICAL
PHASE 2
60.0
53.0
65. 0
12.0
22.0
13.Q
8.0
1.0
8.0
H.O
1.0
11.0
1.0
12.0
12.0
6.0
1*0
82.0
3UAN76
3IJAN76
250CT75
I1AUG76
06NQy76
??HAYl*
I1FEB76
!OApR76
08MAY76
2«UPR76
3|JtiL76
3|JuL76
06NQV76
11*1)676
06NnV76
3|Jl)L76
I5FE87S
23MAR77
Q2FE877
19JAN77
03NOV76
06ApR77
I5SEP76
07APR76
QSMAY76
30JUN76
28JUL76
2SAUG76
03NQV76
OIOEC76
23FEB77
03NOV76
5.8 PUR3
12.a PUR3
11.8
19.8 A
2SAUG76
08SEP76
22.8
22.8
33.8
33.8
33.8
35.8
35.8
37.8
37.8
37.8
11.8
11.8
t7.e
61.8
A
PUR3
FS2
A
PUR3
A
PURS
FS2
A
PUR3
A
-------
unloading and conveying systems.
SWEC has assessed the cost of dismantling the existing boilers
to provide space for installing the gasifier to be about equal to the
recovered scrap value for the boiler tubes, shell, and steel structure
within the boiler.
Note that the tabulation is based on materials take-off for
instruments, switchgear, piping, and wiring. This permits the use of
as low an order of accuracy as the 15 percent stated in the tabulation.
Costs are tabulated on the basis of materials and equipment quotations
valid for November 22, 1974, for a mechanically complete system ready for
shake-down operations.
SHAKE-DOWN AND COMMISSIONING OPERATIONS
Following mechanical completion, NEES will operate the plant
with the architect-engineer observing and correcting construction errors;
Westinghouse and Esso will serve as consultants. Commissioning operations
will include, but not be limited to
• Operating the retrofitted boiler, utilizing low-sulfur oil
provided by NEES, on oil-firing to prove the reliability
and turndown capability of the new dual-fuel burner system
on oil.
• Operating the oil gasification system, utilizing low-sulfur
oil provided by NEES, to prove the adequacy and reliability
of the solids circulation, particulate removal, oil and
air supply and distribution, and start-up heating systems.
• Operating the oil gasification system at design capacity
and sulfur removal for at least 48 hours continuously,
utilizing oil of design sulfur content (3 percent) to be
provided by NEES, to confirm formally that the work has
been satisfactorily completed. The architect-engineer's
representatives shall be present during the performance of
these tests, and upon their satisfactory completion shall
so certify to the prime contractor.
104
-------
• The prime contractor submitting to the architect engineer,
not later than 90 days after the certification of mechanical
completion, a list of additional work deemed necessary for
acceptance of the construction from the architect-engineer.
Upon completion of the listed work, the construction shall
be considered accepted by the prime contractor.
EXPERIMENTAL TEST PROGRAM
NEES will operate the plant with Westinghouse and Esso serving
as consultants. The experimental test program will include, but not be
limited to, operations to investigate the following variations:
• High-sulfur oil (5 to 6 percent by weight) to be gasified at a
reduced capacity to determine the sulfur removal efficiency
for such feedstock
• Operating rates down to 25 percent of design capacity to be
tried, to determine the practical limitations on system turndown.
• Air/fuel ratios above and below design to be attempted to de-
termine the effects of variation from design ratio (20
percent) at full capacity and at reduced capacities.
• The regeneration reactor to be operated under sulfate gener-
ation conditions to explore the practicality of once-through
operation to produce an Innocuous solid waste containing
all the sulfur derived from the oil feed.
• The gasifier fluidized bed to be operated at various levels to
determine the effect of bed level on sulfur removal
efficiency
• The gasifier to be operated with alternative temperature
control methods, e.g. water/steam, to determine the effect
on performance.
During the above operations, the plant will be monitored to
observe the following performance parameters:
• Variation in boiler performance (steam production, boiler
efficiency, boiler tube erosion/corrosion, and particulate
105
-------
removal effectiveness) determined during all stable periods
of test operation.
• Coking in the fuel-gas system systematically monitored during
all periods of stable test operation to determine, if
possible, those conditions conducive to minimum coke laydown
• Pollution control effectiveness of the gasification system
monitored to determine nitrogen oxide, sulfur oxide, partleu-
late, and metal contents of the flue gases during all
periods of stable test operation
• Materials of construction checked periodically to determine
their effectiveness in resisting erosion from fluidized
solids and corrosion from sulfldes, chlorides, trace metallic
elements, and so on.
Following delineation of demonstration plant operating parameters
from the tests and observations described above, the gasification system
will be operated as follows to provide design data for additional oil
gasification installations:
• Tests will be made to determine the maximum output of fuel
gas available continuously from the gasification system,
using temporary additional fuel-gas combustion apparatus
such as a smokeless ground flare if necessary
• An attempt will be made to operate the plant continuously
for a period of 90 days at minimum fuel-gas coke laydown
conditions to demonstrate reliable continuous operation
for base load application during periods of heavy system
demand or of large unit outage.
• Various limestones and fuel oil types will be used to
determine their suitability for use in the CAFB system.
• Solids circulation and control systems will be operated
to determine the most efficient parameters for pulse flow
gas/solids circulation rate variations.
• The dry stone sulfation system will be operated to determine
sulfur recovery and waste stone size relationships.
106
-------
VII. COMMERCIAL PLANT DESIGN AND COST ESTIMATE
BASIS
The demonstration plant preliminary design and cost estimate
presented in Sections V and VI of this report present the basis for the
design of that 50 MW prototype facility. On the basis of certain
assumptions the 50 MW design was prorated to a capacity of 200 MW. The
main assumption was that a 200 MW gasifier, 12.2 m (40 ft) in diameter,
was mechanically feasible and would be readily operable. Oil refinery
experience, where fluid cokers and fluid cat crackers of 12.2 to 18.3 m
(40 to 60 ft) diameter are operated routinely, was used as a basis for
the 200 MW gasifier design. Other assumptions for the 200 MW design are
that:
• All aspects of operation of the 50 MW demonstration unit
will function as projected with no more than minor modifi-
cations, 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 approximate 12.2 m (40 ft) diameter required, so more
expensive stainless steel grids must be used; these have an
underside temperature limitation of 593°C (1100°F), which
requires added provision for charging the gasifier with
precalcined stone for every start-up following a bed removal
and that such grids of this size are technically feasible.
• A regeneration/dry sulfation system with sulfur removal
efficiency 90 percent, turndown ratio 4:1, same boiler
efficiency, waste solids suitable for landfill disposal will
be used
107
-------
• There are normal air/fuel ratio, limestone makeup design
rate stoichiometric, suitable limestone supply at 483 rail
km (300 rail miles), waste disposal area at 48.3 truck km
(30 truck miles)
• Outdoor installation is adjacent to the boiler; no change
in heat conservation is allowed for
• An existing boiler using No. 6 fuel oil, with all storage
and handling facilities for oil, also rail siding and car
spotting equipment, is available
• Eight gas burners of double the capacity of the 50 MW plant,
also eight gasifier cyclones and eight Sturtevant transfer
systems for hot fines>will be used
• Real estate will be provided at no charge to the unit
• There is a clear site, with no overhead or underground
obstructions; soil bearing pressure of 144 kPa (3000 psf);
water table below all foundations, and seismic zone zero
• Limestone is one (1) times stoichiometric equal to 4/1.5 or
approximately 2.7.
DESIGN
Except for the use of metal alloy construction in the gasifier
and absorber grids, the 200 MW design closely approximates the design for
the 50 MW plant as presented in Appendix E. Single unit blowers and
compressors were used, and new concrete silos were included for fresh and
waste stone storage. These silos were not required for the 50 MW unit
because of the availability of existing coal bunkers for fresh stone
storage and the use of existing ash silos for spent stone storage at the
Manchester Street Station.
The 200 MW design assumes that fuel-gas line carbon burnout
will be demonstrated in the 50 MW plant without the need for fuel-gas
line shutoff values. It is further assumed that fuel-gas line shutoff
valves will not be required for boiler safety shutdown. Fuel oil valves
and gasifier steam purge piping will be used to stop the flow of fuel to
108
-------
the boiler and to purge with steam to eliminate the possibility of a
boiler explosion. This steam purge option is based on a similar design
being used by CE in the design of a 4.5 Mg/hr (5 ton/hr) coal gasification
boiler fuel supply system.
COST
The estimate tabulation in this section (Table 18) is for the 200
MW size oil gasification prorated from the 50 MW estimate tabulation presented
in Section VI of this report. The 200 MW costs have been prorated along
conservative lines. The compressors, for example, have been estimated
at the same S/W ($/hp) cost as for the one-fourth size compressors in
the 50 MW estimate; and the instrumentation costs have nearly tripled
for the 200 MW unit compared to the 50 MW unit. The 25 percent accuracy
stated in the estimate tabulation could, because of the conservative
nature of the cost proration, permit a real cost as much as 25 percent
less than the $21,583,500 shown in the tabulation. This becomes
especially true when the possibility of a 17 percent air/fuel ratio
operation is considered, compared with the nominal 20 percent used in
the 200 MW design. The lower air/fuel ratio could reduce the 200 MW bare
cost to $18,300,000. Further, less conservative and possibly more
realistic cost proration could reduce the $18,300,000 to $17,100,000,
considering only the instrumentation and compressors. This is 15.7 percent
less than the bare cost in shown in Table 18.
PERFORMANCE
Reference should be made to the discussion of performance in
Section V of this report. The large 200 MW unit can expect some very
nominal improvement in performance due to lower heat losses and to the
slightly more efficient use of power by the larger blowers and compressors
when compared to the 50 MW design. This better performance, however, is
insignificant when compared with the savings in power possible with the
use of a 17 percent air/fuel ratio in the 200 MW unit. An improvement
109
-------
Table 18
PLANT COST SUMMARY
(200 Mi; Commercial Plane)8
Item
PROCESS EQUIPMENT
A. Towers
B. Boilers, S- heaters
F. Process furnaces
G. General equipment
L. Reactors
H. Drums
Q. Storage tanks
P. Pumps (Including drives)
R. Compressors (Including drives)
S. Stacks
T. Heat exchanges
TOTAL PROCESS EQUIPMENT
PROCESS MATERIALS
C. Piping
D. Structures
E. Electrical
11. BuUdingsb
J. Civil
K. Instruments
N. Insulation15
N. Painting1*
Refractory lining
TOTAL PROCESS MATERIALS
TOTAL DIRECT COST (DM +DL)
DISTRIBUTABLE ACCOUNTS
VI. Insurance (excluding all risk)
V2. Federal and state taxes
X. Temporary construction facilities
Y. Field office (including insurance c
Z. Construction tools and equipment
0. Other distributable items
TOTAL DISTRIBUTABLE
SUBTOTAL-COST OF WORK
U. INDIRECT ACCOUNTS
Engineering
Design
Other headquarters office
Taxes—headquarters payroll
Overhead allowance
TOTAL INDIRECT (HEADQUARTERS OFFICE)
TOTAL, PRESENT DAY PRICES (BARE COST)
Fee, Escalation, and Contingency
Order of Accuracy
"prom original SUEC table.
Subcontract.
•laterlal ($) I Labor ($)
90,
520,
1,800,
1,300,
5,
1,350,
50.
1,720,
—
ISO,
6,985,
3,300,
900,
530,
250,
240,
350,
300,
320,
180,
6.370,
13,355,
ind taxes)
__
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
I Total (S)
._
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
25%
110
-------
in efficiency of at least 1 percent, to 2 percent, can be projected for
the 200 MW unit as compared with the 50 MW unit.
OPTIONS
The 50 MW plant operation will be used to explore process
options for such things as bed height, air/fuel ratio, stone size,
type and rate, and so on. The 200 MW commercial plant design could be
modified to:
• Use multiple train (either two, three, or four gasifiers) to
provide a more flexible or a more reliable installation.
A four-gasifier 200 MW design has been estimated to cost
34 percent more than a single gasifier 200 MW design.
Thus, the $17.1 million bare cost discussed in the cost
section above would become about $22.9 million for a four-
gasifier design. A single gasifier design should have
reliability comparable to that of fluid cokers in the oil
industry, where one year's operation or more is common. It
is possible that with sufficient experience the system
reliability could approach that of fluid cat cracking, where
three years between shutdowns for maintenance is expected.
• Have a larger than 12.2 m (40 ft) gasifier diameter, with
a concommitant increase in capacity, such that an 18.3 m
(60 ft) gasifier would provide 450 MW of fuel gas. The
commercial fluid bed designs in the oil industry have gone
to 18.3 m (60 ft) diameter successfully. The market survey
discussed earlier in this report, however, would indicate
that the market for very large single-train units would not
be great. The 200 MW size or smaller would be adequate for
the majority of potential installations.
Ill
-------
VIII. DEVELOPMENT REQUIREMENTS
Development work in support of a demonstration plant program
would be required in several areas in order to assure successful commercial-
ization and to develop the full capabilities of the fluidized bed residual
oil gasification process. Areas where further work is recommended include:
• The utilization of vacuum bottoms on other high metal content
fuels
• The sulfur removal/spent stone processing system
• The assessment of the environmental aspects of stone
disposal options
• The utilization of spent sorhent
• The demonstration of process components
• The review of safety requirements
• The development of alternative and advanced subsystems
for the process.
Two subsvsterns are recommended for pilot-scale tests prior to
completion of the demonstration plant: the pulsed flow solids transport
system and the dry sulfation spent stone processing svstem. A full-scale
cold model test of the solids transport system proposed for the demon-
stration plant should be carried out. This test program vrill
provide full-scalp, data on the transfer system which was successfully
demonstrated in the ESPO pilot plant. The facility would provide design
information on:
• The parameters associated with dense phase transfer of
solid particles. These parameters include: particle size
and size distribution, fluidization requirements (air
3 3
velocity kg solids/m /[lb solids/ft air]), mode of operation
(continuous or pulsed) solids transfer rates, and distance
of particle transfer.
• Transfer duct geometry including: size and shape of
opening, vertical position and angle, duct material
112
-------
(surface roughness - degree of corrosion) and orientation
of the air nozzles
• Transfer performance with respect to variation on the
operating conditions. This includes the conditions of
fluidization and various conditions associated with the
generation of fines (in other words, location of the fines
return line, quantity of fines generated, and so forth).
• Operating control methods dependent on the solids transfer
system. This includes both pressure and temperature
difference induced control schemes and would involve
assessing both the degree of control with respect to the
dynamics of the particle transfer and an evaluation of
specific equipment that could be used to effect this control.
The spent stone processing system selected for the demonstration
plant should be tested in pilot-scale fluidized bed reactor
units. These tests are required to check projections from laboratory
tests and to develop operating experience.
Experimental work is scheduled to continue to investigate
• Stone selection and stone performance — investigation of
reactivity, attrition, trace element release, potential for
high utilization for once-through operation
• Spent stone processing — investigation of alternative processes
• Environmental impact — investigation of spent stone analysis,
leaching, and emissions; stack emissions including sulfur
dioxide (S02), nitrogen oxide (NC^), particulates, trace
elements.
Support work recommended on the demonstration plant design
includes evaluation of recommended equipment (for example, carry out
pulverizer tests with spent stone from the pilot plant to obtain
performance data), assessment of boiler modification options developed
during the preliminary design, assessment of the cost sensitivity of
critical operating and design parameters, review of safety requirements
for the gasifier and gaslfier-boiler interface, review of burnout and
113
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turndown procedures in light of recent pilot-plant tests carried out by
Esso, and a continuing evaluation of the environmental impact and applica-
ble regulations. Work is also recommended to develop improved understanding
of the process operation:
• Continued operation of the Esso pilot plant to investigate
areas such as carbon burnout procedures, control of carbon
laydown, control of sorbent attrition and elutriation,
temperature control schemes, improved emissions control
levels, critical operating parameters (such as air/fuel
ratio, bed depth, fluidization velocity, etc.), alternative
limestones, operation with vacuum bottoms.
• Identification of potential alternative fuels for utilization
with CAFB technology (fuel availability, market, processing
techniques, economic projection, environmental impact).
In order to maintain perspective and to develop the maximum
utilization of the technology, market studies must be continued. The
availability of fuel and sorbent for the process must be monitored to
project the impact on the market. Competitive technology must be assessed
to determine its impact on the application of the oil gasification technology
to the utility and industrial markets.
114
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IX.
1. Archer, D. H., et al. Evaluation of the Fluidized Bed Combustion
Process. Reoort to Office of Air Programs. Environmental Protection
Agency. Westinghouse Research Laboratories. Pittsburgh, Pa.
Contract 70-9. NTIS PB 211494 and PB 212916. November 1971.
Vols. I and II.
2. Keairns, D. L., et al. Evaluation of the Fluidized Bed Combustion
Process. Vol. IV. Office of Research and Development. Environmental
Protection Agency. Westinghouse Research Laboratories. Pittsburgh,
Pa. EPA-650/2-3-73-04Rd. Contract 68-02-0217. NTIS PB 233101.
December 1973. 322 pages.
3. McGlamery, G. G. and R. L. Torstrick. Cost Comparisons of Flue Gas
Desulfurization of Systems. (Presented at the EPA Flue Gas
Desulfurization Symposium, Atlanta, Georgia, November 1971).
4. Craig, J. W. T., G. L. Johnes, G. Moss, J. H. Taylor, D. E. Tisdall.
Study of Chemically Active Fluid Bed Gasifier for Reduction of
Sulfur Oxide Emission. Office of Air Programs. Environmental
Protection Agency. Esso Research Centre. Ablngdon, England.
Contract CPA 70-46. June 1972.
5. Craig, J. W. T., G. Moss, J. H. Taylor, D. E. Tisdall. Sulfur
Retention in rluidized Beds of Lime under Reducing Conditions.
Esso Research Centre. Abingdon, England. (Presented at Third
International Conference on Fluidized Bed Combustion. Hueston Woods.
November 1972).
6. Availability of Limestone and Dolomites. M. W. Kellogg Company.
Task 1, Final Report, RED-72-1265. 1972.
7. Malhotia, 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.
115
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P.. Kett, T. K., G. C. Lahn, V. L. Schuette. Flexicoking - A Versatile
Residuun Conversion Process. (Presented at the 67th Annual MChF
Meeting, Washington, T\C. December 1^74).
9. Devitt, T. W. and F. K. ^ada. Status of Flue Gas Desulfurization
Systems in the United States. (Presented at the EPA Flue Gas
Desulfurization Symnosium, Atlanta, Georgia, November 1974).
10. Borgwardt, R. H. F.PA/RTP Pilot Studies Delated to Unsaturated
Operation of Lime and Limestone Scrubbers. Control Systems Labora-
tory . Environmental Protection Agency. (Presented at the EPA
Flue Gas Desulfurization Symnosium, Atlanta, Georgia, November 1°74).
11. Van Ness, R. P. Operational Status and Performance of the Louisville
FGD System at the Paddy's Run Station. (Presented at the EPA Flue
Pas PesulfurizatJon Symposium, Atlanta, Georgia, November 1974).
1?. Billon, A., G. Heinrich, A. Hennico, and P. Bonnifay. Cleaning up
the Bottom of the Barrel. (Presented at the 67th Annual AiChE
Meeting, Washington, D.C., December 1974).
13. Hydrocarbon Processing. Refining Handbook Issue. September 1974.
p. 312.
14. Private communication. Stanford Research Institute. Palo Alto,
CaJifornia. 1972.
15. Guthrie, K. M. Capital and Operating Costs for 54 Chemical Processes.
Chemical Engineering. June 15, 1970. p. 140-156.
16. Lund, H. F. Industrial Pollution Handbook. McPraw Hill Co. New York.
116
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LIST OF ABBREVIATIONS
A/F
BCURA
CaC03
CAFB
CaO
Ca(OH)2
CaS
Ca(SH3)2
CaSO,
CE
CH
CH,
1/2
CH3OH
CO
co2
Consol
COS
cs2
EPA
GJ
HDS
H2S
H2S04
kg
kPa
kw
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
117
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MgC03
MgO
ml
mm
m/s
MW
NEES
2
OCR
ORD
ORM
Pa
PTG
rms
sio2
so2
so3
SRE
SWEC
TG
- magnesium carbonate
- magnesium oxide
- milliliter
- millimeter
- meters per second
- megawatt
- New England Electric Systems
- nitrogen dioxide
- oxygen
- Office of Coal Research
- Office of Research and Development
- Office of Research and Monitoring
- pascal, S.I. unit of pressure
- purge transport gas
- root mean square
- silicon dioxide
- sulfur dioxide
- sulfur trioxide
- sulfur removal efficiency
- Stone and Webster Engineering Corporation
- thermogravimetric
118
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TECHNICAL REPORT DATA
(Please read luunictiuns on the reverse before completing)
\ REPORT NO
EPA-650/2-75-027-a
3.
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. fl
5 REPORT DATE
March 1975
6. PERFORMING ORGANIZATION CODE
7 AUTHORIS
'D. L. Keairns, R. A. Newby, E. J. Vidt, E. P.
8'Neill.C.H. Peterson,C.C.Sun,C.D. Buscaglia, and
. H. Archer
8. PERFORMING ORGANIZATION REPORT NO,
9. PERFORMING ORGANIZATION 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. TYPE OF REPORT AND PER
Phase I; 5/73-12/74
RT AND PERIOD COVERED
14. SPONSORING AGENCY CODE
IS. SUPPLEMENTARY NOTES
is. ABSTRACT This voiume of tne report summarizes results of 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-
fications, vessel drawings, equipment arrangement drawings, a site plan, an elec-
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 50 MW 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. cos AT I 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
3. DISTRIBUTION STATEMENT
Unlimited
19. SECURITY CLASS (Tin's Report)
Unclassified
21 NO. OF PAGES
137
20. SECURITY CLASS (This page)
Unclassified
22. PRICE
EPA Form 2220-1 (9-73)
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