EVALUATION OF THE
FLUIDIZED BED
COMBUSTION PROCESS
Submitted to:
Office of Air Programs
Environmental Protection Agency
Contract No. CPA 70-9
By:
Westinghouse Research Laboratories
Pittsburgh, Pennsylvania
Volume I TECHNICAL EVALUATION
-------
EVALUATION OF THE FLUIDIZED BED COMBUSTION PROCESS
SUMMARY REPORT
Contract'No. CPA 70-9
November 15, 1969 - November 15, 1971
Prepared for
Office of Air Programs
Environmental Protection Agency
Research Triangle Park,
North Carolina 27711
Proiect Officer:
P. P. Turner
By
Westinghouse Research Laboratories
Pittsburgh, Pennsylvania 15235
Authors
D. H. Archer, D. L. Keairns
. J. R. Hamm, R. A. Newby
W. C. Yang, L. M. Handman
L. E1ikan
-------
PREFACE
The Office of Air Programs (OAP), of the United States
Environmental Protection Agency, has organized and is sponsoring a
fluidized bed fuel processing program. Its purpose is to develop and
demonstrate new methods for utilizing fossil fuels -- particularly
coal and oil -- in industrial boilers and in utility power plants.
These methods should:
.
Meet air pollution abatement goals for S02, NOx' ash and
smoke emissions
.
Compete economically with alternative means .for meeting
these abatement goals.
Westinghouse has been assigned process evaluation tasks in
the fluidized bed fuel processing program.
These tasks have included:
.
Conducting studies predicting fuel availability and usage
costs; market surveys forecasting industrial boiler and
utility power plant installations; and analyses projecting
air pollution control improvement. The purpose of these
investigations was to provide a basis for boiler design
specifications -- capacities, steam conditions, pollutant
emission restrictions, etc. -- and to determine the possible
improvements in environmental and economic impact of various
improved steam/power generation processes.
.
Designing an industrial fluidized bed boiler and two utility
boilers -- one carrying out combustion at one atm, the other
at 10 atm pressure.
All of the boilers use a limestone/
dolomite sorbent in the fluidized bed for sulfur removal.
Systems for regenerating and recycling the sorbent and
for recovering the sulfur in useful form have also been
i
-------
designed.
All necessary auxiliaries for the industrial boiler
and complete power plant equipment for the utility boilers have
also been specified so that a complete system can be visualized.
The purpose of these boiler system designs is to evaluate the
performance and costs of industrial boilers and utility power
plants employing fluidized bed combustion.
Promising areas
for future development have been identified.
Westinghouse has
been assisted in this work by Erie City Energy, Foster Wheeler,
and United' Engineers.
The design and evaluation process has
also pointed up technical areas where more information is
needed and further development is required to meet the goal
of a non-polluting, economical means of producing steam/power
from foss'il fuels.
.' .' Conceptualizing a fluidized bed combustion boiler development
. .
plant. Preliminary designs have been produced for a development
plant which will make. possible an attack on the remaining tech-
nical problems in those areas where fluidized combustion has
grea.test potential for reducing air pollution and for economical
generation of steam/power. Preliminary estimates of the cost
for such a plant have also been made.
.
Providing technical consultation and assistance on the OAP
fluidized bed fuel processing program, including both combustion
and. gasification processes. Technical and economic comparisons
have been carried out on various fluidized bed. fuel processing
systems and various. conventional means of steam/power generation.,
An atmospheric pressure oil gasification/desulfurization process
has' been assessed in detail as an adjunct to a utility boiler.
The r.esults of the various surveys, designs, evaluations, and
comparisons are presented in this three-volume report.
The report
identifies fluidized bed fuel processing systems which should meet both
market requirements and air pollution abatement requirements and which
are.likeiy to be cheaper than alternative, conventional systems.
it recommends a program for commercializing promising processes.
Finally,
ii
-------
This volume (Volume II) contains an evaluation of available
data, an evaluation of fluidized bed combustion concepts for steam/
I .
power generation, a comparison of alternative concepts, concl~sions
and recommendations.
Volume I is a summary of the complete report,
and Volume III contains the detail surveys, design reports, and support
.s tudies.
iii
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ABSTRACT
The effectiveness and economics of fluidized bed combustion
boilers in pollution abatement and stearn/power generation have been
evaluated.
A 250,000 lb/hr coal-fired, factory-fabricated, industrial
boiler has been designed along with all of its auxiliaries. Utility
boilers have been designed for operation at atmospheric pressure in
conventional steam power plants and for operation at 10 atm in a com-
bined stearn and gas turbine power plant.
Overall capital and operating
costs have been estimated for both .300 and 600 megawatt plants.
Of all the systems studied, the pressurized fluidized bed
combustion boiler operating in a combined cycle power plant appears
most effective in meeting projected emission standards and in reducing
total S02' NOx' and particulate emissions, and most economical in power
generation. Such a plant has an estimated capital cost 20% less and
energy cost 10% less than a conventional utility plant with a stack
gas scrubber; this plant also has the greatest potential for increased
generation efficiency and reduced costs.
A fluidized bed oil gasification-desulfurization system has
also been designed and evaluated as an add-on unit for reducing S02
emissions from utility boilers burning high-sulfur oils. The estimated.
capital cost of such a unit is as much as 50% less than an add-on wet
scrubbing system.
It is recommended that a developmental pressurized fluidized.
bed boiler unit of 10 to 30 megawatt capacity be designed and constructed
to provide necessary technical information for a prototype power plant.
It is also recommended that the installation of a 50 to 200
megawatt demonstration fluidized bed oil gasifier-desulfurizer at a
utility site be pursued.
v
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TABLE OF CONTENTS
VOLUME II
PREFACE.
............
. . . .
. . . .
. . . . .
ABSTRACT. .
. . . . . .
. . . .
. . . . .
. . . . .
1.
ASSESSMENT OF FLUIDIZED BED COMBUSTION
PRELIMINARY SURVEYS. . . . . . .
. . . . . . . .
. . . . . . . .
FLUIDIZED BED COMBUSTION DATA. . .
. . . . .
. . . .
Sulfur Dioxide Removal. . . . . . . . . . . . .
Sorbent Regeneration and Sulfur Recovery Systems. . . . . .
Minimization of NOx Emissions. . . . . . . . . . . .
Combustion Efficiency . . . . . . .
Heat Transfer. . . . . ". . . . . . . . . . . . . . .
Particle Carry-Over. . . . . . . . . . . . . . . . . . . .
Boiler Tube Corrosion, Erosion and Fouling. . . . . . . . .
Gas Turbine Erosion and Corrosion. . . . . . . . . . . . .
INDUSTRIAL BOILER APPLICATION. . .
. . . . .
. . . . . . .
Design Basis. . . . . . . . . . . . . . . . . . . . .
System Specifications. . . . . . . . . . . . . . . . .
Fuel and Limestone Specifications. . . . . . . .
Design. . . . . . . . . . . . . . . . . .
Boiler System Design. .
. . . .
. . . .
. . . . . . .
Boiler System Schematic. . . . . . . . . . . . . . . .
Energy and Material Balances. . . . . . . . . .
Industrial Boiler Design Concepts. . . . . . . . . . .
Selected Steam Generator Design. . . . . . . . .
Coal and Limestone Feed Systems. . . . . .
Particulate Removal System. . . . . . . .
Pollution Control System. . . . . . . . . . . .
Selection of Draft Equipment. . . . . . .
Overall Design Layout. . . . . . . . . . . . . . . . .
Boiler Operation and Performance. . . . .. . . . . .
Operating Procedures. . . . . . . . . . . . . .
Performance Characteristics. . . . . . . . . . . . . .
Controls and Instrumentation. . . . . . . . . .
vii
Page
i
v
2
5
6
48
72
81.
99
119
132
156
164
165
165
165
165
174
174
177
177
183
191
196
196
198
200
207
208
217
221
-------
Boiler Cost. . .
00...............
. . . .
Capital Cost of Once-Through Dry Solid and Wet
Scrubbing Industrial Fluidized Bed Boilers. . . . .
Comparison of Capital Cost Between Fluidized Bed
Boilers and Conventional Boilers. . . . . . . . . .
Comparison of Total Cost Between Fluidized Bed
Boilers and Conventional Boilers. . . . . . . .. .
Effect of Operating Variables on Cost and Design of
Industrial Boiler. . . . . . . . . . . . . . . . .
Recommendations. . . . . . . . . . . . . . . . . . .
ELECTRIC UTILITY APPLICATION.
. . . .
. . . .
.....
. . . .
Introduction. . . . . . . . . . . . . . . . .
Pressurized Fluid Bed Boiler System. . . . .
"-
"-
. . . .
. . . .
System Concept. . . . . . . . . . . . . . . . .
Specifications. . . . . . . . . . . . . . . . . . . .
Power Cycle Analysis. . . . . . . . . . . . . .
Growth Evaluation. . .. . . . .\. . . . .
Boiler Subsys tem . .'. . . . . . . . . . . . . .
Regeneration/Sulfur Recovery System. . . . . . . . . .
Auxiliary Equipment. . . . . . . . . .. .' . . . .
Power Generation Equipment. . . . . . . . . . . . . .
Stack Gas Coolers. . . . . . . . . . . . . . . . . . .
Operation and Controls. . . . . . . . . . . . .
Control and Instrumentation. . . . . . .. . . .
Plan t Layout. . . . . . . . . . . . . . '. . . .
Plant Performance. . . . . . . . . . . . . . . . . . .
Recommendations. . . . .
. . . . .
. . . . . . . . .
Atmospheric Pressure Fluidized Bed Boiler Design. . . . . .
Specifications. . . . . . . . . . . . . . . . . . . .
Cycle Selection. . . . . . . . . . . . . .
Boiler Design. . . . . . . . . . . . . . . . . . . . .
Regeneration and Sulfur Recovery System. . . . .
Systems Analysis. . . . . . . . . . . . . . . .
Economics. . . . .
. . . . . .
.....
. . . . .
. . . .
Capital' Costs.
Energy Costs.'
Evaluation. . .
. . . .
. . . . .
. . . . .
........
. . . . .
. ." . .
. . . . .
. . . . .
.........
Comparison. . . . . . . . . . . .
Poten tia1. . . . . . . . . . . . . .
. . . .
. . . . .
. . . .
viii
Page
224
224
227
227
234
240
241
241
242
242
242
246
277
277
287
292
297
312
313
320
320
325
348
349
349
351
. 352
361
376
389
390
398
404
404
413
-------
2. ASSESSMENT OF A FLUIDIZED BED
OIL GASIFICATION-COMBUSTION POWER SYSTEM
.INTRODUCTION . . . . . . . . . . . . . . . . . . . . . . .
GASIFICATION/DESULFURIZATION CONCEPTS. . . . . . . . . . . . . .
DATA EVALUATION. . . . .". . . . . . . . . . . . . . . . .
The Esso (England) Program. . .
Gasification. . . . . . . . . . . .
. . . . .
. . . . . .
. . . . . . .
. . . . .
Carbon Deposition. . . . . . . . . . . . . . . . . . .
Gasifier Product Analysis. . . . . . . .. . . .
Desulfurization.. . . . . . . . . . . . . . . . . . .
Regeneration. . . . . . . . . . . . . . . . . . . . . . . .
Particle Carry-Over." . . . . . ." . . . . . . . . . .
Metal Retention. . . . . . . . . . .
Burner Design. . . . . . . . . . . . . . . . .
. Preliminary Results of Continuous Operation. . . . . . . .
. . . . .
MATERIAL AND ENERGY BALANCES. .
. . . . . . .
. . . .
. . . . .
Modes of Operation. . . . . . .
Design Parameters. . . . . . . . . .
Temperature Control. . . . . . . . . .
. . . . .
. . . . . . . .
. . . . . .
. . . .
Regenerative Operation. . . . . . . . . . . . .
Once-Through Operation. . . . . . . . . . . . .
Overall Material and Energy Balances. . . . . . . . . . . .
Power Requirements. . . . . . . . . . . . . . . . . .
DESIGN CONCEPTS. . . . . .
. . .. . .
. . . . . . .
. . . .
Retrofit Concepts.
......
.....
.. . . .
. .. .. .. .
Retrofit Location. . . . . . . . . . . . . .
Boiler Performance and Boiler Modifications.
. . . . .
. . . . .
New Boilers
Burners. .
Turn-Down .
. . . . . .
. . . . .
. . . . .. . .
. . .. .:.
. .. .. ..
. . . . .
.. . .. ..
. . . . . .
.. .. .. . .
. . .. .. .
.. .. . . . .
. .. . . .
ECONOMICS. . . .
.. .. .. .. .
. .. . .. .. .. .
. .. .. .
. . .. .. ..
Cost Factors. .. .. .. .. . .. .. .. .. .. .. .. .. . . .. .. .. . .
Regeneration Operation Economics--New and Retrofit. . . . .
Once-through Operation Economics--New and Retrofit. .
TECHNICAL EVALUATION. . . . . . . . . . . . .
. . .'. .
. .. . ..
Comparison Between Regenerative and Once-Through Operations
Comparison with Other S02 Control Methods. . . . . . . . .
Preliminary Design Conclusions. . . . . . . . . . . . . . .
RECOMMENDATIONS. . . . . . . . . . .
. .. .. .. .. .. .. .
ix
Page
423
427
431
432
436
436
438
444
452
454
455
455
455
456
457
458
463
463
482
488
492
495
496
496
501
509
5iO
513
514
515
518
522
525
526
527
529
531
-------
3.. FLUIDIZED BED COMBUSTION DEVELOPMENT PLANTS
DEVELOPMENT PLANT ALTERNATIVES
. . . . . . . .' . . . . . . . . .
PRESSURIZED FLUID BED COMBUSTION BOILER DEVELOPMENT PLANT.
Objectives. . . . . . . . . . . . . . . . . . . .
Plant Concept. . . . . . . . . . . . . . ~ . .
Conceptual Design. . . . . . . . . . . . . . .
Flow Diagram and Material Balance. . . . . . 0 . . . .
Equipment. . . . . . . . . . . . . . . . . .. 0 . . . .
Space Requirement. .. . . . . . . . . . . ... . . . . .
Power Requirement. . . . . . . . . . . . . . . . . . .
Experimental Program. . . . . . . . . . . . . .
Cost Estimate. . . . I,) . . . . . .' . . . . . . . . . .
Development Plant Schedule. . . . . . . . . . . . . .
. . . '0 .
ATMOSPHERIC PRESSURE FLUIDIZED BED OIL GASIFICATION-COMBUSTION
. DEMONSTRATION PLANT . . . . . . . . . . . . . . . . . .
Objectives. . . . . . . . . . . .
Plant Concept. ..
Conceptual De~ign . . . . . . . .
. . . . .
. . . . .
. . . .
. . . . .
. . . . . .
. . . . . . .
Flow Diagrams. . . . . . . . . . . . . . . . . . . . .
Equipment. . '. . . . . . . . . . . . . . . . . . . . .
Boiler Capacity. . . . . . . . . . . . . . . . . . . .
Experimental Program. . . . . . . . . . . . . . . . .
Demonstration Plant Schedule. . . . . . . . . .
REFERENCES. . .
ACKNOWLEDGEMENTS
. . . . . .
. . . .
. . . . . . . .
. . . .
..........
. . . .
x
Page
535
.535
535
537
539
539
539
545
545
545
546
548
550
551
552
553
553
553
559
560
560
563
573
-------
Number
1-1
1-2
1-3
1-4
1-5
1-6
1-7
1-8
1-9
1-10
1-11
1-12
1-13
1-14
1-15
1-16
LIST OF FIGURES
Title
Points where calcium and sulfUr can be measured
Effect of Ca/S on sulfur retention for various coals
Effect of bed temperature on sulfur removal
Effect of superficial gas velocity on S02 removal
Phase diagram for limestone and dolomite
Effect of MgO/CaO ratio on stone reactivity
Sulfur removal effectiveness
, '
Differences in effectiveness of sulfur ,removal with
same limestone, but different coals
Regenerated sorbent activity
, '
Variation in sulfur dioxide reduction with sorbent
par~icle size, bed depth and bed temperature
Effect of stone size on sulfur dioxide reduction
Rate constants for Dolomite 1337 as a function of
calcium oxide utilization
Ye/Ys vs particle diameter for limestone 1359
Theoretical S02 removal efficiency using partially
sulfated stone
Sulfate-sulfide reduction regeneration schematic
Sulfate-oxide reduction regeneration schematic
xi
Page
14
15
17
22
23
25
27
28
30
33
34
42
43
46
51
60
-------
I-
1-17
1-18
1-19
1-20
1-21
1-22
1-23
1-24
1-25
1-26
1-27
1-28
1-29
1-30,
1-31
1-32
1-33
1-34
1-35
Equilibrium constants
.66
66
Equilibrium constants
Equilibrium constants
67
68
Equilibrium constants
Equilibrium constants
68
68
Equilibrium constants
Equilibrium constants
68
69
Equilibrium constants
Equilibrium constants
69
69
Equilibrium constants
Equilibrium constants.
69
Equilibrium constants:
70
Equilibrium constants
TO
Equilibrium constants'
70'
Equilibrium constants
70
Heat content (enthalpy) above 770F datum'
71
Heat content (enthalpy) above' 77°F'datum
71
Effect of oxygen concentration on NO' gener.atfon:
7'5,
78
Temperature effect on NO generation
xii
-------
1-36
1-37
1-38
1-39
1-40
1-41
1-42
1-43
1-44
1-45
1-46
1-47
1-48
1-49
1-50
1-51
1-52
1-53
1-54
Solid carbon loss from bed without recycle
Solid carbon loss from bed without recycle
Combustion efficiency in carbon burn-up cell
Solid carbon loss from bed with internal recycles
CO loss from bed with or without internal recycles
Effect of fluidizing velocity on elutriated solid
carbon loss
Solid carbon loss with fines recycles
CO loss with fines recycle
85
85
86
88
89
91
92
Effect of bed temperature on heat transfer coefficient 102
92
Effect of particle diameter on heat transfer
coefficient
Effect of fluidizing velocity on heat transfer
coefficient to walls
Effect of fluidizing velocity on heat .transfer
coefficient to immersed surfaces
Variation of heat transfer coefficient with position
of heat transfer surface
Change of bed-tube heat transfer coefficient with
pitch/diameter ratio
Change of heat transfer coefficient with tube spacing
Fluidizing velocity curves for 1 atm and 10 atm
operation
Effect of fluidizing velocity on particle size and
carbon distribution of e1utriated solids
Size distribution of coal feed
Projected particle size distribution for material
elutriated from fluidized bed combustor and carbon
burn-up cell .
xiii
104
106
107
110
114
114
120
125
126
127
-------
,.
1.,.55 '
1-56
1-57
1-58,
1-59
1-60
1-61
1-62
1-63
1-64
1-65
1-66
, 1-67
"1-68,
1-69
1-70
, i-7i'
i-72
1-73
Vapor pressure of alkali compounds pr~~ent in coal
Effect of probe temperature on rate of deposition
Effect of time on weight loss
Cumulative weight loss per unit area as a function of
time for medium Csteel
Cumulative weight loss per unit area as a function of
time for 2-1/4% chrome-ferritic steel
Cumulative weight loss per unit area as a function of
time for 12% chrome-ferritic steel
Cumulative weight loss per unit area as a function of
time for austenitic type 316 steel
Cumulative weight loss per unit area as a function of
time for austenitic type 347 steel
Cumulative weight loss per unit area as a function of
time for austenitic esshete 1250 steel
Size consists of coal crushed by Koppers Reversible
Hammermill
Fluidizing velocity and terminal velocity for 1 atm
system
Projected particle size distribution for material
elutriated from fluidized bed combustor and carbon
burn-up cell
System schematic--once-through scheme
System schematic--wet scrubbing scheme
Heat and material
balance--dry solids S02 controi
Heat and material balance--wet scrubber system
, Boiler arrangement
, Boiler sections
Fuel system development
xiv
137
140
145
151
151
151
152
152
152
168
169
173
175
175
179
181
185
187
193
-------
1-74
1-75
1-76
1-77
1-78
1-79
1-80
1-81
1-82
1-83
1-84
1-85
1-86
1-87.
1-88
1-89
1-90
1-91
Proposed arrangement--once-through dry solid system
Proposed arrangement--wet scrubber system
Turn-down scheme of four-bed industrial boiler
Pressurized fluid bed boiler power system
High-pressure fluid bed boiler power system
Effect of first compressor pressure ratio and
overall pressure ratio on net plant heat rate
Effect of first compressor pressure ratio and
overall pressure ratio on plant output
Effect of first compressor pressure ratio and
overall pressure ratio on plant cost
Effect of first compressor pressure ratio and
overall pressure ratio on energy cost
Effect of first compressor pressure ratio and
overall pressure ratio on net plant heat rate
Effect of first compressor pressure ratio and
overall pressure ratio on plant output
Effect of first compressor pressure ratio and
overall pressure ratio on plant cost
Effect of first compressor pressure ratio and
overall pressure ratio on energy cost
Effect of first compressor pressure ratio and
overall pressure ratio on net plant heat rate
Effect of first compressor pressure ratio and
overall pressure ratio on plant output
Effect of first compressor pressure ratio and
overall pressure ratio on plant cost
Effect of first compressor pressure ratio and
overall pressure ratio on energy cost
Effect of first compressor pressure ratio and
overall pressure ratio on net plant heat rate
xv
203
205
213
243
247
249
249
.249
249
250
250
250
250
252
252
252
252
253
-------
1-92
1-93
1-94
1-95
1-96
1-97
1-98
1-99"
1-100.
1-101
1-102
1-103
1-104
1-105
1-106
1-107
1-108
1-109
1-110
Effect of first compressor pressure ratio and
overall pressure ratio on plant output
Effect "of first compressor pressure ratio and
overall pressure ratio on plant cost
Effect of first compressor pressure ratio and
overall pressure ratio on energy cost
Combined cycle plant performance diagram
Discharge-flow percent design value
Part-load characteristics with operating modes 1 and 2
in sequence"
Steam cycle T~S diagram
Temperature-energy diagram for pressurized utility
boiler
Feedwater temperature in high-pressure fluidized bed
boiler
Temperature of steam to reheater in high-pressure
fluidized bed boiler"
"Operating lines for high-pressure fluidized bed system
with variable bed temperature load control
Turn-down characteristics of combin~d cycle with
fluidized bed boiler operating on mode 3
Elevation--pressurized fluidized bed steam generator
for combined cycle plant .
Mass balance--design load
Mass balance--70% plant load
Energy balance--design load
Energy balance--70% plant load
Pressurized regeneration system--flow diagram
Pressurized regeneration system--plant layout
xvi
253
253
253
255
262
264
269
270
272
272
274
278
282
285
285
286
286
288
290
-------
1-111
1-112
1-113
1-114
1-115
1-116
1-117
1-118
1-119
1-120
1-121
1-122
1-123
1-124"
1-125
1-126
1-127
1-128
1-129
1-130
1-131
1-132
Coal handling and feed system schematic
Petrocarb pressur~zed coal feeding system
Schematic of particulate collection system
Operation of Aerodyne particulate separator
Plot plan--steam generators and accessory equipment
" Westinghouse gas turbine--internal arrangement
TG3000 Compound-cycle gas turbine--schematic cycle
arrangement and operating diagram "
Gas turbine V95.
Section through combustion chamber
Plan and elevation of W50l gas turbine modified for
external combustion
Detail of piping arrangement for modified W501 gas
turbine
Installation drawing of modified WSOl gas turbine
Flow diagram for plant start-up
Combined cycle plant--composite flow diagram
Combined cycle plant--site plot plan
Combined cycle plant--plan at grade floor
Combined cycle plant--operating" floor plan
Combined cycle plant--elevation
Combined cycle plant--mezzanine floor plan
1000 MW coal-fired unit--mechanical plot plan
1000 MW coal-fired unit--plan at operating floor
1000 MW coal-fired unit~-plan at grade floor
1000 MW coal-fired unit--elevation
xvii
293
294
296
298
299
302
303
304
307
309
311
314
321
323
327
329
331
333
335
337
338 "
339
-------
1-133..
1-134
1-135
1-136
1-137
1-138
1-139
1-140
1-141
1-142
1-143
1;...144
1-145
1-146
1-147
1-148
1-149
1-150
1-151
1-152
1-153
1-154
1000 MW oil-fired unit--mechanical plot plan
1000 MW oil-fired unit--plan at operating floor
1000 MW oil-fired unit--plan at grade floor
1000 MW oil-fired unit-~elevation
Part-load performance of , Hammond #4 unit
Elevation--atmospheric pressure fluidized, bed steam
generator
Once-through fluidized bed steam generator circuitry
Atmospheric pressure regeneratio flow diagram
Atmospheric pressure regenerator plant layout
Coal handling plant
Coal drying and injection system
Air flow diagram
Dust collector
Mass balance--particulate removal system
Overall material balance--full load
Overall energy balance--full load
Temperature-energy diagram for the atmospheric
utility boiler design
Part-load boiler efficiency for conventional and
fluid bed boiler
Heat rate vs plant load for conventional and fluidized
bed boiler
Boiler load reduction
New plant cost vs. capacity--coal-fired plant
New plant cost vs. capacity--oil-fired plant
xviii
341
343
344
345
354
356
359
363
364
367
369
372
373
375
377
378
379
384,
384
386
394
401
-------
1-155
1-156
2-1
2-2
2-3
2-4
2-5
2-6
2-7
2-8
2-9
2-10
2-11
2-12
2-13
2-14
2-15
2-16
2-17
2-18
2-19
2-20.
Fluidized bed gasification-combustion power plant
Fluidized bed gasification-combustion power plant
Modes of operation
Carbon deposition data.
Gasification-desulfurization heat generation rate
Effect of particle size on sulfur removal
Sulfur removal correlation
Batch cyclic bed tests
Regenerator temperature control
Regenerator design
Regenerator operation
Gasifier temperature control
Gasifier temperature control
Gasifier product heating value.
. .
Regenerative operation flow diagram
Once-through operation flow diagram
Boiler retrofit concepts
Space requirements for internal design--20% stoichio-
metric air
Space requirements for internal design--25% stoichio-
metric air
20MW gasifier add-on
150MW gasifier add-on to a coal-fired boiler
600MW gasifier external to a coal-fired boiler
xix
418
419
. 428
439
443
446
448
450
467
469
471
478
479
481
489
489
497
498
499
502
503
504
-------
2-21
3-1
3-2
3...,3
3-4
3-5
3-6
3-7
8 .
Design of 2.5 X 10 Btu/hr burner
Flow diagram for development plant facility
10-30 MW pressurized fluid bed boiler development
plant schedule
Elevation of regenerative design-~200MW retrofit.
Plan of regenerative design--200MW retrofit
Elevation of once-through design--200MW retrofit -
Once-through design--200MWretrofit
Demonstration plant schedule
xx
51?
540
549 .
554
555
.556.
557
562
-------
Number
1-1
1-2
1-3
1-4
1-5
1-6
1-7
1-8
1-9
1-10
1-11
1-12
.1-13
1-14
1-15
LIST OF TABLES
Title
Factors affecting S02 removal during fluidized bed
combustion
. .
Summary of the results of combustion losses from
experiments conducted by OAP contractors and NCB
laboratories at atmospheric pressure
Effect of primary variables on combustion efficiency
and the mechanisms involved
Combustible losses assumptions for design of the
industrial, atmospheric pressure, and pressurized
fluidized bed boilers .
Summary of heat transfer measurement in fluidized beds
at atmospheric pressure performed by OAP contractors
and NCB laboratories
Operating conditions in fluidized bed boiler designs
'''-,
Analysis of coal ash, coal, and bed material
Experimental work on fluidized bed boiler tube
materials
Typical analyses of metal specimens
Average tube specimen weight loss under pressurized
operation condition, mg/cm2-hr.
Summary of rate of weight loss, test series 1
Summary of rate of metal loss, test series 2
Observed rate of weight losses on specimens, duration
147 hours
Observed rate of weight losses on specimens, duration
500 hours
Fluidized bed combustion boiler designs, boiler tube
materials
xxi
Page
8
83
84
98
108
134
135
138
142
143
146
147
148
149
154
-------
1-16
1-17
1-18
1-19
1-20
1-21
1-22
1-23
1-24
1-25
1-26
1-27
1-28
1-29
1-30
1-31
1-32
1-33
1-34
Fluidized bed combustion boiler designs, chemical
composition of selected boiler tube materials
Anticipated particle size distribution of dust
entrained in gas turbine feed
Ash particle size and concentrations used to establish
permissible loadings in Bureau of Mines coal-fired
gas turbine research project
Analysis of limestone no. 1359
Designs for industrial fluidized bed boiler
Comparison of boiler performance
Coal-fired industrial boiler pollution control
alternatives
Summary of draft losses
Draft fans for industrial fluidized bed combustion
boilers
Comparison of effectiveness of turn-down techniques
for industrial boiler
Operating conditions during turn-down operation
for industrial boiler
Comparison of boiler performance
Summary of draft losses, in. H20
Comparison of power requirements
Air pollution control capability of industrial
fluidized bed boiler
Installed cost of once-through dry solid system
Installed cost of wet scrubber system
Comparison of installed costs for equivalent-capacity,
field-erected, steam generators firing different fuels
Capital cost of coal- and gas/oil-fired industrial
boiler pollution control alternatives
xxii
155
156
159
170
171
176
197
199
201
210
214
218
219
220
222
225
226
228
231
-------
1-35
1-36
1-37-
1-38
1-39
1-40
1-41
1-42
1-43
1-44
1-45
1-46
1-47
1-48
1-49
1-50
1-51
1-52
1-53
1-54
1-55
1-56
Industrial boiler cost summary
Effect of gas turbine pressure drop on design point
plant performance
Effect of condenser pressure on design point plant
performance
Part-load performance under operation modes 1 and 2
Performance data for variable bed temperatures
Effect of gas turbine and steam superheat temperature
'on the high-pressure system capacity and heat rate
600 MW high-pressure regenerator material balance
Fluidized bed parameters--design load
Fluidized bed parameters-~70% plant load
Steam generator performance summary
Fluidized bed parameters--design load
Fluidized bed parameters--75% boiler load
- Fluidized bed parameters--50%boiler load
Boiler plant equipment costs
Coal-fired power plant cost breakdown
Steam turbine plant equipment costs
Air pollution control equipment cost
Boiler plant equipment costs
Oil-fired power plant cost breakdown -
Energy generation costs for coal-fired power plants
Energy generation costs for oil-fired power plants
Heat transfer characteristics
xxiii
232
257
259
265
266
279
289
317
318
319
380
381
382
391
392
393
395
~99
400
402
403
407
-------
1-57.
1-58
Air pollution abatement
408
Fuel utilization
410
1-59
Advanced fluidized bed combustion power plant costs
with. sulfur recovery
1-60
Coal-fired power plants
417
421
2-1
Fuel oil properties
433
2-2
Limestone compositions
434
440
2-3
2-4
Product analysis
Estimated .chemical analysis
441
490
2-5
2-6
Regenerative operation material and energy balances
Once-through operation material and energy balances
491
493
2-7
2...8
Pressure drops
Power requirements
494
2-9
Regenerative operation capital costs
519
2-10
Operating costs for regenerative operation
520
2-11
Once-through operation capital costs
523
524
2-12
Operating costs for once-through operation
2-13
Assessment of fluidized bed oil gasification-
desulfurization
528
3-1 Material balance--development plant 541
3-2 SOasorbent analysis used for preliminary fluidized
be boiler analyses 543
3-3 Development plant cost estimate 547
3-4 Development plant cost estimate 561
xxiv
-------
TABLES OF CONTENTS
FOR
VOLUMES I AND III
xxv
-------
TABLE OF CONTENTS
VOLUME I
ABSTRACT.
. . . . 0
.........
. . . . .
. . . . .
SUMMARY. . .
. . . .
.......
. . . .
. . . . .
INTRODUCTION.
. .. . .
.........
. . . .
. . . .
. . . .
FLUIDIZED BED COMBUSTION BOILER CONCEPTS
............
Fuel. . . . . . . . . . .
. . . . .
. . . . .
. . . .
Fluidized Bed Processing. . . . . . . . . . . . .. .
Boiler Concepts. . . . . . . . . . . . . . . . . . . . . .
Design of a Fluidized Bed Boiler System. . . . . . . . . .
PRELIMINARY SURVEYS. . . .
. . . . . .
. . . . . . . .
. . . . .
Market Surveys. . . . . . . . . . . . . . . . . . . . . . .
Fossil Fuel Survey. . . . . . . . . . . . . . . . . . . . .
Survey of Alternative Means of S02
Pollution Control.
Emission Standards Survey. . . . . . . .. . . . . '. . . .
.......
. . . .
"' . .. . .
ASSESSMENT OF FLUIDIZED BED COMBUSTION.
. . . .
. " . . . . . .
Spec~fic~tions. . . . . . . . . . . . . . . . ., . . .
Functional Boiler Specifications. . . . . . . . . . .
Operating Specifications. . . . . . . . . . . . . . .
Design Specifications. . . . . . . . . . . . . .
Industrial Boiler Application . . . . . . . . .
Des ign . . . . . . '. . . . . . . . . . . . . . . . . .
Controls and Instrumentation. . . . . . . . . .
Operation. . . . ". . . . . . . . " . . . . '. . ",.. . .
Performance Characteristics. . . . ',' . . . . . . . .
Boiler Plant Costs. . . . . . . . . . . . . . .
Evaluation. '. . . . . . '. . . . .
. . . . .
. . . . .
Electric Utility Application. .. . .. . . . . . . . .. .
Pressurized Fluid Bed Boiler
Power Plant Design. . ~ ~ . . . . . . . . . ~
System Concept. . . . . . . . . . . . . . . . ~ .
Power Cycle Conditions. . . . . . . '. . . .
Boiler Plant Equipment . . ',,' . . . . . . .
Auxiliary Equipment. .. . .. . . '. . . . ,. . .
Power Generation Equipment. . . . . .
Controls and Instr.umentation. . . . . . . . .. .
Plant Layout. . . . . . . . . . . . . . . . . . .
Performance. . . '. . . . . . . " . .
xxvi
Page
1
111
1
5
5
,5
8
11
21
23
35
42
43
45
45
45
49
51
53
53
63
64
65
67
72
78
80
80
80
83
90
91
101
101
110
-------
TABLE OF CONTENTS (Cont'd.)
Atmospheric Pressure Fluidized Bed
. Boiler Power Plant Design.' . . . . . . . ..
Boiler Plant Equipment. . . . . . . . . . . . .
Auxiliary Equipment. . . . . . . . . . . . . .
Sulfur Removal/Recovery System. . . . . . . . .
. Performance . . . . . . . . . '.' .'
Economics. . . . . . . ~ . .'. . . . . . . . . . . .
Evalua tieD. . . . . . . . . . . . . . . . . . . . .
Comparison. . . . . . . . . . . . . . . .
Potential. . . . . ," . " ,". . . . . . . . . .
ASSESSMENT OF ATMOSPHERIC PRESSURE FLUIDIZED
BED OIL GASIFICATION. . . . . . . . . . . .
. . . . . . . .
Specifications. . . . . . . . . . '0' . " . . . . . . . .
Gasification/Desulfurization Concepts. . . . . . ..
Design. . . . . . . . . . . . . . .'. . . . . . . . . . .
Economics. . . . . . . . . . . . . " . . . . . . . . . .
Evaluation. . . . . . . . . . . . . . . . . . . . . . . .
Comparison Between Regeneration and
Once-Through Operations . . . .
Comparison with Other S02
Control Methods;
. . . .
FLUIDIZED BED FUEL PROCESSING
DEVELOPMENT PLANTS. . . . .
. . . . .
. . . . . . . . . . .
Pressurized Fluidized Bed Combustion
Boiler Development Plant. . . . . . . . . . . .
Atmospheric Pressure Oil Gasification
Demonstration Plant. . . . . . . . . . . . . . . . . .
CONCLUSIONS. . . . . . .. . . . . . . . . . .
. . . . . . . . .
Fluidized Bed Combustion Boilers. . . . . . . . . . . . .
Atmospheric Pressure Fluidized Bed
Oil Gasification. . . . . . . . . . . . . . . .
RECOMMENDATIONS. . . . . . . . . . . . .
. . . .
. . . . . . .
Industrial Fluidized Bed Combustion Boilers. . . . . . .
Utility Fluidized Bed Combustion. Boilers. . . . . .
Atmospheric Pressure Oil Gasification. . . . . . . . . .
REFERENCES. . .
. . .. .
. . . . . . . . . . . . . . . . . . .
ACKNOWLEDGEMENTS. . .
. . . . . . . . . .
. . . .
xxvii
Page
114
114
117
118 .
120
124
129
129
136
139
139
143 .
148
153
158
162
162
167
167
170
175
175
178
. 181
lf3l
181
182
183
185
-------
TABLE OF CONTENTS
Volume III
APPENDIX A - Electric Utility Market Survey
APPENDIX B - Industrial Boiler Market Survey
APPENDIX C - Development of Fluidized Bed Combustion Boilers
APPENDIX D - Industrial Boiler Design Report
APPENDIX E - Turndown Techniques for Atmospheric
Fluidized Bed Boilers
APPENDIX F - Dynamics of Atmospheric Fluidized Bed Boilers
APPENDIX G - Optimization of Heat Trap System Cost
APPENDIX H - Pressurized Boiler Design Report
APPENDIX I - Regeneration/Sulfur Recovery System Cost
APPENDIX J - Pressurized Boiler Combined Cycle Plant Report
APPENDIX K - Atmospheric-Pressure Boiler Design Report
APPENDIX L - Boiler Burner for Low Btu Gas
APPENDIX M - Gas Turbine Corrosion, Erosion, and Fouling
APPENDIX N - Stack Gas Cooler Design
1QCvi ii
Page
A-I
B-1
C-l
D-l
E-l
F-l
G-l
H-l
I-I
J-l
K-l
L-l
M-l
N-l
-------
L
ASSESSMENT OF FLUIDIZED BED COMBUSTION
Specifications and design concepts were established for
industrial and utility fluidized bed boilers.
Preliminary designs,
performance projections, and cost estimates were prepared on the basis
of the specifications and compared with conventional systems~
The design specifications for industrial and utility fluid,ized
bed combustion boilers are based on
. Preliminary surveys - market. requirements, environmental
control requirements
. Fluidized bed combustion experimental data and related
technology experience.
1
-------
PRELIMINARY SURVEYS
In order .to carry out boiler designs and evaluations, it was
necessary to gather market information on boiler applications and
technical information on fluidized bed combustion.
Market and other
background data were collected for three general purposes:
. To establish functional specifications as a basis for boiler
(or gasifier) designs.
Boiler capacity, steam conditions,
efficiency, operational features such as percent turn-down,
and dynamic response are all established by customer require-
ments.
In addition, a fuel is chosen on the basis of
availability, cost, and convenience.
Finally, government
regulations establish maximum levels of pollutant
and particulates) emissions for boilers employing
selected fuel.
(S02' NOx'
the
. To establish a basis of comparison.
The performance and cost
of the fluidized bed combustion boiler systems must be compared
with alternative systems in order to evaluate the likelihood of
their use.
. To establish the impact of a successful fluidized combustion
boiler on air quality and on economical steam/power generation
Fluidized bed combustion is a technolo~y appropriately applied
to new or replacement boilers; it is not applicable to the
modification of existing boilers.
It must be established
that there is a sufficient market for new boilers to justify
the effort and expense of developing a new boiler technology
with improved pollution abatement potential.
Boiler systems have been built, operated, or proposed which
incorporate. fluidized bed combustion or gasification.
Fluidized bed
2
-------
boiler concepts go back to 1928 when Stratton developed a spouting
fl~idized bed 'boiler (See Appendix C). The early boiler concepts
incorporating fluidized bed combustion were generally developed for
burning low-grade fuels. These designs did not recover heat or consider
sulfur removal in the fluidized bed combustor.
In the last decade
fluidized bed boiler designs have been conceived, in England and in the
United States, which incorporate heat recovery and sulfur removal in
the fluidized bed combustor.
Some of the concepts which have been
proposed are reviewed in Appendix C.
Several surveys have been conducted to provide the information
necessary to boiler design specifications, performance evaluations, and
, impact determination.
Results from these surveys are summarized in
Volume I.
The survey data are presented in Appendices A and B in Volume
Ill.
These surveys include:
. Market surveys for industrial boilers ,and for utility power
plants. These surveys have provided both information on past
boiler sales and projections of future sales. Information on
boiler capacities and on supply pressure and temperature has
also been obtained..
. Fuel survey. Historical data on fuel usage in both industrial
and utility boilers have been gathered. Evidence concerning
fuel availability and estimates of present and future fuel
prices have been collected and examined.
. Survey of S02 removal ~ystems for cleaning combustion products
of conventional boilers. Several proposed systems have been
evaluated with regard to effectiveness and economy.
A
conventional boiler with a limestone wet scrubber for removal
of S02 and particulates was selected as a basis of comparison
for the performance of the fluidized bed combustion boilers
designed in this work.
3
-------
repor.t.
. Emission standards.
Government standards are being
established for limiting 802' NOx' and particulate emissions
from boilers. Projected standards have been used in setting
functional specifications for the fluidized boiler designs.
Results from the various studies are utilized. throughout this
4
-------
FLUIDIZED BED COMBUSTION DATA
Fluidized bed combustion data on sulfur dioxide removal, limestone
and dolomite regeneration, nitrogen oxide minimization, combustion
efficiency, heat transfer,particle carry-over, boiler tube corrosion/
erosion, and gas turbine blade corrosion/erosion were evaluated in order
to establish operating and design specifications.
5
-------
Sulfur Dioxide Removal
Sufficient limestone must be added to a fluidized bed combustor
to reduce its S02 emission below that specified by the air pollution.
control laws in effect. The S02 emission level required to achieve
pollution control objectives was discussed in the Preliminary Surveys
section of the Summary Report.
This section examines the relationship
between the design selected for the fluidized bed combustor and the
reduction in S02 emission which is likely to be achieved. This information
is pertinent to the selection of operating conditions, since it is unlikely
that control objectives will be selected independent of cost factors.
Work to determine the S02 removal achieved in fluidized bed
combustors at various operating conditions has been carried out by a
number of organizations, including Argonne National Laboratories (M~L);
Esso Research and Engineering (Esso R & E); Pope, Evans and Robbins (E>ER);
the National Coal Board (England), (NCB); Consolidation Coal Company (Consol);
the U. S. Bureau of Mines; and Esso (England). (See [26,49,51,69,70,71, and
84] for details of the results obtained by each of these investigators to
date.) No effort has been made in this study to tabulate all of the data
acquired by these investigators with regard to S02 removal.
an attempt has been made to:
Instead,
t Identify the important factors affecting the degree of S02
removal
. Determine, at least semiquantitatively, the effect each of
these factors has on the degree of S02 removal, recognizing
that the effect of a change in any particular variable will be
a function of the levels of many of the other variables
6
-------
, Outline methods of extrapolating existing data to predict
the S02 removal which might be achieved at operating conditions
of commercial interest, including higher pressures, velocities,
and temperatures.
Factors Affecting S02 Reduction
Thirteen variables which influence the percentage of S02removal
achieved in fluidized bed combustion of coal with limestone addition are
listed in Table 1-1. Those variables at the top of the table are believed
to exert stronger effects than those near the bottom. ~o claim is made,
however, that the variables are listed in order of decreasing effect.
The magnitude of the effect of each is a strong function of the level of
the remaining twelve.
Also shown in 'Table 1-1 are. the ranges for each of
the factors for which a significant amount of data has been reported and
the approximate operating conditions for the industrial boiler and the
atmospheric and pressurized utility boilers which are discussed in detail
in later sections of this report.
One method used to summarize experimental data is the
formulation of a mathematical model. A number of models for S02 removal
by limestone and dolomite during fluidized bed combustion of coal have
been proposed [49,70,84,100]. The model proposed by Koppel [49] can be
used to illustrate the effects anticipated when some of the variables
listed in Table 1-1 are changed. Specific mention will be made of those
variables which are not clearly accounted for in this model.
The basic assumptions of Kopp~l's model [49J include:
. Plug flow of gas through bed
. Uniform S02 concentration in the direction ~ransverse to
the direction of the flow
. Perfect mixing of the solids within the bed
. Uniform temperature throughout the bed
. S02 generation throughout the bed with a generation function:
7
-------
ex>
TABLE 1-1
Factors Affecting S02 Removal During Fluidized Bed Combustion
1.
2.
3.
4.
Calcium-to-Sulfur Mole Ratio, Ca/S
Bed Temperature, of
Fluidizing Velocity, ft/sec
Stone Type & Source
L = Limestone
D = Dolomite
5.
Factors Affecting Gas-Solids Mixing
(distributor design, placement of
internals, fluidizing velocity/
minimum fluidizing velocity, L/D
ratio, location of stone feed)
6.
Factors Affecting Regenerated Stone
Activity
number of regeneration cycles,
regeneration conditions, C02 partial
pressure in combustor, amount of
7.
8.
9.
composition of ash
Stone Size
Bed Depth, ft.
Coal Type
Swelling index, amount and
10.
composition of ash.
Location of Coal Feed and Means of
11.
12.
13.
Injection
Combustion Pressure, atm.
Excess Air
Sulfur Level of Coal, wt%
Range for Which
Significant Amount
Of Data Are Available
1-4
1400-1600
2-4
L-1359, L-T18,
L - United Kingdom
D-1337, D-Tymochtee
D - United Kingdom
25-1000 1.1
1-4
SI = 1 to 8
1-5
1-4.5
Industrial
Boiler
6
1650
12.5
1-1359
- 1/4"
2.5
5-5.5
1
4
Atmospheric
Utili ty B~iler
6
1600
6-10
L-1359
- 1/4"
2.5
5-5.5
1
4
Pressurized
Utility
Boiler
6
1750
6-9
L-1359
- 1/4"
10-16
5-5.5
10
4
-------
where
where
G(x) =
FC
o -ax/L
Te
1 - e-a
(1-1) .
F
= mass flow rate of gas through bed, lbs/min (includes
C02' N2' 02' etc. after combustion);
= lbs of S02 generated/lb of gas (after combust~on)
. (assumed equal to the concentration of S02at which
to.63 was determined);
C
o
L
= bed height;
x
= distance above the distributor plate;
= arbitrary parameter which describes where S02 is
released (a = ~ corresponds to all S02 being generated
at the base plate; a = -~ corresponds to all generation
at upper surface of bed; a = 0, to uniform generation).
a
. The reaction at the particle is first order (irreversible)
with respect to S02 concentration and at a fixed temperature
is a function only of the degree of utilization of the stone,
y, and particle size'.
The effect of particle size distribution
is accounted for by using key) for the average bed particle
size,
.9.Y. = kC
dt
Ye
k(y) = (1 - y/y )
Coto.63 e
(1-2)
Ye
= equilibrium loading of the sorbent, lbs S02/lb of
CaO (not a function of gas composition);
9
-------
to.63 = reaction time for a particle to reach 63% of its
equilibrium loading (CotO.63 is assumed independent
of C ).
o .
The following implicit relationship was derived for R, the
fraction of the total generated 802 that is removed by the sorbent:
R =
[1 ...,
1. -ex
- e
1
ex exp[-(l - R)H]- (1
r
R
ex - (1 - -)H
. r
R -ex
- -)He
r .
(1-3)
where
WYe
r =~, the effective ratio of additive feed to 502 feed;.
o
. .
w = feed rate of additive to bed, 1bs of CaO/time;
and
aLy
H = e, the normalized bed height;
FCotO.63
1bs of CaO
a = concentration of additive particles =
ft of bed height'
Results generated from the above model formulation and previously
published by ANL [49] indicate how both the ~ractiona1 S02 reduction,
R, and the particle consumption, C, can be expected to vary with the
parameters'H and r.
equation:
C is related to stone utilization, U, by the
Ys
C = U-
Ye
(1-4)
Qualitatively, one anticipates the following effects based
on the model:
1.
Stone utilization decreases as the degree of S02 removal
increases.
10
-------
2.
At values of r less than 0.5, the effect of GalS completely
overshadows all other effects.
corresponds to Ga/S = 1).
(If Y /y = 0.5, r = 0.5
e s
3.
At values of H greater than 100, the degree of removal is
also essentially determined by the value of r.
4.
At low values of H (which might be caused by using shallow
. beds, beds dilute in GaO content, unreactive stone, or high
gas velocities), the point of sulfur release determines the
maximum degree of removal achievable.
At high H values, the
sulfur release pattern (unless a. ~ -00) is unimportant.
5.
. The equilibrium loading achievable, y , may be the single
. e
most important characteristic of a given sorption system.
. .
This parameter enters into the calculation of both Hand r.
Its effect on H is especially important when H is low (see
4 above), and its effect on r appears important for all
systems studied to date.
Note in the following matrix that
many of the operating variables (stone type, source~
structure, and particle size, ash properties, and T) influence
y .
e
Difficulty in distinguishing between these effects
can be anticipated.
6.
Reaction kinetics for the stone-SOZ reaction (quantified
by the parameter GotO.63) ~lay a dominant role in determining
degree of removal only if H. is low (i.e., little gas-solids
contacting time) or S02 is released near the top of the.bed.
It is possible to estimate the change in SOZ reduction caused by
a change in any of the operating variables (except combustion pressure,
excess air,and gas-solids contacting effects) if values of to.63 and Ye
are known. The variations in to.63 and Ye caused by changes in temperature,
stone type and source, particle size, and processing factors affecting
stone structure must be correlated, and provision must be made for the
unincluded factors in order to arrive at a complete picture based on
Koppel's model.
11
-------
EFFECT OF OPERATING CONDITIONS ON MODEL PARAMETERS
Factor
Mod~l Param~t~rs Aff~cted
Ca/s.
r = CalS x ye/ys
where y = Stoichiometric stone loading
s
Temperature
Accounted for by changes in to.63 .and Ye that occur with
temperature
Fluidizing velocity
F, Ct., V for fixed cross-sectional area
Fla, Ct., V for fix~d bed composition and variable at fixed r
b~d area increasing V will decrease H
Stone type and source
to.63 and Ye
t-'
N
Factors affecting gas-solids
contacting and solids mixing
Not includ~d (See assumptions list 3 in text)
Stone size
Affect to. 6 3 and Y
e
Affect to.63 andy
e
L, IX
Factors affecting stone structure
Bed depth
Coal type
IX, also ash prop~rties affect to.63 and Ye
Location of coal feed
Ct.
Combustion pressure
May aff~ct to.63 and Ye (Howev~r, this would be inconsistent
with other model assumptions)
Excess air
Not included
Sulfur level in coal
No effect if all other variables maintained constant (This
one is tricky, but remember that Co to 63 is constant
by the assumptions of the model) .
-------
Summary of Reported Data
Calcium-Sulfur Ratio.
Increasing the mole ratio of calcium to
sulfur in the feed to the combustor, the Ca/S ratio, increases the degree
of sulfur removal achieved.
There are many points in the system where calcium and sulfur
might be measured, as is indicated in Figure 1-1. In the experimental
investigations reported, Ca/S has most often been defined as the calcium
content of the fresh stone feed (Stream A), in moles per unit time,
divided by the sulfur content of the coal in Stream B, also in moles
per unit time.
The small amount of Ca which enters as coal ash is
usually ignored because it is small in comparison with that added' in
Stream A. In this study Ca/S has been thought of as the mole ratio
which includes calcium from A and C to the boiler. As such, it
includes both Ca and S from the recycle as well as from the fresh feed.
The two definitions are essentially identical when no regeneration is
employed. The sulfur to calcium mole ratio at Point D is referred to as
the degree of stone utilization, U.
The effect of Ca/S on the percentage of sulfur originally
contained in the coal which is retained in the fluidized bed is illus-
trated in Figure 1-2. Percent retention is a nearly linear function of
Ca/S at Ca/S ratios below 1.0. The Ca/S ratio required to achieve 90%
sulfur removal is seen in Figure 1-3 to vary from approximately 2.3 for
Welbeck Coal to greater than 3 for Humphrey Coal, all other conditions
being similar.
Bed Temperature.
Maximum sulfur removal has been obtained at
bed temperatures between l450°F and l600°F for most coals, limestones, and
operating conditions (other than temperature) tested thus far.
Data on
the effect of temperature on S02 removal reported by the National Coal
. Board, Argonne, and Esso (England) are plotted in Figure 1-3. Both the
magnitude of the temperature effect and the absolute level of S02 removal
achieved varied from investigator to investigator as a result of differences
13
-------
Dwg. 2960A47
D
Fluidized -
Stone
Bed
Regenerator
Boiler C
-
4~ ~
Waste
B A Stan e
Coal
Make-up
Stone
Fig. I-I-Points on the system where calcium and sulfur
. levels can be measu red
14
-------
2; 60
......
c
Q,)
Q)
a:::
(/)
~ 40
100
80
20
00
Welbeck
Peabody
Park Hill
Hu mphrey
1 . 2. 3 .
Cats Mole Ratio From Add~ Limestone
Fig. 1-2-Effect of Cats on sulfur retention for
various coals. [Data from National Coal Board
. 69, 70 or 71]
15
Curve 644303-A
%S
SW #
'" 1. 3 l..;1
"'4.0 6-6l
'" 2. 5 7l
'" 2. 7 8-8l
4
-------
in fluidizing velocity, fuel characteristics, stone characteristic~, .
reactor design, etc.
The data of all these investigators, however, show
one common feature -- that a maximum in removal efficiency was observed
rqther than a monotonic increase or decrease in efficiency with temperature.
Theprese.~ce of a maximum in the l450°F to l600°F range is not
explained by available equilibrium and kinetic data on the primary
reac'tion of concern':
-+
CaO+ S02 + 1/2 02 + CaS04'
(1-5)
The'equilibriumpar.tial pressures for S02 range from approximately 10-8
atm at 800°C to 6x 10-7atm a.t 900°C (for 10% excess air). These values
correspond (at 1 atmtotalpressure) to 0.01 and 0.6 ppm respectively.
Since the actual exit gas compositions are several orders of magnitude
higher (300 :to. 1500 ,ppm for the data points shown in Figure 1-3), the
increase in equilibrium partial pressure with temperature does not
explain the drop-off in '8~2 removal at high temperatures.
Kinetic data collected and analyzed by Skopp and co-workers [84]
-as well as by Coutant and co-workers [19] indicate that the rate constant
(for reaction 1-5 preceded in some cases by calcination of CaC03) increases
monotonically with .temperature with a reasonable straight line fit to
an Ahrenniusplot. (1) Skoppinvestigated temperatures from l200°F to
1600°F;Coutant's data were taken between l500°F and 2l50°F.
(1) Once again caution should be taken before reaching sweeping conclusions.
Harrington, Borgwardt, and Potter [40] characterized the reactivity
of ten limestones and found that reactivity increased between l600°F
and 1800°F for four of them, but decreased for the other six.
16
-------
o 0 Hu mphrey coal, T18 Ii mestone; 4 ft/sec;
. CalS ratio 2. (Data from National Coal Board)
Welbeck coal, UK limestone; 8 ft/sec;
Cal S ratio 3. (Data from National Coal Board)
-- - - Illinois No.6 coal, -14 mesh; dolomite 1337; 2.5 to 3 ft/sec;
. CalS = 2.2; data from Argonne National Laboratory (49) .
. --Illinois No.6 coal, -14 mesh; limestone 1359, 490 \..l;
. 3 ftlsec; CalS = 2.5; data from Argonne National Laboratory
. C:.irve 61+4311-A
90
80
"
"
,
,
,
,
,
. '\ .
\
\
,
. . f5 60
......
c
(1)
1i>
c:::
V)
~ 40
,(
x Resu Its reported by Esso Research Center
(England) . . . ..
. velocity ""'lft/see; fuel oil and stone.
unknown; CalS ""'3 .
20
o
700 .
1300
800 850
1500 1600
Bed Temperatu re
Fig.1-3-Effect of bed temperature on sulfur removal
in fluidized bed combustor .
750
1400
900 °C
1700 .
OF
17
-------
Two e;xplanations for . the reduction in S02 removal effectiveness
at temperatures .above l600°F (Argonne [50] reported only 17% reduction
at l800°F) merit further consideration:
. An interaction between the coal ash and the stone which
causes the formation of an inactive surface layer or the
closing of the pores in the stone
. The presence of a localized reducing atmosphere in the
emulsion phase of the lower portion of the fluidized bed
combustor which results in a release of S02 from already
sulfated stone.
Information to be considered in assessing the first explanation
includes the following:
. Early work at Esso R & E indicated that "coal fly-ash
does not stick or .agglomerate with lime particles in the
.desulfurization reactor at 1600°F"[84, p.38].
. Later work at Esso R & E [38,39] indicated that the
presence of coal ash during regeneration (~ 2000°F)caused
a dramatic reduction in the effectiveness of the regenerated
stone for S02 removal. The decrease was attributed to
"interference by. components of the fly ash, possibly Fe203...."
Unfortunately, the above results shed no light on the importance
of ash-stone interaction$ at 1700°F to 1800°F.
18
-------
The second explanation has been proposed by workers at Argonne
National Laboratories [50] and at Pope, Evans, and Robbins [80] and
supported by some interesting experiments utilizing limestone to remove
S02 during the combustion of natural,gas.
Argonne, is:
The theory, as expressed by
In any fluidized bed, an emulsion phase and a gas
bubble phase are present. Gas in excess of that required for
minimum fluidization is in the bubble phase but also circulates
between the bubble and the emulsion phase. The gas in the,
bubble phase is essentially not available for reaction with
solid particles until it moves into the emulsion phase. In
the lower portion of a fluid~bed coal combustor, oxygen in
the emulsion phase reacts very rapidly with coal. All of the
oxygen in the emulsion phase is consumed, forming carbon monoxide.
As the bubbles rise through the bed, air exchanges between the
bubbles and the emulsion phase; in the upper part of the bed,
the emulsion phase contains excess oxygen.
Thus, in the lower part of the combustor, the following
reaction might occur:
-+ '
CaS04 + 4 CO + CaS + 4 C02
In the upper part of the combustor, one or more of the following
reactions might occur:
(1)
, -+
CaS + 3/2 02 + CaO + S02
(2)
-+
CaS + 3 CaS04 + 4 CaO + 4 S02
(3)
-+
CaS + 2 02 + CaS04 .
(4)
Reactions 2 and 3'would result in S02 being :e~e~erated in
the upper part of the combustor, ser~ously l~m~t~ng the
effi~iency of the system for retention of sulfur. These 802
regeneration reactions are highly temperature dependent and
might account for sulfur retention efficiency being lower at
higher temperatures. '
19
-------
While the reaction sequence proposed above is certainly
speculative, the theory that 502 is released as a result of localized
reducing conditions in the lower portion of the fluidized bed is supporta~le.
Investigators at Pope, Evans, and Robbins have demonstrated that CaS04
. can be decomposed .to CaO and 802 at 1800°F to 1900°F,even when' the overall
fe.ed rates of ,fuel and air indicate an oxidizing atmosphere exists [26],
and have proposed an in-bed regeneration concept based on tlrls information.
Argonne [50] acquired data which showed that essentially no 502 was liberated
.from CaS04.under slightly oxidizing conditions at 1625°F, but an appreciable
amount was liberated a.t l720°F and above. At l850°F, the partial pressure
of the 502 liberated was a strong function of the 02 partial pressu~e in the
flue gas, varying from 2700 ppm at 1.8% oxygen to 4900 ppm at 0.5% oxygen.
Summarizing, the poor 502 removal efficiency obtained at
temperatures near i'800°F seems due primarily to the liberation of S02
asa result of local reducing conditions in the. bottom section of the
bed.
5tone~ash interactions may also contribute to the loss of effectiveness,
especially in the case of high-ash coals [96].
If stone-ash effects do
not dominate, the use of deeper beds and higher excess air is expected to
result in bet ter removal at high temperatures.
The experimental results of Zielke, et al. [106] can be cited in
discussing the effect of bed depth at high temperatures. They reported
S02 reductions ranging from 75% to 97% at l800°F, in sharp contrast with
all others who have reported high temperature results. Their tests were
carried out in a 4-inch diameter bed with a depth of 36 in,
an LID of 9.
In contrast, Argonne National Laboratories operated with an LID of 4
and
Pope, Evans, and Robbins with an LID of approximately 1.
The authors
[106] attributed their good results to good solids~gas contacting, but
Ehrlich [26] pointed out that the high LID might have reduced the degree
of solids mixing and thereby have produced the encouraging results.
According to the latter theory, localized reducing conditions exist in
the lower section of the deep bed, but the gas composition has evened out
20
-------
in the top section so that the dolomite there is effective.
It is impossible
to choose between the above explanations on the basis of current data.
In fact, it seems likely that both factors contributed to the improved'
performance.
Fluidizing Velocity.
If all other factors remain unchanged,
increasing the fluidizing velocity decreases the effectiveness of sulfur
removal. Typical results are shown in Figure 1-4. Although increasing
the Ca/S ratio diminishes the effect, even at Ca/S = 4 the experimental
results indicate a decrease from 90% removal efficiency at 2 ft/sec to
70% at 6 ft/sec.
It should be noted that all Ca/S ratios shown are for
fresh stone and that higher Ca/S ratios will be required to achieve the
same effect with regenerated stone.
Increasing the gas residence 'time in the fluidized bed by
increasing bed height can be anticipated to compensate for increased
velocity, provided that gas-solids contacting remains effective.
Stone Type and Source.
Both limestone and dolomite have been
utilized as S02 sorbents in fluidized bed combustion systems. The
distinction between the two, for purposes of this report, is in their
CaC03 content: a limestone contains more than 80% CaC03 while the
dolomites investigated have had nearly equal molal amounts of CaC03 and
MgC03' At a fixed value of Ca/S, dolomites have been found typically
to achieve a higher sulfur removal; since only about half of the dolomite
~
consists of Ca compounds, however, dolomites typically achieve a lower
degree of sulfur removal than limestones at the same value (lbs of stone/lb
of S).
A more important and more fundamental difference between the
two stone types is based upon the phase diagram shown in Figure 1-5.
MgO is the stable Mg-compound at almost all temperatures and operating
\
CaO is stable
pressures being considered for fluidized bed combustion.
only at the conditions in the lower right-hand corner of the plot -- that
is, at relatively high temperatures and low pressures.
At a total
combustion pressure of 10 atm and l600°F, for example, limestone fed
21
-------
100
90
80
70
60
50
~
o 40
E
Q,)
a:::
-
~ 30
u
10...
Q,)
c..
20
10
Curve 544307-A
CRE Data (1470°F)
£r - -iJWelbeek Coal, UK Li mestone (440 ~m)
0- - -<)Welbeek Coal, UK Li mestone (440 ~m)
~-... Humphrey Coal, US #18 Limestone (-10 BSS to -1/8 ineh)
ANL Data (15500F)
o--oCom. Ed. Coal, Limestone No. 1359 (> 1000 .~m)
6 Com. Ed. Coal, point taken from eu rve for
limestones and dolomite (ANL!ES/CEN-F022,
~ 630 ~ m, CatS =4)
~ -- --- Cat S = 3
~- ---
-0 - -.... --
-- -....
CatS = 2 -- ---
---
~
"
....
~
Cat S = 1
Cat S= 4
4 6
Superficial Gas Velocity, ftlsee
8
Fig.l-4-Effect of superficial gas velocity on S02 removal
22
-------
Curve 640926-A
11
10
9
E
10 8
~
~
Q) CaC03
- 7
.-
a
CD
Q) 6
~
::) MgO-CaC03
V)
V)
N Q) 5
w ~
c...
ro
15 4
I-
3
2
11000 1100 1200 1300 1400. 1500 1600 1700.
Temperature, of
Fig. 1-5-Phase diagram for I imestone and dolomite, assu ming a boiler atmosphere
containing 15% C02 .
-------
to the combustor would not calcine and, therefore, would not be an .
effective S02 sorbent. At the same temperature and pressure, dolomite
would be readily converted to the "half-calcined" state: MgO-GaG03.
This provides the necessary stone porosity for effective S02 sorption in
high-pressure boilers, unless the operating temperature (even during
turn-down) would be maintained at l680°F or higher.
Although there is no evidence that MgO reacts with S02 in the
fluidized bed, its presence appears to account for the better utilization
ofGaO in dolomite (as compared to limestone) which was noted previously.
Data obtained by Skopp and co-workers [84] and shown in Figure 1-6
illustrates this point. The reaction rate constant at a given utilization
increased monotonically as the ratio MgO/GaO was increased in these
. experiments.
Apparently, the presence of unreacted MgO helps to maintain
an open structure so that more of the Ga can be reacted.
This conclusion is supported by analysis of partially reacted
limestone and dolomite by Jonke and co-workers [49] and by the National
Goal Board r 71J .
NGB utilized mercury porosimetry to examine the porosity
of the particles and reported that an impermeable sulfate shell was formed
around limestone 1359 particles but not around particles of dolomite
(either that from the United Kingdom or U. S. No. 1337). Jonke and
co-workers at ANL confirmed this result when 802 .and air were injected
into a fluidized bed of limestone. When S02 was generated by combustion
of coal in the fluidized bed, however, no shell was observed around the
limestone particles. (ANL utilized electron microprobe analysis in their
work.) This difference between results obtained with S02 generated by
combustion and S02 injected with air was attributed to the limestone's
cycling between oxidizing (absorbing) and reducing (desorbing) conditions
as it circulated through the combustion bed.
The alternate adsorption-
desorption, it was postulated, permits the S02 to penetrate throughout
the particle. (It also reduces overall percent S02 reduction achieved
in shallow beds.) The above result may explain the fact that no
statistically significant difference in 802 reduction (at GalS ~ 2.3)
24
-------
T
c.
.-
E
("t\ <:) 6
......
x
-
~
...
04-'
~ 4
04-'
en
C
o
U
Q)
04-'
C'c:J
~ 2
o
u
m
e:::
Cur\l~ 644306-A
8
Ho = 2 inches
16000F
6/8 Mesh
MgOI CaO =
. .53
.. .16
. .02.
. .01
o
. 0
0.1 0.2 0.3 0.4
CaO Utilization, X
Fig. l-f>-Effect of MgOI CaO ratio
on stone reactivity
0.5
25
-------
achieved with dolomite 1337 and limestone 1359 was observed in the ANL
6-inch diameter fluidized bed combustor [51J, even though significant
differences in reaction rate constants have been observed [49] on
synthetic flue gases.
In addition to the differences in S02 removal efficiency
obtained with limestone versus dolomite, stones of the same type but from
different sources show wide variations in their reactivity with S02
[19~4l,63,7l,80,84]. In fluidized beds operating at l450°F to l600°F,
it has been found that U. S. Limestone No. 18 has a higher reactivity than
U. K. Limestone~ which has a higher reactivity than U. S. Limestone No.
1359' [71].
Tymochtee dolomite has been found more effective than U. S.
Dolomite 1337, especially at temperatures higher than l600°F [49,51].
As illustrated in Figure l-7~ the magnitude of the differences between
s'tones is large. These differences may be due to differences in the rates
of calcination~ differences in the pore size distribution present after
calcining~ subtle differences in chemistry, interactions between the coal
ash and the stone (see Figure 1-8) or to a combination of these and/or
other effects.
Numerous studies have been made [lO~19,70,7l,80] and are
in progress to elucidate the reasons, but no clear picture has emerged
at this time.
Another important difference between stones is their resistance
to attrition.
This. is especially important when a regenerable system is
considered.
The best results reported to date have been by Consolidation
Coal Company for Tymochtee dolomite [106]: 0.25% to 1.5% calcium attrition
per regeneration cycle. Other results include 3% to 5% attrition per cycle
for limestone 1359 [84] and 10% per cycle for U. K. Limestone. (1)
In view of its high sulfur retention, better regeneration
characteristics, and low attrition rate, Tymochtee dolomite appears to
be an attractive sorbent and should be studied further in any future large-
scale fluidized bed combustion experimentation. The success of future
(l)Esso Research Centre~ England, private communication.
26
-------
Cu rve 64430Lt-A
100
20
80
s::: 60
o
<1.>
""-
"-
+oJ E
s::: 0
Q.) -
"ti) 0
Q:: c;)
(/") 8;
~ 40 ~
~
~
I-.
o
o
1 2
Ca/ S Ratio
3
Fig. }-7- Su Ifu r removal effectiveness
27
-------
100
c
.0
.......
c
OJ
Ci) 60
a:::
L.
:J
-
:J
V)
C 40
OJ
U
L.
OJ
a..
Curve 644305-A
80
20
°0
1 .2
Cal SRatio
3
Fig. 1-8-Differences in effectiveness of su Ifu r
removal with same limestone, but different coals
[Data from National Coal Board 69]
28
-------
fluidized bed combustors in effectively reducing SOZ pollution may
depend upon finding stones with properties similar to Tymochtee dolomite
near the proposed power plant site.
Regenerated Sorbent Activity.
Available data indicate that
sulfated limestone and dolomite can be regenerated at atmospheric pressure
for subsequent sulfur absorption duty.
Calcium sulfate is reduced to
calcium oxide while driving off an SOZ-enriched gas stream -- 3% to 10%
SOZ. No data are available for sorbent activity following the two-stage
regeneration forming CaC03 plus HZS for regeneration at high pressures.
PER performed simultaneous absorption-regeneration runs in
their fluid bed module equipped with carbon burn-up cell (CBC) OMonthly
Progress Report 18, March 1971). Although the CBC would not accomplish
simultaneous carbon burn~up and sorbent regeneration (PER concluded that
three reactors are needed: a boiler, a CBC, and a regenerator), stone
activity remained high.
One bed operated in excess of 100 hours achieved
90% sulfur removal.
Limestone addition was needed to replace analytical
samples and attrition losses -- approximately 5% of the coal rate.
This
method of operation corresponds to that expected in an actual boiler
system.
Sorbent is continuously fed to and withdrawn from the boiler
bed, regenerated and recycled back.
All beds contain some fresh stone and
some stone that has undergone many absorption-regeneration cycles.
Regenerator operating conditions, however, will have to be modified, since
the S02 level in the regenerator off-gas was usually below 3%.
EssoR&E [84] examined the cyclic capacity of 1359 iimestone and
1337 dolomite in batch absorption-regeneration experiments. In the
first set of experiments a bed of solid sorbent was sulfated by a simulated
flue gas. Simulated regeneration gas converted CaS04 back to CaO, driving
off SOZ. The reSults from these experiments are shown in Figurel-9 as
curves I and 2. The limestone activity increased after the first regeneration
cycle and then steadily decreased. Dolomite activity measurements were
difficult to obtain because of high attrition losses.
29
-------
Curve G44308-A
0.9
0.8
1. Esso. 1337 dolomite. 300-700~. si m. flu e gas.
2. Esso. 1359 limestone. 300-700~. sim. flue gas.
3. Esso. 1359 limestone. 460~. coal.
4. Esso. 1359 Ii mestone. 460~. coal.
5. Esso. 1359 limestone. 930~. coal.
6. Consol. Tymochtee dolomite. 6oo-1000~. coa I
+ S 02 enrich.
6
C 0.6
2
-
-
::::J
..9
N
o
V) 0.5
~
~
-
"'
N
0.1
4
5.
01
2
3
8
9
10
4 5 6 7
N, cycle nu mber
Fig.l-9-Regenerated sorbent activity
30
-------
In Esso's second set of experiments, a bed of 1359 limestone
absorbed sulfur released from combusting coal in the bed.
After sulfur
absorption is complete, the bed is regenerated by reacting with simulated
regeneration gas.
The bed is then sent back to the combustion chamber
for another sulfur absorption cycle.
Sorbent activity for these runs is
much lower than when simulated flue gas was employed (see Figure 1-9,
curves 3, 4, and 5).
An interaction between the Fe of the coal ash and
the stone was postulated as the cause of reduced activity.
Zielke and co-workers [106] examined the cyclic activity of
Tymochtee dolomite. Their procedure was somewhat different from Esso's
in that they maintained a continuous sorbent flow through the combustor
and regenerator beds.
Combustion and regeneration were not done
simultaneously, but rather in batches.
In absorption runs the fluidizing
gas was enriched with S02 to achieve the desired sulfur absorption rate --
80%~ Thus, the S02level in the absorption bed is much higher than in
the Esso or PER experiments. Zielke's results are shown in Figure l-9.as
curve 6.
For a conservative design basis, one could select the CaO
utilization following regeneration bounded by curves 1, 3, and 4. This
would suggest that the sorbent could undergo five to six regeneration
cycles and still retain sulfur absorption activity.
Curves 2 and 6 seem
to give an upper limit on the CaO utilization that might be achieved
following regeneration.
The regeneration processes have to be studied in more detail to
determine the effects of regeneration on the sulfur absorption properties
of the sorbent.
Conceptually, various regeneration conditions may alter
the size, structure, chemical composition, surface properties and strength
of stone entering the boiler.
At this point in time we can only speculate
on which regeneration factors are significant.
The ~t Ht&trix gives a
preliminary listing of regeneration conditions we believe will have a
significant effect on sorbent properties.
31
-------
Regenerator temperature
REGENERATOR FACTORS
Average and maximum to which sorbent is
exposed
Heat transfer to sorbent and partial
pressure of gaseous components
Regenerator pressure
Gas composition
Oxidizing or reducing; calcining or
carbonating; steam-hydroxide melt formation;
local and overall
Solids residence time
Exposure time to high temperature; degree
of regeneration
Attrition
Gas velocity and bed depth
ReRenerator fuel
Ash level in bed; carbon coating on particles
Stone Particle Size.
Stone particle sizes utilized in removing
S02 during fluidized bed combustion can be divided into two classes:
. Coarse -- stone particles large enough so that they are not
elutriated and thus make up the dominant material in the
fluidized bed.
Such stones generally have a mean particle
diameter greater than 300 ~.
. Fine -- stone is completely elutriated from bed; bed consists
primarily of an inert material.
Such stones generally have a
mean particle size less than 100 ~.
There is no lack of disagreement in the data reported:
. In shallow beds (10-18 in) at high velocity (13 ft/sec), PER
[80] reported that fine stone gives higher S02 removal than does
coarse stone.
(See Figure 1-10)
. In deeper beds (3 ft) at low velocity (3 ft/sec) the NCB [70,71]
reported higher removal with coarse stone than with fine:
(See Figure 1-11)
, To cover all combinations, ANL[44] reported comparable
results with coarse and fine stone.
Closer examination shows that the fine particle data from all
experimenters is in agreement, despite the differences in velocity. The
low S02 reduction obtaine~ by PER is believed to be a direct result of
their experimental technique and equipment geometry. First, their
32
-------
20
--
c:
Q)
U
I-
Q)
c..
c: 40
o
--
U
:J
"C
Q)
0::::
Q)
"C 60
w
w x
.2
a
I-
:J
-
:J
V') . 80
1000
o
200
10"
18'
400
600
800 1000 1200
Top Particle Size, microns
1800
Curve 645931-A
Test Conditions:
Sorbent: 1359 R, Sized as Shown,
Fed at Stoich iometric
Ratio of 2.6 (Cal S)
Ohio #8 Pittsburgh Seam,
Washed, 3.0% S
Bed: Sintered Ash, Depth as Shown
Superficial Velocity: 13 FPS
Coal:
1400
1600
Fig. 1-1O-Variation in sulfur dioxide reduction with sorbent particle size, bed depth and bed temperature [80]
-------
100
90
. .5 80
u
:::J
~ 70
e::::
Q)
"'0
'x 60
o
.-
o
~ 50
:::J
-
:::J
V) 40
~
30
20
10
Cu rye 645839-A
PER Trend Lines, 12 - 14 fps, 1500 - 1600°F,
12 - 15" Expanded Bed (80)
v NCB, -1/811 Stone, 8fps (70,71)
o 'NCB, - 150~m Stone, 8 fps (70,71)
-7 + 14 Mesh
(1337 Dolan ite)
-325 Mesh
(1337 Delon ite).
v
325 Mesh
(1359 Limestone)
- 7 + 14 Mesh
(1359 Li mestone)
v
00
1. 0 2. 0
Cal S Stoichiometric Ratio
Fig. l-ll-Effect of stone size on su Ifu r dioxide reduction
34
3.0
-------
experiments were started with an ash bed and were not run long enough
to achieve a predominantly stone bed. Thus, referring to Equation 1-3,
the product 'aL' in the definition of normalized bed height was low.
Also, because of the extremely low aspect ratio employed, much of the
S02 would be expected to be released near the top of the bed in their
apparatus.
Coarse particles are recommended primarily because regenerable
systems seem necessary. The ash from most American coals investigated
is elutriated from the bed.
Therefore, if coarse stone which remains
in the fluidized bed is utilized, separation of the stone and ash can
be achieved in the fluidized bed combustor, thus simplifying the regen-
eration process. In addition, if the results of PER are discounted for
the reasons noted above, it appears that the increased residence time
of the coarse stone in the system more than compensates for its reduced
reaction rate constant so that the 802 removal achieved at a given CalS
can be expected to be as high (or higher) with coarse stone as with
fine (except where low LID and high temperatures are employed).
Bed Depth.
Altering the bed depth in a combustor of fixed
cross-sectional area changes the gas residence time in the bed, solids
circulation, and the aspect ratio (LID) for the bed. Reported data
are of little value in predicting the effect of bed height on 802 re-
moval in contemplated boiler designs.
It is useful, however, in
illustrating that both residence time and aspect ratio effects are
important.
Essoreported [38,39] that increasing settled bed height (from
4 in to 17 in) decreased the s.tone utilization fixed value of 802 removal
(from 20% to 7%). Both their calculations and experimental observations
indicated that the bed was slugging badly. In this case the effect of
poorer gas-solids contacting completely overshadowed the apparent increase
in residence time.
Pope, Evans, and Robbins [80] reported an increase in 802
removal from 35% to 47% at l550o~ and from 24% to 30% at l675°F when the
. .
static bed height was raised from 10 in to 18 in (See Figure 1-10).
These increases are approximately 70% of those predicted frqm a simplified
model (i.e., R = 1 - e-yL). Once again, however, it must be pointed out
that these experiments were started with an ash bed and were of short
duration.
35
-------
The inventory of stone in the bed (and thus the product 'aL')
was, therefore, more a function of run time than bed height. Further-
more, aL cannot be calculated from the data published. The mechanistic
reason for the improved removal remains undefined.
The small improvement observed in Figure 1-10 by increasing L
with fine particle systems is not surprising, since most of the stone
is carried along with the gas rather than maintained in the bed. The
gas contact with stone is determined more by the total reactor size
(including freeboard) than by the bed depth.
Bed Area.
The cross-sectional area of the fluidized bed is
also an important factor
for the proposed designs
to consider in data evaluation. The bed areas
range from 35 ft2 for each bed in a 300 MW
to over 300 ft2 for each bed in a 600 MW
pressurized boiler plant
atmospheric pressure plant.
2 2
0.1 ft to 9 ft .
Data have been acquired on beds ranging from
Wall effects, solids circulation patterns, and gas
flow characteristics change with bed diameter.
Preliminary indications
of the effect of scale-up are encouraging: essentially the same 802
removal was obtained in a 3 ft x 3 ft square bed as had been o~tained in
a 6-in diameter bed using the same materials at identical operating
conditions.
Coal Type. The degree of 802 removal achieved may be influenced
by many coal properties: swelling index, ash composition, percent ash,
ratio of inorganic to organic sulfur, total volatiles content, and
particle size, to name a few. Experimental evidence has been presented
to support the effect of the first two properties.
Figure 1-2 illustrates the fact. that better 802 removal is
achieved with coals having low swelling indices than those having high
ones.
This can be explained by the faster release of volatiles from
the coals with low swelling indices. Thus, a is higher for coals with
low swelling indices. Argonne [49] has documented the real effect of
sulfur release patterns on the" removal.
However, the importance of
36
-------
the sulfur liberation pattern from the coal itself would seem unimportant
at bed temperatures exceeding 1600°F where sulfur is constantly being
released and readsorbed by the adsorbent.
effect of temperature.)
(See previous section on the
The effect of coal ash components on the activity of the
stone after regeneration was discussed in previous sections.
The Bureau of Mines [95,96] reported the formation of a glassy
coating resulting from the reaction of silicate minerals in the coal
ash with lime.
The run was made at lSOSoF using a high-ash (Lower
Kittanning), unwashed coal.
The sulfur removal obtained in this run
with Ca/S'= 1.4 was only 15% compared to results near 70% for low-ash
coals in the same series of runs.
In subsequent runs with washed
Lower Kittanning coal, however, 'high sulfur removals (89% to 94%) were
obtained. Further investigation of the interaction between limestone
and silicate minerals seem justified, but no firm conclusions with
respect to its effect on S02 removal can be drawn at this time.
Excess Air.
The amount of excess air appears to playa minor
role in determining S02 removal at lower (below l500°F) combustor
temperatures, but a major role at higher temperatures. This effect is
believed to result from inhibiting the re-release of S02' which occurs
when locally reducing conditions exist (see discussion of temperature
earlier in this section). However, no quantitative data on the effect of
high excess oxygen (laY. to 50%) at elevated temperatures (1600° to 19000F)
were found.
Such data may be vital to effective design of a carbon burn-up
cell with low S02 emissions.
Other Factors. Other factors influencing S02 removal, but not
discussed as separate topics in this subsectio~include:
1.
2.
3.
Location of coal and stone feed and their method of injection
Gas-solids contacting effectiveness
Extent of solids circulation within the bed
4.
5.
Combustion pressure
Sulfur level of the coal.
37
-------
The effects of factors 1, 2, and 3 above have been alluded to in .
earlier parts of this subsection.
Sufficient data are not available to
quantify these effects on s02 removal, 'but a summary of the qualitative
picture will be made here.
Coal Feed.
The coal should be fed in such a way as to minimize
the amount of S02 released near the top of the bed. This can be done by
injecting i't near the bottom of the bed and by preventing its being swept
upward by the circulation patterns in the bed immediately after injection.
Stone Feed. The stone should be fed near the top of the bed to
achieve maximum counter-current action.
.Gas~Solids Contacting Effectiveness.
Good contacting is
achieved by preventing channelling and the formation of large gas bubbles,
i.e., by maximizing the interchange of gas between the emulsion phase and
the bubble phase of the bed. Accomplishing good contacting is one of the
keys to designing fluidized beds and is the subject of an entire section
of chemical engineering .literature. (See, for example:
Kuniiand Levenspiel [60], Zenz and Othmar [105].)
Leva [61],
Solids Circulation.
This factor is likely to be much more
important at high-temperature (i.e. >1550°F) bed operation than at low-
temperature operation.
Since the combination of high temperature and
locally reducing conditions results in the release of S02 from CaS04'
it is desirable to reduce the probability of locally reducing conditions
and also the amount ofCaS04 present near the top of the bed when high
temperatures are employed. This can be achieved by reducing the degree
of solids mixing, if the fresh stone is fed near the top of the bed
and coal near the bottom.
Those factors which improve contacting
effectiveness, unfortunately, also tend to improve solids circulation
unless special ,baffles or other intervals are utilized. Once again, the
method employed to achieve a balance ,is best left to the bed designer.
order in
Combustion Pressure. If the reaction of CaO with S02 is first
PSO ' then operating pressure will affect the degree of S02
2
.38
-------
removal achieved only via its effects on the calcination rate of the
sorbent and on the quality of fluidization.
British results [69,
Fig. 7, Task IIJ in .which greater than 90% reduction in S02 emission was
obtained with an average Ca/S of approximately 2 and an operating pressure
of 5" atm are encouraging. The low velocity (~ 2 ft/sec) and temperature
. (14800F) employed, however, and the limited data available make extra-
polation of these results to the proposed conditions difficult.
Sulfur Level of the Coal.
Once again, no effect is anticipated
if the reaction is first order.
Data reported (see Figure 1-2) have
shown no correlation between sulfur level in the coal and the degree of
removal achieved. One caution: at high operating temperatures where the
amount of S02 released from CaS04 is high compared to the amount released
by the coal, it seems likely that the degree of removal will be directly
proportional to the sulfur level in the coal. In other words, the S02
level in the gas and, therefore, the emission will be determined completely
by the temperature, gas composition, stone loading, etc.
Extrapolation of Existing Data
Despite the complexity of interaction among the effects of the
variables which influence sulfur remova~ it appears feasible to extrapolate
a limited amount of existing data.
The general form of the Koppel model outlined previously is rec-
ommended for this use. Although it is relatively simple mathematically
and contains only three experimental parameters (Ye' to.63~ and a), this
model form appears to predict correctly the direction and order of magnitude
. .
of change in S02 removal which result from changes in Ca/S, fluidizing
velocity, bed depth, and coal type. It has the flexibility to project the
in-bed effects of temperature, stone type and particle size, regeneration
(and .other factors affecting stone structure), pressure, and excess air on
the basis of a limited amount of bench-scale experimental work.
S Ii gh t
modifications are suggested to account for contacting effectiveness,
the re-release of S02 from CaS04 owing to locally reducing conditions, and
the use of partially sulfated stone. Also suggested is a method for
39
-------
treating. deep beds and beds with internals where solids circulation
. cannot. be expected to be. "perfect."
The degree of S02 removal projected by the model for the
industrial boiler design operating conditions illustrates the approach.
The industrial boiler design has the following operating conditions.:
T
= l650°F
L
=, 2..5. ft.
v
= 12.5 ft/sec
d ::<. 2000 ~.
p
limestone, 1359
P
= 1 atmosphere
Pittsburgh. #8 seam coal
4.3% sulfur
SwellIng. index 5 to 5.5.
Argonne:' s data. point at~ l600°F for Illinois No.6 coal, 490 ~
1359 Limestone, 3. ft/sec., Ca/S = 2.5,. static bed height = 15 inches,
and utilizing partially sulfated stone as the bed material (shown in
Figure 1-3) was. chosen. This data point. was chosen because it employed
the same. sorbent.,. the same bed. material, a similar coal,. and nearly the
same temperature.
Also, this run had c'on.tinuous: feed and withdrawal of
re.actants, it was allowed: to reach a steadY. operating condition., and its
result.s are consis.tent with other well-d.eflned. runs at similar conditions.
determine
A plot. of rate. cons:tant versus' C:aO utilization can be used to
(y /y ). The in.tercept of the. straight line with the utilization
e s . .
the value oE Ye/y s' consis.tent with Koppel's expression for
axis gives
k (y) :.
Ye.
key) =. C' t
o 0...63
(1 - y/y ).
e
(1-6)
40
-------
Data for dolomite 1337 [41] are plotted in Figure 1':"12. Similar plots for
limestone 1359 kinetic data obtained by Esso [84] indicated values of
y /y of 0.4 for -16/+18 mesh material (~ 1000 ~) and y /y = 0.25 for
e s e s
-6/+8 mesh (~ 2400 ~). All of the above data were taken at l600°F. A
plot of Ye/Ys'versus average particle diameter is shown in Figure 1-13.
If one assumes that particles in a bed of burning coal have the
same value for Y as those tested. in synthetic flue gases, then it is
e. .
proper to utilize y /y = 0.65 for the 490 ~ particles used in Argonne's
. . e s .
experiment and 0.27 for the industrial boiler which employs stone with an
average diameter near 2000~. If localized reducing conditions exist,
however, it has been shown that sulfur is uniformly distributed throughout
the particle, implying the possibility that y /y + 1. As a compromise,
e 's
used for the Argonne experimental data and
a value of y /y = 0.9 was
e s
Ye/Ys = 0.7 for the commercial unit.
Equation 1-3 simplified for a = 0 (i.e., uniform sulfur
generation'throughout the bed) was utilized in the calculations
(1 - R/r)H + e-(l - R/r)H - 1
R = (1 - R/r)H
(1-7)
Substituting R = 0.86 and r = 2.5 x 0.9 into Equation 1-7, H can be
calculated for the Argonne run as H = 11.5. If H were calculated
. e~,
directly from its definition and available kinetic data, a value of '58
is obtained.
The difference between Hand H can be explained in terms
exp
of gas-solids contacting effectiveness and the reduction in net S02
pickup due to locally reducing conditions~ i.e.,
H
exp
=Hx EX
c
o
- c*
c
o
(1-8)
41
-------
~ 0.20
o
..-
VI
~
.a 0.16
~
.c
.p-
N
-
.= 0.12
E
"""N
a
V>
VI 0.08
.c
.
..-
c:
co
~ 0.04
o
u
Q,)
..-
to
0:::
Curve 644309-A
Dolomite-1337
0.3 vlo 5'02 at 1600°F
~ 0.20
o
..-
VI
dp = 96 ~
o
o
20
40 60
CaO Utilization, %
80
~
a 0.16
to
U
.c
- .
.= 0.12
E
-
N
a
V>
~ 0.08
.
..-
c:
to
..-
~ 0.04
o
u
Q,)
..-
to
0:::
100
dp = 282 ~
o
o
20
40 60
CaO Utilization, %
80
100
Fig. 1-12-Rate constants for Dolomite 1337 as a function of calcium oxide uitilization
-------
~ 0.5
--
~
Curve 645932-A
1.0
o 0
Fig. 1-13 -Ye I Y s vs particle diameter for limestone 1359
43
3000
-------
where
£ = contacting efficiency,
C
industrial boiler and in the Argonne expe~iment, then
Co = S02 concentration in the gas, and
C* = effective equilibrium partial pressure of S02 at the
stone surface due to locally reducing conditions.
C - C*
If one assumes that £ and 0 .
will be similar in the
H (industrial boiler) = H
exp
2..5 ft
x
1.25 ft
3 ft/see
x
12.5 ft/sec
x ~:~ = 4.25
r = 6 x O. 7 = 4,. 2
R = 1/2(5.2 - j(3~2)2 4 x,4.2) = 0 71
+ 4.25 ..
An alternative computation for the same boiler could be based
on PER data r80, FBC test tl72l:
R = .347
dp= 840 ~ (1359 limestone)
Ca/S = 2.6
1640°F
v
= 13 ft/sec
18" static bed depth (assume 50% of bed as stone, 50% ash).
The experimental H is 1.8 and
2.5 13 7
H(boiler) = 1.8 x .5 x 1.5 x 12-.5 x 9" = 4.9
r = 4.2
R = 1/2(5'.2, -')('3.2)2 + 4 ~.:.2) = 0.75.
44
-------
The closeness between the two results is encouraging, although it should
be noted that the assumption of percentages of stone and ash in the bed
were necessary in the second estimate since sufficient data were unavailable
to calculate this quantity.
The assumptions used for the model and the
small-scale fluidized bed data give a conservative estimate of the S02
removal which can be expected for an industrial boiler.
Modification of Koppel's Model for Partially Sulfated Stone
Koppel's model is readily modified to account for partially sulfated stone
being fed to the boiler for S02 absorption duty. This would occur physically.
when regenerated stone is recycled back to the boiler beds. Regenerated
stone will be partially sulfated, especially if regeneration is accomplished
in well-mixed, fluidized beds.
Koppel's equation for sulfur removal becomes
R =
1
{l-
R Yi
-( 1- -- -)H
[a.a r Ye - (1
R.
a. - (1 - -
r
R y.
1. -a.
--r-W)He ]}
- ~)H
(1-9)
-a
1 - e
where
Yi is the sulfur
Silb of GaO).
loading of boiler inlet sorbent (lbs of
For fresh stone Yi = 0.
All other terms
are defined on pages 9 and 10.
The average sulfur loading of sorbent leaving the boiler is given by
FL R
. 0
Y = Yi +-;-
(1-10)
Figure 1-14 shows the variation of sulfur removal with GalS ratio (r)
for different inlet sulfur loadings (y./y ) and SO generation profiles
1. e 2
(a) .
45
-------
1.0
0.9
0.8
0.7
0.6
R
0.5
0.4
0.3
0.2
0.1
a = 5.0; y./y = 0.05;
I e
a = 5.0; y./y = 0.20;
I e
a = 2.0; y./y = 0.20;
I e
a = 2.0; y./y = 0.05
I e
o
10-2
1
100
r
Curve 644312-6
Calculated 502 Removal R
H = 10
c. -c
R = I n out
c.
In
W YE
r= -
F c.
In
101
. Fig. l-14-Theoretical 502 Removal efficiency using partially sulfated stone
10-1
46
102
-------
Deep Beds with Internals. In deep beds with internal heat
transfer surface, circulation of solids may be impeded. If stone is fed
to the top of such a bed and removed near the bottom, a staged counter-
current effect may be achieved.
In such a case, the assumptions of well-
mixed solids and uniform bed temperature made in deriving Koppel's model
cannot be justified. In this case 502 removal may be computed by dividing
the bed into a number of segments with each being considered as well mixed.
47
-------
Sorbent Regeneration And Sulfur Recovery Systems
A number of processes have been proposed for converting a
sulfated limestone or dolomite sorbent from a fluidized bed boiler back
to an oxide or carbonate.
Thus regenerated, the sorbent can be returned
to the boiler for further capture of sulfur as 50Z from the combustion
gases. The sulfur emerges from the various proposed regeneration
processes as either HZSor SOZ mixed with gaseous reactants and reaction
products. Further processing of this mixture then recovers sulfur as
a final product, either elemental S or sulfuric acid, H2S04'
The optimum choice of a sorbent regeneration and sulfur
recovery process is a technical and economic problem. Its solution,
depends on the operating conditions of the boiler, on the type of sorbent
used, on the market for sulfur and/or sulfuric acid, and on a number of
chemical kinetic, and operational factors which have not yet been
thoroughly explored.
A listing and brief de~criptions of some of the proposed
sorbent regeneration processes follows.
1.
Reduction of calcium sulfate to calcium sulfide and
subsequent regeneration with steam and COZ'
{4 HZ}
CaS04 + 4 CO
-+
CaS + {: ~6~}
(1-11)
CaS + HZO + COZ
-+
CaCO) + HZS
(l-lZ)
This process has been investigated by Consolidation Coal and the City
College of New York.
To produce an effluent gas with a sufficiently
48
-------
high concentration of HZS to allow it to be fed directly to a Claus
plant for sulfur recovery requires: 1) sufficiently rapid kinetics for
Equ~tiori 1-12 at 800°F so equilibrium between the gas and solids is
approached; or Z) a C02-H20 regenerating gas essentially free of nitrogen.
Such a mixture could be produced from a steam-oxygen blown gasifier, a
steam-methane reformer. or a hot carbonate flue gas scrubber. Alterna-
tively, an air-blown gasifier could be used, and the regenerated HZS
gas concentrated with amine wash towers.
2.
Direct reduction of calcium sulfate to calcium oxide and
502 employing partially oxidized fuel as the reducing agent.
CaS04 + {~6}
~
CaO + SOZ + {H20}
COZ
(1-13)
This process is thermodynamically favored at low pressure and high
temperature. But Esso and Consolidation Coal have both shown that
ash agglomeration problems may be encountered with this system if there
is a significant ash fraction carried into the regenerator with the
sorbent or with the fuel.
3.
Reduction of calcium sulfate to calcium sulfide and
subsequent oxidation to form calcium oxide and S02'
{4 HZ} ~ CaS + {~ ~~~}
CaS04 + 4 CO .
3 CaO + SOZ
CaS + 2 02 ~
(1-14)
Consolidation Coal and Esso (England) have investigated this technique
experimentally at atmospheric pressure with encouraging results.
However, the equilibrium of the SOZ rejection reaction is unfavorable
for the pressurized boiler plant; the maximum S02 concentration is
about 1% in the gases when regeneration temperatures lower than ZOOO°F
are used.
49
-------
4.
Reduction of calcium sulfate to calcium sulfide and
subsequent regeneration by reaction
-------
V1
t-'
Sulfated Stone
from Boi 1 er
Coal
Ai r
Steam
Solid
Surge
Tank
Gas
Producer
Ash
Regenerated
Stone to Boiler
1st Stage
Regenerator
1/4 CaS04 + (~~)----
1/4 CaS + (CO 2)
H20
2nd Stage
Regenerator
CaS + C02 + H20
- CaC03 + H2S
Waste
Waste
H2S Rich Gas
To Claus Recovery
Compressor
Dwg. 6161A03
To Final Particulate
Remova 1
Turbine Expander
To Stack
CO
2
Enri cner
Fl ue Ga$
Fig. 1-15- Su Ifate-su lfide redu ction rege neration sche matic
-------
Fuel combustion
C + 0z
-+.
COz
(1-15)
1
HZ + "2 0z
-+
HZO
(1-16)
Fuel gasification
C + {HZO}
COZ
-+
co + {HZ}
CO
(1-17)
Water gas shift
CO + HZO -+
COz + HZ
(1-18)
In the first regenerator stage, calcium sulfate from the
boiler is reduced to calcium sulfide.
This regenerator reactor can be
a bed fluidized by the reducing gases from the gas producer, a moving.
bed, or a panel bed. The choice has to be made on the basis of reaction
rates and solids handling processes. Preliminary process designs have
been based on a fluidized bed operating at about 1500°F and 8 atm.
The reduction reactions occurring 'in the first regenerator.
stage with a dolomite sorbent are
Sulfate reduction
. [CaS04. + MgO] +{4 HZ}
4 CO
-+
[CaS + MgO] + {4 HZO}
4 COZ
(1-19)
Carbonation
[CaO + MgO] + COZ -+
[CaC03 + MgO]
(l-ZO)
Water,gas shift
CO + HZO
-+
COz + HZ
5Z
-------
Some undesirable side reactions may also occur in the first regenerator
stage.
Side reactions
CaS04 + {~~ J -+ ( H20 1 (1-21)
CaO + C02 +S02
CaS + H20 -+ CaO +H2S (1-22)
CaS + C02 -+ CaO + cas (1-23)
1 1 (1-24)
CaS + "2 COz -+ CaO + "2 CSz
1, 3 Z CaO + Sz (1-25)
"2 CaS04 + "2 CaS -+
The sulfate reduction reactions are thermodynamically favored
by low temperature and unaffected by pressure. Below l600°F the
. . 8
equilibrium constants for these reactions are greater than 10
(COZ/CO! HZO/H2 > 100); thermodynamicallYt these reactions virtually
proceed to completion. In actuality, the degree of conversion of
CaS04 will be limited either by the supply of reducing gas or by the
kinetics of the reduction reaction.
Low temperature and high pressure suppress the equilibrium
formation of SOZ' COSt HZS, CS2' and S2 while enhancingrecarbonation
of calcium oxide. Recarbonation should terid to drive the reduction
reactions by removing C02 from the gas.
One caution must be exercised in selecting the first-stage
regenerator operating temperature and pressure.
The stearn partial
pressure must be kept below the level at which Ca(OH)2 will form; the
binary system Ca(OH)Z-CaO forms' a low melting eutectic at l450°F, and
Ca(OH)2-CaC03 forms a low melting eutectic at ll80°F. Curran [21]
reports that melts form at temperatures of l500-1600°F whenever the
53
-------
steam pressure exceeds 13 atm.
Within these temperature-pressure
limitations, the actual operating temperature and pressure will depend
on chemical kinetic, equilibrium, and overall process considerations.
Useful energy may be extracted from gas leaving the first.,-stage
regenerator vessel. This gas is at high temperature, l500°F, and at
. high pressure, if the regenerator is used in conjunction with a high-
pressure boiler.
Passing the gas through a turbine expander can recover
some of this energy. .
Reduced solids flow from the first-stage regenerator through
a stand. pipe leg into the second-stage regenerator vessel, where steam
and carbon dioxide convert calcium sulfide to calcium carbonate.
This
vessel contains a fluidized bed operating at about llOO°F and 12 atm;
it produces a gas with 23% H S on a dry basis. A hydroxide-carbonate,
2
Ca(OH)2-CaC03' eutectic melt may form at temperatures above l180°F.
Primary reactions occurring in the bed are:
H2S formation
[CaS + MgO] + H20 + C02
-+
[CaC03 + MgO] + H2S
(1-26)
Carbonation
[CaO + MgO] + C02
-+
[CaC03 + MgO]
(l~Z7)
Increased H2S concentration is thermodynamically favored by
the combination of low temperature and high pressure, which also
suppresses competing side reactions.
Although an equal molar ratio of
steam to COZ in the regenerator top gas yields the maximum theoretical
HZS concentration, regenerating with excess steam produces a higher
HZS concentration on a dry basis. COZ requirements are also reduced
by this approach. A feed gas comprising primarily steam and COZ is
required to obtain high HZS concentration.
54
-------
Feed gas for the second-stage regenerator is obtained by
stripping C02 from boiler flue gas in a regenerable hot-carbonate or
amine scrubber solution. A C02 enricher unit, commercially available,
consists of two large towers -- one an absorber and the other a
regenerator -- a reboiler, a solution cooler, and solution pumps.
A
cost estimate for a Benfield designed system is included in Appendix I
A compressor is required to provide the C02-H20 reactants at the
pressure of the second-stage regenerator.
Pure C02-steam mixtures can also be obtained from a steam-oxygen
blown gasifier operating at, elevated pressure. An on-site oxygen plant
for such a regeneration process, however, would be an expensive
alternative.
A C02-H20 gas from an air-coal-steam combustor can also be
used to generate H2S, The HZS concentration in the off gas would be
much lower than when C02-H20 undiluted by N2 was employed. A selective
,
solvent could then be employed, however, to enrich the HZS concentration
before further treatment.
The two-stage sulfate-sulfide sorbent regeneration process
appears particularly adapted to high-pressure boilers.
Although high
pressure does not enhance the first-stage reactions, it decreases the
formation of SOZ and other undesirable byproducts. Vessel diameter
for a particular gas flow rate is likewise reduced. High HZS concentra-
tion in the acid gas leaving the second stage is also thermodynamically
"favored by high pressure. Compressing the steam and COZ feed to this
reactor and the air to the first-stage reactor, however, increases
capital and operating costs for the process.
Before entering the Claus plant for conversion to elemental
sulfur, the HZS-rich gas may be expanded to approximately Z atm. In
addition, Claus plant operation will be improved if steam is condensed
from the acid gas prior to enter~ng the Claus plant.
55
-------
This pressurized process scheme, or reaction sequence, can also
be used in conjunction with low-pressure boilers. Spent stone would be
sent to a lock hopper either on leaving the boiler or between the first
and second regeneration stages.
&
Regenerated stone would likewise be
held in a lock hopper before being recycled to the boiler.
The remainder
of the regeneration system will be the same as for the high-pressure
boilers.
Should reaction kinetics for the second-stage.~eaction of
calcium sulfide with steam and C02 be su~ficientlyrapid at 800°F, low-
pressure operation of the second stage would be feasible. Negligible
increased heat duty would be placed on the boiler even though stone
would be removed at l600°F and returned below 800°F.
Process Alternatives
A number of process alternatives can be considered for the'
sulfate-sulfide reduction regeneration process.
These alternatives are
concerned primarily with methods of modifying or eliminating the gas
producer and with methods of handling gases supplied to or removed from
the second-stage regeneration.
If natural gas is available, a steam-hydrocarbon reformer
could be used in place of a gas producer to generate CO and HZ for
reduction. Only the reformer step would be required, eliminating the
CO, conversion and C02 removal steps.
Another possibility is to use natural gas directly in the
first-stage regenerator for the reducing gas.
There is some doubt
that metnane will react with hot sorberit particles to produce CO and
HZ or that natural gas will reduce CaS04' Esso has performed experi-
ments using methane to reduce calcium sulfate to CaO while giving off
802' 80Z concentrations produced, however, were very low, and it is
not clear that methane is a suitable reductant to produce CaS at the
lower temperature.
56
-------
A solid or liquid fuel -- coal, char, or oil -- might also be
added directly to the first-stage regenerator along with ~he CaS04
sorbent. The sorbent particles would be fluidized by steam and suffi-
cient air to provide any heat required for regeneration. The C02 and
H20 gases from the first-stage regenerator would then be used in the
second-stage regenerator. If larger quantities of C02 were required for
H2S production in this second stage, it would be provided by increasing
tile flow of CaS04 from the boiler to the first-stage regenerator. More
carbonaceous fuel added to this stage would convert this CaS04 to CaS
and the desired C02'
In the second-stage regenerator, only a fraction of the CaS
would be converted to CaCO) to match the sulfur removal rate. The
remainder of the CaS would be circulated back to the boiler where it
would be reoxidized to CaS04' Solid sorbent, in this case, would serve
a dual purpose, namely, as a sulfur removal agent and a coal oxidizer.
,
Experimental verification °is required to prove this alternative
process technically feasible.
This process has possible problems
connected with ash agglomeration, sorbent deactivation, and high solids
flow rates. But it has the potential promise of eliminating a separate
gas producer and C02 enricher systems.
Sulfur Recovery
Two main choices are available for the sulfur recovery step,
depending on the final sulfur product desired -- elemental sulfur or
sulfuric acid.
A Claus plant would be the most reliable process for
converting H2S to sulfur. The split-flow process is usually used for
gases containing less than 50% H2S or with gas streams containing high
concentrations of hydrocarbons. A three catalytic reactor, split-flow
Claus plant will have a sulfur recovery efficiency of 92 to 94%, emitting
a tail gas containing 7000 to 8000 ppm S02' Ford, Bacon and Davis has
estimated the installed cost and utility requirements for such a plant.
(See Appendix I.)
57
-------
The Ralph M. Parsons Company suggests a different flow scheme
(1)
for HZS recovery. Their recommendation is presented below.
"If the H2S-rich gas is produced, I would recommend a scheme
which produces COZ for the HZS generator by treating a flue
gas with MEA or hot carbonate. I think it would be preferable
to treat the gas. from the HZS generator at 10 atmospheres
pressure with Purisol to extract both HZS and COZ together,
then regenerate the Purisol in two stages; the first stage
would produce H2S-rich feed gas for the Claus unit, and the
second stage would produce C02 for recycling to the HZS
generator. This scheme would be more economical, I think,
than the one Westinghouse proposes and I believe that it has
a very substantial operating advantage in making the HZS
concentration in the Claus feed controllable at a level which
would assure good operation of the Claus unit, regardless
of variations in the sulfur content of the fuel."
Should this option be selected, the flue gas C02 scrubber
system capacity would be reduced. In fact, this mode of operation
could allow regeneration with a dilute COZ-H20 gas stream.
plant cost should also be reduced.
The Claus
An HZS-rich gas stream is a suitable feed for a sulfuric acid
plant. HZS is first oxidized to SOZ' catalytically converted to S03'
and absorbed in water. Browder [14] describes the Ralph M. Parsons Co.,
"Double Contact/Double Absorption" HZS sulfuric acid plant for this
service. A 15% H2S gas stream when combusted with stoichiometric air
will produce an 8.7% 80Z gas stream. This is well within the feed
limits for the Parsons DC/DA process.
Regardless of which sulfur recovery process is selected, the
dust loading of gas entering the unit will be a key factor in the overall
(l)personal communication, July 20, 1971.
58
-------
unit cost.
The catalyst bed reactors cannot tolerate high dust loading,
so a gas purification system may be required.
Sulfate-Oxide Reduction to Produce SOZ (Atmospheric Pressure Operation)
This limestone/dolomite regeneration scheme is a single-stage
process in which the calcium sulfate from the fluidized bed boiler is
converted to calcium oxide by treating it with reducing gases, CO and
HZ. Sulfur is released
reductants and with COZ
elements of the process
from the process as SOZ mixed with unreacted
and H20 oxidation products. The essential
are shown in Figure 1-16. Esso, Consol,
Esso (England) and Pope Evans and Robbins have all reported experimental
results with this process operating at atmospheric pressure and at a
temperature in the range 1900 to 2000°F.
Esso [31] has sulfated batch beds of solid sorbent with
simulated flue gas and with actual combustion gases burning high-sulfur
coal in the bed.
The bed of spent sorbent was then regenerated using
reducing gas. The S02 concentration in the regenerator off gases was
8 mole % or greater at atmospheric pressure; gas-solid equilibrium ~as
approached.
Esso (England) [29] performed similar experiments.
They
sulfated a batch bed of lime by burning high-sulfur fuel oil. Regen-
eration was accomplished by increasing the fuel rate, keeping the air
rate constant.
At 1050°C (19Z0°F) and one at~ Esso was able to pro-
duce a 9% SOZ gas from their regenerator.
Consol [106] continuously regenerated sulfated dolomite, using
partial combustion of carbon monoxide gas at1950°F. Regenerator off
gas contained 3 to 7% S02. Dolomite was sulfated by combusting coal
in a bed of dolomite. The fluidizing gas was enriched with SOZ to provide
the desired sulfur absorption rate. A continuous stream of dolomite was
fed to and removed from both the regenerator and combustor.
59
-------
C1\
o
Steam
Sulfated Stone
from Boi 1 er
Sol id
Surge
Coal
Air
Regenerated.
Stone to Boiler
50
Rich Gas,
CaO
5 pen t
Stone
Gas
Clean
Up
Waste
Fig. 1-16-Su Ifate-oxide reduction regeneration schematic
Stone Fines
Regeneration
CaS04 +(~~ ) -
(C02
aO + SO 2 + H 0
2
Distributor
Dwg. 6l61A04
S02 Rich Gas
to Sulfur Plant
-------
ments.
PER [75] conducted simultaneous combustion-regeneration experi-
They attempted to use their carbon burn-up cell for .simultaneous
regeneration and elutriated carbon burn-up.
Although this dual function-
ing did not prove successful) they were able to maintain S02 levels of
3 to 4% in the CBC off gas by adding coal to the carbon burn-up cell.
PER's current plans call for a three-reactor boiler system - a combustor,
a regenerator, and a carbon burn-up cell.
Spent stone discharges from each bed of the fluidized bed
boiler into a hold vessel -- one hold vessel per boiler module.
The
stone is then pneumatically transported to the regenerator vessel using
air as the carrier gas.
The regenerator vessel is a fluid bed reactor operating at
one to two atmospheres and 2000°F.
Coal is charged into the reactor
bed.
An air-steam mixture fluidizes the reactor bed and combusts
coal to provide reducing gas for reduction of the CaS04'
The primary regeneration reaction is
sorbent sulfate-oxide reduction
CO CO 2
CaS04 + (H ) + CaD + (H 0) + S02
2 2
The fuel combustion and gasification reactions also are carried out in
the regenerator.
Fuel combustion
C + 02 + CO 2
H2 + 1/2 02 + H20
Fuel gasification
H20 H2
C + {CO} + CO + {CO}
2
61
-------
Only enough air is used to provide the necessary CO and
HZ reducing gases in the reactor. The reaction of air with the fuel
also provides the heat required to carry out the high-temperature :.
sulfate-oxide reduction process.
Steam is added to the regenerator
both to control the reactor temperature and to provide hydrogen for
reduction.
The exothermic heat of combustion must balance the
endothermic requirements of reduction and gasification, sensible
heat loss from solids circulation, and heat loss from the reactor.
Competing with the reduction to calcium oxide is the reduc-
tion to calcium sulfide according to the following reaction.
Sorbent sulfate-sulfide reduction
. CO C02
1/4 CaS04 + (H ) ~ 1/4 CaS + (H 0)
z 2
This reaction does not produce S02 but it does consume large quantities
of reducing gas for every mole of calcium sulfate reacted. This
reaction can be minimized by judicious selection of reducing gas compo-
sition and reactor temperature and pressure. Some speculation is
that sorbent is regenerated by first reducing some of the CaS04 to CaS
according to reaction 1-19. The CaS then reacts with CaS04 according
to reaction. below to form CaO and drive off S02.Thiswould be a
solid state reaction.
Sorbent sulfate-sulfide reaction
3/4 CaS04 +1/4 CaS ~ CaO + SOZ
(1-28)
Chemical equilibrium in this reaction can establish the
maximum concentration of S02 which can be attained in gases from the
regenerator. A high percentage of S02 is favored by high operating
..
temperatures -- around 2000°F, and low operating pressure -- about
1 atm.
The SOZ fraction in the stream to the sulfur recovery plant
62
-------
can be increased from 8-10% to 90% by volume by a dimethylamine, DMA,
absorber-stripper process operating on the regenerator top gases sub-
sequent to their cleanup and cooling.
Regenerated sorbent flows from the reactor vessel to a dis-
tributor vessel. Twenty-four lines return regenerated stone from the
distributor to each bed in the boiler modules.
The problems involved in designing this sulfate-oxide sorbent
regeneration process are selecting gas composition and flows and
designating temperatures and pressures to
.
Obtain high S02 concentrations (6 to 10% by volume)
.
Prevent production of CaS
.
Provide the heat required by the endothermic, S02 -
producing reactions and by the process of heating
the sorbent.
The size of the regenerator vessel must also be selected so that com-
bustion-gasification of the fuel and conversion of the sorbent sulfate
to oxide can be carried out to the extent required by the sulfur
sorption rate in the boiler.
This single-step regeneration process is.most effective
with low-pressure boiler operation. At high pressures the S02 con-
centrations produced appear too low for effective recovery of the
sulfur. A high-pressure boiler operation would entail the use of
high-temperature lock hoppers to step the stone down from boiler
pressure to regenerator pressure and then to step it back up to
boiler pressure.
Process Alternatives
Variations in the sulfate-oxide reduction regeneratlon
process are concerned with alternative methods of providing reducing
gases.
Adding coal or char directly to the regenerator brings
63
-------
quantities of ash in contact with the sorbent. Such contac't may cause
problems of agglomeration and sorbent deactivation. A separate
gasifier followed by an ash particle removal system can 'be used ,to
furnish a reducing gas to the regenerator.
Sulfur Recovery
The S02-richgas str,eam from .theregenerator C'an .be con-
verted to sulfuric acid or .elemental sulfur. A suitable acid produc-
.tionplant is connnerciallyavailable at this .time.
Appendix 1
con-
tains capital cost and operating cost estimates .for sulfuric acid
plants to ,treat .the S02-.richregenerator gas. 'The sources of ~these
es timates are Monsanto .Enviro-Chem Sy.stems Inc. and the Ralph .M.
Parsons Company.
. A .key design fac.tor in any sulfur recovery 'process is
the dust loading of the gas fed to the sulfur recovery plant. High
dust loading can cause rapid catalyst failure. '.Approximately half
the .capital cost for Parsons' .sulfuricacid plant is in the gas
cleaning section of the plant.. Monsanto .specifies a maximum dust
,loading of 0.5 gr/scf for the gas ent.ering the purification sect.ion.
No .commercial process is readily available to convert the
S02 ,to elementa'l sulfur. Allied Chemical [2] has investigated
several'dir.ect reduction processes ,i:ncluding one deve'loped by
Asarco,to reduce gases containing 0.3%, 3.0-16%, and 90% 502. . No
process appeared practical for the 0.3% case. Although gases
containing 3 to 16% S02 can be reduced directly, using coke or
methane as a fuel, Allied concluded that enriching .the SO:2 con-
centration up to 90% and then reducing .the S02 is more economical
and reliable.. This .conclusion is .shared by the Ralph M.Parsons Co.
and Ford, Bacon and Dav.is. (See Appendix .1 ). A selective solvent,
. such as DMA, ,to absorb 502 and produce an enriched S02 gas stream
is currently .being developed. Again gascleanliness-- less than
,64
-------
0.2 gr/scf particulates -- is important to the technical feasibility of
this process.
This absorption-enrichment reduction scheme appears promising,
particularly in conjunction with a high-pressure iegenerator, provided
,that the S02 concentration in the regenerator remains above 3%. The
sulfur recovery plant will be less susceptible to boiler and stone
regenerator fluctuations since it will be fed a concentrated S02 stream.
For the technical and economic feasibility estimations, sulfur recovery
was carried out by the Asarco process, as described by Allied Chemical,
using natural gas as a fuel.
Thermodynamic Data for Various Regeneration Processes
Equilibrium constants for the various regenerator reactions
are given in this section.
Data points referenced as "Squires" were
obtained from [89].
Data points referenced as "Esso" were calculated
from free energy data in [32].
"Esso (England)" to [29].
Similarly, "Conso1" refers to [21] and
Parenthesis, "( )", is defined as mole fraction, Le.,
"(H2S)" = mole fraction H2S, "'/T"is total pressure in atmospheres.
Equilibrium constants are either dimensionless or given in units of
atmospheres.
Also presented graphically is the enthalpy above a 77°F
datum for the solid and gaseous species.
Solids data were obtained
from Kubaschewski and Evans [59], and Strassburger [94].
data were obtained from JANAF [46].
Gaseous
65
-------
10Z
8
6
4
Z
101
8
6
4
Z
100
8
6
K 4
'" Z
'"
10-1
8
6
4
~
10-2
8
6
4
Z
10-3
600
700
800
1000
1200 1400 1600 2000 Z400 of
I K I f
f---- -
f-- caC03 -. CaO + COz I
I
K = (C02) 1T r
f
. . Consol !
- Squires
8 Esso
f
I
"(
r
/
I
/
/
I
I
/
I I I 1 I 1 I I I
E
~ 10.0
~
O. 6 x 10-3
(urvt: e..42~9-B
1000
700
900
1000
lZ00 1300 1400
OF
800
HOO
100
I I I
/
/
K /
MgC03= MgO +C02
- V
K= (COZ)1T
Data from K. K. Kelly /
/
/
1/
/
/
/
/
/
/
J
/
/
/
/
/
/
I I " I
1.0
0.1
1.6
.90 10-3
1.6 1.4 1.2 1.0 0.8 1.5 1.4 1.3 1.2 1.1 1.0
!(OK-l) lIT (OK-I)
T
Fig. 1-18
Fig. 1-17
-------
Cur.... 6It7bS6.A
-2.00
Cas04 ~ Cao + S02 + ~ 02
Kp=PS02~
-10.00
-4.00
-3.00
-5.00
~ -6.00
d"
...
0'" -1.00
II>
!:
~
-8.00
o Esso Data
-9.00
-12.00
2400
&XI
1000
1200 1400 1600 2000 2400 Of
600 100
loi
6
4
KI
CaO +H2S - CaS + H20
~ K'" (H2011 IH2S)
\
\
\\
~
\
~"\
\\
,\
. Esso '\
t--- . Squ Ire~ ~
. Con~ol
.. Es~o England
..
~ '\..
\
\,
\\
'\
I I I I I I I I
104
8
6
4
K
103
8
6
4
102
1.6 1.4 1.2 1.0 0.8
1/T (OK-I,
Fig. 1-19
61
0.6)( ur3
-------
:,,, 'd' ",:,.",tiU-II
"XI
!OJ
SOlI
1200 IdOO I(,IX). 2000 24oo"f
7lltJ
1000
,-,
-K-I
o -~CaS04 + H, = 1/4 Ca5 + H20
K = tHZOJ IIHZJ
102
g
101
g
i
. Esso
. Squires
- " Cansol
. Essa England
100
0.8
0.6XIO-3
1.4
1.2 1. 0
lIT IOK-I,
l.O
fig. 1-20
102
8
6
700
800
1200 1400 I(,IX) 2000 2dOO
(,IX)
1000
"f
10-1
8
6
4
f---I. K
~ Ca504 +co - CaO + C02 + 5°2
I---
'C02" 502'" 7
K = -,co;--
I 1
1---. Esso
~. [sso [nglaM ,
f
--I
r 7
/ f
_: --
===r.
-i-. I
I ,
I /
( I
I / , " I
;o'f
101
8
6
100
8
6
10-2
8
6
4
10-3
1.6
1.4
0.6X 10
1.0
~'OK-II
1.2
0.8
fig. H2
.102
8
6
1000
1200 1400 I(,IX)
2000 2400 "f
. 700
800
(,IX)
K
- Ca504 +H2 CaO + H20 +5°2
----= .
- IH201 '5°2'" J
_K=,~.
.
J
r
~.Esso
- . Esso Enaland
1 I I
I I J
IfA f
, I
J
/ J
J
I " ,I f I ,r ,
101
8
6
4
100
8
6
103
10-1
8
6
102
10-2
8
6
4
101
100
10-3
1.6 1.4 1.2 1.0 0.8
~ ,oK-I,
fig. HI
0.6 x 10-3
C"rv~ ""'''582-9
600
104
8
6
700 800
1000
1200 1400 1600
2000 25oo"f
103
I-- 1-1 -'1 -I 1 ~ -1-' -1-1. 1---':
1/4 Ca504 + CO - 1/4 Ca5 + C02
, K = IC021/1COI
\.
\
\
"
'\
'-- . Esso \
. Consol \
- ~ [sso [ngland
10.'"\.
\
I I , I I
103
8
6
102
102
8
101
100
101
0.8
0.6X 11)""3
1..6
1.4
1.2
1.0
liT 1'1<-1,
fig. H3
68
-------
101
8
6
4
000
1011 ~
urn
1200 1400 1600
2(1('01 2400 "F
K .-
~S04 + 1/4 CaS: CaO + S02
- I
- K=IS02'n I
r
-8 Esso
- 8 Consol
8 Esso England
A
'H
II
I
I I I 1 1 1 1 1
-!
100
8
6
4
10-1
8
6
4
2
10-2
8
6
10-3
8
6
10-4
1.6
1.4
1.2
1.0
0.8
0.6 x 10
!(OCI,
T
Fig. 1-24
IO-~
6
4
600
100
1200 1400 1600
2000 2400 "F
800
1000
10-4
8
6
4
I::=' Ca S + C02 = CaO + COS
PI
II
/ II
t=' 8 Esso
t-- 8 Esso England Ir
t-- 8 Consol If
I f
I
,
1/
r
I I 1 1 I I 1
-5
K 10 8
6
10-6
8
6
4
10-1
0.8.
0.6X 1IT"'3
1.6
L4
1.2 LO
! 1"1(-1,
T .
Fig. 1-26
10-3
8
6
4
600
100
800
1000 1200 1400 1600 2000 2400 "F
10-4
8
6
4
K -- I-'-
==- CaS + 1/2 C02: CaO + 1/2 CS2
.rS21 I
- K' TCOZT ;
1
~
- eEsso ./
~ 8 Esso England
7
7
/
1 1 1 7, 1 " I I
10-5
8
6
4
10-6
8
6
4
10-1
1.2
1.0
0.8
0.6' 1IT"'3
1.6
1.4
} ,oK-II
Fig. 1-25
tu,.~t f,l.l~S-8
I
10
8
6
4
000
100
800
1000
1200 1400 1600
2000 24000 F
10-2
8
6
4
~ I
CaS +H20 + C02 - CaC03 +H2S
IH2S1
\: K =,H201Ic02In
\\
\ \ -
\y ..
" ----
\ \.
,\
\\
f---8 Esso
I---
~. Squires \ \
~
\\
\'\.
'\
r 1 r , I I 1 \ 1
100
8
6
4
10-1
8
6
4
10-3
8
6
4
10-4
L6
1.4
1.2 LO
! 1"1(-11
T
0.8
0.6 x 1IT"'3
Fig. 1-27
"
-------
10-1
S
01(1
1200 1400 1600 2000 2400 'F
700 &XI
1000
K
10-3
8
6
--='==t=t K
H!S + C02 CDS + H20
- IC05. tHzOI ./
-- K- --- ---
~0211H2SI
/
/'
Y1 . F(so
,
, "
I
I I
1 1 1 1 1
1.6 1.4 -j
10-2
8
6
4
10-4
8
6
4
10-5
1.2
1.0
0.8
0.6 x 10
f jOK-l,
fig. 1-28
600
1000
1200 1400 1600
2000 2400 'F
700 800
100
. Esso
. Hougen et al
C02 +HZ ~ co + H20
- (H20' ICOI /
K = I'H2' IC02'
I
J
/
.;
/
I
1/
i 1 1
10-1
1.6
1.4
1.2
liT (")(-1,
f~.HO
1.0
0.8
[6XI0-3
10
r,.,r-,.. {A/~~'J'"
10-:
6
600
1100 1400 1600
10-3
8
6
700
HYIO
2000 2400 'F
800
K
=H25 + 2H20 - 502 + 3H2 4
=-: 15°2' (Hi-
~ /H2S1 ~ I J
1/ I
"
/ 1
8
4
-8[550 I
I I
I
I II
f
I I
I / ,/
x -j
10-:
6
IO~
8
6
4
10-8
8
6
4
0-5
10-9
8
6
10-0
8
6
4
10-10
8
6
4
10-7
8
6
4
10-\1
10-8
1.0
0.8
0.6 Hr
1.6
l.4
1.2
f (OK-l/
fig. 1-29
Cur". ~2588-B
100
8
6
600
700
1200 1400 1600 2000 2400 'F
800
1000
10
8
6
4
~=K =1=
~OHI2jS) - CaDIs) +H?O!gl
e--- K = {H20"
/
/ I
i
/
I J
I l . ICT - Dragert
.J . ICT - Johnston
I / Curran
/
1 I 1 1 I
1.6 1.4 1.2 1.0 0.8 0.6 x 10-3
1.0
8
6
4
0.1
8
6
4
0.01
8
6
4
0.001
1I71'K-l,
fig. HI
-------
(Uf...~ (}tZ'J/8-A
80
35
70
60 30
50
Q) -6H, Mbtu/lb- mole ~ 25 -6Hj' Mbtu lib - mole
o 0
E CaC03 521 E CO 47.5
, 40 CaO 273 C02 169.3
:e :e
-- CaS 198 COS 59.0
::J CaS~4 616 ~ 20 CS2 - -49. 6 -
~ ~ 30
- ::E H20 104.0
H2S 8.7
S02 127.7
20 15
10
10
00 400 800 1200 1600 2000 2400
T"F
Fig. 1-32-Heat content (enthalpy) above 77"F datum mbtu Ilb- mole 5
(Kubaschewski & Evans, Strassburger)
tllf'Jt; Ofi.'../':J-f'.
40
0600
800
1000
1200
1400
TOF
1600
1800
2000
2200
2400
Heat content (enthalpy) above 77"F datum mbtu lib - mole UANAF thermochemical data)
Fig. 1-33
-------
Minimization of NO Emissions
. x
Nitrogen oxides (NO ), predominantly nitric oxide (NO) with
x .
lesser quantities of N02' react photochemically in the atmosphere to
generate smog and noxious compounds which have undesirable effects on
plant and animal life.
These gases are, therefore, air pollutants whose
emission must be minimized.
NO can be formed in the combustion of coal and oil either by
x
fixation of atmospheric nitrogen or by the oxidation of the nitrogen
content 01 the fuel. At temperatures above 27000F the N2 of the air
reacts directly with excess 02 remaining after combustion to form NO
N + 0 + 2NO (1)
2 2+
(1-29)
In a conventional high-temperature combustion process, the quantity of
NO formed can usually be estimated from equations involving the kinetics
and the equilibrium of this fixation reaction. At temperatures below
24000F, however, the rate of the reaction forming NO from N2 and 02 is
very slow; the equilibrium quantity of NO which can be formed is also
small.
In a low-temperature fluidized bed combustion process, therefore,
NO is formed as a combustion product from nitrogen contained in the fuel.
No way has been developed to predict the percent of fuel
nitrogen converted to NO , but preliminary experiments indicate that it
x .
may be 20 to 40%. Concentrations of NO higher than .those predicted by
equilibrium of the fixation reaction can occur. Once formed, the NO
can decompose by either of two reactions
2NO + 2CO + N2 + 2C02
(1-30)
2NO + 2502 + N2 + 2503
(1-31)
which may occur in the gas phase or, more likely, on solid surfaces in
the combustor.
(1) ..
Free energy changes 'for the reaction can be obtained from the JANAF .
. Tables [46].
72
-------
NO
x
Emissions From Conventional Boilers
The concentration of NO in stack gases from convent-ional power
x
plants depends on the power output of
and the nitrogen content of the fuel.
in gas- and oil-fired plants [87, 92]
the plant, the design of the boiler,
Emission levels of 400 to 700 ppm
and of 500 to 1000 ppm in coal-fired
~lants [68] have been measured in installations. of 200 to 500 megawatts.
Techniques such as low excess air, two-stage combustion, and flue gas
recirculation can be used individually and in combination to reduce the
NO emissions by 50% or more. (1) Emissions of 42 ppm from a 315 MWplant
x
are predicted for boiler operation with flue gas recirculation and sub-
stoichiometric operation at the burners with overfire air. [88]
NO
x
Emissions From Fluidized Bed Combustors
Measured NO concentrations from pilot fluidized bed combustors
x
operating at l450-l7500F and 1 atm range from 250-1200 ppm depencH.ng on the
operating conditions and coal size [4, 32, 69, 70, 71, 75, 76]. Itl~ high-
est NO concentration occurred with pulverized coal. At l5000Fand elevated.
x
pressure (4-6 atm) , NO emissions decreased to 50-150 ppm, with higher
. x .
values corresponding to larger quantities of excess air.
It might be postulated that the temperature in the neighborhood
of a burning coal particle within a fluidized bed is sufficiently high,
above 2400oF, to promote the fixation reaction (1-29) above. There is no
indication, however, that such high temperatures are reached. Experiments
[4] indicate rather that NO is produced by oxidation of the nitrogen in the
. x .
fuel rather than by fixation. The effect of pressure and air quantity on
NO emissions can then be explained by the effect of these operating vari-
x
ables on the rate of equation (1-29) above.'
A variety of experiments have been carried out to determine the
effects of various operating conditions on NO emissions from fluidized
x
beds and to clarify the mechanisms involved in NO formation and reduction.
. x
The results of these experiments and possible interpretations are described
in the following sections.
(1) k f d .
Barto, W., A. R. Craw or. , A. R. Cunningham, H. J. Hall,.E. H. Manny,
and A. Skopp, "Systems Study of NOx Control Methods For Stationary
Sources," Final Report to National Air Pollution Control Administration,
GR-2-NOS-69, Contract No. PH-22-68-55, November 20, 1969.
73
-------
. Combustion Gas Composition.
Combustion gas composition has been
modified in two primary ways: argon has been substituted for nitrogen
'in the combustion air to discover the source of ' the N.in the NO
x
emissions; the air/fuel ratio has been varied to determine the effect
of excess oxygen on NO emissions.
. x
direct demonstration that the nitrogenous content of the coal is the
source of NO. [4] Argon, an inert gas, was substituted for the active
Argonne conduc ted a simple and,
component of nitrogen in the fluidizing gas.
Sin~e this substitution
did not significantly affect the NO concentration in the flue gas,
most of the NO obviously originated from the nitrogenous content of
x . .
the coal. Additional evidence was obtained by combustion of natural
gas, CH4' also conducted by Argonne National Laboratory [4']. In
these experiments, in which the only source of nitrogen was the
fluidizing air, the observed NO concentration in the flue gas was
, x
less than 100 ppm, a value predicted by the nitrogen fixation equi-
librium.
From the accumulated experimental evidence, there is no
doubt that the nitrogenous content of the. fuel plays a key role in
NO generation in fluidized bed combustion. The amount of nitrogenous
x. ,
components converted to NO may depend on the temperature of combusting
, '
particles and the rate of combustion, which in turn depends on opera-
ting conditions such as bed temperature, excess air, f1uid~zing
velocity, 502 sorbent injection, and coal particle size. These same
factors also may influence the subsequent reduction of the NO
x
formed to N2 and combustion products.
Emission of NO from the fluidized bed combustor was found
to increase with 02 content in the flue gas as' determined by the
excess air rate. The correlations obtained by PER in their FBC and
FBM test series [75] are presented in Figure 1-34.
The lower NO emissions at low excess air are believed to
result from the reduction of NO by CO, which is more abundant during
operation with low excess air. Substoichiometric operation gave even
74
-------
500
'-I
1I1
E
c..
c.. 400
-
V)
ro
Co:)
Q,)
:J
L.L..
c:
J--I
c: 300
0
--
ro
!....
-- Curve
c:
Q,) No:)
u
c:
0 1
u 200
o 2
Z
Bed
Size
Curve 643508-A
1
2-
Bed
Temperature
(OF)
Bed Fluidizing
Height Velocity
(i nches) (ft/see)
Average Coal
Size
Reference:
100
0.0
Bed
Material
FBC 1 211 X 1611
FBM 2011 X 7211
. 1750
1770-1840
Sintered Ash
Sinteredl\sh
8
13
-1/4",Ohio !f8,unwashed
-1/411 ,Oh i 0 #8, washed
14
1 2-14
(2)
1.0
2.0 3.0
Oxygen Concentration In Flue Gas, %
4.0
5.0
Figo 1-34-Effect of oxygen concentration on NO generation
-------
lower' NO
x
emissions'.
emissions: at 68% stoichiometric air Esso found zero NO
x
Esso was unable to find any correlation with' excess air,
,how.ever, because of scattered data [371.
Bed Temperature.
Both Esso and; Argonne' observed a bedt'empera-
ture effect on NO emissions,. The experimental d'ata are- summar:Lzed' in'
x .
Figure 1-35. Decreasing bed' temperature secreased NO emiss-ions. It, is;
, x '
not clear whether this temperature affects primarily the conversio~ ox'
fuel nitrogen to NOx or the' reduction of the NOx after its formation.
PER found no correlation between NO concentration and bed
temperature.
This contradictory observation, is not, surprising, because
the meas:ured' NO concentrations are well above those predicted by thermo-
dynamic' equilibrium.
Coal Size. Coal size plots of available data (Figures. 1-34 and
1:-35) indicate that' the NO emissions from fIuidtzed' bed combustion are,
x
d'ependent upon the' coal particle size. At a, bed' temperature of 1600~F' and
. with limestone or d'olomite as bed material, PER []5", 761 reported' a; NO con-
centration in flue gas of 300-360 ppm' using' - 1/"4" coal particles at 8-10
ft/sec fluidizing velocity; Argonne [4T obtained 380-500 ppm using - 14
mesh coal particles at 3 ft/sec;. and Esso: [32.] measured' 700-740, ppm' using
pulverized coal particles of size '\..200\1 at 3 n/sec'. These results may
be somewhat confounded: because Esso used coal with a higher nitrogen: content
(1.4 wt %) than Argonne's (].l wt %).
But observations seem to confirm
that smaller coal feed' particles' cause an increase in NO
x
emissions "
Superficial Gas Velocity. At a given bed temperature, Esso,found'
no effect of superficial fluidizing velocity' on NO' " while, Argonne showed that
x: .
76
-------
the degree of reduction in NO concentration was inversely proportional
to fluidizing velocity [4].
The conflicting results may be due to the
fact that different coal sizes and reactor sizes were used.
If the
degree of NO decomposition controls the final level of NO concentration
x
reduction to be
in the flue gas, it is reasonable to expect the NO
x
directly proportional to the gas residence time and thus inversely
proportional to the superficial fluidizing velocity.
Bed Material. Esso observed consistently lower NO emissions from
x
Ca804 than from alundum beds (see Figure 1-35). This fact was presented as.
evidence that a heterogeneous reaction between NO and CO might occur over CaS04'
but not over the inert alundum. Argonne also measured low NO emissions
x
from beds of limestone. A reduction of NO concentration in the flue
x
gas up to 40% of its original value simultaneously with the reduction of
S02 concentration upon addition of 802 sorbents to the fluidized bed was
observed [4]. However, PER concluded that in general, NO emissions were
. x
not affected by addition of sulfur control sorbents.
The use of sorbents as bed material also decreases the amount
of S02 in the fluidizing gases. In turn, NOxemissions may be increased
due to the effect of decreased r02 concentrations on the NO reduction
reaction (1-31) above. Both Esso, operating at 1 atm, and the NCB [69, 70,
71] operating at 3.5 atm, reported that NO emissions decreased with
x
increases in 802 concentration in flue gas. In batch operation carried
out by Esso, much higher NO (1000-1200 ppm) was observed with the fresh
. x.
stone when S02 removal was essentially complete. In the continuous long
duration operation at NCB, the NO concentration was lowest (37 ppm) at
the beginning of the test when the S02 concentration was at a maximum.
For most of the test the concentration was 100-120 ppm, a value substan-
tially lower than that obtained in atmospheric operation.
But it also appears that the sorbent catalyzes the reduction
reaction (1-31). At l630oF, Esso found that NO and S02did not react to an.
77
-------
Curve 64~521-B
80Q
~ . _. .-. . -.
..--"M ChC~ C"" 1
.
.
. .-'.-' ''"--..'- ".
700
"-,---- p" ,.- :,_. '-.:.- ~
E
c..
q,
-
>-.
~
'?
, VI
10
<.J?
Q.)
:::\'
!"I"
c:
~
~OO
500
C!
.2
:-J
()O
'-'
'10
L.;
~
C
Q.)
o
g 400'
<".? . ','
o
?
"..,'-, " --> 4
.-'. .,.-...
~.
,,'. ;--,'
.----~;/ Curve
.y'-- N'o
/ ~-
1
2.
3
4
c()~! 02 in
Siz~ f!~~ Gg~
3QO IJ - 4%-' ,
~OO~ 4%
""14 rn~~h ~:Q~?:~ro
""'~4 m~~h 3:6~4:D"!o
R~gc;tQr ,
Si~~
'31! d i~
3." gia.
6" di~
~" di~
~~q
Mat~ri~!
AI!Jn~LJm
CP~Q4
!-M s- 1359
PMT' n~7 '
f'wiqi~ing
Wlo(:Uy (ftl ~~G)
, -'''6 ,. ,
~
2: ~ =- 3. 1
?: ~~~: 1
?QO -
R~fer~n@~;
20Q , "
~4QO
. .- -_.- - - "..-. ~..
.. .' ~'-' - .,. .-- - -
- - - ._. - . -- -.-
-
19QQ
1600 1700
...~~~ T~mperqtLJr~, !>f""
~~QQ
umQ
Fig. t- 35 - Ternperatw re effec:t on NO g~l1eration
-------
appreciable extent in gas phase with gas residence time of about one
second. However, when S02 was injected into a
sulfated stone with N2 as a transport gas, NOx
Ppssible mechanisms for this process have been
but these mechanisms have not been confirmed.
fluid bed of partially
emissions were reduced.
postulated by Esso [32],
Conclusions
The experimental observations on NO. emissions from fluidized
x
beds indicate that NO is formed from fuel nitrogen in the burning process.
This NO can subsequently be reduced in the bed by CO and/or S02 according
to reactions (1~30 and 1-31), which may be catalyzed by active surface such
as that provided by the sorbent.
Emissions of NO from fluidized bed
. x
combustors are complex functions of coal composition and size, air/fuel
ratio, bed temperature and pressure, gas velocity, and sorbent usage.
Three methods for minimizing NO
x
emissions from fluidized bed
combustors can be proposed:
. Modifying operating conditions.
Decreasing excess air and
temperature and increasing pressure and coal particle size
can decrease NO emissions appreciably. Pressurized operation
x
appears particularly effective, probably because of the
increased rate of the NO reduction reactions (1-30) and (1-31)
and the greater probability of CO concentrations in the
vicinity of coal feed points.
. Carrying out two-stage co~bustion. Stoichiometric and sub-
stoichiometric combustion of coal in the fluidized beds has
been shown to reduce the NO emissions tremendously.
x
A
two-stage combustion process where the first stage is operated
close to reducing condition with secondary air injection to
complete- combustion reduces NO emissions. Rearrangement of
x
coal feed and air distribution points can probably achieve
the same effect.
79
-------
. Using decomposition catalysts or sorbents for NO in the
'fluidized beds:. Argonne has injected metal o~ides, Al203'
ZrOZ' and C0304' into fluidized beds. AlZ03 and ZrOZdid'
not changeNOx emissions; when C0304 was tested, NO ,
concentration in the flue gas increased. Presumably there
is a ,catalyst which will effect a reduction in NO
x
emi'ssions.
Catalyst poisoning by sulfur oxides in the bed may be a
serious problem.
Some sorbent ,might be found which is
effective for NOx as limestone and dolomite are for 501'.
However~ most nitrates, nitrites, and nitrides are not
stable at the operating conditions of the fluidized ~bed.
80
-------
Combustion Efficiency
To be able to utilize fully the high heat transfer rate in a
fluidized bed and to take advantage of the air pollution control
capability of fluidized bed boilers for future power generation, the
combustion efficiency of the fluidized bed boilers should be at least
as attractive as that of conventional p.f. boilers.
Experimental
evidence shows that the combustion of coal in a fluidized bed is a complex
process which depends on many operational and design variables.
In
addition to the usual boiler losses (such as sensible heat loss from
the flue gas and moisture in the fuel), the major combustion losses from
a fluidized bed combustor are elutriated solid carbon and unburned
combustibles such as carbon monoxide (CO) and hydrocarbons.
To optimize
the combustion efficiency, a careful selection of ,operating conditions --
bed temperature, excess air, fluidizing velocity, size distribution of
fuel -- and design variables -- freeboard height, fuel feed point
separation -- is essential.
The unburned carbon loss from fluidized bed combustion occurs
entirely through elutriation.
Thus, the amount of the carbon loss may
be expected to be inversely proportional to the percentage of excess
air, bed temperature, and the residence time of the coal particles in the
bed, which in turn depends on the fluidizing velocity.
In most cases the
volatile content of the carry-over carbon is about 1.5% to 5.5% [42,55].
Consequently, the elutriated unburned carbon can be treated as solid
carbon with heating value of ~14550 Btu/lb.
When the coal is screw fed or the fuel feed
points are widely
spaced, experimental evidence shows variation in the flue gas compositions
over the'cross-sectional area of the fluidized bed.
The CO concentration
is usually higher over the coal feed point than at any other location, and
combustion of combustibles above the bed is observed by the fact that
the flue gas temperature measured five to eight inches above the bed surface
shows a temperature up to 200°C higher than the average bed temperature [54].
Measurements also indicate that combustion of CO and carbon occurs up to
81
-------
two feet above the bed [17], and much of the CO in the gases leaving the
bed is formed by reduction of C02 by fine carbon particles. Thus, the
interpretation of data for CO loss will depend on the distance above the
bed at which the gas samples were taken. Methane (CH4) did not appear
in the gas signifi.cantly until the excess air was below 5% [75,101].
existence of other higher hydrocarbons is probably negligible because
The
they burn more readily than methane.
Aside from the cross-sectional variations caused by the mal~
distribution of fuel and air, the losses from elutriated solid carbon
and unburned combustibles generally relate to the other independent
operational and design variables as described below.
The quantitative
results and the qual~tative effects are summarized in Tables 1-2 and 1-3.
Fines Recycle
Ash withdrawn from the bed contains < 0.1% carbon [78,70].
Thus, most of the carbon loss occurs by elutriation of fine particles
from the bed.
Depending on the operating conditions and the particle
siie distribution of the fuel, the carbon carry-over from the bed without
recycle can be very large [70,77,101], as shown in Figures 1-36 and 1-37.
The type of bed material does not seem to affect carbon loss.
Recycle of
the fine particles back to a separate bed running at a higher temperature
and higher excess.ai~ (as done by PER), or to the original bed (as done
by NCB and the U. S. Bureau of Mines), are the most effective ways to
reduce the carbon loss.
The experimental data obtained in a separate carbon burn-up cell
for the carbon burn-up efficiency of fly ash containing 23% to 65% carbon
have been correlated by regressional analysis with operating variables
such as bed temperature, air rate, static bed depth, carbon rate, and
inert rate [27].
The results indicate that combustion efficiency is
improved by increasing bed temperature, air rate, and bed depth, and by
decreasing fuel feed.
The selected experimental data are plotted in
Figure 1-38, and the curves shown in Figure 1-38 are reasonable approxima-
.tions for design purposes.
Integrated operation of a combustor and a carbon
, .
82
-------
TABLE 1-2
SUMMARY OF THE R,:SULTS OF COMBUSTION LOSSES FROM EXPERIMF.NTS
CONDUCTED BY OAP CONTRACTORS AND NCB LABORAIIJRIES AT
ATruSPHERIC PRESSURE
-----:f- -
BED
TEMPERATURE
ORGANIZATION PARTICLE SIZE - (OF)
PER 12" x 16" Limestone Bed:-l0+20 mesh 1510-1630 No No 0_4~0.5% at
BCR 1359 Coal:-l/4 x 0 51 excess air; (0.1% [77]
negligible at
excess air ::1'10%
PER-Carbon 12" x 16" Coal Ash 600 - 1400 1750-2000 10-65 10-15 No No Comb~stion efficiency ranges from (27]
Burn-up Cell 65% to 90%
Tests (Burn up Cell by itself)
PER- Inte- FBM 20" x 6' Coal Ash Bed :-8+16 mesh FBM: 1590- 5-27 Recycle No 1.6-6.4 FB!I: 0-370 ppm [76]
prated FBl11 CBe 20" x 20" Coal: -114 x 0 1700 overall to CBC CBC: 0-600 ppm
CBe Tests CBC: 1800-
2050
USBM 18" diameter refractory Bed :-16+48 mesh 1450-1650 8.6-57.0 2.4-3.4 Recycle Yes ~1.3 Carbon burn-up efficiency [15]
Coal: 1/8" x 0 to or1g1- ranges from 86% to 99.8%
oal bed
ESSO )" diameter limestone Bed: 450..-,930,,", 1500-1800 5-30 2.6 .No No 3-13 450-750 ppm [31]
BCR 1359 Coal: 200""
Argonne 6" diameter alumina Bed: 30 mesh 1600 10-20 No No 3-7 [3]
Coal: -14 mesh
BCURA 12" diameter refractory Bed:-l/811 x 30 1380-1800 0-90 7.2-10.6 No Water _20 [54]
mesh; Coal: -1/411+ cooling
~ 1116"; 1/8" -0 jackets
BCURA 2711 diameter refractory Bed:-1/4"+ 30 1300-1830 0-90 7-14 Partial Yes 4-14 0.4% at 6% [42,55]
Or ash mesh; Coal:-l/4" recycle excess ai r
x 0; -1/8" x 0 to origi- 0.1% at 90%
nal bed excess air
NCB-CRE 6" diameter ash Bed:-lO mesh 1300-1470 0-55 1-3 With & cool- ~7% 0.2-0.9% at [7,8,18]
Coal: -10 mesh without ing without 1300. F
recycle coil recycle (,0.3% at
back to ~% 1470. F
original with
bed recycle
NCB-CRE 12" diameter ash Coal :-1/16" 1560 4-6 3 Partial
recycle Yes 6-7% 0.07-0.09 in [71]
flue gas
NCB-CRE 3 f t square ash Coal: -10 mesh 1300-1560 +30 to -30 2-4 No No 10-25 ~0.2 [70,101]
NCB-BC~ (a) 48" x 24" ash Bed: -1/16" 1470 10-35 Recycle Yes 1.3 ~0.05% [69J
Coal: -1/16" to origi-
oal bed
NOTE:
PER - Pope, Evans, and Robbins, Consulting Engineers, Alexandria,
USBM - U.S. Bureau of Mines at Morgantown, West Virginia
ESSO - ESSO Research and Engineering Company, Linden, New Jersey
Argonne B Argonne Na t ional Labora tory
BCURA - British Coal Utilization Research Association, England
NCB-CRE = National Coal Board Research Establishment, England
(a) Operating pressure at 3.5 atm.
Virginia
-------
VARIABLE
Fines recycle
Excess air
Fluidizing
velocity
Bed temp.
):)
"'"
Freeboard
height
Coal feed point
separation
TABLE 1-3 EFFECT OF PRIMARY VARIABLES ON COMBUSTION
EFFICIENCY AND THE MECHANISMS INVOLVED
CHANGE
LOSS IN CO
LOSS IN SOLID CARBON
Increase
Increase
Increase
Increase
Decrease
Decrease
Increase
Negligible effect
at l300°F - 1550°F
Decrease
Increase
Increase (for.
internal recycles)
Decrease
Negligible Effect
No effect at all in
large unit; large
effect at 1300°F -
1470°F in small unit
No effect
Increase
I
MECHANISM
Further combustion;
reduction of C02 by
carbon fines
Reaction equilibrium
and kinetics
Elutriation
Slugging in small units;
effect of fine recycles
overshadow other effects;
C02-solid carbon reaction
kinetics
Entrainment and pneumatic
transport characteristics
Concentration gradient of
coal in the bed
-------
"C
C1.>
.2!
c:
~ 20
'"
u
'0
...
"C
C1.>
co
E 15
~
~
o
.....
c:
.8
~ 10
u
:E
'0
VI
15
Curve 643503-A
I I
.
. .
. .
.
... -
.
"C
'"
.2!
.
c:
.8
~1O-
<0
E
'0
...
V>
V>
o
.....
c:
~ 5 -
'"
<.)
:E
<5
VI
12" X 16" Combustor
. 15% Excess Air
... 25% Excess Air
Velocity 8-12 It/see
3% S Coal in 1359 limestone Bed
Reference: PER Monthly Rept. .(76)
Note: Data taken alter at least two hours
operation. Carbon loss is much higher
I in first two hours qf operation.
1550 1600
Bed Temperatu re, of
. -
o
1500
1650
Fig. 1-36-Solid carbon loss from bed without recycle
30
25
6" Combustor (ash bed)
14 70°F
&
&,
3 It Square Combustor (ash bed)
12900F 1470°F 156O°F
Bed Temperature
Fluidizing
Velocity (ft/sec)
cv
rn
2
3
4
(l)
4
[]
References: Baileyet aI., (20) Williams. (23)
NCB rept. gives carbon loss of 3-1270
for a 6" combustor operating at 14700F
and 3 It/sec fluidizing velocity. 091
~
o
-40
50
-30
Fig. 1-37-Solid carbon loss from bed without recycle
85
-------
Curve 643519- B
100
A .
-
-
A
.
. 80
~
~
u
c::
.~
u ~
-
-
LLJ
c::
0
- ~
VI
00 ::J
a- .c .
E
o.
u
60
Superf icial
Excess Air % Carbon Content (%) Velocity (ft/sec)
. 20 - 45 30 - 60 10-15"
6 10 - 20 30 - 60 10 - 15
A 60 - 65 55 - 60 10 - 15
References: Ehrlich, (26) PER Monthly Repts. (25)
40
1700 1800 1900 2000 2100
Bed Temperature, of
Fig. 1-38-Combustion efficiency in carbon burn up cell (ash bed)
-------
burn-up cell was also tried by PER in their integrated FBM/CBC system [76].
The observed combustible losses, which included hydrocarbon and elutriated
carbon, ranged from 1. 6% to 6.4%, of which up to 10% was due to hydro-
carbon loss.
The high solid carbon loss was due to the short freeboard
'in the carbon burn-up cell necessitated by the limitations of the FBM
design.
The overall solid carbon loss with internal recycle back to
the same bed is shown in Figure 1-39.
NCB [71] also reported a carbon
loss of 6% to 7% at excess air of 6% to 4% respectively in a l2-inch'
diameter combustor operating at 3 ft/sec with recycle.
A reduction of
> 60% of carbon loss from that in Figures 1-36 and 1-37 is realizable.
Experiments performed in a six-inch combustor indicate that
recycle of fines back to the original b~d tends to increase CO concentra-
tion in the gas due to reduction of C02 by carbon fines [17]. The results
from a six-inch combustor with internal recycle and a three-foot square
unit without recycle are compared in Figure 1-40.
of recycle effect on the hydrocarbon losses.
There is no indication
Excess Air
Combustion in the fluidized bed is governed more by the excess
air than by the ~luidizing velocity or bed temperature in the range
examined.
However, appreciable combustible losses do not occur until
the stoichiometric conditions have been reached.
Increasing excess air
tends to decrease 'the overall carbon content in the e1utriated solids
and the coal concentration in the bed as observed in a six-inch combustor
[8], a 27-inch experimental rig [42], and a three-foot square combustor
[101]. As a result, the overall solid carbon loss also decreases
(See Figures 1-37 and 1-39).
Concentration of CO in the gas also tends to decrease with
increase in excess air, as shown by experiments conducted in a six-inch
combustor [7,8,18], (Figure 1-40). Up to 3% CO was observed to be present
at bed surface, but most of it was burned within two feet above the ,bed
87
-------
"C
Q.)
Q.)
-
c:
0
..c
...... 1.0
ro
u
-
0
~
"C
Q.)
cc
00 E
00 0
......
-
VI
VI
a
-I
c:
.8
......
ro 0.5
u
"C
0
V1
1.5
Curve UI3~23-[j
Curve
Unit Size
Bed Temperature
Velocity
Coal
1
6" diameter
1470°F
2 ftl sec
3 coals with ash
conte nt 10% - 25%
Data from a pressu r ized combustor
/(48" x 24") operating at 3.5 atm,
A 1470oF, 20% excess air, and.
~ 2 ft/ sec. (long du ration ru n:
100 h r )
Bed Material
ash
2-.
18" diameter
1450°F-1650°F
2.5- 3 ftl sec
6 coals with 2
widely different
characteristics
ref ractory
.
.
.
.
Data Spread
..
References: Baileyet. aI., (22)
Cooke et. aI., (271
Coates & Rice, (18)
NCB Rept. I (38)
o
-10
o
10
40 .
20
30
Excess Air, %
Fig. 1'"39-Solid carbon loss from bed with internal recycles
-------
0.3
. 0.2
o
>
'
'"
-
'0
c::
a
u
0.1
o
-10
1
2
3
o
.
.
Curve
~~i.!.~~"
6" Diam~tcr
Curve 643524-B
2
3 fl. Square
3 - . 4.- .
12" Diameler 48"x 24" \
(long duration run,500 hr) (lon9 duration run,100 hr)
12900F - 15600F
2 - 4 fl/see
No ReC'fcle;
One Coa I vii lh
Ash Conlent -30/
Ash
1 560" F
3 ft/sec
Wilh Recycle. U.S.
Humph rey No.7
Ash
Bed Tcmpcfature
~ocily
Operal i on
14700F
2 ft/sec
Inlernal Recycle;
Three Coals ~ilh
Ash Contenl 10-25/
Ash
Baileyetal.(7). (8)
Williali1set al. (18). (101)
NCB rept. (691. (70), (]1)
.
30
Fig. HIO-CO loss from bed with or without internal recycles
Bed t1alerial
. .- - -
References:
4
10
Excess Air. %
14700F & 3.5 atm
2 ft/sec
With Recycle; Welbeck Coal
& U.S. Humphrey No.7
Ash
40
-------
. .
if enough excess air was present. . Measurements taken at two feet above
the bed in a 27-inch diameter combustor burning 1/4 inx 0 untreated coal
with partial recycle show CO concentrations of 0.4% at 6% excess air and
0.1% at 90% excess air, a value almost three times larger than that in
the six-inch combustor [42]. PER [75] also reported a CO loss of 0.5%
and 0.4% at 5% excess air from their experimentscar-:;:ied out in 12 in
x 16 in and 18 in x 6 ft combustors,respectivelY,at 8 to 14 ft/sec without
recycle; CO loss was negligible at excess air greate:r than 10%.
However,
results. from the three-foot square unit measured at secondary cyclone
and without recycle show much less CO'loss, as shown in Figure 1-40.
The carbon and CO losses from a 48 in x 24 in bed operating
under 3.5 atm pressure are also presented iri Figures 1-39 and 1-40 for.
comparison.
Operating pressure does not seem to have any significant
effect on either carbon or CO losses for similar operating conditions.
Similarly, up to 0.8% CH4(which existed at bed surface) was
completely burned within one foot of the bed surface when excess oxygen
was present.
The evidence of combustion taking place in the freeboard was
indicated by progressively increasing concentration of C02 and decreasing
concentration of 02 in the freeboard.
Fluidizing Velocity
For a fixed particle size distribution of the bed material, with
other operating conditions 'constant, increasing the fluidizing velocity
tends to increase .the carbon carry-over [42] (Figure 1-41), based on the
terminal velocity consideration.
A partial reduction of the carbon loss
can thus be achieved by reducing the fluidizing velocity.
The effect
of velocity in a six-inch combustor [7] is also presented in Figures 1-42
and 1-43.
Although the results from the six-inch combustor show a slight
dependence of flue gas CO 90ncentration on fluidizing velocity with fines
recycle, the data from the three-foo't- square unit operated without recycle
indicate no effect of fluidizing velocity on either the concentration of
90
-------
-,:j 20
Q)
.-
-
Co
Co
::J
V)
--
n:J
Q)
.c.
-
0
~
~
V) 10
V)
\0 0
t-' ....J
c:
.8
~
n:J
U.
-,:j
Q)
--
n:J
.-
~
--
::J
L.LJ 0
7
Curve 643514-A
Unit Size
Bed Temperature
Recycle Rate
Coal.
27" D ia meter
--.J lS000F
1/2 Recycle to 5/6 Recycle
Three c~aJ it Swelling Nu moor 1,3
ard5l/~) .
Reference: Hickman etal (42)
Data Sp read
8
9
10 11 12
Superficial Velocity, ft/sec
13
15
14
Figo 1-41-Effect of flu idiz i ng velocity on elutriated sol id carbon loss
-------
1.0
"0
>
>-
.c
8 0.8
Ii'-
",'
to
<.:)
'0 06
c:: .
o
u
0.4
0.2
Curve 643515-A
6" Diameter Combustor
1.0
cv
Q)
.~
Superficial Velocity
2 It! sec
3 It! sec
1 It! sec
2 It! sec
Bed Temperature
1470°F
1470°F
l290°F
l290°F
'C
'"
~
c::
o
.c
3 0.8
~ & , Reference: Bailey et ai, (7)
&~ ~.
~
'0
Ii'-
'"
:3 0.6
c::
.8
"-
to
U
'C
:.s 0.4
VI
3- c#:>
~
ffi
0.2
o
-10
~
20 30 40 50
Excess Air. %
60
Fig. 1-42 -Solid carbon loss with fines recycles
Curve 643520-6
6" Diameter Combustor
Superficial Velocity Bed Temperature
CV 2 It! sec l470°F
Q) 3 It!sec 1470°F
& 1 It! sec 12900F
it 2 ft/sec 12900 F
~
Reference: Bailey et af.(7J
-
£
ffi
&
CR.. ~
......
........... ............ -... Lt
--
. --Q) (2)
Q)~
Q)
o
-10
70
Fig. 1-43-CO loss with fines recycle
92
-------
CO or hydrocarbons [101].
Here the slugging nature of the small unit may
have something to do with the discrepancy; the fines recycle may prove
to be a more important factor.
Bed Temperature
Within the experimental range of 1290°F to 1560°F [101], bed
temperature was found to have little effect on carbon loss
and no effect
at all on CO concentration and hydrocarbon loss in the three-foot square
unit (see Figure 1-36). However, in the smaller unit (six-inch diameter)
[17], the amount of CO leaving the bed is'smal1 at bed temperatures above
1470°F and appreciable at 1290°F. Again, the discrepancy here may be due,
to a difference in the operation of the units: the former was operated
without recycle and the latter with recycle. ,It may also result from
differences in the kinetics of the C02-solid carbon reaction at these two
temperature levels [17].
Particle Size of Fuel
Coal for fluidized bed combustion can be of coarse particles, as
is generally the case, or of pulverized powder, as used primarily by Esso
in their bench studies [31]. The combustion efficiency for these two types
of coal feed will depend very much on the operating conditions.
Pulver-
ized coal particles can easily be e1utriated; their rate of combustion
, is much higher than that of coarse particles because of their much.higher
surface/weight ratio. The comparison of coal combustion efficiency between
these two types of operation is difficult because of a lack of common
basis.
However, a comparison of fluidized bed compactness can be made on
the basis of the heat release rate per unit bed area.
Specifically, 13%
carbon loss was reported [31] for pulverized coal feed at a superficial
velocity of 6 ft/sec, 1600°F bed temperature, and 10% excess air. The
same amount of carbon loss was attained at about 12 ft/sec superficial
velocity for the coarse coal feed of top size 1/4-in particles at the
same bed temperature and excess air (Figure 1-36). Assuming that the
93
-------
carbon burn-up cell will be. equally efficient in both cases, twice the
amount of the primary fluidized bed area will be required for the
pulverized unit to achieve the same overall heat release rate.
Freeboard Height
In comparing the results of solid carbon' loss from different
experimental units, the freeboard height of each unit must be distinguished
from its respective transport disengaging height (TDH).' When gas bubbles
erupt at the surface of the bed, solid particles are splashed far into
the freeboard. If the gas exit is located too close to the bed surface,
the solid particles entrained in the gas stream usually are high. Hence,
it is desirable to locate the gas exit beyond the so-called TDH, where the
entrainment of solid particles becomes constant and depends solely on
pneumatic. transport characteristics. The empirical relationship for
determining TDH for vessels of different diameters operating at different
velocities is given by Kunii and Levenspiel [60]. Clearly, comparing the
results of solid carbon loss from different experimental units without
regard to th~ir freeboard height can lead to error.
Unfortunately,
however,the correlation of solid carbon loss with freeboard height and
TDH has not yet been established.
Coal Feed Point Separation
It is predicted that a wide spacing between coal feed points
will cause variation in coal concentration across the fluidized.bed,,
resulting in a decrease in the combustion efficiency within the bed due
to high combustible losses [44].
A mathematical 'model coupled with some
experimental data shows no net increase in the elutriation rate for
coal feed spacings up to three feet apart.
For spacings greater than that,
the elutriation rate of solid carbon starts to increase, and reducing
conditions start to develop around the area' immediately next to the coal
feed points.
Thus the boiler tubes will have to withstand both reducing
and oxidizing conditions in the same'bed.
94
-------
The value of the coal diffusion coefficient is proportional 'to
fluidizing velocity. Thus, for higher operating velocity the op~imum
coal feed point spacing may actually be greater than three feet. The
value of the diffusion coefficient is also smaller when tubes are inserted
in the bed. The optimum coal feed point spacing will therefore have to
be determined individually for each boiler at specific operating conditions
2
and with specific designs. Without actual experimental data, lO.ft bed
area per coal feed point is recommended as a reasonable and conservative
choice [44].
Other Secondary Variables
. Concentration of coal in the bed.
Increasing the concentra~
tion of coal in the bed or increasing the coal feeding rate
tends to increase the elutriation rate of solid carbon and
decrease the combustion efficiency.
, Bed height.
Increasing the bed height will increase the
combustion efficiency by decreasing solid carbon loss and CO
loss because of increased residence time of coal particles
in the bed.
. Heat transfer surface arrangement above the bed.
Too much
heat transfer surface immediately above the bed, to serve as
I
a splash screen, tends to quench the flue gas too fast and
prevent .further combustion of fines and CO above the bed.
This will result in higher carbon and CO losses.
Careftil
design in this respect is necessary, and more experimental
data are required to provide adequate design information.
. Unit size.
Experimental data obtained from smaller units
tend to give higher carbon loss due to slugging, while the
coal concentration gradient in the bed, which is important
in larger units, is relatively unimportant in smaller units.
The manifestation of the experimental data must be carefully
analyzed for design purposes.
95
-------
Summary
Combustion efficiency for fluidized bed combustion boilers
depends on operating' conditions, e.g., bed temperature, excess air,'
fluidizing velocity, size distribution of fuel; and design variables, e.g.,
bed depth, freeboard height, fuel feed point separation. The major
combustion losses from a fluidized bed combustor are elutriated 'solid
carbon and unburned combustibles,such as carbon monoxide (CO) and
hydrocarbons. Recycle of fines back to the original bed or to a separate
carbon burn-up cell operating at higher temperature and excess air is
essential for efficient operation of fluidized bed combustors.
Carbon monoxide and hydrocarbon losses are most sensitive to
the excess air used for combustion.
However, appreciable combustible
losses do not occur until the stoichiometric conditions have been
reached.
Solid carbon loss occurs primarily by elutriation of fine
particles from the bed, and this is most sensitive to changes in
fluidizing velocity. Presumably the selection of particle size and its
distribution will also be important. Operating pressure and temperature
do not seem to have a pronounced effect on carbon and CO losses within
the ranges tested.
Design parameters such as freeboard height and coal feed
point separation are also important in improving overall combustion
efficiency.
Unfortunately, few data are available, especially for cases
where tubes are inserted in and above the bed.
Most of the studies concentrated on coal combustion in small
units at atmospheric pressure with low velocities and shallow beds.
Although the PER experiments were run at higher velocities (12 to 14 ft/sec),
there were no immersed tubes in the bed.
Extrapolation of available
data for design purposes is possible but risky.
More data collected
under pressurized operation with high velocities (up to 15 ft/sec) and
deep beds (up to 15 ft) are needed.
96
-------
The results of experiments conducted by OAF contractors
and NCB laboratories were summarized in Table 1-2.
The effect of
changing primary variables on combustion efficiency and the mechanisms
involved were presented in Table 1-3.
The assumptions on combustible
losses, derived from evaluation of the available data presented in Table
1-2, for designing industrial, atmospheric pressure, and pressurized
fluidized bed boilers are presented in Table 1-4.
97
-------
TABLE 1-4
COMBUSTIBLE LOSSES ASSUMPTIONS FOR DESIGN OF
THE INDUSTRIAL, ATMOSPHERIC PRESSURE, AND
PRESSURIZED FLUIDIZED BED BOILERS
INDUSTRIAL BOILER
PRESSURIZED
UTILITY BOILER
ATMOSPHERIC PRESSURE
UTILITY BOILER
SOLID CARBON LOSS
From primary bed
13%
From CBC
90% carbon burn-up
efficiency
\D
co
CO LOSS
Included in carbon
1088
HYDROCARBON LOSS
Negligible
10 ft2 per feed
point
COAL FEED"POINT SEPARATION
13%
85% carbon burn-up
efficiency
Included in carbon
1088
Negligible
10 ft2 per feed
point
6%
90% carbon burn-up
efficiency
<0.5% in flue gas
Negligible
" 10 ft2 per feed
point
-------
Heat Transfer
One of the major advantages of. fluidized bed boilers is the
high overall heat transfer coefficient between the bed and the immersed
heat transfer surfaces, which permits a boiler design of higher volumetric
heat release rates. A heat release rate on the order of 500,000 Btu/hr/ft3
furnace volume or higher is possible in fluidized bed boilers, compared to
~20,000 Btu/hr/ft3 in conventional boilers. Consequently, a more economical
boiler with smaller boiler size and less tube surface can be constructed
using fluidized bed combustion technology.
Pressurized operation, where
deeper beds are feasible, offers even further size reduction and economic
\
incentives.'
The literature contains a wealth of data on heat transfer between
fluidized beds and immersed surfaces.
Valentik and Williams [99], of NCB,
have reviewed the recent experimental data from the literature and from the
work done by NCB laboratories. The Massachusetts Institute of Technology
(MIT), under contract to OAF, has compiled published investigations with
a listing and discussion of the various theoretical, empirical, and semi-
empirical expressions for predicting heat transfer coefficients in
fluidized beds [13].
Numerous other articles and books review heat
transfer in fluidized beds.
This discussion is a review of the experimental
studies performed by OAF contractors and by NCB laboratories in England
which are directly related to fluidized bed combustion. Wherever adequate
data from other sources are available, they are also presented for comparison.
A reasonable assumption of the heat transfer coefficient is
essential for design purposes. However, the reported heat transfer
coefficients range from as low as 30 Btu/ft2/hr/oF to as high as 170
2 .
Btu/ft /hr/oF, because of varying operating conditions and design variables.
In the case of high-temperature combustion in fluidized beds where the
radiant contribution is important, the observed heat transfer coefficients
can be several times higher.
Therefore, our primary objective here is to
evaluate the effects of operating variables and design parameters on the
overall heat transfer coefficient in fluidized beds so that the
99
-------
knowledge can be utilized in designing fluidized bed boilers for
power generation and air pollution abatement.
Two principal resistances in series must be considered in
evaluating the overall heat transfer coefficient:
bed-tube heat transfer
For nucleic boiling
2
on the waterside, heat transfer coefficients on the order of 1000 Btu/ft /hr/
resistance and tube-water heat transfer resistance.
OF may be expected.
If the water-side or steam-side conditions are forced
circulation, as may be expected in economizer or superheater and reheater
designs, the water-side coefficients are usually less (300":'600 Btu/ft2/hr/oF)
but can .be reasonably estimated from the existing empirical relations,
such as the Dittus-Boelter equation [43].
The bed~tube heat transfer
coefficient, however, depends to a great extent on the operating ~onditions
(bed temperature', flli.idizing velocity, bed particle size, solid circulation
patterns in the bed); on the design parameters. (bed size, tube pitch/
diameter ratio, number of rows of tubes, triangular or. square. pitch,
vertical or horizontal tube); and on. the properties of reactants (thermal
conductivity, thermal diffusivity, density, and heat capacity of the bed
material and gas in the bed).
The operating and design variables are
classified into primary and secondary variables, depending on their
relative importance in influencing the heat transfer coefficients.
Primary Variables - Operating Variables
Bprl Temperature.
Changing the bed temperature affects the
overall heat transfer by altering:
. The thermal conductivity of gas and thermal diffusivity of
the bed material
. The radiation contribution portion of the total heat transferred
. The overall temperature difference between the bed and the
immersed heat transfer surfaces.
100
-------
The heat transfer between the bed and the immersed tubes
cons!itutes two portions, the convective and the radiant heat transfer.
The magnitude of convective heat transfer depends on the particle size
distribution, residence time of the particles at the heat transfer surface,
the density of particles in the bed, and the physical properties of the
gas medium.
Botterill [12J showed that particle-to-surface contact areas
are too small for significant heat transfer to occur by conduction through
points 'of solid contact;
Experimental results show that the heat transfer
takes place primarily by conduction through the thin gas film between the
surface and the particles. In other cases, the residence time of the particles
near the heat transfer surfaces can be a limiting factor.
Hence, the heat
transfer coefficients were found to be higher when gases with higher thermal,
conductivities [13J and solids with higher heat capacities were used at
given experimental conditions.
Experimental data indicate that the heat
transfer coefficient, h, depends on the thermal conductivity of the gas,
k , and the heat capacity of the solid, C , in the following ways:
g s
h a k 0.33 ~o k 1.0. haC 0.25 to C 0.8.
g g' ~ s
Experimental measurements at bed temperatures between lOOO°F
and l800°F suggest that radiation in a fluidized bed may be treated as a
black body and is additive to the convective heat transfer [42]. At
~1500°Fbed temperature, the radiation contribution ranges from 10% to 50%
of total heat transferred in the bed, depending on the sizes of bed
particles and the fluidizing velocity.
Hence, the heat transfer coefficient increases with an increase
in bed temperature because of the increase in thermal conductivity of the
gas, the increase in thermal diffusivity of the particles, and the
increased contribution of radiation. The quantitative effect of the bed
temperature on the heat transfer coefficient, based on the studies by
Kharchenko and Makhorin [56J and Peel [73J, is presented in Figure 1-44.
The bed temperature effect is complicated and depends on additional
101
-------
160
u..
0
I
I...
-7= 140
N
-
-
-
::J
I-
co
- 120
-
c::
Q)
u
-
'Q) 100
0
U
I...
Q)
~ -
0 VI 80
N c::
ro
I...
I-
-
ro
Q)
:r: 60
E
::J
E
>< 40
ro
~
20
200 400
5
2
800
Particle
Material Size (inches)
Quartz Sand 0.013
Fi re-Clay 0.017
Sand 0.012
Coal 0.010
Char 0.010 I
1000 1200 1400 1600
Bed Temperature, of
Cu rve
1
2
3
4
5
Curve 643511-A
1
Reference
Kharchenko & Makhorin (56)
Kharchenko & Makhorin (56)
Peel (73)
Peel (73)
Peel (73)
2000
Figo 1-44-Effect of bed temperature on heat transfer coefficient
-------
variables such as bed material and particle size.
Note that the data are
not obtained in a fluidized bed combustor.
In fact, few data from
fluidized bed coal combustion which elucidate the bed temperature effect
are available.
Fortunately, a comprehensive study of bed temperature
effect is not necessary for the fluidized bed combustion of coal because
the temperature range of interest is quite narrow (1400°F to 1700°F in the
primary combustor and 1800°F to 2000°F in the carbon burn-up cell) and
can probably be accounted for satisfactorily by a radiation contribution.
Operating Pressure. Running at higher pTessure will provide
better fluidization and heat transfer because the density of the gas is
higher.
Because of the scarcity of experimental data, the quantitative
improvement is not clear.
Particle Size.
Experimental evidence [54,56,105] indicates that
decreasing the particle diameter causes an increase in the heat transfer
coefficient due to an increase in particle mobility and reduction in the
thickness of the gas film between the immersed heat transfer surfaces and
the aggFegates of particles.
The particle size distribution will also
affect the magnitude of the heat transfer coefficient:
a wide size particle
range tends to give better fluidization and higher particle mobility and
agitation.
The reported dependence on average particle diameter ranges
from d-O.17 to d-O.96 [13,43]. The effect of average particle size is
. p P
shown in Figure 1-45 according to Wright et al., [102]. . The data are for the.
convective heat transfer component only; the radiation contribution is
subtracted from the total heat transfer by assuming the fluidized bed
to be a black body at its bed temperature [42].
The data show that the
convective heat transfer coefficients start to level off at a mean
particle size larger than ~O.07 in. The design particle size in the
fluidized bed combustor is -1/4 in with a mean particle size around 0.07 in
(~1800)J) .
Variation of the heat transfer coefficient due to changes of
mean particle size from attrition, e1utriation, and screening operations is
thus estimated to be less than 10%.
effect may be more pronounced.
At sma11er'mean particle sizes, the
103
-------
.01 0.02 0.03 0.04
o
70
L.J....
0
I
L..
~
I
N 60
+-'
-
-
::J
I-
co
- 50
+-'
C
.~
u
-
-
Q) 40
0
u
...... L..
a Q)
-I'- -
VI
c 30
"'
L..
I-
+-'
"'
Q)
:J:
Q) 20
.~
+-'
U
Q)
>
c 10
0
u
Curve 643510-A
Average Particle Diameter, inches
O. 05 O. 06 O. 07 O. 08
0.09
0.10
Reference
Wright et al. (102)
Kharchenko & Makhor i n (56)
Baerg, Klassen & Gishler (6)
Kim, Kim, Chu n & Choo (57)
o 50 40
16 14 12 10
Equivalent U. S. Mesh Size
30 25 20
8
18
Fig. 1-45-Effect of particle diameter on heat transfer coefficient
0.11
7
-------
Fluidizing Velocity. As the fluidizing velocity is increased
, ,
over the minimum fluidizing velocity, the heat transfer coeffi~ient first
increases rapidly due to increased agitation in the bed until a maximum
is reached at about two to nine times the minimum fluidizing velocity
[56,58,74], depending on the particle size in the bed. TIlen the heat
transfer coefficient remains essentially constant for a short range.
Further increases in fluidizing velocity cause a decrease in the heat
transfer coefficient, because the decrease in particle concentration in
the bed overtakes the effect of increasing particle agitation.
In a bed
of larger particle size, where the large particles tend to segregate out
at the bottom of the bed and become sluggishly fluidized at lower
fluidizing velocities, the effect of superficial velocity is especially
clear. Figure 1-46 shows the experimental measurements [42] of heat
transfer to the containing walls,in a fluidized bed burning an untreated
coal of 1/4 in
x 0 size at different bed depths.
In the lower part of
the bed, where a sluggishly fluidized bed of bigger particles is observed,
the heat transfer coefficient is lower and highly dependent on the
fluidizing velocity for both horizontal and vertical tubes in the bed. In
the upper part of the bed, where the quality of fluidization is good,
higher heat transfer coefficients are obtained with less d~pendence on
velocity.
In general, stronger dependence on velocity was observed with
vertical tubes in the bed.
For coarser particles and higher bed
temperatures, the region for a constant heat transfer coefficient can
cover quite a wide range of gas stream velocities at fully fluidized conditions.
The rate of heat transfer to the immersed surfaces was found
to be generally higher than that to the walls and was also dependent on
fluidizing velocity, as shown in Figure 1-47 where only the measurements
at the fully fluidized section were shown.
The operating conditions and
the heat transfer coefficients obtained are summarized in Table 1-5.
It is important that, in order to avoid segregation of particles
in the bed, the design fluidizing velocity must be at least larger than
the minimum fluidizing velocity of the biggest particle in the bed.
For example, the minimumf'luidizing velocity of 1/4-in particles at the
105
-------
L.J....
o
I
.....
.r:
N~ 10
-
::3
~
ca
-- 0
c:
Q)
:g 40
-
~
o
u
2 30
V')
c:
«:I
.....
~
16 20
Q)
:I:
Curve 643525-8
30
.
,- -.. . , .
-. '....~~..L-----
1--. 8A .
20
.
10
Depth 18 5/8" - 24 3/4"
o
40
30
. .. - -
-.. . --~
- .-, --------;;
.
.
.
20
Unit Size: 27" Diameter Unit
Coal Size: Untreated, 1/4"-0
Bed Depth: 22 - 24"
. Horizontal Tubes in Bed
. Vertical Tubes in bed - - - --
Reference: Hickman et al.. (42)
. Depth 12 3/8" - 18 5/8"
.
.. ~
.. ~ .
.-. -: . . ,. ./" ~
... ./"
. -./"
J ./"
.-/
.....- . .
-
10
Depth 6 1/8" - 12 3/8"
o
30
20
14
.
..
.
.
.
.
10
o
5
8 9 10 11
Superficial Velocity,ft/sec
Fig. 1-46-Effect of fluidizing velocity on heat transfer coefficient to walls'
7
12
13
106
-------
60
50
u..
o
I
L.
~
I
N
~ 40
~
......
co
I-'
o
--oJ
--
c
a>
:~ 30
-
-
a>
o
U
L.
a>
-
V)
c 20
co .
L.
......
--
co
a>
::::I:
10
05
.
.
Curve 643504-A
.
... /'
.~/
.. /
/
~
6
7
8
9 10
Superficial Velocity, ft/ sec
.
Key
Unit Size: 27" Diameter Unit
Coal Size: Untreated 1/4" -0
Bed Depth: 22" - 24"
. Horizontal Tubes in Bed
. Vertical Tubes in Bed -
Reference: Hickman et ai, (42)
11
13
14
12
Figo 1-47-Eftect of fluidizing velocity on heat transfer coefficient to immersed surfaces
-------
TABLE 1- 5
SUMMARY OF HFAT TRANSFER MEASUREMENT IN FLUIDIZED BEDS AT
ATMOSPHERIC PRESSURE PERFORMED BY OAP CONTRACTORS AND NCB LABORATORIES
BED FLUIDIZING
REACTOR TEMP. BED PARTICLE VELOCITY TUBE AT
ORGANIZATION SIZE (OF) MATERIAL SIZE (IT/SEC) IN BED IN BED WALLS
ABOVE
BED
REFERENCE
BCURA 27t1 diam. 700°C-lOOO°C Refractory Inert: 6-14 Vertical 25-40 35
or ash - 1/4" + 30 mesh hor izon tal 40-50 35-40
coal:
- 1/4" - 0
BCURA 12" diam. 7500C-9820C Refractory Iner t: 7-11 Horizontal 40-57 20-30 [54]
18-30 mesh
coal:
- 1/8" - 0
.....
0 NCB-CRE 3 ft. sq. Room temp. Coal ash - 10 mesh 1-4 Horizontal 20-50 30% less [43]
00
NCB-CRE 12" square 8000C-8500C Coal ash - 10 mesh 2-2.5 Horizontal 78-83 [85]
tube & loop
USBM 18" diam. l450-16500F A1203 - 16 + 48 mesh 2.4-3.4 Horizontal 54-67 Avg. 53 [15]
PER 12" x 16" 1450-1650°F Coal ash or Ash: -7+ 14 mesh 12-14 Bayonet 47 [75]
limestone limes tone: heat
BCR 1359 - 10+20 mesh exchanger
"
NOTE: BCURA = British Coal Utilization Research Association, England.
NCB-CRE = National Coal Board - Coal Research Establishment, England.
USBM = U.S. Bureau of Mines at Morgantown, West Virginia.
PER = Pope, Evans, & Robbins, Consulting Engineers, Virginia.
[42]
-------
experimental conditions is '\110 ft/sec; the fluidized bed with the
same particle size distribution and operating at a velocity> 10 ft/sec
will have better particle mixing and consequently a more uniform heat
transfer coefficient in the bed (Figures 1-46 and 1-47). Segregation of
bed particles will result in a lower overall heat transfer coefficient,
temperature gradient in the bed, and possible agglomeration due to
sluggish agitation.
However, this is not a totally detrimental effect.
Segregation of bed particles which results in lower overall heat transfer
coefficient will be helpful during boiler turn-down (see Appendix E)
if particle agglomeration can be prevented.
The relationship between particle segregation and fluidizing
velocity is not very well understood and should be further investigated,
especially in pressurized operation with deep beds, where the phenomenon
is expected to be enhanced.
Cold model studies in this area are being
conducted to gain more understanding of the problem and to provide some
information for design.
Primary Variables - Design Parameters
Position of Heat Transfer Surface.
Many investigators have
reported variations in the heat transfer coefficient relative to the position
of heat transfer surfaces in the fluidized bed, particularly the
differences in coefficients for vertical and horizontal tubes in the bed
[65,99] and the differences between containing walls and immersed heat
transfer surfaces [42,55,56], (Figures 1-46 and 1-47).
Morgan [65] reported an increase of heat transfe~ coefficient
,,'
up to 40% greater for vertical tubes than for horizontal tubes, depending
on the average particle size employed (Figure 1-48). This contradicts
the data presented in Figures 1-46 and 1-47. Although the vertical tube
arrangement usually gives worse fluidization because of a possible
channeling effect, the consistently higher heat transfer coefficient for
vertical tubes here is probably due to the stagnant layer of particles
on the upper surface of the horizontal tubes which renders part of the
109
-------
lJ...
o.
I
~
.r::.
I
N
.......
-
-..
::J
I-
co
Curve 643507-A
100
90
80
70
60
50
-
.......
c:
(])
u
40
-
-
Q)
o
U
30
Key
. Horizontal Heater
& Vertical Heater at Axis
Sand 328~
Flu idizing Velocity 0.50- O. 85 ft/ sec
Reference: Morgan (65).
~
Q)
-
V)
c:
ro
~
I-
20
.......
ro
Q)
:c
10 1 2 3 4 5 . 6 7 8 9 10 20
Effective Air Flow Rate, (actual air flow-air flow at minimu m flu idization),
. . ft3/ min .
Fig. 1-48-Variation of heat transfer coefficient with position of
heat tra nsfer su rface
110
-------
surface inactive.
Westinghouse cold model study confirms this hypothesis
[53] .
However, the study by BCURA burning 1/4 in
x 0 untreated coal in
an ash bed with vertical and horizontal tubes sized at 2-3/8 in D.D.
had the opposite result (Figures 1-46 and 1-47), where the heat transfer
coefficient to vertical tubes was shown to be up to 30% below that of
horizontal tubes. The explanation of this appa~ent contradiction may.
be that the stagnant layer of particles on the upper surface of the
horizontal tubes was absent in BCURA's studies because of the high
velocities and the large particles employed. The stability of this
stagnant layer depends on particle size/tube diameter ratio and the
violence of agitation. Presumably, the tube pitch/diameter ratio also
plays an important role here.
The difference in the heat transfer coefficients at walls and at
the internal heat transfer surface, however, is caused by an entirely
different mechanism.
For a gas-solid fluidized bed operating at atmospheric
pressure, the solid agitation is controlled almost entirely by bubbling
phenomena of gas bubbles which carry the solid particles with them in the
wake until they finally burst at the bed surface. In addition, the gas
bubbles tend to pass from the bottom to the top of the fluidized bed
through the lowest resistance pass, i.e., clear of the containing walls.
To maintain the balance in the bed, particles have to move downward in
the particulate phase and thus set up a mass circulating flow pattern [52]
where the solid particles are swept rapidly upward in the middle of the
bed by the gas bubbles and then move slowly downward near the walls.
Hence,
the same particles near the walls remain in close contact with the walls
for a relatively long period of time without replacement and cause the heat
transfer coefficient to drop (Figures 1-46 and 1-47).
Conflicting results were also reported for the heat transfer
coefficient of the horizontal tube banks at the top and bottom of the
fluidized bed.
Highley
et al. [43] found, in their studies of heat
transfer in a three-foot square cold ash bed with a bed depth of three
feet, that raising a bundle from the lower to the upper half of the bed
increased the heat transfer coefficient 20% and also observed some
111
-------
segregation of particles in the bed at about two times the minimum
fluidizing velocity. Using 1/4 in x 0 size particles at low fluidizing
velocity, Hickman et al. [42] reported a difference of heat transfer"
coefficient for heat transfer surfaces located at different bed depth
due to segregation of bed particles, which resulted in a sluggishly
fluidized portion at the bottom of the bed (see Figure 1-46).
There
is, however, no other strong evidence that the heat transfer coefficient
should be different for tubes at different bed depth at good fluidization
~
conditions.
For tubes exposed immediately above the bed and still in the
splashing range of the bed material, a drop of 30% in the heat transfer
coefficient was noted in room-temperature ash bed experiments [43].
However,in combustion of coal at bed temperatures in the range l450°F to
l650°F, Coates et al. [15] reported approximately the same overall heat
2
transfer coefficient, 53 Btu/ft /hr/o~ both for the tubes immersed in the
bed and for the tubes a few inches above the bed.
The discrepancy may come
from either the confusing definition of bed height or from the additional
combustion of combustibles above the bed and the radiation effect in
the "hot" experiments.
In summary, the "heat transfer coefficient for the immersed
surfaces in the bed is at least 10% greater than that for the immersed
surfaces at walls.
Efficiency in heat transfer for vertical tubes versus
horizontal tubes can be quite different depending on other operating
and design variables such as particle size/tube diameter ratio, superficial
fluidizing velocity, and tube pitch/diameter ratio. For tubes exposed
immediately above the bed and still in the splashing range of the bed
material, a drop of up to 30% in the heat transfer coefficient is
possible if the additional combustion of combustibles above the bed and
the radiation effect are absent.
More data are required to permit
accurate design of heat transfer surfaces above the bed and in the
freeboard.
112
-------
Tube Arran~ement.
Tube pitch/tube diameter ratio is another
important variable to be considered for design.
adjacent tubes tends to reduce the mobility of
the heat ,transfer coefficient. Increasing the
Reducing the gap between
the particles and hence
number of ,the staggered
rows has the same effect.
tubes in stages from 11 in
In reducing the distance between adjacent
to 0.6 in,
a red~ction of 25% in the heat
Adding an extra row to increase
transfer coefficient was observed [43].
the staggered rows from two to three rows when the gap was 2.5 in reduced
the coefficient by a further 15%. The range of tube pitch/tube diameter
ratio employed is two to eight with tube sizes of 1-3/8 in
O.D. and
2-3/8 in
O.D.
The result of the measurements at room temperature in a
vessel three feet square is shown in Figure 1-49.
The results can also
be correlated in terms of the narrowest gap between tubes in the bundle,
as shown in Figure 1-50 [64]. The heat transfer coefficient decreases
continuously for about 25% as the gap is reduced from 11.1 in to 0.59 in
without any sharp drop indicating a marked decrease in particle mobility.
The particle size/narrowest gap (between tubes) ratio is probably
important; however its effect has not yet been studied.
The extent of the effect of using square pitch rather than
triangular pitch is not clear.
It is advantageous to keep the pitch/diameter ratio as small
as possible. For example, to lower the pitch/diameter ratio from eight
to two, an increase of 16 times of heat transfer surface per unit volume
of bed,decreases the heat transfer coefficient by only 18% (Figure 1-49).
Thus, the compactness of the boiler design, rather than the heat transfer
coefficient, is the major consideration in determining the pitch/diameter
ratio.
However, there is a narrowest gap between tubes, depending on
the particle size, below which a sharp drop in the heat transfer coefficient
will occur. Determination of this narrowest gap requires further experimental
studies.
Secondary Variables
Fines Recycle.
E1utriated fines can be recycled back to the original
113
-------
I..L..
o
I
....
..c:
I
N
~ 40
::J
l-
ce
c:
'"
:~ 30
a;
o
U
....
2
~ 20
'"
.... .
I-
ro
'"
:J:
~ 10
::J
7
"C
'"
cc
I..L..
o
I
....
..c:
I
N;:::: 40
--
::J
l-
ce
c:
.~ 30
a;
o
U
....
'"
~ 20
'"
....
I-
..-
'"
'"
:J:
~ 10
::J
l-
I
"C
'"
cc
50
50
00
" -
....-
...
60 mm 35 mm No. of Pitch/ Diameter
Tubes Tu bes Rows Ratio
0 . 2 2
. 2 4
l::" ... 3 4
0 3 6
. 4 8
Reference: Highley et al (43)
2 4
Pitch/ Diameter Ratio
6
Curve 643505-A
-----
8
Fig. 1-40-Change of bed-tube heat transfer coefficient with pitch/diameter ratio
"w
~
...
00
60 mm 35 mm No. of Pitch/ Diameter
Tu bes Tu bes Rows Ratio
o . 2 2
.. 2 4
l::" ... 3 4
o 3 6
. 4 8
Reference: Williams et al (43), (64)
2
4 6
Narrowest Gap Between Tubes, inches
8
Fig" 1-50 -Change of heat transfer coefficient with tube spacing
114
Curve 643509-A
~
10
12
-------
bed or to the carbon burn-up cell.
transfer coefficient in two ways:
The fines recycle affects the heat
recycle of fines changes the particle
size distribution in the bed; large recycle .of fines at a temperature
different from that of the bed may create a local temperature gradient in
the bed.
Coal Characteristics.
Coal of different swelling numbers will
change the particle size distribution in the bed after combustion.
Different ash characteristics (elutriated or retained in the bed) will
also change the particle size distribution.
In general, the effect
should be small, because coal represents <0.5% of the total bed material.
For coal whose ash will accumulate in the bed after combustion, the
effect will be more important.
probably be important, too.
Caking properties of the coal will
Air Distributor Design.
Different air distributor designs will
generate air bubbles of different sizes and frequencies and will cause
different pressure drops.
This affects the quality of fluidization, which
is directly related to the effectiveness of mixing and heat transfer in
the bed.
Unit Size.
Slugging and poor mixing in the bed and, thus,
poorer heat transfer, are more likely to occur iri smaller units; larger
units are relatively free of these problems~
However, larger units
with improperly designed air distributors are apt to cause by-passing of
fluidizing air and affect the quality of fluidization. The major problem
associated with this difference in unit sizes arises from applying the
data on heat transfer coefficients obtained in smaller units to the design
of larger units.
Bed Material.
affect the heat transfer
Bed material of different heat capacities will
coefficient (h a C 0.25 to C 0.8). PER found
. s s
limestone bed and an ash
the same heat transfer coefficients for a
bed, i.e., 47 Btu/ft2-hr-oF [75].
115
-------
Summary
Data on heat transfer coefficients obtained in fluidized bed
combustion experiments by OAP contractors and NCB laboratories were
presented in Table 1-5. The effect of changing different primary va~iables
on heat transfer coefficients in the bed and the possible mechanisms of
these effects are summarized in the foJlowing matrix.
From the information
shown in Table 1-5 and from the estimation (based on this matrix) of poss~ble
changes in heat transfer coefficient due to changes in operating and design
2
variables, a heat transfer coefficient of 50 Btu/ft /hr/oF was chosen for
design of the industrial, the atmospheric pressure, and the pressurized
fluidized bed boilers.
This heat transfer coefficient is considered
reasonable and may be somewhat conservative in light of presently available
information.
Heat transfer in fluidized bed combustors can be understood
qualitatively with the data presently available, but quantitative design
correlations are still lacking.
Although considerable effort has been
spent on understanding the heat transfer properties in the atmospheric
pressure fluidized bed combustors, no comprehensive study has been
performed for pressurized combustors with deep beds to investigate the
effect of particle segregation and possible temperature gradients in
the bed.
More data must be collected to confirm the meager information
presently available on the effects of fluidizing velocity and tube
arrangement in the beds, especially at higher fluidizing velocities (up
to 15 ft/sec), smaller tube pitch/diameter ratios, and larger numbers
of rows of tubes.
Scaling-up is still the most important .problem for designing
full-scale fluidized beds from the data accumulated in laboratory
experimentation.
Until the mechanics of gas-particle motion in a
fluidized bed and its relation to the operating and design variables are
better understood, no theoretical design correlation describing the heat
transfer properties of fluidized beds can be formulated.
The solution to
116
-------
EFFECT OF DIFFERENT PRIMARY VARIABLES ON HEAT TRANSFER
COEFFICIENT IN THE BED AND THE MECHANISMS INVOLVED
VARIABLE
HEAT TRANSFER
COEFFICIENT
MECHANISM
CHANGE
Bed temperature
Increase
Increase
Increase in thermal conductivity and
thermal diffusivity; increase in con-
tribution of radiation
Particle size Decrease Increase
Fluidizing Increase Fir st increase
velocity beyond the rapidly then
minimum decrease slowly
fluidizing
velocity
Geometric
factors
walls or Higher to the
inunersed inunersed tubes
tubes
vertical or contradictory
horizontal results
tubes
tube bank in No change.
different bed Higher close to
depth surface of the bed
in splashing Cold 30% drop from that
zone experiment in the bed.
hot No change
experiment
tube pitch/diam. Decrease Decrease
number of staggered Decrease Decrease
tube rows
square or tri- Not clear
angular pitch
Increase in particle mobility and
decrease in thickness of the gas film
Increase in particle agitation
Decrease in particle concentration in
the bed
Bubbling phenomena and solid circulation
pattern in the bed
Possible particle deposit on upper half
of the horizontal tubes; possible
channeling effect for vertical tubes;
tube pitch/diameter ratio may be important
At good fluidization.
Segregation of particles in the bed results
in sluggishly fluidized particles at botto~
Decrease in particle concentration.
Radiation contribution; additional
com~ustion above the bed.
Decrease in particle mobility
Decrease in particle mobility
117
-------
the problem of scaling-up will mostly rely on the empirical or semi-
empirical correlations obtained through small-scale laboratory
experimentation.
118
-------
, Particle Carry-Over
For design purposes, knowledge of the rate of entrainment of
solids from a fluidized bed, of the size distribution of the elutriated
solids in relation to that in the bed, and of the variation of both
these quantities with the height of the freeboard is very important.
In the case of fluidized bed coal combustion for air pollution abatement,
where limestone or dolomite is recommended as bed material, the increase
of dust loading from limestone fines and by limestone decrepitation and
, attrition may be expected to be the primary consideration for partic-
ulate control.
When limestone is the bed material, the solids elutriated
from the fluidized bed consist of coal ash, CaS04' unreacted CaO, and
uncalcined CaC03' When dolomite is used, the elutriated solids will also
contain the low reactive portion of MgO and MgC03'
Empirical investigations of the elutriation rate from an inert
bed have been reviewed by Kunii and Levenspie1 [60].
Studies of the
elutriation rate from an ash bed were also conducted by Williams [78]
in a one-foot square cold bed at 2.5 ft/sec fluidizing velocity. All
studies show that the elutriation of the selected size fraction is
directly proportional to its concentration in the bed and the excess
fluidizing velocity (flu1dlzing velocity minus minimum f'luidizing velocity -
see Fig. 1-51), with the proportional constants being different for dif-
ferent materials and different particle sizes., These correlations are
available [60, 78]. Thus it is possible to estimate the elutriation
rate of a particular fresh limestone or dolomite once its attrition
and decrepitation rates in a particular environment are determined.
For limestones and dolomites which repeatedly go through absorption-
regeneration cycles, the attrition and decrepitation rates are much
more difficult to determine and will depend very much on their temper-
ature history and on the reactor operating conditions. For the
elutriation of carbon fines, the situation is also complex because
the particles are no longer inert; the e1utriation of carbon fines
depends now not only on their concentration in the bed and the fluid-
izingvelocitybut also on bed temperature, excess air, and their relative
combustion rate in the particular operating conditions.
119
-------
104
103
::a.
~
1>
E
'"
<5
..
~
c..
102
101
10-3
10,000
1000
~
c
e
....
'E
..:
'"
1>
E
'"
<5
..
"
t
'"
c..
100
10
0.001
,,.c\\'\
\It''''''
,~\'\.\(\I\
~\\I\
.~\I~
*\(\\
Basis
Temperature
Pressure
Solid Density: Ps
Gas Density, Pg
Gas Viscosity, II
Basis
Temperature
Pressure
Solid Density, p
Gas Density, p s
9
Gas ViscositY,1I
Curve 61+6566-8
17000F
I atm
125 Ibltt3 (IimestoneJ
O. 189 Iblft3
O. 045 cp
1/2
ud3.lg I:: - PgIDp)-}
j/ ',A~sh
~/A /' /
/ \ /' f'lo Mesh
u "'? Iop'Ps -~I.?/, ~/
mf =~ - 245 Pg ,16 Mesh
/' ,~
/ ~
.... / '
./
~ ./' --
0;. "/'/:/", 60 Mesh
9~ ~ ..../"' [I - 122~1/3
.......-7 \1.'> ~ U = ~ PsPq 9 D
...../// ~~<,I) . t 225 PglI P looMesh
~/ ~/
/ ~....--~
/.... ~/"' 2
....-- ~ \le\i.\~,~ U = 9 Dp IPs - p)
~./ \\\.'1>\ ~ t 1811
~/ ~~/
..../' ~/
..../ ~/
./ '
D 2 I P - P Ig
U -...L.:.L-JL.:
mf - 165011
3 1/2
II \[ D P (p - Pig]
Umf = DpPg l m.7J2 + 0.0408 P 9 / 9 -33. 7
28 Mesh
200 Mes h
10-2
10-1
Actual Gas Velocity, ft/sec
102
100
101
Curve 594086-0
D2( ) 'I 1/2, {,
U = P Ps-Pgg [3.19/P -PIDJr
mf 165~ /U = 5 9 P
t Pg . I 6 Mesh
3 112 ~ '/ II 1/8"
II 2 DpPq (ps-Pq'g -33.7 /, ~
Umf=Dr~33.7) +O.OQ 2 J ~ 1/ / 10 Mesh
pPg ~ II A' / II
/' ( ,~
. ~ /\ ,/~'
u 0 ?pIPS-Pq,g ,/ //
mf 245 p / ~
9 /
17000F
10 atm
125Ib/ft3(1imestone)
O. 189 Ib/ ft3
0.045 cp..
16 Mesh
28 Mesh
// / [
,/' ~Ut= 2~5
/' ....-",,-/
Ps =50 ~/" ~/
/' '
./' 125 ....~
,/' ,~
,...-/ .~ ./'/
./' \It\Y/
/' \t~~\ /"
..../ ~'
,/ /"'
/
Ip -p ,2 2J 1/3
s 9 gn D
Pg II P
60 Mesh
,~\\'\
\It'
,~\i\(\\\
~\ \1\
,~\I~
*\(\\
100 Mesh
200 Mesh
D 2 (Ps - pi
ut = 9 P
1811
0.01
0.10
Actual Gas Velocity, ",see
1.0
10.0
100
Fig. 1-51- Fluidizing velocity curves for I 111m. and 10 111m. operation
12D
-------
rhe experimental data for e1utriation of carbon fines were correlated and
:'" :;;1 . ~ " . - . ~ ; . .
discussed in the Combustion Efficiency section. Our primary concern
".,
here :is~t6 determ:i:ne the rate of e1utriation or dust loading at dif-
fere,nt.'operating conditions and their respective particle size dis-
" "\'" .
t iibutions .
'." ,." .:,;
Since the evaluation of the combustion of coal in different
,': '.~~', . .~ :' I. .
beds (ash', refractory, or limestone bed) shows no clear indication that
the type of bed material will affect the coal combustion efficiency, it
is considered reasonable to assume t~at the elutriation of ash and lime-
stone are independent processes and that their elutriation rates are
additive.
For certain coals ,with ash content less than ~ 15%, notably
the Ohio Pittsburgh No.8 seam, experimental evidence [15] shows that
most of the ash is elutriated out during operation. Thus the total
e1utriation rate is the rate of ash input from coal (which can be
determined accurately through coal analysis) plus the independent
e1utriation rate of limestone or dolomite (which can be experimentally
determined for each stone at specific operating conditions).
1)
The elutriated limestone or dolomite comes from three sources:
attrition of the particles in the bed followed by subsequent
elutriation; 2?
decrepitation of particles due to their history-
regeneration process, temperature cycles, etc.; 3)
the size fraction
of the feed material which has terminal velocities lower than the
operating velocity.
Unfortunately, very few data are ,available in
these areas, except the limited attrition data from Esso, PER, and
Consol[37, 75, 84, 107] obtained at low velocities and in small-batch
laboratory units.
The picture is also confused by the fact that the physical
integrity of the stone depends very much on the origin, preparation
and temp~rature history of the individual stone. Argonne [3] reported
in their studies, that:: afte,I' calcination at l600°F of once-screened 30 ~ .
. dolomite 1337 particles, ,40% of the material decrepitated to fine
powder.
Esso [84] studied the rate of attrition, followed by'
121
-------
elutriation, of calcined dolomite 1337 and limestone 1343 and 1359 up
to excess superficial velocity of 5 ft/sec.
They found that the
attrition rate of dolomite was up to eight times larger than that of
tested limestone.
Assuming that extrapolation of Esso's data to
higher velocities is possible, a fluidized. bed of limestone 1359
particles of average size ~ 2500 V and running at 15 ft/s~c will hRve
an attrition rate of ~1.5% of bed weight per hour; for dolomite 1337
this would be 12% of bed weight per hour. In another study 01],
where fine coal (200 \1) was burned in a coarse limestone bed (930 \1),
. . .
the stone 135ge1utriation rate was found to decrease from the initial
particles "rounding off" period of 'V 5% bed' weight per hour to < l%/hr
in about two hours.
PER [75] studied the attrition of limestone 1359 in a batch
reactor of 12 in x 16 in at fluidizing velocities 8 to 12 ft/sec
and found the attrition rate during ~a1cination was 5% to 7% of initial,
charge per hour; during sorption 3% per hour; and during regeneration
4% per hour4 In Conso1's studies [107] ~uperficia1 velocity 2 to 3
ft/sec), where Tymochtee dolomite was used, the attrition rate for
hi~lly sulfated dolomite was about 0.5% per hour.
However, for lightly
sulfated dolomite, an attrition rate as high as 4% per hour was noted.
In addition, the size degradation was also found to depend on 'the type
of coal or char burned in the fluidized bed.
Thus, in a coarse bed
operation, where the terminal velocities of all the particles are
greater than the operating velocity, the rate of e1utriation can be
assumed to be the rate of attrition alone.
To provide this narrower
particle size distribution, the as-crushed material usually has to be
single- or double-screened.
The available e1utriation data obtained
. , "
by OAP contractors and British laboratories are summarized on page 130~
At steady-state, onc'e-through operation where the bed
, .
material is continuously fed and withdrawn, the elutriation rate will
depend not only on the bed weight but also on the material feed rate
as well, if the feed material consists of fine particles., In addition
to the attrition followed by e1titriation, the size' fractions of the
122
-------
feed material which have terminal velocities smaller than the operating
fluidizing velocity will also be e1utriated. The total amount of these
fractions e1utriated will depend on the residence time of the particles
in the bed or the change-over rate of the bed material, because part of
these solids will be withdrawn from the bed rather than e1utriated
from the top.
The worst case here would be that in addition to the
attrition and decrepItation contributions, the size fractions with
terminal velocities smaller than the operating fluidizing velocity
are completely elutriated.
When spent stones are to be regenerated and recycled with
make-up stones, the attrition in the regeneration and in the pneumatic
transport lines also has to be taken into account.
The temperature
cycling of the stone particles by going. back and forth between the
combustor and regenerator will also affect the physical integrity of
the particles. Studies by Esso [31] on 1359 limestone at 2 to 3 ft/sec
velocity showed little decrepitation up to six absorption - regeneration
cycles. Previous studies [84] in a similar system using the same stone
showed that less than 0.5% bed weight per cycle was e1utriated after
the first cycle (3.5% for the first cycle).
How to extrapolate these
data to higher velocities in a continuous absorption-regeneration
operation is not clear; more studies should be done.
Changing the fluidizing velocity usually changes not only the
rate of elutriation but also the size distribution of the e1utriated
solids. When fluidizing velocity is increased, the smaller-size
fractions will change little because of their low terminal velocities.
However, the larger-size fractions will progressively increase in
amount w~th increasing velocity so that the weight fraction of the
smaller particles tends to decrease and that of the larger particles
tends to increase, as shown in the studies by BCURA in a 27 -'in
diameter rig [42] (Figure 1-52).
In addition, the smaller-size
fractions tend to contain more unburned carbon than the large-size
fractions. As the fluidizing velocity increases, the carbon content
in elutriated solids increases accordingly, because of decrease in
123
-------
residence time of the coal particles.
But the carbon content in the
large-size fractions tends to increase more rapidly than that in the
smaller-size fractions (Figure 1-52). This may be because of two
factors:
first, the amount of the smaller-size fractions elutriated
changes little with the fluidizing velocity; and second, smaller-size
particles have higher combustion rates due to their larger surface
area/weight ratio.
In a steady-state operation, the elutriation rate and the
elutriated particle size distribution depend on the size distribution
of feed material as well.
feed has to be specified.
To be meaningful, the size distribution of
In the industrial application, the feed
material can be fed as crushed or it can be single- and double-screened,
depending on various application considerations. The end product of
these operations may be a vastly different size distribution. The
typical as-crushed coal or limestone size distribution from a Koppers
reversible hammermill designed for crushing 5 in x 0 to 95% through
1/4 in with a minimum of minus 100 mesh is shown in Figure 1-53, along
with the size di~tribution of coal feed used by different research
organizations.
With the coal feed size distribution as shown, the
elutriated particles from the combustors at BCURA (27 in - diameter unit)
and PER (18 in x 6 ft unit) using -1/4 in coal and at 10 to 15 ft/sec
.fluidizing velocity give comparable size distribution of elutriated
solids (Figure 1-54). The elutriated particle size distribution from
the carbon burn-up ~ell shown in Figure 1-54 is based primarily on
PER's experiments [76], where the fly ash of 30% to 60% carbon content
with a size distribution comparable to the solids elutriated out from
the primary combustor was injected into the simulated carbon burn-up
cell.
It is cautioned that the data presented here are obtained from
combustion of coal. in either an ash bed or an inert bed such as
refractory.
When limestone or dolomite is used as bed material, the
rate of elutriation and the particle size distribution of the elutriated
particles will depend on the type of stone used--limestone or dolomite
of different classifications; operational factors-- whether the stone
124
-------
Curve 6'13516-8
60 8
2r' Diameter Unit
50 2/3 Recycle 6
Reference: Hickman ..
et al. (42) ..
4D 4 .
.
30 .", 2
.!!!
- 200 Mesh Fraction co.
>f<. co. -200 Mesh Fraction
~
V> c
~ 20 .8 0
0
VI 40 ~ 6
'"
.", u
.:!! . 0
.!!! >f<
~
:; 30 ... CO
W .5;! 4
"0 . ~ .
c . .
N .5;! -70 +200 Mesh Fraction . .
~ u '"
~ 20 .!::!
VI 2 .
LL 50
.~ c
+ 70 Mesh Fraction c
VI ~ .
o -70 +200 Mesh Fraction
40 u
. 6
+70 Mesh Fraction
30 . 4
20
W7
.
13
07
Fig. 1-52 - Effect of fluidizing velocity on particle size and carbon dislribution
of el utriated solids
turl<';' U"~!,I /-P.
.
.
..
.
.
.
.
2r' Diameter Unit
Excess Air 20 - 50%
Reference: Hickman et aI., (42)
..
. .
.
. .
~
.
8
W II
Superficial Velocity, Itl sec
13
14
12
-------
Curve 643522-B
.1 ,/"
J V 5
1 5 y ~ .'
"" /3 ~" 3
./ 4/ V
5 / ~ / ~,/ /
./ / ' L" 10
15 / ~ """ ~l. ~ 20
25 ./ / /' ~'-I1
./ ,/ / ' ./" .. 30
/' /' /' ./~ / --
35 1/ ,/ ,/ ,/ 40..
- 45 ./ /. /,,, v
"" ./ ,/ .A 50
..c / / ...../ -. ----f-----
.S:> 55 ,/
Q) /' 1/ "" ~ ~ ,/ 60
:5: 65 / /I / ./.1
~ / V / "/ /./ ~ 70
J-' .~ 75 /' / / ~ //-/
N Q) / / ./~./ V'/ /
a-
N Curve Reference ~ - 80
VI ~/ / "' " ./~~/
.... 1 (42), (55)
Q) 85
> /' // :0V'l
a ,/ 2 (15)
I "" 3 (101)
~ /' - 90
,," J ~ 4 Size Distribution for Elutriated -
93 /" f~
5 Solids from Primary Combustor
./ ./ & Carbon Burnup Cell (see fig. 95
96 / ,/" 1- 54)
// - 97
," Range of Size Distribution
98 " "" -- Obtainable from Koppers -
// Reversible Hammermill - 98.5
99 '" I II I I I I i II II
10
100
1000
Particle Size Microns
10,000
Fig. 1-53-Size distribution of coal feed
-------
......
N
--...J
. 1
1
5
15
25
+-' 35
~
.~ 45
~ 55
~ 65
~
, 75
Q)
N
en
~ 85
>
o
I
c=:: 93
Curve 643513-A
96
I I I I I I I I I I I I I I I
~ ~ -
) v'
-1 r-
f-- ~- ~ -
/ " '\ ~
f-- " >,; -
" > "
f-- / "',," >< <:, -
f-- ./ >< ~ ,,' -
V
f-- ~ r'\.... ........ -
f-- ~ ......,- ..... )K r"-. -
- "/, .... ,,~ f"-.
- ~ )0 -
- ~'" > ~ " -
........ ' ,
~,l ~ ,I"" < K ' "'""
- /< '" -
" " "
CBC ~ ~ ,, " Coal: -1/4" x 0
1/ / ~ 1'\
- " -
~ Gas Velocity: 10-15 fps
V ~ XXXX Represents Data Spread-
- /' / I
~ /' I . I I I I I I I
- /f / If References: (42),. (55), (76)! -
,
, !
I
FBC '
- -
I I I I I I I I I I I I I I I I
98
99
10 100
Particle Size, microns
1000
1
Fig.. 1-54-Projected p'article size distribution for material elutriated
from fluidized bed combustor (FBC) & carbon burn-up cell (CBC)
-------
will be used once-through, regenerated on-site and recycled, or regen-
erated off-site; and particle size distribution of the feed stones~-
as-crushed, single-, or double-screened.
For example, if as-crushed stone with a size distribution
similar to that shown in Figure 1-53 (from Koppers' reversible. hammer-
mill) is used once-through in a fluidized bed operating at 15 ft/sec,
the size fractions with lower terminal velocities can be elutriated
out.
The total amount which will be elutriated depends on the change-
over rate of the bed material, with the maximum rate ~ 35% of the total
feed rate at the size distribution and fluidizing velocity specified.
In the absence of experimental data, the particulate removal equipment
can be designed for the most pessimistic case, where all-size fractions
with terminal velocities lower than the operating velocity are com-
pletely elutriated - that is, 35% of the total feed rate in the present
example.
When the as-crushed stones are merely used as make-up supply
for the regenerated stones, the elutriation rate can be taken to be
35% of the make-up stones plus the attrition rate of the recycled
stones. For example, if the make-up stones constitute 10% of the
total stone input, and the attrition of the recycled stones represent
1.5% of the total weight, then the total elutriation rate would be
about 5% of the total stone feed rate.
Hence, the use of as-crushed
stones for once-through application will generally give much higher
dust loading than the regeneration case when the attrition rate of
the stones is small.
In. this case, single- or double-screening the
as-crushed stones will effectively cut down the dust loading.
When
the specified particle size distribution is such that the terminal
velocities are larger than the operating velocity, the elutriation
represents the most optimistic case and is only due to attrition of
particles in the bed.
No elutriation data are available for combustion of coal in
limestone or dolomite beds operating at high velocities, in larger units,
or in continuous absorption-regeneration operations.
Further work is
needed to provide the required data for accurate estimation of the
elutriation rate for design purposes.
128
-------
Summary
When limestone or dolomite is used as bed material, the
increase in dust loading from stone fines and by stone attr~tion and
decrepitation may be expected to be the primary consideration for
particulate control. The design of particulate removal equipment is
also complicated by the fact that the physical integrity of the stones
depends very much on the origin, preparation, and temperature history
of the individual stone. Unfortunately, most of the elutriation
studies performed so far have been confined to batch combustion of
coal in ash. beds at low velocities. The available information is sum-
marized in the following matrix.
Based on the available information,
design bases were chosen for preliminary design of the industrial,
atmospheric pressure, and pressurized fluidized bed boilers as shown
in the second matrix.
Data are needed from combustion of coal in
limestone or dolomite beds operating at high velocities (up to 15
ft/sec), in larger units, and in continuous absorption-regeneration
operations.
129
-------
ORGANIZATION REACTOR SIZE BED MATERIAL
USBM IS'! diameter refractory
Argonne 6" diameter alumina
E'ssq
Esse)
,...
w
o
PER
Conso1(a)
NCIj-CRE
3" d~ameter
1337
1343
1359
SUMMARY OF AVAILABLE ICLIiTRIATlON DATA AT ATMOSPHERIC
PRESSURE FROM (JAP CONTRACTORS N'IIJ NCB LABIJRATOIO ES
l-
BED
TEMPERATURE
PARTICLE SIZE (OF)
Bed:-16+48 mesh
coa1:1/8" x 0
Bed: 30 mesh
coal: -14 mesh
sorbent: 25.Ao'
100~, 300...-,.
Coarse pas ticks
for studying
rate of attrition
Bed: 93CWf;
coa 1: 200.......-
Bed:-10+20 mesh
coal: 1/4" x 0
-16+28 mesh
-10 mesh
1450-] 650
1600
Room temp.
Combustion:
1600
r'2generntion:
2000
Combustion:
1500-]600
regeneratiqn:
]900-2000
1700-1900
Room temp.
--J::---
FLUIDIZING
VELOCITY
(FT/SEC) ELUTRIATION
2.4-3.4
3
Excess
superficial
velocity:
0-5
Combustion:
3
Regeneration:
2
8-12
2-3
2.5
USBM = U.S~ BUF~a:u of Mines at Morgantown, West V~rgiI1~a
Argonne = Argonne Nattona1 Laboratory .
Esso = Esso Research and Engineering ComP4oy, Linden, Ne\v Jersey
PER'= Pope, Evans, anq Rob~ins, Consulting Engineers, Alexandria, Virginia
Consol ='Consolidation Coal Company
NCIj-CRE = N~ti~na1 Co~l Board - C~al Researcl) Establishment, England
NOTE:
3" diameter
1359
12!1 X 16"
1359
4" diam~t~F
Tymqeh~ee
Dolomite
12'! square
Ash
(a)Operating press~re 8 psig.
OR ATTRlTlON RATE
R2fractory: 0.1% of bed weight per hour.
ash: 50%-]00% for 6 ~ifferent coals.
Ash: 100%; 80%-90% uf sum of ash and sor-
bent with 100~sorbent; 60-80% with 300~
sorbent. - .
Attrition Rate:
0-3.5% hed weight/hr for
0-0.5% hed weight/hr for
0-0.7% bed weight/hr for
BCR 1337
BCR 1343
BCR 1359
Rounding-off period: 5% bed weight/hr
steady state: ~l% bed weight/hr
cyclic absorption-regeneration: 3.5%
for Is t cycle; < 0.5% for 2nd to 10th
cycle.
Calcination: 5-7% initial charge/hr
sorption: 3% bed weight/pr
regeneration: 4% bed weight/hr
0.5% bed weight/hr for highly sulfated
dolomite; 4%/hr for +ightly sulfated
dolomi~p..
R = KC
R = el~~riation rate
C = concentration of fines in th~ bed
K = ~ proportional constant depen4s 00
opera~ing velocity and terminal velo-
city of fines frac~~o~.
REFERENCE
[15]
[3]
[84]
[31]
[75 ]
[107]
[78]
-------
DESIGN BASES FOR PARTICULATE REMOVAL EQUIPMENT FOR THE
INDUSTRIAL, ATMOSPHERIC-PRESSURE, AND PRESSURIZED FLUIDIZED
BED BOILERS
INDUSTRIAL BOILDER
ATMOSPHERIC-PRESSURE
UTILITY BOILER
PRESSURIZED
UTILITY BOILER
ASH
100% elutriated
100% elutriated
100% elutriated
SORBENTS
(-1/4" with
minimum of fines)
3% of limestone
feed rate(a)
3% of limestone
feed rate(a)
5% of limestone
feed rate(a)
I-'
W
I-'
PARTICLE SIZE
DISTRIBUTION
From Primary Bed
From CBC
Figure 1-54
Figure 1-54
Figure 1-54
FigUJ;e 1-54
Figure 1-54
Figure 1-54
(a) Limestone feed rate is the combined feed rate of the recycled stream and the make-up
stream. The recycled stream may contain a smaller amount of fines than that in the.
make-up stream but usually has higher attrition and decrepitation rates.
-------
Boiler Tube Corrosion. Erosion and Fouling
Introduction
The economics ~nd performance of fluidized bed combustion
boilerswfth'l~ersed heat transfer surface depend on the resistance to
corrosion, erosion, and fouling of the heat transfer surface in the
fluidized bed and in the freeboard.
Theoretical analyses have been made
to assess potential corrosion, erosion, and fouling of boiler tubes at
projected operating conditions.
Experimental data have been obtained
on tubes in fluidized beds of limestone and dolomite burning various
coals.
An evaluation of the available theoretical and experimental
information indicates that conventional boiler tube materials can be
used in the proposed fluidized bed combustion boiler designs without
encountering corrosion, erosion, or fouling problems.
The fluidized bed boiler may also'enable power plants to
operate at steam conditions in excess of present practice.
In the case
of conventional boilers, the steam cycles have hitherto been limited to
a temperature of lOSO°F because of fouling and corrosion of superheater
tubes.
Fluidized bed combustion of coal conducted at a much lower
temperature (1400°F to 1800°F) compared to a flame temperature of up to
3000°F in conventional boilers promises to cut down fouling, corrosion,
and erosion of the boiler tubes, and permits operation at higher-temperature
steam cycles in addition to providing a higher heat transfer coefficient
in the bed.
Theoretical Evaluation
Metal wastage in boiler tubes may occur in maay different ways.
It can occur from internal erosion and corrosion by steam and its impurities,
such as CI and 02' This is especially true when low-alloy tubes are used.
It can also occur through external abrasion, fouling, and corrosion. The
internal water/steam-side corrosion/erosion has been pretty well characterized
through the accumulation of experience in operating steam power plants.
132
-------
Thus, the present discussion will be restricted to the fouling, corrosion,
and erosion of tubes immersed in a fluidized bed.
The operating conditions
in the beds which affect corrosion, erosion, and fouling are summarized
in Tables 1-6 and 1-7.
Erosion of metal caused by dust particles is generally a
function of particle size, particle shape, particle physical properties,
impact velocity, impact angle, and the metal physical properties, as
discussed in the next section, on gas turbine blade erosion.
Although
the problem of erosion in the conventional coal-fired boiler is not as
severe as that in the gas turbine, the ash erosion problem does exist
in the conventional boilers and must not be ignored even though coal ash
particles may be exceedingly fine. The "hard" nature of p.f. coal ash,
cenospheres, makes the particles particularly erosive.
Where ash particles
are concentrated, as at turns formed by baffles within boiler banks,
erosion is a potential problem which can be remedied by limiting gas
veloci ties.
The maximum allowable gas velocity recommended for the
conventional pulverized coal-fired boiler is ~100 ft/sec, and that for
the coal-fired spreader stoker is ~60 ft/sec [90].
External erosion of the immersed tubes in the fluidized bed may
result from the vigorous particle motion which continuously scours the
tubes' surface.
Erosion per se is not considered important in the fluid
bed application where the maximum gas velocity in the boilers is 15 ft/sec,
and the particle velocity is considerably less than the gas velocity
due to the high frequency of inelastic particle collisions in the bed.
In addition, the vigorous particle motion may also be effective in
scrubbing off deposits formed on the heat transfer surfaces.
If the
fouling rate is higher than the eroding rate, the particle erosion
actually has an advantageous effect because it keeps scale from building
up at a high rate.
Corrosion will occur predominantly by chemical attack with
different mec~anisms such as surface oxidation, removal of the protective
scale on metal surfaces through chemical reactions with corrodents, and
133
-------
TABLE 1-6
OPERATING CONDITIONS IN FLUIDIZED BED BOILER DESIGNS
ATMOSPHERIC
PRESSURE PRESSURIZED
INDUSTRIAL UTILITY UTILITY
BOILER BOILER EOILEH
PRESSURE, ATM 1 1 10
BED TEMPERATURE, of 1300-2000 1300-2000 1300-2000
FT /SEC < < <
SUPERFICIAL GAS VELOCITY, -13 .-10 -10
MAXIMUM GAS VELOCITY BETWEEN 15 24 14
THE TUBES IN THE BED, FT/SEC
EXCESS AIR IN PRIMARY BED, % 10 10 8.5
DESIGN MEAN TUBE WALL TEMPERATURE,
of
Water walls 570 975 975
Pre-evaporator 710 732
Evaporator 570 850-890 975
Superheater 790 875-1175 900-1150
Reheater 1100 1122
134
-------
Table 1-7
ANALYSIS OF COAL ASH, COAL, AND. BED MATERIAL
COAL ASH. (Same for all Designs)
Compound
---c-
Si02
Fe203
P205
MgO
K20
Al203
Ti02
CaO
Na20
S03
COAL
Compound
-_J
Wt %
45.3
27.3
0.11
0.60
1. 80
21. 2
1.0
1.9
0.2
0.7
Wt %.
Chlorine
Sulfur
BED MATERIAL
0.05
4.3
Compound
Wt %
Dolomite 1337
Limestone 1359
Si02.
Al203
Fe203
MgO
CaO
Ti02
SrO
Na20
K20
Mn02
135
0.85
0.30 .
0.17
1.07
97.
<0.05
0.07
<0.02
<0.1
<0.05
. 0.78
0.15
0.25
45.0
53.0
0.02
<0.03
<0.02
<0.1
<0.03
-------
direct chemical attack on the metal surfaces.
The corrosion of the
immersed tubes in the fluidized bed is expected to be much less severe
than that obtained on the steam tubes of modern. high-temperature p.f.
boilers. due to the relatively low bed temperature employed. Figure 1-55
shows the variation with temperature of the vapor pressure of some of
the alkali metal salts usually present in coal ashes [9,28,45,79].
vapor pressure of the salts at normal flame temperatures of ~2500°F
is several orders of magnitude higher than that in a fluidized bed operating
at l400°F to l800°F. Also, the only alkali compounds having significant
The
vapor pressure at the fluid bed operating temperature are the chlorides
and oxides. Therefore, it is considered unlikely that carbonates or
sulfates, which are common causes of fouling and corrosion in conventional
boilers. will be a problem.
In addition, an operating temperature of
l400°F to 1800°F is considered too low to allow formation of sintered ash
deposits on the boilers through reactions of the volatilized alkali compounds
with aluminosilicates.
However. deposits of alkali chlorides on heat transfer
surfaces in fluid bed combustors burning high chlorine content coals can
lead to corrosion.
Chlorine has been known to contribute significantly
to the fouling and corrosion tendencies of coal when present in concentrations
higher than 0.3% [79].
The theoretical analyses indicate that corrosion. erosion. and
fouling will not be a problem for the proposed design conditions and that
higher steam temperatures are practical for fluidized bed boilers.
Experimental Data
Experimental data on corrosion, erosion, and fouling of boiler
tube materials are available from work conducted in England [34.69.70.71]
and the United States [76.97]. Work sponsored by OAP is being continued
by NCB and the U. S. Bureau of Mines, Morgantown, West Virginia. A
summary of the experimental work is presented, in Table 1-8.
Dainton and Elliott [22] immersed a stainless steel probe in a
fluidized bed burning Thoresby coal of high chlorine and sulfur content.
136
-------
Curve 643913-A
6
7
104/TO K
9
11
10
8
~
~
.s
"
"
." '\.
\~'\ KOH
\'\' ~
" ""
" """
\. \. ""
\\ '\ ~NaO~
, \\ '\.'\.
" " "
1\ ' "'' I\.
\ \. '\. ' '\.
\ \ \KCl\. '\.
\ NaC 1\ \ ~ '-
\ " "
\K2S04 \ ""
\ \. "''''
\\ \Na2S04 \ ~ \.
\\.K~03 \ \ ~ \.'\.
'n , " " "
~ ., \ , ""
~'\. \ \ '\.
\\ \. \ ~
\\ \ \
" " , "
\. " \ "
, '\. \ ,
~ \\ \
Na2COj\ ' \1 \
0).
~
:J
VI
VI
0)
~
c..
~
8-
ro
>
1400° C
900° C 800° C
Temperature
Figo 1-55-Vapor pressure of alkali compounds
present in coal
137
12
103
102
101
100
10-1
10-2
10-3
600° C
-------
TABLE 1-8
EXPERIMENTAL WORK ON FLUIDIZED BED BOILER TUBE MATERIALS
OPERATING
PRESSURE & NUMBER OF NUMBER NUHBER OF DURATION
FLUIDIZING BED t1ETAL MATERIALS OF COALS L I:rES TONES OF
INVESTIGATOR APPARATUS VELOCITY TEMP. of TEMP. of STUDIED STUDIED STUDIED TESTS, HRS.
Dainton & Ellio tt [22] 8" dia. 1 atm 1300 900-1200 1 1 0 120
Gliddon [34] 8" dia. 1 atm 1300 930-1+10 1 8 0 30-150
NCB [69, 70, 71] 48" x 24" 3-5 atm 1470 620-1470 7 2 2 100
2 ft/see
12" dia. 1 atm 1560 735-1290 6 1 1 100-500
3 ft/ see
t-' 27" dia. 1 atm 1522. 750-1290 6 1 1 500-1000
w (test not 8 ft/see
00
completed
yet)
PER [76] 72" x 18" 1 atm 1600 '\,600 (not primarily intended for eorrosio~ studies)
12-15 ft/see
USBM [97] 18" dia. 1 atm 1500-1600 600-1200 4 (work. not started yet)
-------
(Cl = 0.86 wt%; S = 1.06 wt%) at a temperature of 1300°F for 120 hours
and found that the metal surface at a temperature above 1100°F was free
of any deposit or corrosion. The fouling became progressively worse when
the metal surface temperature decreased below this level, due to the
increased temperature differential between the probe and the bed.
The
deposit was predominately sodium chloride. The same coal burned in a
test rig at BCURA simulating p~f. firing gave massive deposits on
superheater tubes which almost bridged across the tube bundle in a
24-hour experiment [28]. In the same series of experiments, Gliddon,
of the Central Electric Generation Board, England (CEGB) [34], combusted
four coals and four chars -- Thoresby, Babbington, Pye Hill, and
Oller ton -- of different sulfur and chlorine contents. He found that
the maximum fouling deposition rate for the high-chlorine Thoresby coal
(0.86 wt% Cl, 1.06wt% S) was approximately six times that for the
Babbington char (0.33 wt% Cl, 0.48 wt%S). The majority of the deposit
for Thoresby coal is sodium chloride, and the rate of deposition falls
off sharply at probe metal temperature above ~950°F, suggesting that
the dew point of sodium chloride at the operating condition is between 950°F
and 1100°F.
This may necessitate special start-up and shutdown procedures to
prevent deposition of sodium chloride.
Traces of sulfate were found in
Babbington deposits but none in any of the Thoresby deposits.
The results
also show that the rate of fouling can be qualitatively predicted by the
vapor pressure curve of sodium chloride. The typical fouling rate for
Thoresby coal is presented in Figure 1-56.
A 30-hour test with Pye Hill
coal (0.42 wt% Cl, 2.71 wt% S) indicated no visible corrosion of the probe.
The probe also showed no marked signs of corrosion after 150 hours in the
bed burning Ollerton coal (0.3 wt% Cl, 1.33 wt% S), although a very thin
adherent layer of scale observable under magnification was forming in
places along the probe.
The deposits contained significants of alkali
metals, but no chloride or sulfate was detected.
The experimental data
indicate that the initial rate of attack is often much higher (about 8.5 x 10-3
inches per 1000 hours for Thoresby coal) than that which occurs after a
long period of operation under constant conditions.
139
-------
Cur'le 643512-A
40
o
350
Potass iu rn (K20)
I
I
I
~ 30
N--
E
u
--
~
I-' .
~ .
o
~
c::
o
;;: 20
V)
o
c..
Q.)
o
-
o
Q.)
.....
~ 10
400
450 500
Probe Temperature °C
550
Thoresby Coal
Reference: 04)
I
I
600
.650
Figo 1-56-Effect of probe temperature on rate of deposition
-------
The most comprehensive and systematic studies on corrosion and
deposition of boiler tube materials are being conducted by NCB under a
contract with OAP [69,70,71]. The corrosion/erosion studies are divided
into three parts:
high-pressure operation, short-duration atmospheric
pressure operation at low velocities, and long-duration atmospheric pressure
operation, to assess the effect of limestone addition.
Corrosion/Erosion\of Boiler Tube Materials under Pressurized
Operation Conditions. The investigation was performed by combusting a
U. K. coal with a U. K. dolomite and by combusting a U. S. coal with a
U. S. dolomite in a 48 in
x 24 in
pressurized fluid bed at 3 to 5 atm
pressure. The U. K. coal chosen was Welbeck coal, with 0.53 wt% Cl and
1.25 wt% S; and the U. S. coal was Humphrey No.7, with 0.08 wt% Cl and
2.75 wt% S. Specimens of seven boiler tube materials (Table 1-9) formed
into continuous tubes were immersed in the bed and air-cooled to the
temperature range 620°F to l470°F. The test specimens were descaled and
weight losses examined after each run of ~100 hours duration at a
fluidizing velocity of two ft/sec.
The average weight loss for the
2
U. K. coal and dolomite experiments ranged from 2 to 445 ~g/cm hr,
depending on the temperature of the specimen I and the specimen tube
material. The weight losses for the U. S. coal and dolomite test were
2
9 to 361 ~g/cm hr.
The detailed data are presented in Table 1-10.
In
all cases, visual observation indicated that the color, adherence, and
type of deposits depended'on the material and temperature of the specimens.
Metallographic examinations found no evidence of intergrannular penetration
in any of the tested specimens, and the higher weight losses of the 2-1/4%
Cr-l% Mo, 1% Cr-l/2% Mo ferritic specimens were associated with a general
surface roughening. The nonuniformity of the weight losses can, in some
instances, be related to the position of the specimens relative to the coal
feeding nozzles.
Corrosion Studies of Boiler Tube Materials at Atmospheric
Pressure and Low Fluidizing Velocities.
Three series of test runs were
planned, two of which have been completed. U. S. Humphrey No.7 coal
was combusted in a l2-in rig at three ft/sec fluidizing velocity with
141
-------
TABLE 1-9
TYPICAL ANALYSES OF METAL SPECIMENS
Designation NOMINAL COMPOSITION %
Cr Ni Mo Mn Ti Al Nb Fe
12% Cr 12 88
Rf 36 18 12 1 69
SF 316 17 12 2.5 2 .66.5
Esshete 1250 15 10 1 6 1 67
P.E. 16 18 37 5 1.2 1.2 37
1%Cr 1/2%Mo 1 0.5 0.5 98
2 1/4% Cr 1%Mo 2.25 1 0.5 96.25
142
-------
TABLE 1-10
AVERAGE TUBE SPECIMEN WEIGHT LOSS UNDER
PRESSURIZED OPERATIO~ CONDITION, mg/cm2-hr.
t-'
~
v..>
TUBE MATERIAL
TEHPERATURE Esshete 1% Cr 2 1/4% Cr
RANGE, OF (a) 12% Cr RF 36 SF 316 PE 16 1250 1/2% Mo 1% Mo
Series B(b) Series U(c) B U B U B U B U B 'U B U B U
1350- 1470 - - 16 145 15 166 7 50 50 223 - - - -
1460 1450 63 80 37 207
1320- 1280- 5 18 8 48 8 54 5 36 28 210 - - - -
1380 1380 4 I
1040- 1100- 5 15 4 28 23 3 13 5 48 - - - -
1220 1180
730- 830- 6 29 13 29 11 33 6 17 9 - 373 254 445 319
920 1090
660- 620- 2 18 4 9 4 22 3 8 5 15 24 99 35 151
990 920 i
!
(a)
Temperature ranges are the temperature differences
between the ends of the composite tubes.
Series B is the result obtained from combusting a
U.K. coal and a U.K. dolomite.
(b)
(c)
Series U is the result obtained from combusting a
U.S. coal and a U.S. dolomite.
-------
'or without limestone addition.
Test series 1 was run for 100 hours
without limestone addition; test series 2, for 500 hours without limestone
addition; test series 3, for 500 hours with limestone addition.
Specimens
of tube form and coupon were used in the bed and in the freeboard.
The
available results (test series 1 and 2) are summarized in Tables 1-11 ~nd
1-12.
There was no marked difference between these results and the
similar tests under pressurized conditions reported above.
The high
chromium content alloys appeared satisfactory within the conditions
tested. However, the medium carbon and 2-1/4% Cr-l% Mo steels suffered
a higher rate of metal loss in 500 hours.
findings described in the last section.
This confirms the earlier
In all instances, substan.tially
less metal wastage was observed for the tubes and coupons located in the
freeboard.
Comparing the present results with those from the earlier work
burning a U. K. coal (Newstead coal with 0.6 wt% Sand 0.3 wt% Cl) for
100 hours and 1000 hours, it is evident that most of the metal loss
occurs in the first 100 hours.
The variation of metal loss with operating
time for these two coals is about the same and is shown in Figure 1-57.
..
The comparison also implies that there is no measurable difference in
the corrosive potential of these two coals.
Long-term Corrosion Tests at Atmospheric Pressure to Assess
the Effect'of Limestone Addition on Corrosion of Boiler Tube Materials.
Four series of tests with duration of operation up to 1000 hours were
run.
At the same time, the particle size of limestone and the combustion
conditions were changed to evaluate their respective effect on corrosion.
The available data using U. S. Humphrey No.7 coal without limestone
addition at a fluidizing velocity of 8 ft/sec are presented in Tables 1-13
and 1-14.
These tests showed that the medium carbon and 2-1/4% Cr-l% Mo
ferritic steel specimens had suffered internal attack on the steam side,
possibly by Cl and 02 in the steam supply. Thus, a correction of internal
weight loss is required. Again the data show that the austenitic steels
. .
perform satisfactorily; and of the chrome-ferritic steels, the 12% chromium
steel behaves the best.
The low-chromium steels suffer much higher weight
. losses.
144
-------
l.Ox
c
o
tJ 0.8 x
ro
'-
L.L.
V)
V)
o
.-.J
-
~ 0.6x
Q)
$':
c
Q)
E
u
Q)
~ 0.4x ..-
. 0.2 x
Curve 643506-A
o
100
. I
I
I
I
I
I
I
I
I
I
I
I
I
I
I
I
I
I
I .
I
I
I
I
I
I
Reference: (69)
500
Time, hr
Figu 1-57 -Effect of time on weight 1055
(100 & 1000 hr rates: Newstead coal)
(100 & 500 hr rates: Hu mphrey No.7 coal)
145
1000
-------
TABLE 1-11
SUMMARY OF RATE OF WEIGHT LOSS-~g/cm2h
TEST SERIES 1--DURATION 100 hrs
RATE OF WEIGHT LOSS - ~g/cm~h
LOCATION NOMINAL Type 347 Type 316 Esshete 12% Cr 2-14% Cr Medium
TEMPERATURE Austenitic Austenitic 1250 Ferritic Ferritic C
OF SPECIMEN Steel Steel Steel Steel Steel
of
BED
Tubes 750 1 1 67 74
1 1 83 100
3 109 90
111 101
923 3 153 125
2 3 154 192
2 3 198 143
210 190
1105 8 6 10 696
7 10 11 462
7 9 9
8 7
1275 10 9 17
12 13 21
11 13 24
11 24
11 19
Coupons 1560 12 9 20
FREEBOARD
Tubes 730 1 2 86 63
1 68 .66
2 65 62
78 57
1285 3 3 7
5 3 8
6 3 10
3 10
5 16
Coupons 1500 3 2 4
146
-------
TABLE 1~12
SUMMARY.OF RATE OF METAL LOSS~~g/cm2h
TEST SERIES 2 -- DURATION 500hrs
RATE OF WEIGHT LOSS - ~g/cm2h
LOCATION NOMINAL. Type 347 Type 316 Esshete 12% Cr 2-1/4% Cr Medium Nimonic
TEMPERATURE Austenitic Austenitic 1250 Ferritic Ferritic C PE 16
OF SPECIMEN . Steel Steel Steel Steel Steel
of
BED 735 2 2 1 34 73
Tubes <1 <1 2 29 35
<1 33 36
930 2 2 5 81 65
1 2 8 64 67
1 87 132
1095 4 4 5 459
2 3 5 5 793
3 3 365
1290 3 5 6 2 3
3 4. 9 3 3
4 2 5 5
Coupons
FREEBOARD
Tubes 755 <1 <1 <1 29 36
<1 <1 2 23 21
<1 23 18
1285 1 1 2 1 1
1 1 5 1 1
2 2 .6 3 2
Coupons
---
147
-------
TABLE 1-13
OBSERVED RATE OF WEIGHT LOSSES ON SPECIMENS
DURATION 147 HOURS
RATE OF WEIGHT LOSS (jJg/cm2h)
NOMINAL Type 347 Type 316 Esshete 12% 2 1/4% Medium
TEMPERATURE austenitic austenitic 1250 Cr Cr C
(OF) steel steel ferritic ferritic steel
750 8 14 7 21 248 A(a)
930 5 18 19 67 888 A
1110 11 13 20 38 A A
1290 11 16 15 18 A A
Note:
These data may be affected to an unknown extent by the violent bed
condition$, particularly temperature, pertaining to the shut-down. .
(a) Specimens suffer internal attack on the steam side,possibly by Cl- and
02 in the steam supply, so data require additional correction.
148
-------
TABLE 1-14
OBSERVED RATE OF WEIGHT LOSSES ON SPECIMENS
DURATION - 500 HOURS
RATE OF WEIGHT LOSS 2
(\J.g/ cm /h)
NOMINAL Type 347 Type 316 Esshete 12% 2-1/4% Medium
TEMPERATURE austenitic austenitic 1250 Cr Cr C
(OF) steel steel ferritic ferri.tic steel
750 7 2 2 3 30 41
4 4 2 4 53 55
1 1 4 96 614
8 5 4 111 80
4
7
930 7 4 6 6 301
3 4 9 6 A(a) A
9 6 7 9
7 8
11
1110 6 3 6 2
3 7 3
2 7 2
4 9 2 A A
4 8 5
1290 13 6 11 3
3 8 12 4
8 8 12 4
4 14 21 4
12 13 22 5
15 10 19 6
(a) Specimens suffer internal attack on the steam side, possible by Cl- and 02 in
the steam supply, so data require additional correction.
149
-------
.:A~phbughthe complete data for this systematic study of
corrosion/erosion on boiler tube materials are still not available, some
conclusions can. be drawn from the data summarized in Figures 1-58 to 1-63:
. The austenitic steels (Figures 1-61 to 1-63) give very small
weight loss «10j.lg/cm2 after 500 hours' operation) at. all
tested temperatures for lower chlorine content coals and at
fluidizing velocities up to eight ft/sec. Increasing the
. fluidizing velocity increases the rate of weight loss. .
. A high-chlorine coal is unlikely to cause corrosion of an
austenitic steel heat transfer surface immersed in a
fluidized bed if the metal temperature is kept above the dew
point temperature of sodium chloride in. the gas.
Sodium
chloride dew point temperature is determined mainly by the
chlorine content of the coal and the bed temperature which
determines its vapor pressure.
The dew point would rise as
the bed temperature is increased, thus increasing the
temperature range over which metal corrosion is likely.
. The chrome-ferritic and medium carbon steels are more
sus~eptible to attack, as shown by greater weight loss rates
(Figures 1-58 to 1-60), although the rate tends to decrease
with increasing operation time.
. Metallographicexamination shows absence of intergranular
penetration or pitting for t.ests now completed under normal
operating conditions.
all instances.
Slight oxide formation is present in
, Rate of weight. loss decreases with increase in operation
time under constant operating cond'itions.
. Substantially less corrosion was found for the tubes or
coupons situated in the freeboard as compared to that in the
bed.
150
-------
Curv8 643S01-A
lIXX)
CurvlI 643497-A 280
50
.93O"f at 2 ft/sec
0 75O"f at 3 ft/ see 240
NE 40 NE
~
j E 200
.; ~
.. <[
<
.. c
c: 30 ::> 160
::> f.
~
..
... ~
~
E \20
~ E '"
f20 ~
'"
'" >
> 'i
.. :; 80
::> E
E ::>
::> u
u 10
, 75O"F at 4ft/ see
. 93O"F at 4ft/ see
... llOO"f at 4ft/ see
. 75O"F at 8ttJ see
References: (69), 1101. 17\)
References: 1691,1701,1111
o
o
I!XXJ
500
Operati"9 Time. hr
o
500
Operating Time, hr
Fig. 1-58-Cumulative \Wight loss per unit area as a function 01 time for
medium C steel
Fig. 1-59-Cumulative weight loss per unit area as a function
of time for 2 )j4'j, chrome-ferr~ic steel
"'!:;
~ 5.0
7.0
Curve """,.98-A
6.0
:
<
'2 4.0
~
~
~ 3.0
~
..
>
..
~ 2.0
::>
U
12'10 Chrome-ferrilic Steel
: =] 4 ft/sec
... llOO"f
. 1285"f
~ 1= 18ft/see
D l285"f J
'"
~
...
,
References: 1691,1101, (7\)
---- Q---
--
..-,c-
/1'/" -----v-
,'/
"
,/
1.01:
I'
,
500
Operati"9 Time, hr
lIXX)
Fig. 1-6O-Cumulative \Wight loss per un~ area as a lundion
of time for 12'10 chrome-ferritic steel
-------
6
N
E
u
} 5
,.;
f
""
"E 4
~
~
E
.IZ' 3
N ~
~
]1
~ 2
E
~
u
Curve 64)499-A
Austenitic Type 316
. 750"f)
. 93O"F
AllOO"!' 4ft/see
. 1285"f
9 750"f)
o 930"f
"llOO"F 8ft/see
" 1285"f
References: (69/, 1101. (7D
00
500
~erating Time. hr
Fig. 1-61-{;umulative weight loss per unit area as a function
of time for austenitic type 316 steel
N
E
u
cr.
E 5.0
..
..
.:;:
1000
"E
:: 4.0
!.
~
:g, 3. 0
~
..
>
ii
"5 2.0
E
~
u
6.0
Curve 643S01-A
. 75O"f }
: l~~ 4ft/see
. l285"f
V 75O"f }
o 930"f
"1100"!' 8 ft/ see
" 1285"f
d
I
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,
,
I
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I I
I' ;'
I 'I
" ,'~' V
I ",
I ' I
I ' I
, I
" ,
/;' /
"
1.0
References: 1691. (701,1711
500
Operating Time. hr
Fig. 1-62-Cumulative weight loss per unit area as a function
of time for austenitic type 347 steel
1000
11
N
E
i
10
Curve 64]5QO-A
. 75O"f I .
. 930"f
A llOO"!' 4ft/see
. 1285"f
v 75O"f I
o 93O"F 8 ft/ see
" 1lOO"F
" 1285"f
Reference: (691, (701, (711
00
500
~ratlng Time, hr
1000
Fig. 1-6}-Cumulative weight loss per unit area as a function
of time for austenitic esshete 1250 steel
~ 8
""
. %:
c::
:>
....
:!:. 6
is
....
E
.IZ'
~
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4
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-------
Boiler Tube Materials Selection
Boiler tube materials were selected for the present fluidized
bed boiler designs on the basis of conventional practice.
The selected
tube materials and their respective chemical compositions are summarized
in Tables 1-15 and 1-16.
Since the materials were selected, results from
the NCB have become available, and the materials selected will be
reassessed when results from the NCB and USBM tests are complete.
Preliminary evaluation indicates that the conventional boiler tube materials
can be used in the fluidized bed application if the long-duration runs
, show that the rate of weight loss does indeed level off with increase in
operation time.
Final assessment can only be carried out after more data
are accumulated at higher velocity, higher pressure, high temperature, and
longer duration.
Assessment
Additional data on corrosion, erosion, and fouling of boiler
tube materials are needed to evaluate critically the present fluidized
bed designs.
. Longer-duration runs (up to 2000 hours) should be performed
at a higher pressure (up to 10 atm) , and a higher velocity
(up to 15 ft/sec) to evaluate the rate of weight loss with
respect to operation time.
. The effect of limestone addition and the effect of burning
different coals with different impurities, especially the
high-sulfur coals, should be carefully studied.
The corrosion
rate from burning a U. S. coal with a U. S. limestone has been
up to seven times larger than the corrosion rate from burning
a U. K. coal with a U. K. limestone (see Table 1-10).
. The effect of thermal cycling in the fluidized bed should be
evaluated.
The possibility of increasing corrosion due to
deposition of sodium chloride during start-up and shut-down
should be studied.
153
-------
Table 1-15
FLUIDIZED BED CO}ffiUSTION BOILER DESIGNS
BOILER TUBE MATERIALS
INDUSTRIAL BOILER ATMOSPHERIC PRESSURE BOlLE PRESSURIZED BOILER
Design Temp. Design Temp. Design Temp.
Material of Material of Material of
Water walls (sA-l ~~-A J 570 SA-2l3-T22 975 SA-213-T22 975
Pre-evaporator SA-2l0-Al 710 SA-2l0-Al 732
tA-l92
Evaporator 570 SA-213-T2 850-890
Superheater SA-2l3-T22 790 . SA-2l3-T2 875-930 SA-2l3-T2 900
I-' .
V1 SA-213-T22 1075 SA-213-T22 1058
~
SA-2l3-TP-304H 1175 SA-213-TP-304H 1150
Reheater SA-213-T22 1075 SA-213-T22 1122
SA-2l3-TP-304H 1175
-------
TABLE 1-16
FLUIDIZED BED COMBUSTION, BOILER DESIGNS
CHEMICAL COMPOSITION OF SELECTED BOILER TUBE MATERIALS
(in wt %)
NICKEL
CARBON MANGANESE PHOSPHORUS SULFUR SILICON
SA-2l0-A1 0.27 max 0.93 max 0.048 max 0.058 max 0.10 min
SA-213-T2 0.10-0.20 0.30-0.61 0.045 max 0.045 max 0.10-0.30
SA-2l3-T22 0.15 max 0.30-0.60 0.030 max 0.030 max 0.50 max
~ SA-213-TP304H 0.04 max 0.03 max
V1 0.04-0.10 2.00 max 0.75 max
V1
welded SA-178-A 0.06-0.18 0.27-0.63 0.05 max 0.06 max
tubing (low carbon steel)
seamless SA-192
tubing
0.06-0.18
0.27-0.63
0.048 max
0.058 max
0.25 max
CHROMIUM
0.50-0.81
1. 90-2.60
18.00-20.00
MOLYBDENUM
0.44-0.65
0.87-1.13
8.00-11.00
-------
Gas Turbine Erosion and Corrosion
Introduction
A gas turbine is used in the fluidized bed boiler combined
steam-gas turbine system to extract energy from the high-pressure, hot
combustion gases leaving the boiler. The gas turbine in a typical
application provides 15 to 20% of the net electrical output.
The gas which this turbine expands consists typically of
-2 .
74 mole% N2' 15 mole% C02' 2 mole% 02 and 1.7 x 10 mole percent S02'
It contains dust loadings of 0.1 to 0.15 grains per standard cubic foot.
The dust composition is approximately 80% coal ash, 7% carbon, and 13%
spent limestone sulfur sorbent.
The anticipated particle size
distribution is given in Table 1-17.
TABLE 1-17
ANTICIPATED PARTICLE SIZE DISTRIBUTION OF DUST ENTRAINED
IN GAS TURBINE FEED
AVERAGE'
PARTICLE DIAMETER
(MICROMETERS)
WEIGHT PERCENT
IN- SIZE RANGE
WT%
CONCENTRATION
GRAINS PER SCF
less than 1 ~m
1-2 ~m
2-3 ~m.
3-4 ~m
4-5 ~m
5-6 ~m
6-7 ~m
76.8
16.6
4.7
1.6
0.2
0.1
0.0
0.08-0.12
0.02-0.03
0.005-0.008
0.002-0.003
0.0002-0.0003
0.0001-0.0002
-------------
Two-stage separation, final collector advanced design
Aerodyne cyclone .
156
-------
This dust-containing gas will be at a pressure of approximately
10 atm and at a temperature of l600°F. Approximately 50% of the dust
will have been previously exposed to 2000°F temperature in the carbon
burn-up cell or the sorbent regenerator.
The balance of the ash will have
experienced a maximum temperature of l750°F in the fluid bed.
These
temperature maximums affect both the erosive properties of the solid
dust particles and the corrosive nature of the gas stream.
of turbine operation are of special concern:
Three aspects
.
Erosion of turbine rotor blades and stator vanes or nozzles
by the entrained dust
.
Corrosion of the blades by reaction of the hot sulfur- and
oxygen-containing gas with the blade alloys, accelerated by
sodium and potassium sulfates derived from the ash
.
Deposition of material on the blade and vane surfaces,
resulting in loss of efficiency of turbine stages.
Gas turbine operation on other than clean fuels -- natural gas
or refined fuel oils -- has provided many examples of problems arising
from erosion, corrosion, and blade fouling.
It is important to keep
in mind the differences in particle size distribution, physical and
chemical nature of the ash (resulting from exposure to 1750 to 2000°F
fluid bed temperatures rather than 2500 to 3000°F gas turbine combustor
temperatures), and the smaller excess air ratios (10 to 1 instead of
50 to 1) while assessing the difficulty of efficiently operating gas
turbines in the fluidized bed boiler power plant.
Erosion
Serious erosion of gas turbine rotor blades has been experienced
in the operation of:
.
Aircraft gas turbines which ingest large quantities of
dust-laden air [67,91]
157
-------
.
Gas turbines combusting pulverized coal [5,86,98]
.
Gas turbines used as power recovery units expanding hot
gases from catalytic cracking operations [93]
.
Gas turbines combustingand recovering energy from blast
furnace gases' [47] .
Operating experience with turbines in these applications and
laboratory studies of blade erosion [33,62] suggest several conclusions.
The majority of particles smaller than two micrometers in
diameter can be expected to pass through the turbine without impacting
on the blades.
Those particles that do impact on the blades will have
a polishing rather than an erosive effect.
[62,67,93]
However,
particles larger than 10 micrometers in diameter are erosive.
They
impact on the blading with velocities and impact angles that are
damaging [33,62]. These larger particles can be effectively removed by
well-designed separators.
No particles larger than six micrometers are
expected in the gases leaving the carbon burn-up cell using advanced
design Aerodyne separators.
(Table 1-17)
Particulates entrained in the off-gases of the fluidized bed
boiler have been subjected to lower temperature than those encountered
by the ash in conventional gas turbine combustors.
the ash has seen temperatures approaching 2000°F.
Only 50 percent of
This tempera ture is
below the initial softening temperature of the vast majority of U.S.
cqals [1].
These particles have not fused to hard, potentially erosive
glassy particles.
Electron micrographs of ash particles from the, British
Coal Utilization Research Association's fluid bed pilot units showed that
the larger particles were platelets of low mechanical strength. Although
this bed operated at l500°F, increasing bed temperature to 1750°F would
not be expected to change materially the character of the ash, since the
initial ash deformation and softening temperatures are not approached.
Ash from the fluidized bed boiler is expected to be less erosive than
the fused ash resulting from conventional combustion processes.
The
ash should also be less erosive than the catalyst particle carry-over
158
-------
responsible for erosion in oil refinery gas turbine applications. In this
application, the use of three st~ges of mechanical dust collection resulted
in turbine operating life in excess of four years [93].
Finally, dust loadings of 0.01 grains per standard cubic foot
are projected to give acceptable turbine life, based on testing of the
Locomotive Development Committee's coal-burning gas turbine.
This unit
was operated for over 4000 hours at Dunkirk, N.Y. and later for 2000 hours
at the Bureau of Mines, Morgantown, W.Va. [86];
Tests were carried
out using combusted pulverized coal having the particle size distribution
and concentrations shown in Table 1-18.
TABLE 1-18
ASH PARTicLE SIZE AND CONCENTRATIONS USED TO ESTABLISH PEIDIISSIBLE
LOADINGS IN BUREAU OF MINES COAL-FIRED GAS TURBINE RESEARCH PROJECT
PARTICLE SIZE wr. PERCENT CONCENTRATION
(MICROMETERS) (GRAINS PER SCF)
less than 2 ------
2-5 ~m 73.5% 0.088
5-10 ~m 16.8 0.020
10-20 ~m 7.3 0.009
20 ~m 2.4 0.003
100.0 0.120
Under a loading of 0.12 grains per SCF, the useful life of
rotor blading was projected, based on 2000 hours of testing, to be
20,000 to 30,000 hours. Stator vane useful life was projected to be
5,000 to 7,000 hours.
It was concluded that a factor of ten reduction
in dust loading to 0.01 grains per standard cubic foot should give
acceptable blade and vane life.
159
-------
The Bureau of Mines did carry out erosion tests at their
projected acceptable ash loading of 0.01 grain per standard cubic foot.
Unfortunately, these tests were carried out at gas velocities unrealistically
high for utility gas turbines and were made with 94 weight percent of the
ash larger than 5 micrometers in diameter.
A 1000-hour test under these
conditions resulted in erosion of both the front and back surfaces. of the
blades [98].
The Bureau's original projection of an acceptable dust loading
of 0.01 grains per standard cubic foot, especially in view of the less
erosive nature of fluidized bed ash, is considered valid for the fluidized
bed boiler system.
Erosion damage should not be a serious problem in turbines
used with the fluidized bed boiler system provided that the dust loadings
anticipated in Table 1-17 can be maintained.
The concentration of parti-
cles larger than two micrometers in diameter should be between 0.008 and
0.012 grains per standard cubic foot.
No particles larger than ten
micrometers in diameter are expected to pass the collection system.
Careful design of the turbine to avoid concentrating the ash
through the action of secondary flows within the turbine is essential
to avoid erosion damage.
Entrainment of ash in the wake formed at the
stator vane trailing edges, resulting in ash concentration near the axis
of the machine, caused erosion of rotor blade r,oots at their leading
edges in both U. S. Bureau of Mines and Australian tests [5,86].
As a result of operating experience, the Bureau of Mines had
their coal-fired turbine redesigned to incorporate:
!
.
An annulus immediately following the first-stage stator vanes
formed by omitting the first row rotor blading, to provide
space for centrifuging ash particles to the outer wall of
the turbine
160
-------
.
Thickened trailing edges of the rotor blading to encourage
radially outward secondary flow in the rotor blade wakes
.
Thinned trailing edges of stator vanes, to inhibit radially
inward directed flow in the wakes of these vanes
.
Added vanes to the stator rows, to minimize flow separation
on the faces of the vanes.
These flow separations had the
effect of contributing ash to the secondary flow directed
toward the rotor axis.
.
Stepped side walls so that the tips of the stator vanes ended
below the rotor blade platforms and tips of rotor blades
extended above the stator vane roots
.
Titanium carbide wear rings, to armor the rotor blade and
stator vane roots against wear.
The Australian turbine designers redesigned their turbine to
overcome erosion problems by:
Corrosion
.
Adding a boundary layer fence, mid-span on the trailing edge
of the stator vanes, to break up secondary flows
.
Extending stator vane tips below the platform of adjacent
rotor blades, so that ash leaving the stator blade tips
would be diluted with cooling air and swept under the rotor
blade root platforms
.
Increasing turbine diameter to lower the velocity of gases
through the turbine.
Because erosion damage is a strong
function of velocity (damage proportional to velocity raised
to a power greater than two), lowering velocity was considered
the most important step taken to overcome erosion problems.
Accelerated "hot" corrosion of gas turbine blades results
from the formation on the blade surfaces of liquid films which react to
161
-------
remove the protective oxide scale and to raise the sulfur activity .in
.contact with the blade alloy to the point where chromium sulfide can
form. Eventually the local level of chromium in the alloy drops to a
level that permits the formation of low-melting nickel sulfides.
The
. ~lade loses mechanical integrity and is opened to catastrophic oxidation.
[11,24,36,72,83].
Sodium sulfate liquid films, formed by the reaction on blade
surfaces of sulfur dioxide-trioxide mixtures arid sodium chloride vapors,
cause hot corrosion attack in turbines operating with inlet gas
temperatures above lSSO°F (840°C) [23]. (The melting point of Na2S04
is1623°F, 884°C.) In the presence of gaseous potassium compounds,
reaction on blade surfaces to liquid films could take place at tempera-
tures as low as 930°F (500°C) [9].
The fluidized bed contains both sodium and potassium compounds
which will exert small, but potentially harmful vapor pressures at the
bed operating temperature.
Operation of the bed at l7S0-.l800°F lowers the vapor pressur-es
exerted by the compounds approximately three orders of magnitude below
.those expected in conventional boiler furnaces .
In addition" the 'fluid
.bed provides a large surface area of alumino-silicate ash and long
residence times to encourage in-bed reactions.
These reactions should
act to lock up sodium and potassium in solid compounds, reducing the
quantity of corrosive material leaving the bed.
Detailed studies of the
reactions between the bed materials and volatile sodium and potassium
compounds are needed to provide quantitative data for design. It may
be possible to eliminate the danger of hot corrosion and fouling lfthe
ac tualpartial pressures of the harmful condensable vapors can be reduced
through gas-solid reactions below those allowing condensation in the
turbine.
In the wors.t case,limita.tion of the temperature of the inlet
gas to the turbine may be used to prevent liquid film formation on
162
-------
turbine hardware.
Reduction of gas inlet temperature reduces the
efficiency of the turbine, which in turn lowers the overall plant
efficiency. Should it be necessary to limit the turbine inlet tempera-
ture to l400°F, the turbine efficiency would drop from approximately
28 to 25%, resulting in about a one percent reduction in overall plant
efficiency.
Combining a carefully engineered fluid bed boiler and dust
collection system with an industrial gas turbine employing sulfidation-
resistant high-chromium alloys in its construction should allow
operation of the turbine without serious corrosion problems.
163
-------
INDUSTRIAL BOILER APPLICATION
Better methods are needed for combusting the vast reserve of
coal to meet the ever-increasing energy demand in the United States.
Fluidized bed combustion stands out as a technology with the potential
to combust different grades of coal. efficiently and simultaneously to
conform to increasingly stringent air pollution control regulations.
Pollution from industrial boilers is not to be disregarded, especially
that of SO .
x
total SO emissions and 11% of total NO emissions from power/steam
x x
generating facilities.
As a whole, industrial boilers emit about 14% of the
Industrial boilers are generally considered nonprofit equipment
for which the desirable characteristics are high availability, low
maintenance, and economy. Therefore, an industrial fluidized bed boiler
that can be shop-fabricated and transported by rail is the most promising.
The preliminary design of an industrial fluidized bed boiler was begun
with these characteristics in mind. Different conceptual designs were
evaluated in conjunction with the Erie City Energy Division of Zurn
Industries, a Westinghouse subcontractor.
From this preliminary
evaluation, two systems were selected for detailed design and cost
analysis: one-through dry solid and wet scrubbing. Costs of other air
pollution control alternatives were also derived and compared with these
two systems.
164
-------
Design Basis
System Specifications
.J
The sys tem specifications for the indus trial flliidized bed
boiler are summarized in the first specification list.
The plant
capacity and steam conditions were selected on the basis of the market
survey, which projects that 40% of new units in 1980 will be in this
category. At the chosen capacity, the boiler remains shop-fabricable
and railroad-transportable.
Fuel and Limestone Specifications
The ultimate analysis and proximate analysis of the Ohio
Pittsburgh No.8 seam co~l are presented in the second specification
list. The particle size distribution of feed coal as crushed from
Koppers' reversible hammermill is shown in Figure 1-64.
The ash residual
after combustion in the fluidized beds wasassUilled to be 100% elutriated
on the basis of experimental evidence from the U. S. Bureau of Mines'
laboratories at Morgantown~
The particlesize-fluidizirig velocity
relationship is presented in Figure 1-65.
Limestone No. 1359, with composition analysis as shown in
Table 1-19, was chosen for the industrial fluid bed application.
Design
The designs for the fluidized bed combustor (FBC) and the
carbon burn-up cell (eBe) are summarized in Table 1-20.
The physical
dimensions of the boiler were kept within the railroad transportation
limitations, which are roughly 13 feet wide, 16 feet high, and 40 feet
long for most eastern United States locations.
165
-------
SYSTEM SPECIFICATIONS FOR INDUSTRIAL BOILER
PLANT CAPACITY & STEAM CONDITIONS
250,000 1b/hr steam at 600 psig and 750°F
CYCLE CONDITIONS
Gas
-- air inlet temperature
80°F
stack gas temperature
350°F (once-through)
>130°F (wet scrubbing)
Water -- water inlet temperature to economizer
250°F
water outlet temperature from economizer
350°F
PARTICULATE REMOVAL REQUIREMENT
Flue gas particulate loading
0.01-0.05 gr/SCF
SULFUR REMOVAL REQUIREMENT
>90%
TURN-DOWN REQUIREMENT
Turn-down ratio
>4:1
Dynamic response
~5%/min
166
-------
OHIO PITTSBURGH NO.8 SEAM COAL
(Source of data:
USBM, Pi ttsburgh. Pa.)
SAMPLE:
Run of mine - as received
PROXIMATE ANALYSIS (wt %):.
ULTIMATE ANALYSIS (wt %):
(includes moisture)
GROSS HEATING VALUE:
NET HEATING VALUE:
ASH ANALYSIS (wt %):
FUSIBILITY OF ASH:
PARTICLE DENSITY:
GRINDABILITY (Hardgrove):
FREE SWELLING INDEX:
Mois ture
Volatile matter
Fixed carbon
Ash
C
H
o
N
S
Ash
13000 B tu/ 1b
12500 Btu/1b
Si02 45.3
A1203 21.2
Fe203 27.3
Ti02 1.0
P20S 0.11
CaO 1.9
MgO 0.6
Na20 0.2
K20 1.8
S03 0.7
100.1
71.2
S.4
9.3
1.3
4.3
8.5
100.0
3.3
39.5.
48.7
8.5
100.0
(~60% organic & ~40% pyritic)
Initial deformation temperature -- 2080°F
Softening temperature -- 2230°F
Fluid temperature -- 2420°F
Coal -- ~1.4 gm/cc
Ash --~2.8 gm/cc
SO-60
5-5.5
167
-------
.1
1
5
15
25
- 35
~ 45
Q,) 55
:5:
. ;:; 65
~
Q,) 75
N
~ VI
0\
(X) 10-
Q,) 85
<3
I
~
95
96
98
99 10
100 1000
Particle Size, Microns
10,000
Fig. 1-64-Size consists of coal crushed b{Koppers Reversible Hammermill
.5
3
10
lU
30
40
50
ro
70
80
90
96
97
98.5
-------
"-
Q)
.....
Q)
E
(1:J
CI
~ Q)
\!) -
U
:;:;
"-
. (1:J
a..
10,000
VI
c:
o
"-
u
E
1000
100
10
0.001
Curve 594087-A
Basis
Temperature
Pressure
Solid Density, Ps
Gas Density, Pg
Gas Viscosity, ~
1700°F
1 atm
1251b/ft3 (limestone)
0.0189 I bItt 3
0.045 cp
\i\\'\
'-1 ~\
. ~\i\ ~~
~\'V.....
'V.~
.~\~
*-\"
0.01
0.10 1.0
Actual Gas Velocity, ttl see
10.0
100
Fig. 1-65-Fluidizing velocity and terminal velocity for 1 atm system
-------
TABLE 1-19
ANALYSIS OF LIMESTONE NO. 1359
COMPONENT WT % AS RECEIVED(a)
SiOZ 0.85
AlZ03 0.30
FeZ03 0.17
MgO 1.07
CaO 97.
TiOZ <: 0.05
SrO 0.07
NaZO < O.OZ
KZO < 0.1
MnOZ < 0.05
(a) .
R. W. Coutant, J. S. McNulty, R. E.
Barrett, J. J. Carson, R. Fischer,
and E. H. Lougher; "Investigation of
the Reactivity of Limestone and
Dolomite for Capturing SO from Flue
Gas," August 1968 (Battelte Memorial
Ins ti tute) .
170
-------
. i-'
-...J
i-'
TABLE 1-20
DESIGNS FOR INDUSTRIAL FLUIDIZED BED BOILER
FLUIDIZED BED COMBUSTOR (FBC)
CARBON BURN-UP CELL (CBC)
. 0
Bed Temperature, F
1650
1900-2000
Superficial fluidizing velocity, ft/sec
12.5
10
Bed depth (expanded), in
30
30
Excess air, %
10
50
Carbon burn-up efficiency, %
87
90
Bed material
. Limestone No. 1359
Limestone No. 1359
Gas side heat transfer coefficient,
50 in the bed
50 in the bed
BTU/ft2-hr-OF
40 in stop-splashing screen
Sulfur r=.moval schemes
Wet scrubbing or once-through
Stone
Limestone No. 1359
Ca/S ratio
1.2/6
-------
The size distribution of the particles e1utriated from the
FBC and CBC, shown in Figure 1-66, was also projected on the basis of
experimental data accumulated in the laboratories of the National Coal
Board (NCB) and the British Coal Utilization Research Association (BCURA)
in England, and of the Office of Air Program (OAP) contractors in the
United States. This'data correlation was discussed on pages 119-131.
172
-------
. 1
1
5
15
25
....... 35
.c 45
.~
~ 55
>- 65
..c
~
Q) 75
N
VI
~ 85
I-' Q)
...... >
w a
,
0::: 93
96
98
99
Curve 643513-A
I I I I I I I I 1 1 I I I I I I
e- -
- f--- f--- -- -- - ---.- --- ~ :;; :;;--
e- /-- ~ -
---- -- - .... ~
I-- ~ '" -
I, !)
I-- /' i-" ,,,,, )< ~ -
-- -- L ~ " ~---
I-- )(~ -
-- f-- ~
e- ~ i'\ ~ -
1--
I-- /. ;;'" )...~ -
e- 'Y', ).~ " ~ -
...' "
I-- /" ~ ~ " -
'" '"
,.t1/' ,'" -
tI'" ~
I-- ...k " ~ -
". ' " "
CBC V x. ,I" Coal: -1/4" x 0
1./ / ~}
I-- -
Gas Velocity: 10 -15 fps
V .., XXXX Represents Data Spread-
I-- / /
" V I I 1 I I T T 1
I-- /f / References: (42), (55), (76) I -
FBC ( -
I-- 'i
I I I / I I I I I I I I I I I
1
10 100
Particle Size, microns
1000
Fig. 1-66 -Projected particle size distribution for material elutriated
from fluidized bed combustor (FBC) & carbon burn-up cell (CBC)
-------
Boiler System Design
The overall boiler system design is outlined here.
The flow
sheets and energy/mass balances are summarized on pages 174-177
Different boiler design concepts are discussed on pages 177-183
and the final design selected is described on pages 183-191
Other
auxiliary systems include the coal and limestone feeding system,
particulate removal system, and air pollution control systems; these
are presented on pages 191-195 , and 196-198
detailed design is presented in Appendix D.
respectively.
The
Boiler System Schematic
The overall flow sheets for the two chosen systems -- once-
through dry solid and wet scrubbing -- were simplified and presented in
Figures 1-67 and 1-68.
The detailed schematics are included in Appendix D.
For the once-through dry solid system, six times stoichiometric
limestone is fed to the primary bed and calcined. Although this will
not be favorable in industrial applications, it is included in this.
analysis for comparison for two reasons.
First, the waste solid rate
is comparable to that of the wet scrubbing system as shown in Table 1-21.
Second, for the once-through limestone application large quantities of
stone, on the order of six times stoichiometric limestone, are required
" for 90% S02 removal at design conditions. Ninety percent of S02 will
be removed in the bed; this represents ~16% stone utilization. The
limestone is used once~through without regeneration.
Carbon-rich ash
is recycled back to the carbon burn-up cell.
The flue gas passes through
a superheater, convection zone, economizer, and electrostatic precipitator
with an efficiency of > 99.5%. The particulate loading in flue gas
is therefore expected to be below 0.02 gr/ft3.
174
-------
r-'----
I
I
i
I
I
L------
---------
Economi zer
f--- - D
Super-.
- Convection Heater ..
Zone
----
i-~r
I - I.
I I
t :.
, I
I
1
I
I
I
.-L
Carbon
Bu rnup
Cell
Fluid-Bed
Combus tor
Fig. 1-67 - System Schematic - Once-Through Scheme
,,~Ll m ----
S tack Gas
r----------------
I
I
i
I
!
~<"-""'
'I
!
L....:'----- Convection
Zone
Leqend
Ga.
501 ids
Limes tone
Slurry
----
,',',',',',,','
-.-.-
Heater
------------,
------.-----,
I
L___-
Make-Up Water &
Recyc I ed liquor
r--
I
I
I
)
I
...
I
I
I
i
i
-1..
Carbon
Burnup
Cell
Fluid-bed
Combustor
Fig. 1-68- System Schematic - Wet Scrubbi ng Scheme
175
Dwg. 2957A32
Legend
Ga.
----
$01 ids
Limes tone
,i ,i ,i ,i
Fresh Limestone Feed
To Di sposal
Dwg. 2957A31
,K- '1
.!,
I
!
!
I
,
I
'f'
I
i
'.
I
"
r
..-..t1
...
To Waste
Pond or
Vacuum F i 1 ter
-------
TABLE 1-21
COMPARISON OF BOILER PERFORMANCE
DRY SOLID
WET SCRUBBING
COAL FEED RATE (lb/hr)
LIMESTONE FEED RATE (lb/hr)
WASTE SOLID RATE (lb/hr)
BOILER EFFICIENCY (%)
TOTAL LOSS (%)
36,650
21,450
17,300
83.42
16.58
25,900
. 4,200
16,050 .
85.92
14.08
Loss due to calcination and S02
absorption, % .
Loss due to water and hydrogen
in fuel, %
2.74
0.56
Loss due to unburned carbon, %
Loss due to radiation and unaccounted
for % .
,
3.80
1.93
3.83
2.05
1.50
1.50
Loss due to evaporation in wet
scrubber, . %
Loss due to flue gas sensible
h ea t, %
4.84
6.61 (flue gas.
. temperature
= 350°F)
1. 30 ( fl ue gas
temperature
= 132°F)
176
-------
For the wet scrubbing system, 1.Ztimes stoichiometric lime-
stone is fed to the primary bed and calcined. The calcined limestone
is then mixed with water to form slurry and pumped to the wet scrubber.
The fluidized bed boiler design is essentially the same as that in the
once-through sys,tem.
The only difference is that the flue gas, instead
of passing through the electrostatic precipitator, goes to the wet
scrubber for SOZ and particulate removal. The stack gas heater is
olnitted in the present design, so the stack gas temperature is ~130°F.
Energy and Material Balances
Energy and material balances for both once-through dry solid
and wet scrubbing systems are shown in Figures 1-69 and 1-70. The
comparison of the boiler efficiencies of these two systems is summarized
in Table l-Zl. Heat generation from SOZ-CaO reaction is assumed to be
6 .
3.0 x 10 BTU/ton CaS04 produced; heat absorption from calcination of
6 .
CaC03 is assumed to be 1.5 x 10 BTU/ton CaC03 calcined.
The wet scrubbing system has higher boiler efficiency because
the stack gas temperature is l3zoF, compared to 350°F in the dry solid
system. An additional efficiency loss of about 5% in the wet scrubbing
system is expected if the flue gas is to be heated from l3ZoF to 350°F.
Industrial Boiler Design Concepts
A shop-fabricated and railroad-transportable industrial boiler
is highly desirable. Rectangular beds are therefore more suitable for
industrial boilers, for they make the best use of the available shipping
clearance. A horizontal configuration in which fluidized beds are
placed next to each other has clear advantages over the vertical con-
figuration.
First, during start-up or load change, solid circulation
from an operating bed can be used to light off the idled beds and makes it
unnecessary
to provide large ignitors for individual beds.
Second, main-
tenance is easier, and fewer support structures are needed.
Submerged heat transfer surface as vertical or horizontal
tube design were both considered in the original evaluation.
The vertical
177
-------
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tube design used natural water circulation,while the horizontal tube
design used forced water circulation.
Vertical tube design was rejected
when the preliminary cost analysis showed that its cost was about twice
that of the horizontal tube design, primarily because of the large
number of welds required. (See Appendix D).
The superheater in the selected design is conventional and is
placed in the convection section.
Putting a superheater in the carbon
burn-up cell is considered inadequate for two reasons.
First, to maintain
carbon burn-up efficiency in the CBC, the bed temperature has to be
kept constant, even during turn-down. With the boiler load decreased,
the recycled carbon fines back to the CBC decrease. Therefore, the
bed temperature will drop unless the bed depth is simultaneously
decreased. The regulation of bed depth involves costly equipment and
complex operation.
In addition, for a load reduction to 25% of full
load, the bed depth would be too shallow for efficient operation. Second,
additional fresh coal can be. fed into the CBC to keep the bed temperature
constant without changing the bed depth.
For a load reduction to 25%
of full-load capacity, the steam flow in the superheater tubes is just
about 25% of the full-load rate. Thus, the steam would be ,superheated
to a much higher temperature and the tube metal temperature would be
excessively high, requiring the use of expensive, high-grade alloy
material.
Through the same kind of preliminary concept evaluation for
each individual boiler component, a final boiler design was selected.
Selected Steam Generator Design
After evaluation of different boiler design concepts on the
basis of presently available information, a boiler design with the
overall isometric view shown in Figure 1-71 was chosen. The sectioned
view is presented in Figure 1-72. The boiler comprises two separate
shipping modules: the primary module (overall dimensions, 10 ft x 14 ft
x 40 ft) consists of four fluidized beds with air plenum sectionalized
in~o four sections, but not the bed interiors; the secondary module.
183
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(overall dimensions» 8 ft x 15 ft x 40 ft) houses the carbon burn-up
cell» superheater, and the convection pass.
The carbon burn-up cell is
a fluidized bed design; the superheater and the convection pass are
conventional designs. The two modules will be shipped separately to
the construction site for installation.
presented in Appendix D.
The detailed boiler design ,is
Boiler.
One-inch horizontal tubes are used in the bed with
horizontal spacing of 6 in and vertical spacing of 3 in. in staggered
positions. A recirculation pump provides the forced circulation
of water through the tubes for heat removal. A particle splash screen
three tube-rows in depth is incorporated in the design using one-inch
tubes on a 3 in x 3 in spacing. Free volumes are provided between
the submerged surface and the grid (12 in in depth)>> and between the
top row of the submerged' surface and the stop-splash screen (8 in
in
depth). The freeboard above the stop-splash screen is left empty for
gas passage and for easy access to boiler interior. The use of one-
inch tubes instead of larger tubes in the bed has the following design
advantages: smaller vertical headers can be used; fewer pipes are
required to collect the smaller vertical headers into the steam drum»
and this in turn means fewer on-site welds; larger heat transfer surface
area can be submerged in the bed with a constant tube pitch/tube
diameter ratio.
Two-inch tubes on three-inch centers are used for water walls
with natural circulation.
The same water walls are also used to cool
the grid plate.
The high carbon content ash collected in primary cyclones
will be recycled back into the carbon burn-up cell running at a higher
temperature and excess air.
Carbon Burn-Up Celi (CBC).
The presently proposed carbon
burn-up cell design was based primarily on turn-down considerations.
During turn-down» the bed temperature in the CBC has to be kept constant
for efficient carbon burn-up.
This becomes a problem in actual operation.
189
-------
if the CBC is used for superheating the steam.
When the boiler is
turned down, the total input of carbon fines into the CBC decreases.
There are three ways to maintain the bed temperature constant in this
turn-down condition: 1) lower the bed depths to decrease the total
heat transfer surface in the bed; 2) feed additional fresh coal into
the CBC while keeping the bed depth constant; 3) sectionalize the CBC
so that sections can be turned off corresponding to the operation of the
primary beds. However, methods 1 and 2 raise problems. For example,
for a load reduction to 25% of the full load, the superheater metal
temperature would be very close to the bed temperature because 25% of
the original steam flow would be running through the superheater. Method
1 would require reduction of the bed depth to almost 25% of the designed
bed depth -- too shallow to operate efficiently.
The logical solution is to divide the CBC into four separate
sections (each corresponding to a primary bed) by installing water tube
screens in the bed and by subdividing the air plenum. Turn-down by
shutting down a bed would require shutting down the corresponding section
in the CBC. In this way, each section would only require a load reduction
capability similar to that of each primary bed.
This load reduction can
be provided by designing the CBC at higher excess air.
During load
reduction, the excess air is correspondingly reduced to keep the bed
temperature in the CBC constant.
For 30% load reduction, a reduction of
excess. air from the design value of 50% to ~15% is necessary.
Dividing the CBC into two sections is not considered feasible.
First, each section would require> 50% load reduction capability, so
that the CBC would have to be designed at even higher excess air. That
represents an increase in sensible heat loss from flue gas and a decrease
in overall efficiency.
Second, such a division might complicate the
control problem during operation.
Superheater.
The superheater is a conventional design and is
located in the convection zone right before the convection section.
The
proposed superheater design occupies approximately 14 feet of the 30-
foot convection section, with serpentine superheater tubes arranged
190
-------
for counter-flow of steam to flue gas. Carbon steel tubing of 2--1/2 in
o.d. is proposed with 6-1/2-incDperpendicular(to the flue gas flow
direction) and 4-in parallel spacing, as shown in the inset of
Figure 1-69.
With the present design, the control of the superheated
steam temperature is expected to be minimal because the flue gas flow
. .
rate through the superheater decreases proportionally, as does the hcat
transfer coefficient, as the load is reduced.
Convection Pass.
Two-inch tubes are arranged on 4-1/2
-in
spacing in the convection zone to recover heat from the flue gas to a
level such that only one heat trap (economizer) is needed further
downstream.
Natural circulation is employed in this section.
No valves
are used to parallel flow through various circuits.
Economizer.
An extended-surface economizer was chosen for
this design on the basis of a study of the cost of heat traps, including
tubular and regenerative air heaters, and bare-tube and extended--surface
economizers.
The extended-surface economizer is a convection unit
similar to the bare-tube economizer, with multiple parallel circuits
arranged for counter.-flow of gas to water.
The gas-side heat transfer
surface is increased by applying continuous spiral fins 3/4 in
high,
1-1/2 in
arranged in a 4-1/2 in
The tube size considered here is
square pitch.
with two fins per inch of tubing.
Coal and Limestone Feed Systems
The detailed designs of coal and limestone feed systems are
shown in Figure 1-73.
below.
These two feed systems are discussed separately
Coal Feed System.
Altogether 32 coal feed points, eight
for
each bed, are needed for the four primary beds. This represents about
2 .
10 ft bed area per coal feed point, chosen for the present design on the
basis of the British data {44].
Four additional feed points are provided
for the CBC for feeding the carbon-rich ash collected in primary cyclones.
This corresponds to 16 ft2 bed area per feed point.
191
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Coal crushed to 1/4 in
(see Figure 1-64) is first transferred
to the feeding hopper and then to a moving belt gravimetric feeder
with adjustable belt speed for different feed rate control. (See
Appendix D for a detailed description of the feeder.) The coal from
the gravimetric feeder drops into a four-way splitter to divide the coal
into four equal streams, one for each bed.
An individual shut-off
valve is provided for each stream, but no individual control of the coal
feed rate for each stream is possible.
Each stream of coal is again
spl~t into eight equal port.ions, one fO,r each coal feed point, by using
I I
, a vibrating ta'bl'e (detailed in Appendix D). The coal is then fed
pneumatically into the bed.
, ,
Coal drying is also accomplished in the coal feeding system
by using preheated sealing air. No condensation of moisture in the
coal feeding lines is expected if the lines are kept at a temperature
higher than the dew point of the gas. No 802 or particulate will be
vented into the atmosphere during this drying operation because of
the tight enclosure in the feeding system.
The individual coal feeder is a "straight-in" design forming
a 30° angle with the distributor plate to facilitate the maintenance
and replacement of the coal feeder even while the boiler is in operation.
Limestone Feed System.
Since limestone is mixed with coal at
the discharge end of the coal feeder, no additional limestone feed
point is provided in the fluid beds. The limestone hopper is a preSsure-
seal type, and the limestone feed rate is controlled by a variable
amplitude vibrating table feeder.
In both the once-through dry solid and the wet scrubbing
systems, the limestone feed rate is not regulated according to the coal
feed rate but rather is adjusted to keep the SO
x
emissions in the stack
gas within the regulated limits.
195
-------
Particulate Removal System
Primary.
Four cyclones of 90% efficiency at 10 ~ will be
used as primary cyclones.
They are refractory lined and internally
insulated.
Ash collected in the primary cyclones is carbon-rich and
is recycled back to the carbon burn-up cell for refiring.
At part-
load condition the flue gas flow rate is ~educed and so is the efficiency
of the cyclones.
At part-load operation, the overall boiler efficiency
can either be sacrificed or be maintained by using high-temperature
dampers to bypass some of the cyclones.
The latter would be the better
solution; however, the high-temperature dampers may not be readily
available or may not provide trouble-free service.
Secondary. For the once-through dry solid system, an electro-
static precipitator of 99.5% efficiency (specifications presented in
Appendix D) will be used to reduce the flue gas particulate loading.
In the wet scrubbing system, the final particulate removal system is
incorporated in the wet scrubbers.
Zurn Industries' medium-energy wet
scrubber system is used here.
(The details are shown in Appendix D.)
Ash Handling.
Besides the collection of carbon.-rich ash in
the primary cyclones, the residual ash from the bed sections and from
an ash pick-up point underneath the superheater will be pneumatically
transported to the ash hopper. In the case of the once-through dry
solid system, the ash from the electrostatic precipitator is also
discharged to the same ash handling system. (The positive-pressure
pneumatic ash~handling system is presented in Figures FBB-30 and FBB-3l
in Appendix D.)
Pollution Control System
Four major pollution control systems are under consideration
for the industrial boiler application.
These are outlined in Table 1-22
with estimations of their respective fresh limestone requirements and
total waste solids produced.
A complete cost analysis for all four
systems was conducted and is presented below.
Only two systems, once-
through dry solid and wet scrubbing, were selected for detailed design
and cost analysis.
196
-------
TABLE 1-22
COAL-FIRED INDUSTRIAL BOILER POLLUTION CONTROL ALTERNATIVES
BASIS:
4.3 wt % S COAL
......
\0
'-J
BOILER SYSTEM
FRESH PRESSURE
LIMESTONE SO PARTICULATE DROP
EQUIPMENT FEED WASTE EMISsIoN EMISSION (POLLUTION CONTROL
FOR EMISSION (STOIC. CA/S) SOLIDS PPM GR/SCF SYSTEM,
SYSTEM CONTROL (TONS/TON COAL) TONS/TON COAL (LB/106 BTU) (LB/106 BTU) IN H20)
Desulfurized Electrostatic none '" 0.1 < 0.5
char precipitator
'" 1 wt %S '" 770 0.01-0.05
'" 0.3 wt%S '" 230 (0.02)-(0.08)
(0.52)
Limestone
scrubbing
8 calcine Scrubbing '" 1.2/'" 0.18 0.6-0.7 '" 500 0.01-0.05 12-16
stone in systems (1.13) (0.02)-(0.08)
boiler; Water
once- reclamation
through
8 remove S02 Scrubbing 1-3/'" 0.36 0.6-0.8 '" 500 . 0.01-0.05 < 0.5
in bed, system (1.13) (0-02)-(0.08) I
regenerate Water
stone, reclamation I
scrub Electrostatic i I
effluent precipitator I I I
i I I
I ' I
'" 6/'" 0.84 i '" 0.7 170-350 0.01-0.05 I < 0.5
Once-through Electrostatic i i
limestone precipitator (0.38-0.80) (0.02)-(0.08) I
Larger lime- ' I
1
I
stone han- I
dling facil- i I j
ity i j
i I
Limestone I i
, I I
regeneration I I 0.01-0.05 < 0.5
8 on-site Electrostatic 1-2/'" 0.1 '" 0.3 170-350
precipitator I (0.38-0.50) (0.02)-(0.08)
Regeneration
system ,
8 off-site Electrostatic 1-2/'" 0.1 '" 0.1 170-350 0.01-0.05 < 0.5
central precipitator (assume mixed I (0.38-0.80) I (0.02)-(0.08)
I
facility at central ! I
facili t ) 1
y
-------
In the once-through dry solid system, six times stoichiometric
limestone is fed and calcined in the primary beds. S02 removal is
expected to be 90% in the bed, which represents about 16% stone utiliza-
tion.
The limestone is used only once through.
About 1.2 times stoichiometric limestone is needed in the wet
scrubbing system.
The limestone is first calcined in the bed and
then mixed with water to form slurry for use in the wet scrubber.
parallel units of wet scrubbers with two stages each are used for
Two
scrubbing the stack gas. The design of the wet scrubber is presented
in FBB-29, Appendix D. In addition t9 the contacting stages of the
wet scrubbers, drawing FBB-36 shows the components of the slurry
dewatering system, which includes the slurry thickener and vacuum
filter. The particulate and S02 removal efficiencies for the wet
scrubber are expected to be 98% and 90% respectively.
In addition to the wet scrubber design by Zurn Industries
described in Appendix D, a. budget quote from Norton Company for the
packed bed wet scrubber was obtained.
Although the packed bed scrubber
only costs about 1/3 as much as the wet scrubbing scheme selected, it
may suffer more operational problems due to the intrinsic scaling
problem of the limestone scrubbing systems and thus is not suitable
for the industrial boiler application. (The quote by Norton is included
in Appendix D.)
To avoid calcium sulfate precipitates in lines, vessels, and
pumps, the flue gas may be scrubbed with a caustic solution.
Lime
and spent caustic are mixed i~ a reactor vessel, forming calcium sulfate
and regenerating caustic solution. A calcium sulfite, sulfate filter
cake is discharged. from the process. The Ceilcote Company has prepared
a design and budget estimate for such a process, given in Appendix D.
It costs about three times more than the selected wet scrubbing system.
Selection of Draft Equipment
The total draft losses for the selected once-through dry
solid and wet scrubbing systems are summarized in Table 1-23.
"
A single,
198
-------
TABLE 1-23
SUMMARY OF DRAFT LOSSES
COMPONENT
AIR RESISTANCE OR DRAFT LOSS, IN. WG
DRY SOLIDS SYSTEM WET SCRUBBER SYSTEM
Section dampers
Perforated plate
Grid
0.2 0.2
0.5 0.5
1.0 1.0
5.0 5.0
20.0 20.0
5.0 5'.0
0.5 0.5
2.6 2.6
5.6 5.6
2.0 2.0
0.2 0.2
0.5
0.2
16.0
0.2
43.3 58.8
FORCED DRAFT DUCTS
FLUIDIZED BED CHAMBER:
Bed
MECHANICAL COLLECTOR
CARBON BURN-UP CELL (THRU FLOW)
SUPERHEATER
CONVECTION SECTION
ECONOMIZER,
ECONOMIZER FLUES
PRECIPITATOR
PRECIPITATOR FLUES
WET SCRUBBER
SCRUBBER FLUES
TOTAL
199
-------
forced-draft fan o~ a combination .of one forced-draft fan and one
induced-draft fan can be used.
A cost estimate of both these schemes
is presented in Table 1-24.
A single forced-draft fan provides the
saDIe total pressure drop more cheaply and is also favored in the wet
scrubbing system because there is no downstream induced-draft fan to
be clogged up by possible deposits of CaS04'
Overall Design Layout
The proposed overall boiler design arrangements for the
once-through dry solid and wet scrubbing systems are shown in Figures 1-74
and 1-75 respectively, with the appropriate dimensions.
200
-------
TABLE 1-24
DRAFT FANS FOR INDUSTRIAL FLUIDIZED BED COMBUSTION BOILERS(a)
TOTAL /)'P SINGLE, FORCED-DRAFT DOUBLE, FORCED-AND INDUCED-DRAFT FANS
FOR FANS HP
IN H20 HP COST, MOTOR + FAN INDUCED! FORCED COST, MOTOR + FAN
50.9 800 $33,000 450 500 $46,000
56.9 800 $33,000 450 600 $51,000
59.5 900 $35,000 450 600 $51,000
63.4 1000 .$35,000 450 700 $53,000
N $56,000
o 69.4 1100 $54,000 450 700
I-'
(a)250,000 1b/hr industrial boiler
75,000 CFM air flow at lOO°F
120,000 CFM top flow at 3750F
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203
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Hoiler Operation and Performance
A coal-burning fluidized bed boiler would be operated somewhat
differently from a conventional chain stoker, and certainly differently
from gas or oil-burning boilers.
For example, starting up a fluidized
bed involves preheating the fluidized bed to a minimum temperature of
~800°F with auxiliary burners before coal can be fed. On the other
hand, gas or oil burners for gas- or oil-burning boilers are simply
ignited. Providing a turn-down ratio of 4:1 in a fluidized bed boiler
necessitates turning on and off sections of fluid beds.
Turning off a
fluid bed involves only simple procedures; however, turning on a bed
once turned off requires preheating the bed to a minimum bed temperature
of ~800°F before coal injection, which becomes the limiting factor in
load response rate. This requires a separate subcontrol loop. Although a
gas-
or oil-burning boiler applies the same principles for load control,
. .
ignition is much simpler.
Because coal combustion in fluidized beds involves a different
combustion mechanism and different operational scheme, the performance
characteristics of fluidized bed boilers are necessarily different in
some respects from those of conventional boilers. For example, the
pressure drop in a fluidized bed is higher because of the relatively
high pressure drop through the distributor plate and the fluid bed.
In addition to conventional heat losses through the sensible heat of
flue gas and moisture in the fuel, the solid carbon loss becomes an
important item to consider, due to elutriation.
Although the principles for controlling the fluidized bed
boiler are the same as for conventional boilers, the philosophy is
different in some respects.
The bed temperature is the primary variable
for control:
too high a temperature will cause ash to fuse and limestone
to lose reactivity; too Iowa temperature will decrease substantially
207
-------
the reactivity of limestone and possibly decrease the combustion
efficiency of coal in the bed.
The detailed operating procedures, performance characteristics,
controls, and instrumentation are discussed below.
Operating Procedures
Start-up
Raising Boiler Pressure to On-Line Conditions.
The boiler
is started up first by firing an auxiliary gas burner above the bed in
the freeboard (as shown in Figure FBB-23 in Appendix D), and raising the
boiler pressure to on-line conditions, i.e., 600 psig and 489°F. The
bed temperature at this point will be essentially 489°F because of
recirculation of saturated water through the immersed horizontal tubes
in the bed.
The time required to raise the boiler pressure to on-line
conditions is estimated to be about 1-1/2 to 2 hours.
Heating Bed Material to Minimum Coal Injection Temperature.
Experiments by PER and NCB found a minimum bed temperature of ~800°F
to be necessary for self-sustained combustion of coal particles in the
fluidized beds. Therefore, the bed must be heated up from the saturation
temperature 489°F to at least 800°F by injection of auxiliary gas fuel
into the minimum-fluidized bed before coal feeding is started. Only
the first of the four beds and sections in the carbon burn-up cell have
to be started in this way. Start-up of the other beds can rely on the
solid recirculation from the neighboring operating bed to raise the bed
temperature. (This problem has been mathematically analyzed in Appendix F.)
Raising Load and Putting Other Beds in Service.
Once the bed
reaches the minimum coal injection temperature, the coal feed is started
and the load is increased.
The first section of the carbon burn-up
cell can then be fired in the same way, using the elutriated carbon
fines from the operating section as fuel.
To be started, the next bed
can be fl~idized at minimum fluidizing conditions; and the bed material
is allowed to circulate from the operating bed to raise the bed temperature
208
-------
to minimum coal injection temperature. The second carbon burn-up cell
section can also be started by recirculation of solids from the operating
section.
The third bed and thirdCBC section can be lit off in the
same fashion and so on until the boiler is in full operation.
(The
dynamic response of the bed temperature during start-up is presented
in Appendix F.)
Load Control
Numerous load control techniques for fluidized bed boilers
can be employed to achieve the desirable percentage load reduction.
However, the preferred turn-down technique is optimal for ease of
control, stability and reliability of operation, degree of turn-down
desirable, rate of system response required, efficiency, and economy.
(A comprehensive analysis of turn-down techniques is presented in
Appendix E.)
Load control in fluidized beds involves turning the beds on
and off. In turning off the beds, the load response is fast enough
to meet 10%/min. load change as analyzed in Appendix F. However,
turning on the idled beds to meet the desirable response rate becomes
a big challenge.
(A comprehensive analysis is included in Appendix F.)
, . .
Turn-Down Operation in Single Bed. In atmospheric industrial
boilers, a turn-down capability to less than 25% of full load at a
load swing rate of at least 5% per minute is desirable. To meet these
requirements, a technique which simultaneously decreases fuel and air
input and maintains excess air and bed depth constant is recommended.
The recommendation. is based on the general analysis presented
in Appendix E. Table 1-25 compares the operating conditions of the two
most promising modes: 1) turn-down by decreasing fuel and air input
and maintaining excess air and bed depth constant; and 2) turn-down by
decreasing fuel input and maintaining air input and bed depth constant.
The operating conditions before and after turn-down for both modes are
compared, assuming the load reduction is to be accomplished by lowering
the bed temperature from l650°F to l400°F with bed depth constant.
209
-------
TABLE 1-25
COMPARISON OF EFFECTIVENESS OF TURN-T)OWN TECHNIQUES
FOR INDUSTRIAL BOILER
Steam Conditions: . 750°F & 600 psig
OPERATING CONDITIONS
BEFORE TU~-DOWN
AFTER TURN-DOWN
MODE 1 j MODE 2
BED TEMPERATURE, of
1650
1400
1400
EXCESS AIR, %
FUEL REQUIREMENT (ARBITRARY UNIT)
10
10
40
1
0.69
0.79
SUPERFICIAL FLUIDIZING VELOCITY, FT/SEC
15
8
12
LIMESTONE FEED RATE
(ARBITRARY UNIT) (a)
1
1.4
1.6
LOAD REDUCTION IN THE BED, %
21.5
21.5
GAS TEMPERATURE, of.
Leaving combustion chambers
1530
1250
1293
Entering convection pass
1530
1250
1293
Leaving convection pass
672
580
626
Leaving economizer
350
315
352
HEAT LOSS THROUGH FLUE GAS
(ARBITRARY UNIT)
1
0.96
1.40
OVERALL MAXIMUM TURN-DOWN ACHIEVABLE
WITH ALL BEDS OPERATING,%
31
21
(a) The reaction rate for 1imestone-S02 reaction at 1400°F is
that at 1650°F. This means either more reactive dolomite
absorbant or more limestone has to be fed.
just about half of
has to be used as
210
-------
This reduces the total heat transferred in the bed by 21.5%
after turn-down in both cases, assuming the heat transfer coefficient in
the bed is constant and independent of turn-down operations.
However,
the total heat transferred in the convection pass and economizer is much
lower in Mode 1 turn-down than in Mode 2 due to a reduction of almost
30% of the original flu~ gas mass flow rate, G, which affects the heat
transfer coefficient, h, in the convection section by the relationship
h a (G)0.6l. This results in a maximum achievable turn-down in operating
beds of 31% for Mode 1 versus 21% for Mode 2.
In addition to this
advantage, less fuel and a lower limestone feed rate are required, and
much less heat is lost through flue gas in Mode 1. The temperature
of l400°F is chosen on the basis of the 502 removal capability of lime-
stone and does not represent the minimum operating temperature. For
flexibility and convenience in load swing, the bed can be operated at
temperatures higher than l650°F (operation at1700°F provides ~7% exten-
sion in load) or lower than l400°F (operation at l300°F provides an addi-
tional ~8% reduction in load) for short periods of time without seriously
affecting efficiency of carbon combustion and 502 removal. When wet
scrubbing is used for 502 removal, the bed temperature is more flexible
because it is no longer limited by limestone reactivity in the bed but by
ash fusion temperature (~2000°F) at the upper limit and coal combustion
efficiency (~1200°F) at the lower limit.
Change of bed depth as an alternative means for turn-down was
ruled out on the basis of the following considerations:
.
The designed bed depth in the atmospheric industrial
boiler is already quite shallow, i.e., 2-1/2 ft. A more
than 25% decrease in bed depth would increase the carbon
carry-over and thus decrease overall efficiency.
.
The heat transfer surface exposed by changing the bed
depth w~ll still have a high heat transfer coefficient,
comparable to that of heat transfer surface in the bed.
Hence, the effectiveness of changing the bed depth in
turn-down is doubtful.
211
-------
.
The additional cost in providing the rapid transport of the
solid in and out of the bed with respective flow control
devices does not merit consideration in an atmospheric
industrial unit. The stability and reliability of operating
a shallow bed with rapid flow of solids in and out of the
bed might be poor.
Turn-down Operation in Entire Boiler.
The present industrial
boiler is sectioned into four equal beds with the flue gases from the
individual beds collected before passing through the superheater and a
common convection section. The proposed philosophy for load reduction
is best illustrated in Figure 1-76: all four beds are simultaneously
reduced in load at least 25% before turning off bed number 4 and boosting
/
the other beds to full operation; the remaining three beds are turned
down to their lowest load before shutting down the number 3 bed; and
so on.
The scheme is continued until orily the number I bed is in
operation. The number I bed operating at lowest load represents ~17.5%
or ~lO% of total boiler loadt depending on the pollution control
alternative employed. The operating conditions before and after turn-
down for Mode 1 are presented in Table 1-26 for beds operating in
the temperature range l400°F to l650°F. It is significant that the heat
loss through flue gas is actually decreasing -- or the boiler efficiency
is increasing -- during turn-down; this is because the size of the common
convection section is designed for full-load operation.
Tnis same
amount of heat transfer area is also available during turn-down conditions
with a lower flue gas flow rate, resulting in lower stack gas temperaturet
as shown in Table 1-26.
The fact that the efficiency of the boiler is
actually increasing
for load reduction.
during turn-down provides additional flexibility
Now Mode 2t which achieves turn-down by increasing
be incorporated to give more flexibility in turn-
excess air, can also
down without suffering loss of boiler efficiency.
In the actual boiler
operation, the heat loss through the flue gas might be larger than that
shown in Table 1-26 because of by-passing of some flue gas into the cold
beds which are not in operation.
This is espe~ially true when three of
212
-------
100
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-
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.......
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C"I
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.......
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. Curve 643916-A
80
/
/
1//
I / / //
I / / /
/ / /
I / /
1 / / /
/ / / '
I / / /
20 I I // /
1///
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60
40
3 = 3 Beds Operati ng
i.e., No.1, No.2 & No.3
o
o
20 40 60 80
Total Boi ler Load, %of full load
100
Fig. 1-76-Turn-down scheme of four bed industrial boiler
, ,
213
-------
TABLE 1-26
OPERATING CONDITIONS DURING TUfu~-DOWN OPERATION FOR INDUSTRIAL BOILER
4 BEDS OPERATING 3 BEDS OPERATING 2 BEDS OPERATING 1 BED OPERATING
OPERATING CONDITIONS Full Load Part Load Full Load Part Load Full Load Part Load Full Load Part Load
BED TEMPERATURE, of 1650 1400 1650 1400 1650 1400 1650 1400
EXCESS AIR, % 10 10 10 10 10 10 10 10
FUEL REQUIREMENT 1 0.69 0.75 0.52 0.50 0.35 0.25 0.17
(ARBITRARY UNIT)
N SUPERFICIAL FLUIDIZING VEL. FT/SEC 15 8 15 8 15 8 15 8
I-'
.j::'-
LIMESTONE FEED RATE 1 1.4 0.75 1.05 0.50 0.70 0.25 0.35
(ARBITRARY UNIT)
VELOCITY IN CONVECTION PASS 1 0.59 0.74 0.45 0.50 0.31 0.25 0.15
(ARBITRARY UNIT)
GAS TEMPERATURE, of
Leaving combustion chamber 1530 1250 1516 1281 1516 1281 1521 1281
Entering convection pass 1530 1250 1516 1281 1516 1281 1521 1281
Leaving convection pass 672 580 634 568 592 543 .540 513
Leaving economizer 350 315 335 305 310 290 280 270
HEAT LOSS THROUGH FLUE GAS 1 0.96 0.96 0.85 0.87 0.79 0.76 0.72
(ARBITRARY UNIT)
PERCENTAGE OF TOTAL LOAD, % 100 69.0 75.5 51. 8 50.8 34.8 25.7 17.7
-------
the four beds are shut down; however, the extent of the heat loss is hard
to estimate.
On the other hand, the flue gas velocity in the convection
zone is reduced up to 85% of the original velocity during turn-down
and may cause the deposition of ash in the convection section. Erie
City's experience indicates this will not be an operational problem.
Turn-Down Operation in Carbon Burn-Up Ce~~~(CBC). The bed
temperature in the carbon burn-up cell has to be kept constant during
turn-down for efficient carbon burn-up.
The proposed design divides
the CBC into four separate sections, each corresponding .toa primary
bed, by installing water tube screens in the bed and by subdividing
the air. plenum. Turn-down by shutting down a bed would require shutting
down the corresponding sect~on in the CBC.
Each section in theCBC
thus has to provide only the same load reduction capability as each
primary bed. The load reduction in the CBC can be accomplished by
designing the full-load operation conditions at 1900°F and high excess
air, say 50%. When the CBC is operated at reduced load, the percentage
excess air ,can be decreased to decrease the heat loss to the flue gas
and keep constant the bed temperature of the CBC.
For a 30% load
reduction, the operating conditions of the CBC after turn-down are 1900°F
and ~l5% excess air.
To operate theCBC in this way, individual on-off
control of the spent-coal feed to each section of the CBC is provided.
. Shut-down
Single Bed Shut-Down.
When removing one bed from service,
the coal feed is shut off first while fluidizing air is maintained to
burn off the residual coal in the bed and to prevent deposit of coal in
the pneumatic transport lines.
After the residual coal is burned out,
the bed temperature starts to drop, and the fluidizing air can be shut
off.
However, the recirculation water-flow through the immersed horizontal
tubes will be maintained throughout the operation.
Thus, the heat
. .
. .
removal from the defluidized bed after shut-down depends entirely on
conduction of bed material to the water tubes.
(The change of bed
temperature and the residual steam production after shut-down is analyzed
mathematically in Appendix F.)
215
-------
Entire Boiler Shut-Down.
To shut down the entire boiler, one
fuel-burning bed section is shut down at a time. The start-up burner
should be ignited prior to shutting down the last bed section to maintain
the steam generator well above the flue gas condensation temperature.
The cooling rate is lim{ted so that there are no undue thermal stresses.
A cooling rate of IOO°F per hour, the same as the heating rate, should
be used.
The recirculating pump is kept running until the steam generator
is cooled to below 200°F and is ready for draining.
As long as the
recirculating pump is operating, the water level in the steam drum
must be maintained.
Emergency Operation.
Foreseeable emergency situations and.
their remedial procedures are outlined below:
1.. Loss of combustion air flow
..
Trip all fuel feed to the steam generator
II
.Continue operation of sealing air fan and ash system.
blowers
.
Continue operation of recirculating pump and maintain
water level in steam drum
.
Open superheater vent as soon as boiler pressure drops
below the line pressure.
2.
Loss of recirculation pump
. Trip air and fuel feed
.
Continue operation of sealing air fan, ash-handling
equipment, and gas clean-up equipment
CI
Maintain steam-drum water level
.
Open superheater vent when boiler pressure drops below
the line pressure.
216
-------
3.
Loss of steam-drum water level
8
I
Stop fuel feed and cool steam generator as' rapidly as
possible
8
Continue operation of forced-d~aft fan
8
Observe operation of recirculation pump to prevent
cavitation
. Drum level should be restored by feeding water at a low
rate (5% to 10% of the maximum steam flow).
4. Sudden burn-out of bed tubes
8 'Shut down steam generator as soon as possible.
5. Loss of bed fluidization
.
Trip fuel and then air feed
.
Continue operation of recirculation pump.
6.
Loss of fuel feed is not in itself an emergency condition
as long as the forced-draft fan and other equipment continue
to operate.
However, care should be taken to prevent damage
to the superheater due to over..-firing on restarting.
Performance Characteristics
Overall Boiler Efficiency.
The comparison of total boiler
efficiency between once-through dry solid and wet scrubbing systems is
presented in Table 1-27. The wet scrubbing system has higher boiler
efficiency because the stack gas temperature is 132°F, compared to 350°F
in the dry solid system. An additional efficiency loss of about 5% in
the wet scrubbing system is expected if the flue gas is to be heated up
from 132°F to 350°F.
Draft Losses.
A summary of draft losses for the once...through
dry solid system and the wet scrubbing system is presented in Table 1-28,
and a comparison of corresponding power requirements for both systems is
estimated in Table 1-29.
217
-------
TABLE 1-27
COMPARISON OF BOILER PERFORMANCE
ONCE-THROUGH
DRY SOLID(%)
Loss due to calcination and S02 absorption
Loss due to water and hydrogen in fuel
2.74
3.80
N
......
00
. Loss due to unburned carbon
1. 93
Loss unaccounted for and due to radiation
1.50
Loss due to evaporation in Wet scrubber
Loss due to flue gas sensible heat
6.61
o
(Flue gas temp.=350 F)
Total loss
16.58
Overall boiler efficiency
83.42
WET SCRUBBING(%)
0.56
3.83
2.05
1.50
4.84
1.30
o
(Flue gas temp.=132 F)
14.08
85.92
-------
TABLE 1-28
SUMMARY OF DRAFT LOSSES, IN. H20 (GAUGE)
COMPONENT
DRY SOLIDS SYSTEM
WET SCRUBBER SYSTEM
FORCED DRAFT DUCTS
0.2
0.2
FLUIDIZED BED CHAMBER:
Section dampers
0.5
0.5
Perforated plate
1.0
1.0
Grid
5.0
5.0
Bed
20.0
20.0
MECHANICAL COLLECTOR
5.0
5.0
CARBON BURN-UP CELL (THRU FLOW)
0.5
0.5
SUPERHEATER
2.6
2.6
CONVECTION SECTION
5.6
5.6
ECONOMIZER
2.0
2.0
ECONOMIZER FLUES
0.2
0.2
PRECIPITATOR
0.5
PRECIPITATOR FLUES
0.2
WET SCRUBBER
16.0
SCRUBBER FLUES
0.2
TOTAL
43.3
58.8
219
-------
TABLE 1-29
COMPARISON OF POWER REQUIREMENTS (a)
COAL-FIRED FLUIDIZED BED COMBUSTION
STEAM GENERATOR
DRY SOLIDS
SYSTEM
WET
SCRUBBER
FORCED-DRAFT FAN, HP 700 950
RECIRCULATING PUMP, HP 200 200
COAL CRUSHER, HP 60 60
SLURRY PUMPS, HP 150
TOTAL POWER
REQUIREMENTS, HP 960 1360
POWER REQUIREMENTS
% OF OUTPUT (b) 0.86 1. 21
(a) Only major operating horsepower requirements are shown.
(b)
Where a horsepower hour is 2544 Btu and output is 289 million
Btu per hour.
220
-------
Air Pollution Control Capability. The control capability of
the industrial fluidized bed boiler for NOx' S02' and particulates is
summarized in Table 1-30 from the existing experimental data discussed
previously. NO emissions from the industrial boiler are expected to
x 6 ..
be in the range 250 ppm (0.40 IbN02/10 BTU) to 500 ppm (0.81 IbN02/BTU),
compared to~550 ppm from a conventional spreader stoker. S02 removal
depends primarily on the amount of limestone fed and is expected to
be >90% for the once~through dry solid system. The wet scrubbing.
system is expected to have an S02 removal efficiency of ~90%. For a
coal containing 4.3 wt% sulfur, the S02 emission is ih t~e range of
350 ppm to 500 ppm (0.80-1.13 Ib/l06BTU). For the once-through "dry
solid system, the particulate loading in flue gas is expected to be
less than 0.02 gr/SCF. The particulate removal efficiency of the present
wet scrubber design is ~98%.
Controls and Instrumentation
Overall boiler control has three subdivisions:
combustion
control, feed water flow control, and steam temperature control.
The
basic control principles for the fluidized bed boiler are similar to
those for any subcritical recirculating boiler.
The detailed control
loops are presented in Appendix D, along with lists of control equipment;
the functional aspects of the control logic are outlined here.
Combustion Control System.
Steam drum pressure is the load-
sensitive control element and is used to initiate the fuel feed.
Combustion air is fed to a common plenum by the forced-draft fan, with
the plenum air pressure kept constant by the inlet vanes positioned
by the signal from the fuel:air ratio setter. Individual dampers
control the combustion air flow to each individual primary bed and each
individual section of the carbon burn-up cell.
Flow to each section is
measured and totalized. An oxygen analyzer for flue gas will be provided
to record the excess air, but it will not be in the control loop.
221
-------
TABLE 1-30
AIR POLLUTION CONTROL CAPABILITY
OF INDUSTRIAL FLUIDIZED BED BOILER
DRY SOLID
I
WET SCRUBBING
NO emission
x
ppm 250-500 250-500
Ib/106BTU 0.40-0.81 0.40-0.81
S02 emission
ppm 350 (90% removal) 350':'500
Ib/106BTU 0.80 0.80-1.13
Particulate emission
gr/SCF 0.01-0.05 0.01-0.05
Ib/106BTU 0.02-0.08 0.02-0.08
222
-------
Temperature in each bed is measured and used as a limit in
starting up and shutting down an individual bed.
Bed temperature in the
carbon burn-up cell is always kept constant and is used as a sensitive
control element for adjusting the excess air in the CBC.
For bed start-up a sub-loop, including a transfer switch
and time-delay volume chamber, is used to control the rate of air flow
introduction and bed fluidization.
Feedwater Flow Control System.
The feedwater flow control, or
steam drum level control, has three elements.
Feedwater and steam flow
signals are compared.
If a difference occurs, a signal is sent to a
rate response system and to the feedwater flow control.
The steam
drum level is compared to a set point.
If a difference exists, a signal
is sent through a summer to the feedwater control system.
The signals
from the flow and from the level are compared and sent to a proportional
reset controller, to the feedwater hand-auto station, and then on to the
flow control valve positioner.
Steam Temperature Control System.
control system will be a two-element system.
The steam temperature
The outlet steam temperature
signal is compared to a set point.
The resulting error signal is sent
through a proportional plus reset controller and summed with the steam
flow signal. The combined signal is passed through the steam temperature
attemperator hand-auto station to the spray flow control valve.
223
-------
"
Boiler Cost
. The capital costs of the two designs selected (once-through
dry solid and wet scrubbing systems) were evaluated by using the regular
boiler pricing procedures employed by Erie City. The resulting boiler
capital cost was then compared with the cost of the conv~ntional coal-
fired spreader stoker and gas-
or oil-fired industrial boiler with
approximately the same capacity and steam conditions.
The capital costs of other designs with different pollution
control schemes (shown in Table 1-22) were also derived from these'two
basic designs, as were the capital costs of gas-and oil-fired fluidized
bed boilers.
The total costs of all the alternatives, including labor,
power, solid disposal, fuel, and limestone costs were then compared
to determine the better systems for industrial application.
Capital Cost of Once-Through Dry Solid and Wet Scrubbing Industrial
Fluidized Bed Boilers
The fluidized bed steam generator design described previously
. was subjected to the standard boiler pricing procedures employed by
Erie City. The auxiliary equipment costs for once-through dry solid
and wet scrubbing systems were procured.
individual equipment was also estimated.
The installation cost of the
These were used to determine
the capital costs of the once--through dry solid system, shown in
Table 1-31 and of the wet scrubbing system, shown in Table 1-32. The
steam generator cost represents about 45% to 55% of the overall boiler
system cost.
Aside from the steam generator, the ash-handling system
is the major cost item in the dry solid system (17%); and the wet
scrubber system is the major cost item in the wet scrubbing system (22%).
Surprisingly enough, the fuel feeding system represents only ~6% of the
total boiler cost.
224
-------
TABLE 1-31
INSTALLED COST OF ONCE-THROUGH DRY SOLID SYSTEM
(Shipped As Two Shop-Assembled Modules)
APPROXIMATE
EQUIPMENT COST $
APPROXIMATE
INSTALLATION COST. $
. TOTAL $
% OF TOTAL COST
Crusher 18,000 not included 18 ;000 1.18
Fuel system (not including
bunkers, and system upstream
of bunkers). 74,090 25,000 99,000 6.50
N
N
\.J1 Forced draft fan 35,000 5,000 . 40,000 2.62
Fuel system sealing fan & heater
system 12,000 10,000 22,000 1.44
Steam generator system, recirculat-
ing pump & economizer 753,000 95,000 848,000 55.64
Ash-handling system 170,000 85,000 255,000 16.73
Mechanical dust collector 30,000 15,000 45,000 2.95
Controls 61,000 30,000 91,000 5.98
Electrostatic precipitator 76,000 30,000 106,000 6.96
TOTAL 1,229,000 295,000 1,524,000 100.00
-------
TABLE 1-32
INSTALLED COST OF WET SCRUBBER SYSTEM
(Shipped As Two Shop Assembled Modules)
APPROXIMATE APPROXIMATE % OF TOTAL
COMPONENT EQUIPMENT COST $ INSTALLATION COST $ . TOTAL $ COST
Crusher 18,000 Not included 18,000 0.96
Fuel system (not including 74,000 25,000 99,000 5.35
bunkers, and system upstream
of bunkers)
Forced draft fan 54,000 5,000 59,000 3.19
Fuel system sealing 12,000 10,000 22,000 1.19
Fan & heater system
Steam generator system 753,000 95,000 848,000 45.79
Recirculating pump & economizer
Ash system for slurry system 100,000 50,000 150,000 8.10
for carbon burn-up
Mechanical dust collector 30,000 15,000 45,000 2.43
Wet scrubber system 300,000 100,000 400,000 21.60
Controls 61,000 30,000 91,000 4.91
Dewatering 80,000 40,000 120,000 6.48
TOTAL 1,482,000 370,000 1,852,000 100.00
,
,---------.---..- ..- "--.- .-----. ..' ..---..-
226
-------
Comparison of Capital Cost Between Fluidized Bed Boilers and Conventional
Boilers
The installed cost of the two basic fluidized bed boilers was
compared to that of a field-erected conventional coal-fired spreader
stoker and field-erected gas-
or oil-fired boilers with equivalent
capacity, as shown in Table 1-33.
The bases for comparison are presented
in the description following that table. The once-through dry solid
fluidized bed boiler is cheaper, with an installed cost of $6.l0/lb/hr
as compared to $7.40/lb/hr for the wet scrubbing system and $7.57/lb/hr
for a conventional coal-fired spreader stoker.
However, the field-
erected gas- or oil-fired boiler cost only $2.44/lb/hr. If the gas-
oil-fired steam generator is shop-assembled, the installed cost can be
reduced an additional 16%.
or
The capital costs for other pollution control alternatives
and the gas/oil-fired fluidized bed boiler were also derived from these
basic cases and compared in Table 1-34.
Some schemes, such as once.-
through limestone injection in a conventional spreader stoker and in a
field-erected oil-fired boiler burning high-sulfur oil, are clearly
not optimum schemes because of low sulfur removal capability (~40%),
but they are included in the table for the sake of completeness.
Comparison of Total Cost Between Fluidized Bed Boilers and Conventional
Boilers
The total costs of all the air pollution control alternatives
(Table 1-22) (including labor, power, solid disposal, fuel and limestone
cost) are compared in Table 1-35 to determine the better systems for
industrial application and to evaluate the limitations of each individual
system.
The total cost was obtained from the assumptions that the
capital charge is 16.7% per year with a load factor of 70%; that the
coal cost is $7.50/ton (an equivalent of 30c/106 BTll); that sorbent
cost is $3.00/ton; and that waste disposal cost is $5.00/ton.
227
-------
TABLE 1-33
COMPARISON OF INSTALLED COSTS(a)
FOR EQUIVALENT-CAPACITY, FIELD-ERECTED,
STEAM GENERATORS FIRING
DI FFERENT FUELS
STEAM GENERATOR
(See detailed
description)
A
B
C
D
Fuel & Fi ring.
Technique
Gas-or oil-
firing
Coal-fired,
spreader
stoker
Coal-fired,
fluidized
bed combus-
tion
S02 Emission Control
Low-sulfur
fuel
Not
furnished
Dry solids
system
Wet scrubber
system
Installed Cost
for Boiler System, $
6l0,000(b)
1,375,000
1,524,000
1,852,000
Installed Cost for
S02 Control System,
Add, $
520,000
Ins taIled Cas t
Total for Equivalent
Emission to Atmosphere, $
610,000
1,895,000
1,524,000
1,852,000
Cost per Unit
Steam Generation, $/#/hr.
for Equivalent Stack
Emission
2.44
7.57
6.10
7.40
(a)C b d . . h
osts are ase on experlence Wlt
current prices. Estimated factory
Steam Generators.
similar steam generating units escalated to
and installation prices are used for FEC
(b) If the steam generator is shop assembled, the installed cost will be reduced 16%.
228
-------
DESCRIPTION OF TilE FOUR DIFFERENT STEAM GENERATORS USED II~ CAPrT/\L
COST AND OPERATING COST COMPARISONS
Steam Generator "A"
----------
Steam Generator liB"
----.-- ---.---
Erie City Energy Division, Field-Erected, "Keystone" Steam Generator
Erie City Energy Divisi.on, Field-Erecte:l, "Cross Drum" Steam Generator
OPERATING
CONDITIONS:
Steam flow -- 250,OOOU/hr.
Steam pressure -- 600 psig
Steam temperature -- 750°F
Feedwater temperature -- 250°F
Stack temperature -- 300°F
Boiler efficiency -- 82.2%
System resistance -- 11.8 in wg.
(Pressurized operation)
OPERATING
CONDITIONS:
Steam flow -- 250,000 U/hr
Steam pressure -- 60D psig
Steam temperature -- 750°F
Feedwater temperature -- 250°F
Stack temperature -- 400°F
Boiler efficiency -- 85%
System air resistance draft loss -- 12.4 in wg.
(Balanced draft)
FUEL:
FUEL:
Natural gas -- U2 oil
Bituminous coal sized to no more than 30% through
4 mesh screen
N
N
'"
TERMINALS ARE:
Inlets to: economizer, fuel train and stack
Outlets from: superheater
TERMINALS ARE:
[~lets to: eco~omizer, stoker feeders and stack
Outlets from: ash-handling system and superheater
EQUIPMENT FU&,ISHED:
Burners -- ECED circular burners
Boiler -- bottom supported - no preassembly
Superheater
Economizer -- extended surface
Flues, ducts and dampers (not including stack)
Supporting steel
Soot blowers
Combustion and feedwater control systems
Refractory and insulation
Piping for boiler fuel system, connection betWeen
economizer and boiler and trim piping
SERVICES
FU&'HSHED:
Erec tion
Erection
Start-up
supervision
labor
service
EQUIPHENT FU&HSHED:
Forced-draft and induced-draft fans and drives
Stoker -- Detroit rota grate -- spreader stoker
Boiler -- no pre assembly -- top supported
Superheater
Economizer -- bare tube
Boiler columns and top grid steel
Soot blowers
Combustion and feedwater controls
Refractory and insulation
Mechanical collector
Electro-static precipitator
Flues, ducts and dampers (not including stack)
Ash-handling system
Piping for connection between economizer and boiler
and trim piping
SERVICES
FURNISHED:
Erection
Erection
Start-up
supervision
labor
service
-------
STEAM GI,NERATOR DE~;CIUI'TI(jNS (Continued)
Steam Generator "ell
-------------.-
St~am Generator "I)"
Erie City Energy Division, Modular, Fluidized Bed Combustion, Steam
Generator Using Dry Solids 502 Control
Erie City Energy Division Modular, Fluidized Bed Combustion Steam
Generator Using A Wet Scrubber - Particulate and S02 Removal System
OPERATING
CONDITIONS:
Steam flow -- 250,ODOO/hr.
Steam pressure -- 600 psig
Steam temperature -- 750°F
Feedwater temperature -- 250°F
Stack temperature -- 350°F
Boiler efficiency -- 85.4%
System resistance -- 43.4 in wg.
(Pressurized operation)
OPERATING
cmlDI TI():~S :
S team flow -- 250 ,00011/h r.
Steam pressure -- 600 psig
Steam temperature -- 750°F
Feedwater temperature -- 250°F
Stack temperature -- 132°F (Saturated
vapor)
Boiler efficiency -- 85.8%
System resistance -- 58.8 in wg.
(Pressurized operation)
wi th wate r
FUEL:
Bituminous coal -- 100% through 4 mesh
FUEL:
Bituminous coal -- 100% through 4 mesh
N
W
o
TERMINALS ARE:
Inlets to: economizer, coal feeder, limestone
. feeder and stack
Outlets from: ash silo, superheater
TERMINALS ARE:
Inlets to: economizer, coal feeder, li.mestone
feeder and stack
Outlets from: vacuum filter and superheater
EQUIPMENT FU&~ISHED:
Forced-draft fan and drive
Coal feeding system, including rate feeder,
splitter and injectors
Limestone feeder
Boiler -- bottom supported pre-assembled into
two modules
Superheater
Economizer -- extended surface
Supports for mechanical collector economizer and
precipitator
Combustion and feedwater controls
Refractory and insulation
Mechanical collector
Electro-static precipitator
Flue, ducts and dampers (not including stack)
Ash-handling system and silo
Piping for boiler fuel system, between economizer
and boiler and trim piping
E(IU1 PMENT
FUR:HSHED:
Forced-draft fan and. drive
Coal feeding system, including rate feeder,
splitters and injectors
Limestone feeder'
Boiler -- bottom supported (pre-assembled into
two modules)
Superheater
'Economizer -- extended surface
Supports for mechanical collector, economizer and
precipitator
Combustion, feedwater and scrubber controls
Refractory and insulation
Mechanical collector
2-stage wet scrubber with slurry tank and pumps
Flues, ducts and dampers (not including stack)
Pneumatic transport ash-handling system
Slurry and ash dewatering system
Piping for boiler fuel system, between economizer
and boiler trim piping and scrubber system
convections
SERVICES FURNISHED:
Erection supervision
Erection labor
SERVICES FU&~ISHED:
Erection supervision
Erection labor
-------
TABLE 1-34
CAPITAL COST OF COAL- AND GAS/OIL-FIRED INDUSTRIAL BOILER POLLUTION CONTROL ALTERNATIVES ($/lb/hr).
FLUIDIZED BED COMBUSTION BOILER
(packaged)
CONVENTIONAL BOILER
(field erected)
COAL-FIRED (a)
Desulfurized char
Low-sulfur coal
Limestone scrubbing
Once-through limestone
Limestone regeneration
On-site facility
Off-site central facility
5.25
5.25
7.40(b)
5.50
7.60
7.75
6.10
GAS/OIL-FIRED
N
....,
......
Clean fuel
High-sulfur oil
Once-through limestone
,Wet scrubbing
4.00 - 5.00
5.25 - 5.50(c)
7.00 - 7.50(d)
2.44(e)
3.1O(f)
4.50
(a) Costs do not include cost of equipment before coal feeding system, e.g., railhead, bunkers.
This cost is $.25 - 1.00/lb/hr.
(b) Cost is, identical for two processes: calcine stone in boiler, once-through; or remove S02 in
bed, regenerate stone, scrub effluent.
(c) Either once-through limestone or off-site regeneration.
(d) Limestone calcining - either wet scrubbing or on-site regeneration.
(e) Cost for a packaged system would be $2.00/lb/hr.
(f) The percentage sulfur removal in this case will be low.
-------
TABLE 1-35
INDUSTRIAL BOILER COST SUMMARY
<;/106 Btu
CAPITAL' FUEL SORBENT(b) SOL IDS / SL URRY
CHARGES (a) COST LABOR POWER DISPOSAL (c) TOTAL
FLUIDIZED BED COMBUSTION - COAL-FIRED
Desu1furized char 10.95 (30) 9 1.1 1.90 52.96
Limestone scrubbing 15.40 30 9 2.08 1.5 13.50 71.50
N Once-through dry limestone 12.70 30 9 9.69 1.1 13.50 76.00
u..>
N
Limestone regeneration
On-site 16.20 30 9 1.15 1.5 5.80 63.67
Off-si te 12.70 30 9 3.45 1.5 1.90 58.57
SPREADER STOKER
. Lo\v-sulfur coal 11. 50 60 9 1.1 1.90 83.51
High-sulfur coal - wet scrubbing 15.90 30 9 2.08 1.5 13.50 72.00
OIL/GAS-FIRED BOILER - FIELD~ERECTED, 5.10 65 4 0 I. 74.50
....
BU&.~ING CLEAN FUEL
(a) Capital charge 16.7% fuel 6
per year, 7.0% load factor; cost $7.50/ton ~ 30<;/10 BTL
(b) regenerated stone $9.00/ton
Raw stone $3.00/ton;
(c)$-/ .
) ton dlsposal cost
(cost information has varied from $0-10/ton)
-------
From this comparison, several projections are clearly in
order:
.
A fluidized bed boiler burning desulfurized char is the
most economical case, if the desulfurized char is available
at the cost projected, i.e., 30~/106BTU.
.
The limestone scrubbing scheme and once-through dry soU.d
scheme are comparable in total cost.
.
Limestone regeneration schemes have a cost advantage of
. approximately 15% over the other systems, if a large
sulfur recovery facility is available on-site or off-site.
.
A spreader stoker burning low-sulfur coal is more expensive
than burning high-sulfur coal with wet scrubbing for S02
removal because of the high cost of low--sulfur coal.
.
Although the capital cost of the field-erected oil/gas-fired
boiler is substantially less than other systems, the total
cost is comparable to other systems, again, because of
the high fuel cost.
It is significant that the total cost of the boiler is highly
sensitive to the fuel cost. Thus, the relative advantages of the
alternative systems will very much depend on the future fuel situation.
233
-------
Effect of Operating Variables on Cost and Design of Industrial Boiler
The cost and design of a fluid bed boiler depend mainly on
operating variables such as fluidizing velocity, feed particle size,
operating pressure, bed depth, bed temperature, excess air, and air
preheat temperature.
In setting the ?perating variables, not only the
cost, but also the efficiency of pollutants removal must be taken into
consideration. An optimal boiler design will meet the requirement of
air pollution abatement at a minimum cost.
An optimal choice of the
operating variables as such will depend on the total optimization of
the operating variables involved. This is a meaningless task without
accurate knowledge of the change of heat transfer coefficient in the
bed, 802 removal efficiency, combustion efficiency, elutriation, attrition,
and erosion with respect to change of different operating variables.
Unfortunately, theoretical or experimental correlations which give
accurate predictions are still not available.
Nevertheless, the total
cost of a base design may be perturbed to evaluate the effect of operat-
ing variables.
Variables capable of significantly affecting the cost and
design of industrial fluid bed boilers were analyzed. The design bases
of bed temperature, excess air, and air preheat temperature are usually
restricted by considerations of 802 removal efficiency, turn-down
requirement, combustion efficiency, and/or overall boiler efficiency.
The scope of change of these variables. is limited, as is their effect
on the cost of the boiler design.
The choice of bed depth dictates the
fan power requirements; but once the fan size is chosen, the maximum bed
depth allowable is fixed. The only two variables which affect signi-
ficantly the cost and design of an industrial boiler are fluidizing
velocity and operating pressure.
234
-------
Fluidizing Velocity
With other design variables held constant, the bed area required
for a specific fluidized bed combustor is inversely proportional to
the superficial fluidizing velocity.
If one coal feed point is required
for a finite bed area, the total coal feed points -- and thus the
complexity of the coal feed system required -- will also be inversely
proportional to the superficial velocity. To minimize the elutriation
of bed material and coal, larger particles have to be used at higher
fluidizing velocities; this decreases the crushing cost.
A higher heat
transfer coefficient in the bed may also be expected, due to more
agitation of solids at higher velocities. However, this will probably
be partially offset by the increase of particle size, because the heat
transfer coefficient in the bed is inversely proportional to the bed
particle size.
Increase in fluidizing velocity tends to increase the heat-
releasing rate per unit bed area of boiler.
Thus, more heat transfer
surface can be located .in the bed, where a higher heat transfer
coefficient is possible.
However, there is a limit to how much heat
transfer surface can be put into a unit bed volume before generating
adverse effects on the coal distribution or quality of fluidization.
At some point an increase in bed depth will be necessary to accommodate
all the heat transfer surface in the bed; such an increase in bed depth
requires high operating fan power.
In the case of the atmospheric industrial boiler, the shop-
fabricated and railroad-transportable boiler is highly desirable because
of its lower cost, compared to that of the field-erected boiler. In
this respect, the selection of fluidizing velocity for design is very
important because it dictates the size of the boiler and thus determined
the possibility of shop fabrication. To evaluate the effect of fluidizing
velocity on the total cost of the boiler, the change in boiler cost
over the present Westinghouse - Erie City design
will be estimated.
235
-------
The present Westinghouse - Erie City design for a 250,000 lb/hr
fluidized bed boiler consists of two shipping modules -- a primary.
module with the dimensions 10 ft x 14 ft x 40 ft, and a secondary module
with the dimensions 8 ft x 15 ft x 40 ft, which are within the general
. shipping limitations of 13 ft x 16 ft x40 ft.
velocity in the primary beds is 12.5 ft/sec.
The designed fluidizing
Changing the design value
of fluidizing velocity should not affect greatly the design of the
secondary module, which houses primarily the carbon burn-up cell,
superheater, and the convection zone.
However, increasing the design
value of fluidizing velocity will decrease the bed area of the primary
beds and also require a smaller tube pitch for horizontal tubes in the
bed.
That means the primary module designed at 15 ft/sec should be
cheaper than that designed at 12.5 ft/sec.But higher fluidizing
velocity gives higher elutriation, attrition, and carbon carry-over.
Increase of fluidizing velocity to more than 15 ft/sec is probably not
practical at the present time.
Decreasing the design value of fluidizing velocity will
necessitate an increase in bed area of the primary beds. It can be
estimated that the largest shippable primary module can be designed at
a fluidizing velocity of ~10 ft/sec.
That means any further decrease in fluidizing velocity will
require that
shows that a
the primary module be erected on-site. Cost investigation
change of fluidizing velocity between 10 to 15 ft/sec
total boiler cost from the base case only <5% because only
changes the
the primary module (~25% of total boiler cost) is affected.
Cost analysis
for a boiler designed at 8 ft/sec which requires field erection of the
primary module gave a boiler capital cost of $8.60/lb/hr, compared to
$6.l0/lb/hr for the base case -- an increase of about 40%. This clearly
indicates that the major saving by designing at a higher fluidizing
velocity in the industrial boiler comes from the difference in cost
between shop fabrication ~nd field erection, and not from savings in bed
area and the coal feed system.
Thus, it is advantageous to design the
industrial fluidized bed boiler at the highest fluidizing velocity feasible.
236
-------
The present Westinghouse - Erie City industrial boiler design
is rather conservative. Evaluation of different arrangements for
heat transfer surface showed that the 250,000 lb/hr industrial boiler
can actually be housed in a single module [1081. This is accomplished by
installing the superheater and the convection section in the freeboard
of the primary module. If the boiler can be designed at a higher flui-
dizing velocity (~l5 ft/sec) and the heat transfer coefficient in the
bed can be improved (>50 Btu/ft2-hr-OF), the economizer can also be
situated in the freeboard area. This will reduce the capital cost sub-
stantially.
The potential of reducing the capital cost below the figures
presented here is excellent; .however, this requires further experimental
evidence.
237
-------
Pressurized Industrial Fluidized Bed Boiler
Pressurized operation requires less fluid bed area, permits
operation of a deeper bed, and allows immersion of more heat transfer
surface in the bed, where a higher heat transfer coefficient is possible,
and thus decreases the total heat transfer surface requirement.
Operation
at higher pressure in deeper beds may also increase combustion efficiency
and S02 removal efficiency because of increased residence time of coal
and gas in the bed.
However, gas turbines ar~ required to recover the
energy stored in high-pressure flue gas. Although this permits operation
of a higher-efficiency gas-stream combined cycle, it may also require
stringent particulate removal before the gas turbines and may lead to
erosion, corrosion, and long-term maintenance problems of turbine blades..
High-temperature and high-pressure piping, valving, ducting, particulate
removal equipment, and pneumatic solid transporting lines are inherently
more expensive.
Hence, a potential pressurized application in utility
boilers does not necessarily dictate its application in industrial
boilers, because of differences in scope of design and plant economics.
For pressurized operation, a boiler enclosed in a cylindrical
shell is preferred on the basis of structural strength considerations.
The cylindrical vessel can be placed horizontally or vertically, and
this configuration affects the design of the boiler significantly.
A
vertical arrangement. precludes natural circulation because there is no
simple way that the steam drum can be accommodated and still make the
boiler railroad-shippable. A steam drum can conveniently be put inside
a horizontal shell; however, the bed depth would then be limited by the
diameter of the vessel, which cannot exceed 13 feet because of shipping
restraints, and the advantages of the pressurized deep-bed operation are
completely lost. A vertical configuration permits use of a deeper bed.
For small industrial boilers, a single deep bed would be adequate; however,
in this case turn-down becomes a p~oblem. Control of bed temperature
alone would provide only about 30% turn~down of total boiler capacity
before S02 removal in the bed is greatly affected. The additional
238
-------
boiler turn-down would have to be provided through change of bed depth,
\
which is complicated and expensive. Additional problems for the
pressurized industrial boiler are:
.
More stringent particulate removal is needed before
gas turbines
.
High-temperature duct from boiler to gas turbines is
expensive..
High-pressure coal and limestone feeding
systems have to be used
.
For a vertical configuration, only once-through forced
circulation can be used.
All these necessities are inherently expensive compared to
the atmospheric design.
A small pressurized fluid bed industrial boiler of capacity
250,000 lb/hr steam will thus not be economical compared to the present
atmospheric industrial boiler.
However, the economic potential of the
pressurized industrial boiler increases with increases in capacity.
A boiler capacity of >50 MW will probably be economical because design
and economic analysis of a four-module, 300 MW, pressurized utility
boiler, each module of capacity 75 MW, showed considerable savings over
conventional coal-fired boilers.
239
-------
Recommendations
Fluidized boilers for ,industrial application do not have
significant overall cost and performance advantages over conventional
coal-fired boilers with pollution control.
On the basis of the market
projections for clean fuel for industrial boilers and the technical
,assessment, we recommend that
.
Further development work on industrial fluidized bed
boilers not be carried out at present
.
Fuel usage for industrial boilers be monitored to determine
any change from the use of clean fuels
.
Development of an industrial fluidized bed boiler be
reassessed if low-sulfur char or alternate low-sulfur,
low-grade fuels become available.
240
-------
ELECTRIC UTILITY APPLICATION
Introduction
Concepts for utility type fluidized hed hailers evolved from
the fluidized bed boiler technology which Westinghouse compiled.
Foster Wheelert using the preferred conceptst then made preliminary
designs of utility boilers in the capacity range of 300 to 600 MW.
Boiler subsystem costs were.estimated over this size range.
Using the ~-Foster Wheeler boiler designt, United Engineers
and Constructors made layouts of a power plant and estimated the cost
of generating energy in this plant.
241
-------
Pressurized Fluid Bed Boiler System
System Concept
The, system concept for the high,-pressure fluidized bed. boiler
in a combined.: cycle plant is shown schematically in' Figure L-77.
Combustion air is supplied to the bed. by the gas turbine compressor.
Coal is crushed and dried prior to being fed. into the bed, the material
of which. is. a sorbent.which removes the sulfur from the products of
combustion as the. coal is. burned.
The bed is maintained. at a temp'era-
ture below the agglomeration point of the. ash by heat transfer to the
stearn cycle. The spent dolomite is. regenerated and recycled; and..
make-up dolomite, to replace that blown down, is fed in wi.th the, coaL
All of. the ash from the coal is elutriated. from the bed., along
with a' significant: amount of unburned. carbon- and: attrited sorbent.
A
high percentage, of- the particulates from the primary bed is collected
in the primary separation and fed to. a fluidized bed, carbon burn-up.
cell which operates' at a higher t"emperature and. with a greater per.centage
of excess air than the. primary bed. . The. effluents from the pr.imary
separator- and the carbon burn-up celL are blended and passed through a
secondary particulate removal device to reduce the solids content' .to
the level required for attaining a: low erosion rate in the; gas .turbine.
The exhaust g!1ses' are, cooled,. af ter expansion through the. gas turbine,
by two stages of stack gas' coolers.
A third-stage particulate remova'l
device may be necessary to meet stack emission standards.
The st'eam cycle has one stage of reheat and regenerative'
feedwater heating 'which is p~rtially in parallel with the second-stage:
stack gas cooler.
The. sulfur' released. in the sorbent regeneration is conver.ted
to elemental sul
-------
-Gas FJ()I,'"
Solids Flow
-;-~~ Steam Cycle
Sulfur
Rich Gas
De$_ul fur i z i n9
Agent
Regenerator
"
...
w
Air-
Spent
Desulfurizing
Agen t
Regenerated
Stone
Waste
Stone
Make-Up
Stone
COIuprcssor
- - u- - -- - - - - - - -- ---
Primary
Particulate
Removal
Fluidized
Bed
Combus tor-
Desul furizer
Carbon
Burn-Up
Cell
Coal
(Oi I)
Make-Up
F eedwa t e r
COlllprcs~or
rur~jnc
----EJ
Fig. 1-77 -Pressurized fluid bed boiler power system
P
-------
Plant
Boiler
. Capacity - 300 MW
. Type - intermediate load
. Turn-down requirements - 4:1
. Transient requirements
5% per min.
. Geographic location - Eastern U. S.
. Pollution control
Particulate emissions - 0.01 gr/SCF
Sulfur removal - 95%
6
Nitrogen oxide emissions - < 0.7 lb/lO Btu
. Fuel - Ohio Pittsburgh No.8 Seam Coal
Proximate analysis (wt %)
Volatile matter
Fixed carbon
39.5
48.7
8.5
3.3
100.0
Ash
Moisture (as mined)
Sulfur content - 4.3%
Particle size - 1/4 in x 0
Complete data on coal is given on page 167.
. Sorbent - 1337 dolomite
Particle size - minus 1/4 in with minimum fines
Feed ratio - 6 x stoichiometric
Elutriation from beds - 5%
Loss during regeneration - 10%
Complete data on sorbent is given on page 543.
244
-------
. Bed Design Parameters
Primary beds
Temperature - l750°F design pt.
Size distribution of material in bed
Max. 5000 ~
Min. 500 ~
Mean 2500 ~
Superficial velocity - 6 to 9 ft/sec
Excess air - 10%
Density - 45 lb/ft3
Expanded height - 10 to 15 ft
Elutriated material
Ash - 100%
Carbon - 6%
Sorbent - 5%
Size Distribution - see page 127
CO in flue gas - < 0.5% by volume
Gas side heat transfer coefficients
In bed - 50 Btu/ft2 hr-oF
Above bed - 40 Btu/ft2 hr-oF
. Carbon Burn-Up Cell
Temperature - 2000°F design
Size distribution of material in bed
Max. 5000 ~
Min. 500 ~
Mean 2500 ~
Superficial velocity - 5.7 ft/sec
Excess air - at least 79%
Combustion efficiency
Density - 45 lb/ft3
Expanded height - 10
- 90% design pt.
to 15 ft
Recovery of carbon elutriated from primary beds - > 90%
245
-------
Elutriated material
Ash - 100%
Carbon -10%
.Sorbent - 100% ofsorbent in with recovered coal
plus .3% of .recirculated limestone.
Size distribution - see page 127
Gas-side heat transfer coefficients
.Same as for primary beds
.Stack gas temperature- 27.5°F
Air preheat temperature - .gas turbine compressor
outlet temperature.
Power Cycle Analysis
A computer program 'was wri,t.ten,to calculate plant capacity ~
plant :heatrate~ ':plantcost, ,and energy cost for combined, cycle ',power
systems using fluidized bed boilers.
Figure 1-78 is the generalized
block diagram of such a system.
This program provides for a wide
var,iety of power cycle configurations , including both sub- and
supercritical steam cycles.
Candidate .Steam Cycles.
The steam conditions ,for the ,No. 4
uni't at Hammond Stat.ion of Georgia Power and Light were used for the
subcri,tical case; ,those at the Homer City Station for the 'supercri,tical
case.
These conditions areas follows.:
Stibcritical
(Hammond 114)
Supercritical
(Homer City)
Boiler pressure~PSIA
Superheatertemp.-oF
Reheater 'temp.-oF
2400
1000
1000
3500
1000
1000
Condenser pressure-in. Hg
Feedwater heater 's,tages
3
6
3
7
246
-------
Dump
Dwg. 2924A68
o
@
@) Ai r
Block Diagram
Combined Cycle
Ai r
Ai r
To r-!~ -l +
condensert ~ 10000F
B@ @~LP (4) l~p ;- :. ~:;O~F F~~~d - Fuel- - Note,
~ T -,@@@(2) Boiler ~ l I. Diagram is for 3500
I fO\ 10\ a-. t?\ I PSIA Supercritical
~ "II 1@1~1~~01 I\::!,I I t@ '3r+ Boiler.. For 2400
+- ,T 'T T '<:!!I PS IA Boi I der, #8
'-J I I I I I I I I L - - - - - -l I Heater is Removed.
I I I I . I L - - - -, 2. For Non-Intercooled
I I I I I L_____-, - I GasTurbine,(2)=@
I I I L I I I Upper 3. TI is Compressor
Conden I I I L - - - - - - I I I LS.G.C. Turbine and T2 is
ser I L- - I I I I r Power Turbine
I L -, I I I 4. States @ to @ in
I I I I ~ cll- I Steam Cycle Represents
;, lr-~Q' ;~~~,~~~I (,I~~~ 7 @.~ B - -- - ~ @5. ~~~;:d R*::::n~:eam
- - Final FW Temperature
@ @ @ I @6" I @ ~ @ I to Boiler.
L - - ~ ~ '--...J '~ ---1 '* - ..J L '.J}-e.-J '-.. ..J '--4 ..J J
L -cr- -.- - - - lower - - -
~ S. G. C.
Products
o
T2 @ - G
Inlet
Stack
Fig. 1-78-1-1igh pressure fluid bed boiler power system
-------
In addition, the following assumptions were made for the boiler and the
plant:
Stack temperature
Excess air
1750°F'
l600°F(1)
275°F
Bed temperature
Turbine inlet temperature
Boiler efficiency
Auxiliary load
Annual capital charge
10%
89.6%
3%
Plant capacity factor
Fuel cost
15%
0.8
6
25 cents/10 Btu
Operating and maintenance cost
(including limestone regeneration)
Boiler costs:
0.5 mills/kwh
$20/steam kilowatt
$25/steam kilowatt
3%
2400 psia
3500 psia
Gas turbine cycle pressure drop
The cost of plant components other than the boilers was
obtained from a breakdown of costs for conventional plants over a
capacity range of 300 to 600 MW.
Preferred Cycle Determination.
First, computations were made
to determine the optimum pressure ratio for this combined cycle system.
Figures 1-79 through 1-82 show the results for a boiler pressure of
2400 psia and a gas turbine airflow of 650 lb/sec, (2) and Figures 1-83
through 1-86 show the results for a boiler pressure of 2400 psia and
a gas turbine airflow of 1300 lb/sec. . The effect of compressor inter-
cooling was also determined for both these condition sets.
These
plots indicate that the best performance and the lowest cost occur
with no compressor intercooling at a pressure ratio less than 10:1
for both gas turbine airflow rates.
However, since the typical large
(1) It is assumed that the internal and external heat losses from the
hot gas would result in a difference of 150°F between the bed
temperature and the turbine inlet temperature.
(2) This is the airflow of the Westinghouse W50l gas turbine.
248
-------
Curve' 646689-A
,9600
9500
9400
Boiler Press. = 2400 psia
Gas Turbine Airflow = 650 Ib/sec
Non-IQtercooled (one-shaft I
9300
9200
'i 9100
~
:::>
~ 9000
'"
~ 8900
Boller Effie = 89. ff!o
Auxiliary Load = 3% ,
Stack Temperature = 275°F
Turbine Temperature = 1600
Back Pressure = 3" Hga
~ 8800
8700
9300
i 9200
...
~ 9100
~ 9000
'16
~ 8900
880010
15
20
Overall p/R
Figure 1-79-Elfect of first compressor pressure ratio
and overall pressure ratio on net plant heat rate
Curve 646697.A
190
180
Boiler Pressure-2400 psia
Gas Turbine Airflow = 650 Ib/sec
Non- Intercooled lone-sh
Boiler Elfie = 89. ff!o
Auxiliary Load = 3%
Stack Temperature = 275°F
Turbine Temperature = l0000F
Back Pressure = 3" Hga
150
140
130
180
170
Intercooled To loo~
~IOO
..,..
150
140
13Dl0
Figure 1-81-Effect of first compressor pressure ratio
, and overall pressure ratio on plant cost
30
30
Curve 646691-A
330
320
, Boiler Effie = 89. ff!o
Auxiliary Load = 3%
Stack Temperature = 2750f
Tu rbi ne Te mperatu re = 1600~
Back Pressure = 3" Hga
Boiler Pressure = 2400 psia
Gas Turbi ne Airflow-650 Ib/sec
~ 310,
310
, Non- Intercooled lone-shaft I
,300 '
330
Intercooled, To looOf
",320
E
i 6.2
...
~ 6.0
~
~
~ 6.2
-
;:g 6. 0
:E
249
Fi rst Compressor pI R = 4
3
310
300
10
30
15
Figure 'l-BO-Effect of first compressor pressure ratio
and overall pressure ratio on plant output
Cu rve 646690-A
7,0
6.8
Boi ler Pressure = 2400 psi a
Gas Turbine Airflow = 650 Ib/sec
Non- Intercooled (One-Shaft J
6.6
6.4
5.8
5.6
6.6
6.4
Boiler Elfie = 89. fJ'I.
Axuxiliary Load = 3%
Stack Temperature = 275~
Turbine Temperatu re = 1600~
Back Pressure = 3" Hga
Intercooled to 100~
First Compressor p/R = 4
5.8
5.610
30
Figure 1"82 -Effect of first compressor pressure
'ratio and overall pressure ratio
one nergy cost
-------
Curve 646695-A
200
Boiler Pressure = 2400 psia 6.8
190
Gas Turbine Airflow-13OO Ib/sec 6.6
180 Boiler Effie = 89. (ff.
Auxiliary Load = 3% c:: 6.4
170 Stack Temperature = 275"f :I:
~
:3 Turbine Temperature = 1600"f '" 6.2
-;;;160 Back Pressure = 3" Hga -
V>
~ 6.0
150 Non- Intercooled (one-shaft)
5.8
140
5.6
130
180 6.6
6.4 Intercooled To 100"f
170
'"
160 ~ 6.2
:3 '"
~ 150 Intercooled To 100°F ~ 6.0
~
140 5.8
13010 30 5.610
~ 9100
~ 9000
~ 8900
'"
~ 8800
:I:
~
~ 9200
,~ 9100
~ 9000
'"
~ 8900
:I:
Curve 646694.A
9400
9300
9200
670
650
640
630
:3
E 620
610
600
650
640
630
30 610
600
10
Boiler Effie = 89. (ff.
Auxiliary Load = 3%
Stack Temperatu re = 275"F
Turbine Temperature = 16000F
Back P ressu re = 3" Hga
8700
9300
Intercooled To 100°F
First Compressor p/R =
3
880010
15
20
Overall p/R
Figure 1-83-Effect of first compressor pressure
ratio and overall pressure ratio on
net plant heat rate
Figure 1-85-Effect of first compressor pressure ratio
and overall pressure ratio on plant cost
250
Curve 646704-A
Boiler Pressure = 2400 psia
Gas Turbine Airflow = 13OOIb/sec
Boiler Effie = 89. (ff.
Auxiliary Load = 3%
Stack Temperature = 275°F
Tu rbi ne Te mperatu re = 1600°F
Back Pressure = 3" Hga
Intercooled to loo"F
First Compressor p/R = 4
15
30
Figure 1-84-Effect of first compressor pressure ratio
and overall pressure ratio on plant output
Cu rve 64669B-A
Boiler Pressure = 2400 psia
Gas Turbine Airflow = 1300 Ib/sec
Boiler Effie = 89. (ff.
Auxiliary Load = 3%
Stack Temperature = 275"F
Turbine Temperature = 1600"f
Back Pressure = 3" Hga
First Compressor P
3
15
30
20
OVerall pI R
25
Figure 1-86-Effect of first compressor pressure ratio
and overall pressure ratio on energy cost
-------
utility type gas turbine has a pressure ratio of about 10:1, this is
taken to be the minimum practical value.
Figures 1-87 through 1-94 show the results of the supercritical
steam conditions and gas turbine airflows of 650 and 1300 lb/sec.
Here,
again, the performance and costs are best at an overall pressure ratio
of 10:1 with 'no compressor intercooling.
Comparing the results for
3500 psia with those for 2400 ps{a shows that the heat rate is lower
and the power higher for supercritical steam conditions than for
subcritica1 conditions. However, the specific plant cost and the
energy cost are higher for supercritical than for subcritical conditions.
In view of the above, it was decided that the optimum conditions
for use in making the preliminary design of the combined cycle plant
with a fluidized bed boiler are:
. Simple cycle gas turbine with a pressure ratio of 10:1
and without compressor intercooling
. Subcritical (2400 psia) steam conditions for both 318 and
.635 MW level.
.
Design Point Conditions. .The design point conditions [or
the combined cycle plant with the fluidizec.l bec.l boiler are as foll.ows:
CAPACITY - MW
318 I 635
Boiler pressure-psia 2400 2400
Initial steam flow-lb/hr 1,727,020 3,454,040
Steam temperature-OF
Initial 1000 1000
Reheat 1000 1000
Condenser pressure-in.Hg 3 3
Final FW temperature-OF 578 578
Compressor airflow-lb/sec 650 1300
Gas turbine temperature-OF 1600 1600
251
-------
~
j 9100
~
;S 9000
'"
~ 8900
! 9200 -
:2 9100
co
~ 9000
""
A 100
..,
" 100
~
'";;> 150
9500
9400
Cu rye 6466924A
9300
9200
,
Boiler Pressure = 3500 psia
Gas Turbine Airflow = 650 Ib/sec
Boiler Effie = 89. ff!o -
Auxi liary Load = 3%
Stack Temperature = 275"F
Turbine Temperature = loo0"F
Back Pressure = 3" Huga
8700
9300
Intercooled To 100"F
First Compressor P/R = 4
3
8800
10
Figure 1-87 -Effect of first compressor pressure ratio
and overall pressure ratio on net plant heat rate
Curve 6466994A
190
180 Non- Intercooled (one-sha
170
Boiler Pressure = 3500 psia
Gas Turbine Airflow = 650 Iblsec
Boiler Effie = 89. ff!o
Auxiliary Load = 3%
Stack Temperature = 275°F
Turbine Temperature = 1600°F
Back Pressure = 3" Hga
150
140
130
180
170
Intercooled To 100"F
140
131\0
25
Figure 1-89-Effect of first compressor pressure ratio
and overall pressure ratio on plant cost
30
30
252
Curve 646688-A
330 Non-Intercooled
(one-shaft)
320
Boiler Pressure = 3500
Gas Turbine Airflow = 650Iblsec
Boiler Effie = 89. ff!o
Auxiliary Load = 3%
Stack Temperature = 275"F
Turbine Temperature = lOOO"F
Back Pressure = 3" Hga
"
E
"
E
310
300
330 I
320 Intercooled To 100°F
I
Flrstl Compressor
P/R = 4 '
310 -
300
10
-
I
15
I
25
I
20
Overall P/R
30
Figure 1-88 -Effect of first compressor pressure ratio
and overall pressure ratio on plant output
7.0
Curve 646702-A
6.8
6.6
Boiler Pressure = 3500 psia
Gas Turbine Airflow = 650 Iblsec
Non-Intercooled (one-shaft)
"" 6.4
:I:
~ 62
~.
Boi ler Effie = 89. ff!o
Auxiliary Load = 3%
Stack Temperature 275°F
Tu rbi ne Temperatu re = 1600°F
Back Pressure = 3" Hga
~ 6.0
5.8
5.6
6.6
6.4
0::
~ 6.2
""
~ 6.0
3
~
5.8
5.6lO
30
Figure 1-90-Effect of first compressor pressure ratio
and overall pressure ratio on energy cost -
-------
Curve 646693-A
9500
9400
9300
Boiler Pressure - 3500 psia
Gas Turbine Airflow- 1300 Ib/sec
Boi ler Effie = 89.6%
Auxiliary Load = 3%
Stack Temperatu re = 275°F
Turbine Temperature = lWO"F
Back Pressure = 3" Hga
~ 9200
:;
Q5 9100
Q,
~ 9000
':;j
~ 8900 -
8800
8700
9300
i 9200
...
~ 9100
0.>
~ 9000
~ 8900
880010
Figure 1-91-Effect of first compressor pressure ratio
and overall pressure ratio on net plant heat rate
Curve 646700-A.
200
Boiler Pressure = 3500 psia
190 Gas Turbine Airflow = 1300 Ib/sec
180 Boiler Effie = 89.6%
Auxiliary Load = 3%
170 Stack Temperature = 275"F
:3 Turbine Temperature = 1600"F
.:::: lW Back Pressu re = 3" Hga
...
150
Non- Intercooled (one-shaft I
140
130
180
170
:3 lW
... Intercooled To 1000F
-
-
150
140
13010 30
Figure 1-93-Effect of first compressor pressure ratio
and overall pressure ratio on plant cost
30
:3 630
E
620
610
:3 630
E
0::
~ 6.2
""
~ 6.0
Curve 646703-A
650
640
Boiler Effie = 89.6%
Auxiliary Load = 3%
Stack Temperatu re = 275°F
Turbine Temperature = 1600°F
Back Pressure = 3" Hga
Boiler Pressure = 3500 psia
Gas Turbine Airflow = 1300 Ib/sec
600
650
640
620
610
Intercooled To 100°F
WOW
15
20
Overall p/R
25
30
Figure 1-92-Effect of first compressor pressure ratio
and overall pressure ratio on plant output
7.0
Curve 6/+£696-A
Boiler Pressure = 3500 psi
6.8 Gas Turbine Turbine Airflow = 130 Ib/sec
Boiler Effie = 89.6%
Auxi liary Load = 3%
Stack Te mperatu re = 275"F
Turbine Temperature = 1600
Back Pressure = 3" Hga
6.6
6.4
~
5.8 Non- Intercooled (one-shaft)
5.6
6.6
6.4 Intercooled To 100°F
0::
I
~ 6.2
-
~ 6.0 First Compressor p/R - 4
~
5.8
5.610 20 25 30
Overall p/R
253
Figure 1-94-Effect of first compressor pressure ratio
and overall pressure ratio on energy cost
-------
CAPACITY - MW
318 I 635
Gas turbine pressure ratio 10:1 10:1
Turbine cooling air 5% 5%
Gas turbine pressure drop 3% 3%
Excess air 10% 10%
Stack temperature-OF 275 .275
Gas turbine PWR,...KW 56 ,')')7 II'l,II'!
Steam turbine PWR-KW 270,029 538,400
Net' plant power-KW 318,485 635,362
Plant heat rate-Btu/kwh 8953 8974
Appendix
J , prepared by United Engineers and Constructors,
contains a flow diagram of the combined cycle plant.
It is reproduced
in Figure 1-95. The design point thermodynamic data are given at
each station on this diagram.
Gas Turbine Pressure Loss Effect.
In calculating the plant
performance for the above design points, the pressure loss between the
gas turbine compressor outlet and turbine inlet was assumed to be 3%.
Since many system components contribute to this pressure, 108s, and in
most cases there are trade-offs between component cost and pressure
drop, a sensitivity analys'is was made of the effect of the gas turbine
pressure losses on the plant capacity and heat rate.
The results of
this analysis are shown in Table 1-36 over a range of pressure losses
from 3% to 10%.
This shows that a change of one percentage point in
pressure loss between the gas turbine compressor exit and turbine inlet
causes a 0.1% decrease in plant power and a 0.1% increase in plant
heat rate.
The reasons for the relatively small effect of gas turbine
pressure loss on plant capacity and heat rate are:
. Only the gas turbine is directly affected, and it
contributes a minor fraction of the plant power
. Decrease in gas turbine power results in higher gas turbine
exit temperature and. increased steam power.
254
-------
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SHAM TUR8'NE OUTPUT
IOAS TURBINE OUTPUT
TOTAL PLANT OUTPUT
AUXiliARy POWER
NET PLA.N OUTPUT
PLANT MEAT RAT[
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116.2"
64.lh
CONDENSATE
PUMP
o9P"
635 MW FLUIDIZED BED
BOILER COMBINED CYCLE PLANT
PERFORMANCE DIAGRAM
S... 1I1(IInIB8.-..
-------
Table 1-36
EFFECT OF GAS TURBINE PRESSURE DROP
ON DESIGN POINT PLANT PERFORMANCE
PRESSURE GAS TURBINE STEAM TURBINE NET PLANT PLANT HEAT
DROP (a) POWER -- MW POWER -- MW POWER -- MW RATE -- B TO /KW HR
0.03 113.1 538.4 635.4 .8975
0.04 112.1 538.9 634.8 8983
0.05 111.0 539.4 634.2 8992
0.06 109.9 539.8 633.6 9000
0.07 108.8 540.3 633.0 9009
0.08 107.7 540.3 632.3 9018
0.09 106.6 541.3 631.7 9028
0.10 105.5 541.8 631.0 9037
(a). .
Pressure loss between gas turbine compressor outlet and turbine inlet divided
by compr~ssor outlet pressure.
257
-------
Condenser.Pressure Effect.
A s team condenser pressure "o'f
.3 in. Hg was assumed for the design point calculations. This is
'typical fora power plant with a cooling tower; a pressure of
1-1/2 in. Hg ,is typical for a plant with 'heatrejectionto a body of
water. .Sincecooling .water temperature 'is a major variable among plant
. sites., .designpoint' calculations were made for the antIcipated range
of condenser:pressures.
These are shown in Table 1-37..
. .
Pal'.t :Load .Performance.
The .turn-down requirements for
fluid bed combined cycle ,power plant have 'been specified as 75% of
design .load~ As .in any p.ower sys.tem, the output of -theconibined cycle
plant with fluidi:zed bed boIlers is roughly proportionate to .the fuel
.input.
In m~ny combus.tion systems effici'entburningrequires a nearly
constant fuel-air ratio, so the air rate mus t be maintained approximately
.proportionate .to fuel :rate.
.In. flU:idizedbed combus tion, however., $ood
burning can be obtained over :a 'wide range of 'fuel-air ratios.
Ano.ther 'possible cons:traint on air rate is that imposed by
the hydrodynamic requirements of the bed. Figure :1-51 shows that the
,ratio of 'maxImum to m'J.nimum superficial velocities for a given bed
'ma.ter.ia-l .18 of the-order of 10: 1.
Since there are nu severe cOllsLraints
on either fuel-air :ratio or airflow 'rat.e .in fluidized ,bed combustion"
these two 'variables areef.fective :parameters for .controlling energy
.inputto the fluidized bed boiler..
.The 'steam cycle :power output .is.pr:imari.1,y .a .function of 'the
heat transferred to the working fluid i-n the fluidized :bed boiler, i.e.,
'q .='U S tiT.'
LM
(1-32)
where
q '=total heat transferr.edtothe working fluid~Btu/hr
.U = .ar.ea weighted average .overall heat transfer coefficient-
Btu/(hr) (ft2) (OE)
S= :total .hea.t transfer surface in the :fluidized .bedboiler
both .in and out of 'the ~ed-ft2
tiT = area weighted .averagelogmean .temperature :difference
1M
between the .bed or hot gas andthewo.rking fluid-oF..
258
-------
TABLE 1-37
EFFECT OF CONDENSER PRESSURE ON DESIGN POINT PLANT PERFORMANCE
CONDENSER PRESSURE - in Hg 1 2 3
Boiler pressure--price 2,400. 2,400 2,400
Steam temperature--OF
Ini tia1 1,000 1,000 1,000
Reheat 1,000 1,000 1,000
Compressor airf10w--1b/sec 1,300 1,300 1, 300
Gas turbine temperature--:-°F 1,600 1,600 .1,600
Gas turbine pressure ratio 10:1 10:1 10:1
Gas turbine pressure 10ss--% .3 3 3
Gas turbine cooling air--% 5 5 5
Excess air--% 10 10 10
. Ini tia1 steam f 10w--1b /h r 3,454,040 3,454,040 3,454,040
Final feedwater temp.--oF 578 578 578
Gas turb'ine power--kw .113,114 113,114 113,113
Steam turbine power--kw 575,217 553,350 538,400
Net plant power-~kw 671,075 649,864 635,362
Plant heat rate--Btu/kw hr 8,497 8,775 8,974
259
-------
The averall heat transfer caefficient decreases slightly with decreasing
steam flaw because of Reynolds number effects. The temperature af the
warking fluid in the.Rankine cycle is fixed with respect to. plant laad
fo.r bath constant pressure and.variable pressure systems. . Therefare,
far a fixed heat transfer surface area, the anly independent parameter.
. .
far cantralling steam cycle power is bed temperature.
There are definite limitations an the operating temperature
af a fluidized bed wheri burning coal. The maximum value af 'V1750°F is
set by the a~h saftening characteristics, and the minimum temperature,
abaut 1600°F far limestane and 1400°F far dalamite, is impased by the
reactian.kinetics af the sulfur sarbent.
. When the range af variability af the bed temperature is
. .
. .
limited, the capacity far turn-dawn af steam pawer in a fluidized bed
bailer can be extended by varying the amaunt af heat transfer surface
immersed in the bed. This can best be accamplished by reducing the
bed depth and uncavering a part af the heat transfer surface, expasing
it to. the hat gas effluent from the bed, fram which ~t absarbs heat.
Hawever, the heat transfer caefficient is significantly. lqwer for
surface in the averhead than for surface in the bed, so. there isa
change in the area weighted average heat transfer caefficient.
The power autput af the gas turbine is a function of the
turbine inlet temperature, the compressor pressure ratio, and the
campressor airflaw for fixed camp anent eff.iciencies, Le.,
PGT = ~(Tg, PRc, Wa).
(1-33)
is canstant and PRc = $ (Wa, Tg).
canfiguratian
Therefare, far a
Wg/Ti
number ( P .)
fixed gas turbine
For a given turbine configuration, the turbine flow
PGT = (Tg, Wa).
(1-34)
260
-------
The gas temperature out of the fluid bed boiler is a function of the
bed temperature, which makes the gas turbine power output a function
of the bed temperature and the compressor airflow,
i.e. ,
PCT = ~(Tb' Wa).
(1-35)
The axial flow compressor used in gas turbines has the
pressure ratio versus flow rate characteristics shown in Figure 1-96.
For a fixed speed, the flow rate is essentially constant over a wide range
of pressure ratios.
Airflow, however, can be reduced by approximately
1/3 by the use of variable inlet guide vanes.
It has been shown that the energy input to the boiler, the
steam cycle power, and the gas turbine power in the combined cycle
fluidized bed boiler are a function of the following four parameters:
. Airflow rate
. Fuel-air ratio
. Bed temperature
. Bed depth.
There are three possible modes of load control for the combined
cycle fluidized bed boiler using different combinations of these
parameters:
Mode 1
. Constant fuel-air ratio and bed temperature
. Variable airflow and bed depth
. Fuel rate decreases to maintain constant fuel-air ratio
. Bed depth is reduced simultaneously with airflow to maintain
constant bed temperature; turbine inlet pressure decreases
to satisfy constant turbine flow number
. Steam power decreases due to reduced U; gas turbine power
decreases due to lower Wa.
261
-------
130
120
N
0'\
N
Q)
.2 110
ro
>
C 100
.~
~ 90
"C
-+-'
~ 80
u
!....
:?5.. 70
~
::: 60
ro
e:::
50
!....
::J
~ 40
Q)
!....
a.. 30
20
100
Axial Compressor Field
. PIC = 14.1 psia
T lC = 80. OaF .
Curve 643408-A
Li nes of
Constant
EfficiencY-l1c
. "'~
" \" \ '"
:\~
c.,~
Lines of Constant
Speed - N/(B
10 20 30 40 50. 60 70 80 90
. W
Fig. 1-96- Discharge-Flow ar:- per cent design value
100, 110
120
-------
Hode 2
. Constant Wa and bed temperature
. Variable bed depth and fuel-air ratio
. Bed depth is reduced simultaneously with fuel-air ratio to
maintain constant bed temperature
. Turbine inlet pressure decreases slightly to satisfy constant
turbine flow number with reduced fuel flow
. Steam power decreases due to reduced U; gas turbine power
would be nearly constant.
Mode 3
. Constant bed depth and airflow
. Variable fuel-air ratio and bed temperature,
. Gas turbine inlet pressure reduces to satisfy constant
turbine flow number
. Steam power decreases due to lower bed temperature
. Gas turbine power decreases due to lower gas temperature.
The Westinghouse W50l gas turbine has a design airflow of
650 lb/sec.
By using variable inlet guide vanes, the airflow can be
reduced by 1/3 to 433.5 lb/sec.
Computations have shown that when
load control mode 1 is used, the system output from this combined-cycle
plant is nearly proportional to airflow. Using mode 2 sequentially with
mode 1 can reduce the system output to as low as 45% of the design
'value.
This indicates that the use of mode 2 alone can effect a power
reduction of as much as 34%. The performance data for part-load
operation using modes 1 and 2 in sequence are given .in Table 1-38.
The part-load heat rate for the fluid bed combined cycle system with
operating modes 1 and 2 in sequence is plotted in Figure 1-97, along
with that for a conventional coal-fired plant. This shows that there
is a minimum heat rate at about 75% load which is over 1% lower than
the design point heat rate.
263
-------
12, 000
11,000
10,000
9900
9800
9700
~ 9600
.r::.
::
~ 9500
-...
::3
to 9400
I
Q.)
-
"'
~. 9300
-
"'
Q.)
::I: 9200
9100
9000
8900
88000
Cu rove 64670 I-A
Back Pressure = 3 in Hg
Auxiliary Power = 3%
Overall Comp. PR = 10
Non- Intercooled
Ha m mond #4 2400 A 1000/100
Fluid Bed Boiler 2400A 1000/1000
20
40 60
% Load
80
100
Figure 1-97-Part Load characteristics fluid bed combined
cycle with operati ng modes 1 & 2 in sequ ence
264
-------
TABLE 1-38
PART LOAD PERFORMANCE UNDER OPERATION MODES 1 AND 2
FRACTION FUEL EXCESS HEAT P1AJ.~T
STEAM STEAM FLOW AIRFLOW AIR AIR RATE OUTPUT
POWER LB/HR LB/SEC RATIO % BTUjKWH KILOWATTS
Max. Guarantee 3,454,040 1300 .091937 10 8974 635344
100% 3,386,314 1278.9 .091937 10 8968 625551
80% 2,709,051 1049.8 .091937 10 8873 519033
60% 2,03lt788 867 .086593 16.8 8899 402533
40% 1,354t526 867 .062535 61. 7 9057 285630
Two modes of load control have been considered for the high-
pressure fluidized bed boiler power system wherein the bed temperature
is held constant over the full range of load.
One mode varies airflow
keeping the fuel-air ratio constant; the other varies the fuel-air
ratiot keeping airflow constant.
Any mode wherein bed temperature is
kept constant requires that the bed depth be varied to satisfy the
heat transfer relationship.
The combined cycle performance calculation program was used
to generate a matrix of performance data for a series of turbine inlet
temperatures and fuel-air ratio value sets (see Table 1-39).
A
procedure for determining a system operating line for variable bed
temperature from this matrix was developed.
By definitiont the system
operating line is a locus of condition sets wherein the heat transferred
from the bed to the steam cycle matches the power derived from the
steam cycle.
For the high-pressure system, all of the heat transfer surface
for the steam cycle.--except for feedwater heating--is located in the
fluicized bed or in the transition zone above the bedt where the heat
265
-------
TABLE 1-39
PERFORMANCE DATA FOR VARIABLE BED TEMPERATURES
Steam Cycle Load Fraction (ref.) 1.02 1.00 0.8 0.6 0.4
Turbine inlet temp. - of 1600
Plant power - mw 635.4 627.4 543.3 446.2 325.6
Steam power - mw 538.4 530.5 446.9 350.1 229.3
Plant heat rate - Btu/kw hr 8974 8967 8869 8846 8952
Fuel-air ratio 0.0919 0.0907 0 . 0 T77 0.0636 0.0470
Turbine inlet temp. - of 1500
Plant power - mw 619.5 611. 6 527.4 430.5 311.5
Steam power - mw 537.5 529.5 445.7 348.8 229.3
Plant heat rate - Btu/kw hr 9091 9080 8999 8998 9165
Fuel-air ratio 0.0908 0.0895 0.0765 0.0625 0.0460
Turbine inlet temp. - of 1400
Plant power - mw 603.5 595.6 511. 5 414.9 297.3
Steam power - mw 536.1 528.2 444.1 347.2 229.3
Plant heat rate - Btu/kw hr 9211 9202 9134 9161 9392
Fuel-air ratio 0.0896 0.0884 0.0753 0.0613 0.0451
266
-------
transfer coefficient is nearly as high as it is in the bed.
The heat
transfer equation q. = U.A.~TLM has to be satisfied for each bed. The
1. 1. 1. i
overall heat transfer coefficient varies measurably among the evaporator,
the superheater, and the reheater because of differences in steam-side
coefficients; it also varies with load because of Reynolds number
effects.
Therefore:
q. = U.A.lITLM
1. 1. 1. i
(1-36)
and
T) - T)
q = ~iq1.' - ~.U.A.lITLM .
1. 1. 1. .
1.
(1-37)
Also:
q. = W. (h .
1. 1. 1.out
- h. ).
1..
1.n
(1-38)
Combining 1-37 and 1-38 gives
A =
i
W. (h.
1. 1.
out
U. ~TLM
1. .
1.
- h. )
1..
1.n
For constant bed depth
A = ~~A. = C .
1. 1.
For a constant pressure steam cycle, the q. values for a given
1.
fraction of reference steam power are constant with variable bed
temperature in all beds except the one containing the pre-evaporator.
Therefore, the distribution of heat absorbed by the various beds shifts
with bed temperature, and for this reason the temperatures of all the
267
-------
beds cannot be the same at reduced load.
It was assumed for this analysis
that the functions of the pre-evaporator bed and the two superheater beds
would be lumped and that the temperatures of these beds would be
maintained the same at all loads.
The temperature of the reheat bed
would therefore be different from that of the other beds.
The conditions and assumptions used in this analysis are as
follows:
. Temperature and pressure profile at design point (see
Figure 1-98)
t = 578.30F P = 2632 psig
1 1
t = 675°F
2
t = 671°F
3
t = 10000F P = 2486 psig
4 4
t = 650°F P = 601 psig
5 5
t - 10000F P = 581 psig
6 - 6
t = 1750°F
B
. It has been assumed that the bed temperature is l750°F
and the gas turbine inlet temperature is l600°F. This
difference would decrease with decreasing bed temperature.
It was assumed that the difference between these two
temperatures would vary linearly from 150°F at a bed
temperature of l750°F to zero at a bed temperature of
550°F.
. The carbon burn-up cell performs a portion of the pre-
evaporation function as shown in Figure 1-99.
It was
assumed that the heat absorbed by the steam cycle from
the carbon burn-up cell would be a constant 5% of the
total heat absorbed over the load range.
If a constant
temperature must be maintained in the carbon burn-up
cell, the bed depth will have to be varied accordingly.
268
-------
Dwg. 6169A24
@
(6)
.:::/
T
@
S
Fig. 1-98 - Steam cycle T-S diagram
269
-------
4000
N
"
o
u...
o .
0>' 3000
~
:J
-
C'CI
~
0>
c..
. ~ 2COO
I
I
I
Balance of Boiler I
System Losses:
I Adiabatic Flame
I Temperature
I
I
I
I
I
I
I
I
1750 (FBCD)
I
I
'Tube Waif Temp.
: Steam Temp.
I
I
I
Feeawater
H~ater
I I I
I I I
I : Pre-Evaporator ...1
I I : . --- - -. ....'"
I I --- :.......~
I r------II__.,-_-' ~ i
I I I , Re-
" I Evaporator I I Superheater Heater I
I I I
I I
5 6 7 8 9 10 II 12 13 14
Energy from MAF Fuel - 1000 Btu/lb
Stack
1000 Losses
o
1
2
4
Fig. 1-99-Temperature-energy diagram for pressurized utility boiler
Curve 645594-A
15
-------
. It was assumed that the enclosure wall surface where
evaporation takes place is exposed to the same temperature
as the lumped beds.
. The feedwater temperature as a function of the fraction
of steam power and the gas turbine inlet temperature is
given in Figure 1-100.
. The temperature of the steam returning to the reheater
is given as a function of the fraction of steam power
in Figure 1-101.
. Overall Heat Transfer Coefficient
FRACTION STEAM POWER (REF)
1.02(1) 0.8(2) 0.6(2)
Pre-evaporator 47 46 45
Evaporator 47 46 45
Superheater 43 42 41
Reheater 39 38 37
(1) Calculated by Foster Wheeler
(2)Assumed
271
-------
700
700
650 650
u...
°
600 T - of ~. 600
GT ::J
ro
1600° '-
Q)
<:>.
u... 15000 E
° Q)
I f-
~ 550 14000 ~ 550
c:
f- .....
N
-.J '-
N Q)
ro
Q)
..c:
500 ~ 500
450
450
400
o
0.2
0.4 0.6 0.8
Fraction Steam Power (Ref. I
1.2
400
o
1.0
0.4 0.6 0.3
Fraction Steam Power (Ref.)
Fig. 1-101- Temperature of steam to reheater in high-pressure fluidized
bed boi ler
0.2
1.0
1.2
Fig. 1-100- Feecmater temperature in high-pressure fluidized bed boiler
-------
The operating lines for the lumped beds and the reheater were
determined for the above conditions and assumptions.
These are shown
plotted on the performance matrix in Figure 1-102.
This shows that the
reheater bed will operate at a higher temperature than will the other
beds.
The operating line for the reheat bed is shown for reference
only and is valid only- for bed temperature and fuel-air ratio, since it
obviously can have no significance for heat rate.
No correction was
made for the effect of the higher temperature of the reheat bed or for
the gas turbine inlet temperature.
The minimum operating bed temperature is determined by the
requirements for absorption of the sulfur released from the coal by the
bed material. For limestone the projected minimum bed temperature is
>1600°F; for dolomite it is l300°F to l400°F. Figure 1-102 shows that
the power for a minimum bed temperature of l600°F is 528 MW, or 83% of
design power, whereas for bed temperatures of l300°F and l400°F the powers
are 325 MW (51%) and 395 MW (62%) respectively.
Reducing the compressor airflow by 1/3, with a constant bed
height and a bed temperature of.1300°F, would not affect the steam power
but would reduce the gas turbine power by about 15 MW to about 48% plant
load. The part load performance line for the fluid bed combined cycle
plant with operating mode 3 is also superior to that of a conventional
coal-fired plant as shown in Figure 1-97, but it is not as good as for
operating modes 1 and 2 in sequence.
Since varying the bed depth requires expensive materials
handling equipment to avoid a large response time constant, an operating
mode which uses constant bed temperature is not considered to be the
optimum turn-down method. In addition, the turn-down capability with
constant bed height is nearly as good as that with variable bed height.
Therefore, the preferred method for controlling load in the high-pressure
fluidized bed boiler is to vary the bed temperature, possibly in com-
bination with variation in airflow.
273
-------
11. 0
10.8
10.6
10.4
"" 10.2
10
.....
x
e::: 10.0
:J:
s:
~
~
'" to-
..., CD
"'"" 9.6
QJ
n:;
e:::
n:; 9.4
QJ
:J:
9.2
9.0
8.8
8.6
0 50 100
Curv~ 643383-6
Primary Beds
Reheat Bed (Ref. )
Wf /Wa x 102
\
\ .
06 0.8
. Fraction Steam Power (Ref.)
1. 02
650
700
Fig. 1-102-operati ng lines for high pressure fl uidized bed system with variable bed temperature load control
-------
The turn-down capability of a single fluidized bed boiler
coupled with a single gas turbine is only about 50%, whereas the
specified turn-down for this power plant is 4:1.
Therefore, a modular
plant configuration will be necessary to meet the turn-down requirements.
The 318 MW plant gas turbine requirements could be satisfied by either
one W50l gas turbine or two W25l gas turbines.
In view of the turn-
down requirements, two W25l gas turbines were selected for this size
plant.
Two W50l gas turbines would be required for a 635 MW plant and
one or more boiler modules would be required for each gas turbine module.
is as
The relationship between plant and module turn-down capabilities
follows:
}L =.! X
~L' 1 n m
(1-39)
where
~ .=
1
minimum load fraction of plant with one module operating,
n
=
number of modules in plant, and
X
m
= minimum load fraction of module.
The values of X required to give a plant turn-down ratio of 4:1 are
m
n X
m
1 0.25
2 0.50
3 0.75
4 1.00
In addition to being able to meet the minimum load, it is
advantageous to have continuity in plant load between the minimum and
full load values.
Continuity is satisfied when
275
-------
~i '
= X P(i-1)
~i i 1 X
= . - .
n m
~(i-1)
i-1
=-
n
X
m
=
i-1
i
for i values greater than 1
X
m
:f
o (n)
where
~i = minimum load fraction with i modules operating and
,
XP(i-1) = maximum load fraction with i-1 modules operating.
This shows that a module turn-down capability of 50% is required
to satisfy continuity, regardless of the number of modules in the plant.
For the plant with two gas turbine modules, a module turn-down
capability of 50% will satisfy the requirements for both minimum load
and load continuity. The calculations indicate that this can just be met
with a l300°F temperature in the primary beds of the two boiler modules
operating in unison.
If the minimum bed operating temperature for satisfactory 502
removal is 1400°F, the gas turbine module turn-down capability with two
boiler modules operating in unison is about 60%.
This does not meet
either the minimum load or the continuity requirement.
Operation of the
gas turbine module with one of the boiler modules turned off must be
considered. The idle temperature for the W251 and W501 gas turbines is
about 970°F. When the primary beds in both boiler modules are at 1400°F,
the gas turbine inlet temperature is about 1300°F; if one bed is turned
276
-------
off, the turbine inlet temperature is about 1000°F. Therefore, operating
a gas turbine module with one bed turned off and the other at a primary
bed temperature of 1400°F is feasible.
The power ranges attainable with all possible combinations of
gas turbine and boiler modules operating between primary bed temperatures
of 1300°F and 1750°F are plotted in Figure 1-103. This shows that a
discontinuity exists between 27% and 31% of design power when the minimum
primary bed temperature is 1400°F and that a minimum temperature of
about 1325°F is necessary to eliminate this discontinuity.
Growth Evaluation
Utility gas turbines are currently operating at temperatures
of 1800°F to 1900°F for base load applications. Therefore, there is a
possibility for improvement of the high-pressure fluidized bed system
if the bed temperature is increased above 1750°F or if the temperature
difference between the bed and the turbine inlet is reduced.
In addition,
it is thought that the boiler tube corrosion problem in the fluidized
bed will be substantially less than it is in the conventional boiler so
that it may be possible to increase the steam superheat temperature.
Cycle calculations were made to determine the effect of increasing gas
turbine temperature and/or steam superheat temperature on the plant
capacity and heat rate.
The results of those calculations are given in
Table 1-40.
This indicates a significant growth potential for the high-
pressure fluidized bed boiler system.
Boiler Subsystem
A boiler pressure of 10 atm gives the best plant performance
and cost for an assumed constant boiler cost.
An analysis was made of
the effect of pressure on boiler size and pressure vessel costs which
showed that very little cost reduction could be realized by increasing
the pressure level beyond 10 atm.
A four-module boiler design must be used to meet the turn-down
requirements for the plant if the minimum bed temperature is 1400°F.
Other design features which had to be considered are:
277
-------
Case I
2 G. T. Modu les
4 Boi ler Modu les
Case I I
2 G. T. Modules
3 Boiler Modu les
Case III
1 G. 1. Module
2 Boiler Modu les
N
....
co
Case IV
1 G. T. Module
1 Boiler Modu Ie
I
o
I
20
o-----~
o
0--
o
o
o
I
70
Curve 6431128-B
o
Primary Bed Temp-oF
1750
1400
1300
L-
80
I
90
<>
I
100
Fig. 1-103-Turn-down characteristics of corntJined- cycle with fluidized bed boiler operating on mode 3
Discontinu ity
for 1400°F
I
30
I
40
I I
50 60
Percent Design Power
-------
Table 1"""40
EFFECT OF GAS TURBINE AND STEAM SUPERHEAT TEMPERATURE
ON THE HIGH-PRESSURE SYSTEM CAPACITY AND HEAT RATE
STEAM SUPERHEAT REHEAT GAS TURBINE GAS STEAM PLANT HEAT RATE
PRESS. - PSI TEMP. of TEMP. of TEMP. ° F TURBINE TURBINE CAPACITY
FLOW FLOW %
%
% %
N.
-....J
\0
2400 1000 1000 1600 100 100 100 100
3300 1100 1100 1600 100 95.2 104.2 96.0
3300 1100 1100 1700 100 94.0 105.4 94.8
4500 1200 1100 1600 100 92.2 106.6 93.8
4500 1200 1100 1700 100 91.0 107.8 92.7
-------
. Arrangement of the pressure parts
. Size and shape of the fluidized bed or beds
. Arrangement of the bed or beds in the pressure vessel
. Bed depth
. Orientation of the vessel.
Since the maximum allowable bed temperature is less than the
state-of-the-art gas turbine temperature, there is no need for convection
heat transfer between the bed and the gas turbine.
Therefore, all of the
boiler heat transfer surface is located in the bed or immediately above
the bed where most of the heat is absorbed from the elutriated solids
rather than from the gas.
In order to provide flexibility in the oper-
ation of the boiler module, each boiler function, e.g., evaporation,
was allocated at least one separate bed.
The level of superficial
velocity used in the primary beds results in the elutriation of a sig-
nificant amount of unburned carbon.
Rather than recycle this material
to the primary beds, a separate bed or carbon burn-up cell is used in
each module.
Each module, therefore, would have a minimum of four
primary beds plus the carbon burn-up cell.
Alternative Concepts for Boiler Preliminary Design.
Having
selected a modular boiler with multiple beds in each module, choices
had to be made regarding the orientation of the modules, the shape of
the bed, the bed depth, and the orientation of the boiler tubes.
The vertical module with stacked bed was chosen over the
horizontal bed with side-by-side beds because of its advantages in the
distribution of the air to the beds, the collection of the products of
combustion from the beds, and the steam circuitry arrangement.
The
vertical module with stacked beds also permits the use of one continuous
water wall enclosure for all of the beds.
The use of a horizontal
module places a definite limit on the maximum bed depth for .a given
pressure vessel diameter.
The shape of the bed and the tube orientation are interrelated:
an annular bed precludes the use of horizontal tubes in serpentine
280
-------
bundles; vertical tubes fit well into a circular bed, whereas horizontal
tubes are possible but expensive because of the nonuniformity of tube
bank element; square beds, of course, work equally well with vertical
and horizontal tubes.
Horizontal tubes were chosen over vertical tubes
because of the possibility of poor hydrodynamics in the bed with ver-
tical tubes and because their use permits conventional fabrication
methods.
This led to the choice of a rectangular bed as the preferred
arrangement.
Boiler Design
Modules. An elevation drawing of the final design of the high-
pressure 'fluidized bed boiler module is shown in Figure 1-104. Detailed
drawings of the module are given in Appendix H For a 318 MW plant
module the pressure vessel is 12 feet in diameter and 110 feet high. Both
of these dimensions are within their respective limits for rail trans-
port, so the module can be shop-fabricated.
external steel framing.
The module is supported by
view.
Beds. The primary beds are 5 ft x 6-1/2 ft in the plan.
By using beds of these dimensions and by using the water walls
for evaporation, it was possible to get sufficient heat transfer surface
for the other boiler functions in four primary beds approximately
12 ft in depth.
One bed is for pre-evaporation, two are for super-
heating, and the fourth is for reheating.
The carbon burn-up cell is located near the bottom of the
module in a water-walled enclosure adjacent to the pre-evaporator bed
and has dimensions of 1-3/4 x 7 x 20 ft.
in this bed.
There are no boiler tubes
Boiler Tubes.
Several alternative tube configurations were
considered.
Those selected are as follows:
281
-------
o
~~
Reheater Bed
Superheater Bed
Superheater Bed
t
Pre-evaporator Bed
CBC
Flue
Gas
.. .1
I
I
Carbon I'
Burn-up
Cell i
i
I
Air I
"I
I
10 Ft.
110 Ft.
I
PLANT
SIZE
320 mw
635 mw
VESSEl
DIAMETER, 0
12 Ft.
17 Ft.
Feed Water c::>
Grade Elevation
ELEVATION
PRESSURIZED FLUIDIZED BED STEAM GENERATOR
FOR cor.ml~.JED CYCLE PLANT
FOUR (4) REQUIRED
Fi g. 1-104
282
-------
FUNCTION TUBE DIA. - IN. SPACING - IN. SURFACE - FT2
Pre-evaporator 1-1/2 O.D. 3-1/2 between rows 13,900
3 in rows
Evaporator 2 in. O.D. 3-1/2 on. centers 17,200
in fin wall
Superheater 1-1/2 O.D. 3-1/2 between rows 36,400
3 in rows
Reheater 2 in. O.D. 3-1/2 between rows 16,700
4 in rows 84,200
From the standpoint of fluidization, economy of space would be promoted
by using the smallest diameter tubes possible; but steam flow require-
ments and structural problems must be considered in choosing the tube
sizes.
The mean metal temperatures for the boiler pressure parts were
calculated, and material specifications were made for use in estimating
the boiler cost.
Steam Circulation.
As previously stated, this boiler is once-
through circulation. Feedwater enters the module near the bo.ttom and
flows upward through the pre-evaporator tube bank. Evaporation starts
in the uppermost loops of this bank and the mixture of water and steam
flows into the first of four tube wall circuits. After each circuit the
mixture is transported to the bottom of the next through unheated down-
comers. Thus, the two-phase flow in heated tubes is always upwards,
preventing phase separation.
Saturated steam leaves the fourth wall
circuit and flows through the two superheater beds in series and then
to the high-pressure steam turbine. The steam returns to the reheat
bed between the high- and low-pressure steam turbines.
Air Circuits.
The combustion air is taken from the gas turbine
compressor exit to the pressure vessel and is distributed to the various
beds in the module by an internal manifold which is formed by the volume
between the tube wall and the pressure vessel. Air dampers are provided
to control the flow to the plenum chamber below each of the fluidized
bed distributor plates. The carbon burn-up cell has a separate direct
air supply connection to the plenum below the bed.
283
-------
Products of Combustion.
The products of combustion from the
primary beds leave the tube wall enclosure through tube screens located
near the top of the overhead volume.
These tube screens are made by
offsetti~g alternate tubes in the tube wall. The flue gases are col-
lected in a 1-3/4 x 7 ft exhaust manifold which is formed by tube walls
with one side common with the bed enclosure.
The bottom 20 feet of these
tube walls are the carbon burn-up cell enclosure.
This is the only
convection heat transfer surface in the boiler module.
Near the middle
of the module, an insulated pipe carries the flue gases from the water-
wall passage through the pressure vessel to the first-stage particulate
separator which is located outside the module.
The flue gases from the carbon burn-up cell are segregated
from those from the primary beds. A separate insulated pipe leads from
the carbon burn-up cell overhead out through the pressure vessel to
join the primary flue gas pipe downstream of the first-stage particulate
separator.
Solids Handling. Each primary bed has four injection points
for pneumatically feeding coal into a tube-free space above the dis-
tributor plate.
The make-up sorbent is fed in along with coal.
Each bed is equipped with one nearly vertical pipe at a point
about four feet below the top of the tube bank which removes the spent
sorbent from the bed for regeneration.
Each bed also has one inlet
pipe located about four feet above the air distributor plate for pneu-
matically injecting the regenerated sorbent.
The carbon burn-up cell is similar to the primary beds in
that it has four injection points for the fuel, one withdrawal pipe for
the spent sorbent, and one return pipe for the regenerated sorbent.
Boiler Performance
Energy and mass balances were calculated for the design load
and for 70% of design load. These are shown in Figures 1-105 through
1-108.
284
-------
Flue Gas
631,800 Ib/hr
Solids 1
7216 Ib hr
pri mary
Beds
Air Dolomite
586,5001b/hr 48801b/hr
Coal
53,910 Ib/ hr
1
Carbon
AS!l
Dolomite
Flue Gas
640, 90 Ib/ hr
Solids 1
5, 263 I b/ h r
Dolomite
6308 Ib/ hr
31. 9"/.
63.4%
4.6
FIGURE 1-105-MASS BALANCE
DESIGN LOAD
lONE MODULE)
Solids
722
Flue Gas
676,460 Ib/hr
Solids
53501b/hr
Flue Gas
44,660Ib/hr
Sol ids 1
6494 Ib/ hr
Sol ids 2
46301b/hr
Air
42,800Ib/hr
2
Carbon
Ash
Dolomite
3
Carbon
Ash
Dolomite
4.5%
89.0
6.5
FIG. 1-106-MASS BALANCE
70% PLANT LOAD
(ONE MODULE)
Flue Gas
662,680Ib/hr
SOlids3
3,7801b/hr
Flue Gas
21,200Ib/hr
Solids2
33011b/hr
Primary
Beds
Solids 1
47401b/hr
Dolomite
6127 Ib/hr
Air Dolomite
608.900Ib/hr 48801b/hr
Coal
38,600Ib/hr
1
Carbon
Ash
Dolomite
Air
20,200Ib/hr
31. 2%
62.1%
6.7%
2
Carbon
Ash
Dolomite
4.3%
86.4%
9.3%
285
Owg. 616&\70
Flue Gas
676,460 Ib/hr
Solids
1601b/hr
Solids3
51901b/hr
8.2
85.8
6,0
0\./9. 616&\ 71
Flue Gas
662,680 Ib/ hr
Sol ids 3
114 Ib/ hr
Solids
36661b/hr
3
Carbon
Ash
Dolomite
7. rn.
83.9%
9.1%
-------
Losses
42.9 x 106 Btu/ hr
Primary
Beds
FIG. I-1m-ENERGY BALANCE
DESIGN LOAD
(ONE MODULE)
Flue Gas
279.0 x 106 Btu/hr
Carbon
6.4 x 106 Btu/ hr
Flue Gas
22.9 x 106 Btu/hr
Flue Gas 256.0 x 106
Btu/ hr
Carbon 33.9 x 106 Btu/ hr
Carbon
30.6 x 106 Btu/hr
Steam
448 x 106
D"9. 6168A72
Flue Gas
279.0 x 106 Btu/hr
Carbon
0.2 x 106 Btu/hr
Carbon
3.1 x 106 Btul hr Carbon
Steam. 6.2 x 106 Btu/ hr
10.5 x 106 Btu/h
Air Sensible Coal
80.8 x 106 Btu/hr 700 x 106 Btu/hr
Air Sensible
5.9 x 106 Btu/hr
~ FIGURE H08-ENERGY BALANCE
7rJ'/o Plant Load
(ONE MODUlE)
Flue Gas
228.1 x 106 Btu/ hr
Flue Gas
239.6 x 106 Btu/ hr
Carbon
4.6 x 106 Btu/ hr
Carbon
23.9 x 106 Btu/ hr
Losses
30.4 x 106 Btul h
Primary
Beds
Carbon
21 x 106 Btu/ h
Steam
296 x 106 Btu/ h r
Flue Gas
11. 5 x 106 Btu/hr
Solids
2.2 x 106 Btu/hr
DI"9. 6168A 73
Carbon
0.15 x 106 Btu/hr
Flue Gas
239.6 x106Btu/hr
I .
Carbon
Steam 6
6 4.SxlO Btu/hr
10.5 x 10 Btu/hr
Air Sensible Coal
77.4 x 106 Btu/hr 501 x 106 Btu/hr
Air
2.7 x 106 Btu/hr
286
-------
Steam generator efficiency calculations are misleading in a
combined cycle plant if the gas temperature leaving the steam generator
is used to calculate the stack gas losses.
The stack gas temperature
should be used to calculate a combination boiler and gas turbine com-
bustor efficiency.
The calculated and estimated losses for the high-
pressure fluidized bed boiler on this basis are as follows:
COMPONENT
LOSS - %
Dry gas loss @ 275°F
Loss due to hydrogen and moisture in coal
Loss due to moisture in air @ 275°F
Radiation loss
Incomplete combustion
3.88
4.14
0.08
0.15
1.51
Solids sensible heat
Unaccounted for and manufacturer's margin
0.11
1.50
11. 37
This gives a combustion efficiency of 88.6%.
Regeneration/Sulfur Recovery System
The regenerator system selected for the high-pressure boiler
is based on the "reduction-steam C02 oxidation" reaction scheme. Process
options and thermodynamics were discussed above, and Figure 1-109 is a
flow diagram of the selected process. Table 1-41 gives material balance
information, and Figure 1-110 is a schematic plant layout showing boiler-
regenerator interconnections.
Spent stone discharges from each boiler bed and the carbon
burn-up cells to a hold vessel; one hold vessel is provided for each
boiler module. Four lines pneumatically convey spent stone from each
hold vessel to the first-stage regenerator reactor -- the CaS04 reducer
vessel. Air serves as the carrier gas. The reactor, a fluid bed
reactor, operates at l500°F and 8 atm. The gas producer vessel supplies
reducing gas for CaS04 reduction by partially combusting coal with
air and steam. After particulate removal, the effluent gas from the
287
-------
Boile rs
Hold
Vessels
Va
N
00
00
S2
S3
Dwg. 2950A14
:
~ To Final
I Particulate Removal
~ Turbine
G4 -+r4 Expander
~-- ~Waste
CaSO 4 Sol ids
Reducer
Vessel
G2
Gas Gl S!eam
G3- Air
V) Produce SI . C I
ro oa
u - - - Gl!.. - - ~s~ - -1
H2S'
Generator GI0
Vessel . .
I
I~.. Compressor
G9 Gg
-- --
Spent
S ton e
C02
Scrubber
Fig. 1-109-Pressurized regeneration System-Flow Diagram
I GlO: Tail Gas
to Stack
. Sulfur
G]
---
Slip Stream
From Stack
-------
TABLE 1-41
600 MW HIGH-PRESSURE REGENERATOR MATERIAL BALANCE
T P LB MOLE
STREAM (oF) (PS IA) LB/HR HR H2 H2S
G1 640 150 183,000 6,300 79.1 20.9
G2 640 150 9,000 500 100
G3 1500 135 221,000 8,400 8.0 8.4 16.3 7.4 59.4 0.5
G4 1500 120 182,000 6,800 1.0 19.8 Q.3 5.5 73.4
G5 1500 116 182,000 6,800 1.0 19.8 0.3 5.5 73.4
!\.) G6 'V18 182,000 6,800 1.0 19.8 0.3 5.5 73.4
00 G7(a)
\D 230 135 595,000 20,000 8.6 0.2 16.0 73.4 1.7
G8(a) 220 19.7 266,000 10,000 68.8 30.9 0.2
G9(a) 212 19.0 179,000 6,500 63.7 36.3
G10(a) >212 180 179,000 6,500 63.7 36.3
G11 1100 165 117,000 4,900 73.9 16.1 10.0
Sl Amb 135 33,500
(coal)
S2 1500 135 400,000 'V16% sulfated solids
(spent dolomite)
83 1100 180 450,000 'Vl% sulfated compositions
(regenerated
dolomite)
(a) Flow rate for C02 scrubber design basis.
-------
Regenerated Dolomite
N
\J:)
o
Transfer Li ne
from Coal Feed
System
Transfer Line to
H2S Generator
/' Vessel
Plant
B
Coal
B - Boiler Module
H - Ho'lding Vessel
51(.J, 101 ht One Boiler
R - Regeneration System Modul e
, Vessels: RV-Reducer
vessel, G-H2S Generator
Vessel, C-Particulate
Collector
GP - Gas Producer
S - Dolomite Surge Vessel
1 0 I (.J, 60 I h t
Holding
Vessel
for
Bo i 1 er
Module
Dwg. 2950A12
To
Recovery
Turbi ne
To
Claus
Plant
G
11. 51(.J
541 ht
H
7.51(.J
20
Lines
to
Boi ler
Modules
A Li nes
Shown
One
From
Each Holding
V es s e I
Fig. I-nO-Pressurized regenerator System: Plant Layout
701 Transport
Line ( 30'
Extends into
H2S Generator)
GP
91(.J
501 ht
Coal
Ai r Steam
-------
CaS04 reducer vessel is expanded through a turhine expander which drives
a compressor.
The effluent gas passes through a final particulate
removal step before being discharged to the power plant flue gas.
Reduced solids, CaS, flow through a standpipe leg into the
second-stage reactor, the "2S generator vessel. This fluid.izcd bed
operates at 11000r and 12 atm. The standpipe serves as a pressure seal
between the reducer and the generator vessels.
to CaC03 for recycle to the boiler.
C02 and steam convert CaS
Regenerated solid is conveyed back to each boiler bed through
20 discharge lines from the H2S generator -- one line per boiler bed and
CBC. Feed gas to the second reactor is provided by scrubbing a C02-rich
stream of boiler flue gas in a regenerable hot carbonate or monoethanol
amine (MEA) scrubber system. The C02 compressor is driven by the turbine
expander on the reducer vessel outlet. H2S-rich gas is sent to a conven-
tional Claus plant for conversion to elemental sulfur.
Overall regeneration performance, particularly with respect to
sulfur removal efficiency, is impossible to predict precisely.
Conven-
tional. Claus plants recover 94% to 96% of the sulfur fed to them.
With
added tail gas clean-up this can be improved to 99+%.
Losses from the
regenerator should amount to approximately 2% to5% of the total sulfur
(assuming gas leaving the reducer vessel has the equilibrium S02 concen-
tration). Therefore, the overall sulfur recovery efficiency should be
around 90%.
The cost estimate for the regeneration system for a 635 MW plant
is
$8.90/kw.
The Claus plant system cost is $3.78/kw.
Claus plant tail
gas clean-up will add - $3/kw.
The cost estimates for the various components of this regenera-
tion system were obtained from the following sources.
information is given in Appendix I.
Detailed design
. Solids feed system
. Claus plant
Petrocarb
Ford, Bacon and Davis Texas
291
-------
. C02 scrubber
. Process vessels
Benfield Corp.
. Westinghouse Heat Transfer Div.
Auxiliary Equipment
CaalHandling. . The caal handling system includes receiving,
starage, drying, crushing, and canveying facilities (See Figure 1-111).
Caal .(1-1/2 xO in) arrives at .the plant by unit train twice a week and
is unlaaded into. the 400-tanreceiving happer in abaut five hours.
A
12;OOO-tonsila acts as a buffer between the batch-.ty:pe delivery :apera-'
tian and thecantinuaus feeding aperatian. It also. pravides live
starage far shart-term delays in caal deliveries.
~erm delays is .pravided by a 50,OOO-tanapen pile.
.Dead starage far lang-
An apen pit "far
reclaiming coal fram dead starage is provided aver the canveyar between
the receiving happer and the silo..
Caalfram the silo., to. which make-up dalomite is added, is
dried to. <3% maisture in a fluid bed caal-fired dryer.
It is then
. crushed to. 1/4 x 0 in. in a reversible hammermill.
The size distributian
'af 'the~rushed caal is given in AppendixH.
It is then canveyed to. a
ISO-tan surge-binfram which dual feeders and canveyars carry it to. the
interface with the caal ,pressurizing and feeding system.
Make-up Sarbent Handling. The dalamite sarbent is purchased
crushed to. size (1/4 x 0 in) and is delivered to. a lOO-ton receiving
happer in cavered cars. A 2000-tan silo ispravidedwithsufficient
capacity far bath live and dead starage. The crushed sarbent is fed anta
the caal canveyar just ahead of the caal dryer.
Caal 'Pressurizing and Feeding.
The coal pressurizing and feed-
ing system includes feeder surge bins, lack happer~, fuel injectors, and
pneumatic transpar.t lines (see Figure 1-112). The functian af the surge
bin is to. receive the caal fram the caal handling system and dampen fluct-
uatians in flaw ra:te..
The autamaticallycantralled lack happer pres-
surizes the caal fram aneatmtathe pressure required far feeding the
caal to. the bed. A camprehensive study af pressurizing techniques was
292
-------
D\'/g. 296SAOS
By-Pass to Dead
Storage
Dead Storage
FI uidized Bed Dryer
Coal Bi n
Surge
Bin
Mill
/Belt Scale
Belt Conveyor
Dust Collector
IV
\J:)
w
~
Storage Injector ---
Primary Injector
-
Fig. I-Ill-Coal handling and feed system schematic
-------
2 INJECTION AIR
.CONTROLLER
PROCESS AIR
+
COAL
9 PRESSURIZING AIR
INLET VALVE
5 WEIGH APPARATUS
MEASURES COAL
INJECTION RATE
.--1
.
o
,-
ONE OUTLET FOR
EACH TUYERE
4 COAL INJECTION
RATE INDICATOR
1 COAL INJECTION
SHUT OFF VALVE
3 INJECTION AIR
FLOW INDICATOR
FIGURE 1-112
0.1
DUST
COLLECTOR
10 DEPRESSURIZING
VALVE
6 COAL INJECTION
RATE CONTROLLER
..0.
13 FURNACE SAFEGUARDS
COAL INJECTION
STOPPED BY
A. LOW BLAST PRESSURE
B. HIGH BLAST PRESSURE
C. LOW PROCESS AIR PRESSURE
D. INSTRUMENT AIR FAILURE
E. ELECTRIC POWER FAILURE
PETROCARB PRESSURIZED COAL FEEDING SYSTEM
294
-------
carried out (see Appendix H) before the lock hopper was chosen. Next
comes a fluidized bed fuel distributor which meters the coa1-sorbent mix-
ture to pneumatic transport lines which convey it to the injection
tuy~res.
There is a surge bin, lock hopper, distributor combination for
each primary bed and a feeder outlet for each of the four injection
tuyeres in a bed. The pneumatic transport from the feeders to the
tuyere is controlled by. the amount of transport air.
Compressed air re-
quirements for the coal feed system for each of the boiler modules are
estimated to be 8400 SCFM at 200 psig.
Particulate Removal.
The particulate removal system in the
combined-cycle plant with fluidized bed boilers has three functions:
to
recover most of the unburned fuel in the solids elutriated from the pri-
mary beds for recycle to the carbon burn-up cell; to reduce the concen-
tration of particulates in the gas stream going to the gas turbine below
the level where turbine blade erosion is a problem; and to control the
emission of particulates to the atmosphere.
Figure 1-113 shows the particulate collection flow sheet for
the high-pressure boiler modules.
The anticipated requirement for
erosion-free gas turbine operation is less severe than that projected for
atmospheric emissions.
It is thought that an optimum system will be one where the
second-stage separator would meet the turbine requirements but not the
emission requirements; a third collector in the flue gas duct would meet
the emission requirements.
Therefore, for this preliminary design, collection efficiencies
were specified for the first-stage (90%) and for the second-stage (97%)
collector to meet only the requirements for erosion control.
The first-stage collector is a conventional cyc1one'separator.
Four cyclones in parallel are used for each boiler module, and these are
enclosed in a single pressure vessel. The design point pressure drop for
this collector is 0.7 psi.
295
-------
'2 ~o ~~ A.Ca
Se.P~R.A\~
\ \'f ~ i ~ Ca E
SLPA.~\~
_uP i....\M ~ a.. '< -
.... ~ EoO S
..
o
:i
o
1/'1
o
'"
t-
'0
..
J
'J
~
U
FLUe. C4~~ 4
!.1.\ln..IA.1tO
50 I. It> S .
C A. ... t>D '"
~\)Ii~..~' U P
. Co eo\. \.
FL\l~ Co A.S ...
E.L\lT~I!:\1tO
:!>o\. I I) ~
c. e. co.. NO. ....8.8
i
III
.....
~ .J J In
.& .{
- () .... ~
ct; \J c(
; i. I : I
FIG. 1-113
SCHEMATIC OF
PARTICULATE COllECTION
SYSTEM
Thl. Drowing i. the Property 01 the
FOSTER WHEELER CORPORA TIOH
110 SOUTH ORANGE AVENUE
LIVINGSTON, NEW JERSEY
AND '1 LENT WITHOUT CONSIDERATION OTHER THAN THE 80A"OWE"'1
AO"I[I[M£NT THAT IT SHALL NOT BE REPROOUCED.COPIED. LIENT, 0" Oil.
PO.KD O~ OUU::CTL.Y 0" INDIRECTL.Y NOAUSED FOR ANY PURPOllE OTHE!q:
THAN THAT "0" "HICH IT II SPEC1~IC"'L.LV FuRNISHED. THE APPARATUI
IHOWN IN THE O.....WING'I COYEREO 8Y PATENTS.
296
fl. u e. G ~ 5 "To
Go~~ 'UIL~It-.lE.
11\
o
J
o
VI
SCALE:
= "-0"
-------
The second-stage collector is a tornado separator of the type
shown in Figure 1-114 wherein the dirty gas is split into primary and
secondary streams.
in this collector.
All particles larger than five microns are removed
Two of these collectors are used in parallel for
each boiler module.
Each unit has its own pressure vessel.
The arrange-
ment of the boiler modules with the first- and second-stage particulate
separators is shown in Figure 1-115.
Power Generation Equipment
Gas Turbine.
Two Westinghouse W251A gas turbines match the
requirements for the 318 MW combined-cycle plant.
turbine modules with two boiler modules each.
This gives two gas
The high-pressure fluidized bed boiler is also the combustor
for the gas turbine, so provision must be made in the gas turbine design
for removing the air from the compressor discharge and returning the hot
gases to the turbine inlet. No industrial and utility gas turbines
manufactured in the United States have external combustors; they all use
an integral combustor arrangement similar to that used in aviation gas
turbines.
Gas turbines which are designed for use with external recu-
perators can be adapted to external combustors with minor modifications.
Only one commercially available gas turbine is provided with a recuper-
ator -- the General Electric Series 5000 unit.
In this study of combined cycle plants using fluidized bed
boilers, two Westinghouse gas turbine models have been specified. In the
preliminary design of a high-pressure boiler for a 318 MW plant made by
Foster Wheeler, two Westinghouse W25lA gas turbines are used.
In the
635MW plant for which plant layout and cost estimates were made by United
Engineers and Constructors, two Westinghouse W50lA gas turbines are used.
In both cases, each gas turbine module has two boiler modules. The
Westinghouse Small Steam and Gas Turbine Division made a layout of a
modified W50lA gas turbine for use with a high-pressure fluidized bed
boiler and estimated the cost of manufacturing and developing this modi-
fied gas turbine design. It was assumed that the W50l design and cost
information would be applicable to the W25l unit.
297
-------
Exhaust (Clean Gas)
Secondary
Gas Inlet
Inlet (Dirty Gas)
::::MfO:"::::::::
"" " ""'"
""" """
...... ""'",
~ ~ ~ ~ ~ ~ ~ ~ ~ ~ i i i i ~ i :. : . : i ! : : ! ! ! ! !
"""""'" .......
.................... """"
................... """'"
""""" .
""""""""" .
""""""""" .
Fig. 1-1l4-.-operation of Aerodyne particulate separator
298
Dwg. 6J64AJ2
Secondary Air Pressure
Maintains High
Centrifugal Action
Secondary Air Flow
Creates Downward
Spiral of Dust and
Protects Outer Walls
From Abrasion
Dust is Separated From
Gas By Centrifugal
Force lis Th rown
Toward Outer Wall and
into Downward Spiral
Falling Dust is
Deposited in Hopper
-------
- , ----r-o.~
-
, ,
+-
T" . ,,,,:-~~- "::-+--~-'-~,, --I
I."." - -----------I
..,.,. I
I
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I ,
. I
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r
r-
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ll::!'': - - -
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coc.(..~
--
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1-115
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299
~"La. ""
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--l-----
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1~.:=j
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1 ~OML":' 1»"""
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~"""08'::::;'~ ....
-" -
_K-
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PlOI PLAN
~"Tf."M c.lwfa.A.iOQ.!I>
"'II ~ .. tt! ~
11=0 a.... ~Ol.'f 'tQUI~Io\t.~'
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c.'f c: \..1. PI.."" N 'T
-......--.
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t2. D-710' 10"
!.... <:> o'
--
=-::: :.
- -"
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-.-
--~
---
~.-
~ F~"':.~:;~;
ii =;- -:;.T :.::=
:-.:.:- -=.:: ~~-;--~.: =~.g
-------
In the late 1950s Westinghouse developed a gas turbine for
operation on blast furnace gas.
One unit was built for U. S. Steel and
installed at the Chicago South Works. .This gas turbine had an external
can-type combustor located vertically below the unit which supplied hot
gas to the turbineat1350°F. Figure 1-116 shows the internal arrange-
ment for removing the compressor air and returning the products of com-
bustion to the turbine inlet.
There are two air connections which are
isolated from the hot gas manifold; the hot gas manifold has a single
connection.
This unit has operated for 16,000 hours on blast furnace
gas and for 12,000 hours on natural gas.
In the early 1960s, FIAT [16], an Italian licensee of Westing-
house, developed a40 MW compound cycle gas turbine which uses external
combustors (see Figure 1-117). In 1968 (104), three of these units were
in operation and four were on order. Westinghouse assisted FIAT in the
development of these combustors, which were patterned after the blast
furnace combustor design.
The early units were 1350/ l350°F, and the
later units were 1450/ l450°F.
The low-pressure unit has two air connections and a single hot
gas connection to the internal manifold, the same as the blast furnace
unit.
The high-pressure unit has a single concentric connection for
both the air and the hot gas.
By mid 1968, each of the three units installed had accumulated
about 10,000 hours of operation on residual and distillate fuels.
Brown Boveri [104], .Siemens [81], and Kraftwerk Union AG [82]
also build gas turbines with external combustors.
The Brown Boveri units
are compound cycles with single combustors and operate mostly at l150°F.
The Siemens unit, which is being used in a Velox boiler application, has
a turbine inlet temperature of about l500°F.
Its configuration is not
known.
The Kraftwerk-Union turbine operates at l47SoF and has two exter-
nal combustors with concentric air and gas connections (see Figure 1-118).
While the above-mentioned gas turbines with single or double
connections have operated. satisfactorily, the turbine inlet temperatures
I
used are appreciably. lower than those anticipated for this application.
301 .
-------
Fig. 1-1l6
302
RM-53404
-------
1-4
j;'1~h
II il II III
l..l-.l!..-.-JL
-
-
A - Low-pr...ur. compr...o,
B - AIr Intorcoollng
C - H Igh-pr...ur. compr...or
o - Hlgh-pr...ur. combu8tor
E - Hlgh-pro..ur. turbln.
:'';'''''"j..oII"O....u8 c:ombultor
G - L.o...-or...ur. turbln.
M - Lo'ad
T
13'
. .=-~/\~"
T,.
"
T2.l...2'
TT~'"
._._.""';12
i
I
.-.---........ T,
\
s
t .2 - Air companion
2 . \' - Air cOiJlln.
l' .~' - AI, r.:::omp'..lion
2' . 31 - HOi.:ing c.' .i,.
3' . <" - Expln,;Of. of h~t gal.1
4' . 3 - R'...tat!flZ c,t 61'.'
3 .4 . EJ:~.1r:,':;.!': c.( ~:.t 1111'
4 . I - £0"8..,t of i8'" ,
FIGURE 1-117 TG3000 COMPOUND-CYCLE GAS TURBINE
SCHEMATIC CYCLE ARRANGEMENT AND OPERATING DIAGRAM
303
-------
Dwg. 6169A25
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Fig. 1-118 - Gas turbi ne V95. Section through combustion chamber
304
-------
For gas turbine inlet temperatures of l600°F and greatert distortion
problems are anticipated with internal hot gas manifolds because of non-
uniform temperatures.
Thereforet the Westinghouse Small Steam and Gas
Turbine Division designed a transition system without an internal hot
gas manifold (see Figures 1-119, l-120t and 1-121).
The system has 16 concentric air and hot gas pipes which lead
from the unit housing between the compressor exit and the turbine inlet
to two vertical external manifolds. Separate connections for the air
and the hot gas are provided at these external manifolds.
A crossover
pipe between the two manifolds is provided for the hot gas. The hot gas
pipes terminate at the turbine inlet in 16 sectors which completely fill
the turbine inlet annulust thus eliminating the need for an internal hot
gas manifold.
It has not been concluded that an internal manifold with one
or more connections is not feasible for use at turbine inlet temperature
over l600°F. Such a transition system could be less complex and possibly
less expensive than the one shown in Figures 1-119 and 1-120. It is
thought, however, that a substantial amount of development effort would
be required to obtain an adequate internal manifold design for this tem-
perature levelt whereas the external manifold design presented herein is
considered to be more nearly a state-of-the-art design.
The MaYt 1970 cost of a Westinghouse W50l AA gas turbine is
$4,000tOOO.
It is estimated that the manufacturing cost of a unit modi-
fied for use with the high-pressure fluidized bed boiler will be the
same. The development of the modified gas turbine design is expected to
take two to three years and cost about $2,000tOOO.
Estimated costs for the air and hot gas transition system with
the external manifolds is as follows:
Concentric piping for
air and hot gases
Concentric air and hot
gas manifolds
$ 48t500
3lt600
305
-------
!
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modified for external combustion
307
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Fig. I-I20-Detail of piping arrangement for
modified WSOI gas tu rbi ne
309
-------
r
~
Fig. 1-121- Installation drawing of modified W501 gas turbine
311
-------
Insulated crossover
pipe for hot gases
Total
Round off to
43,600
$123,700
$125,000
'Stearn Turbines.
'The stearn conditions assumed for the 318 MW
plant are 2400 psig/lOOO°F/lOOO°F and 3 in. Hg back pressure. The stearn
turbine for this plant would be a 3600 RPM tandem .compound unit with
, high-pressure, intermediate-pressure, and low-pressure components. The
low-pressure end would be four-flow with 23-inch blades.
All. four boiler
modules would serve one steam turbine with a rating of 265 MW - 300MVA.
Stack Gas Coolers
A ,conventional heat recovery sys:tem is used to recover the
sensible heat from the gas leaving the gas turbines.
coolers are used to preheat the feedwater.
These stack gas
Exhaus,t ,gas leaves the turbine at approximately 830°F and is
cooled to 525°F in an upper stack gas cooler while heating 480°F fe~d-
water from the last feedwater heater stages to 578°F. The gas then
passes through .alower stack gas cooler where it is cooled to 275°F.
The lower stack gas ,cooler hea ts feedwa ter from the feedwa ter hea ters
associated with the low-pressure steam turbine section from '22SoF to
480°F.
Modular designs and costs were prepared by vendors. The
cost, including transition ducting with the gas turbines, is $4.1/kw
for both a 318 MW and 635 MW plant. Design specification and questions
are presented in Appendix N.
312
-------
Operation and Controls
The steam flow circuits of the high-pressure boiler are very
similar to those of a typical once-through unit.
The modular boiler
arrangement with independently fired fluidized beds makes the operation
of the system more complex, but the basic characteristics of once-through
operation are not changed.
In a conventional coal-fired boiler the practice is to operate
the furnace at the minimum amount of excess air consistent with good
combustion efficiency to minimize stack losses.
In a fluid bed boiler,
constraints are imposed on the combustion airflow by the hydrodynamics
of the bed.
In the combined cycle plant, further constraints are placed
on combustion airflow by the gas turbine compressor if the power turbine
which drives the generator is not separate from the compressor turbine.
With a constant speed compressor the airflow can be varied by only about
one-third of the boiler capacity.
In some aspects of the steam generator operation, the fluid-
ized bed boiler has greater flexibility than the conventional pulverized
coal-fired boiler.
The separation of the steam generator functions into
separately fired beds is highly advantageous since it permits specific
control of heat input to each part of the water/steam circuitry.
Start Up. In starting up the 318 MW high-pressure fluidized
bed plant one gas turbine module is started up at a time, and one boiler
module is started up at a time.
As shown in Figure 1-122, each gas turbine module is equipped
with a by-pass line around the two boiler modules. During start-up about
one-half of the compressor airflow would be diverted through this by-pass
line where an oil-fired start-up combustor is located. The combustor is
ignited and the gas turbine is accelerated to the self-sustaining speed
of about 1630 RPM. Here the starting motor is cut out and the unit acce-
lerates to the idle conditions of 3600 RPM and about 1000°F turbine inlet
temperature.
At this point the air temperature from the gas turbine com-
pressor is about 575°F.
313
-------
PRV
Connections
To Second
Module
Dwg. 616&\68
v.>
I-"
-I'-
Prepared
Coal
Hopper
Generato
Steam Turbine
530 MW Tandem Compound
3600 RPM
4 Flow
28-1/2" Blade
I nj ector
Condenser
Feedwater
Pump
Lock Hopper
for Ash Removal
Regenerative
F .W. Heaters
I
I
I
I
I
I
I .
L______- ---- ----- -
2 Stages of
Particulate
Removal
To Stack
Stack Gas
Coolers
Booster
Compressor
enerator
Starting
Motor
ompressor Turbine
W 501 AA Gas Turbine
3600 RPM
650 Ib/sec
10 Atmos.
16000F
Ai r
Inlet
Fig. 1-122 - Flow diagram for plant start-up
-------
The start-up of the first boiler module in the plant is differ-
ent from that of the later modules since there is no steam available for
preheating the beds. All beds are equipped with oil-fired igniters to
heat the beds to the ignition point of the coal, which is assumed to be
750°F. Water is circulated through the pre-evaporator bed, the wall
tubes, and the first superheater. From there it goes to the flash
tank and then to the condenser.
The pre-evaporator and the first super-
heater bed depths are reduced to three feet, preheated to the ignition
temperature, and coal feeding is started. Once ignition is established,
dolomite is fed into the beds to bring the bed depths up to full capacity
in a period of about one hour. The pressure and temperature are built up
in this circulating loop with the steam from the flash tank going to the
condenser and to the second superheater tubes.
When the gas temperature
from the boiler module is sufficiently high, the start-up combustor is
turned off.
Next, the second superheater bed is ignited; and, when the
steam conditions are adequate, the turbine is warmed and rolled.
Then
the reheat and carbon burn-up cell beds are ignited. The steam pressure
is increased until it reaches design value with an output of about 10%
full plant load or 40% of module rating. The t~rbine is then put on
governor control.
The other boiler modules are started up as neerled by preheating
the beds with steam from the first module.
Since feedwater is now at a
temperature of about 500°F, the start-up times for the second, third, and.
fourth boiler modules are much shorter than for the first.
Detailed
descriptions of the boiler start-~p procedure are given in Appendix H
by Foster Wheeler and in Appendix J by United Engineers and Constructors.
Load Control.
The specified turn-down requirement for this
plant is 4: 1.
It has been shown that a modular boiler system with each
module having a capability of 50% turn-down is required to meet this
specification with no discontinuities.
With the arrangement of two gas
turbine modules and four boiler modules which was selected for this pre-
1iminary design study, a 4:1 turn-down is attainable.
However, a minimum
primary bed temperature as high as 1400°F may be required for adequate
315
-------
sulfur removal,- whereas a bed temperature of about l325°F is required to
attain 25% plant capacity with one gas turbine module and two boiler
modules.
Therefore, there may be a discontinuity in steady state operation
between 25% and 31% of plant capacity.
Tables 1-42 and 1-43 list the operating parameters of the
individual fluidized beds for design and 70% plant load. This shows
that the bed temperatures vary significantly at part load because of a
shift in the distribution of heat input among the boiler functions as
load changes, and that the superficial velocities do not vary subs tan-
tially with load change.
Table 1-44 lists the steam generator parameters for design
and 70% plant load.
The specification for load change rate for this plant is 5%
p~r minute. . Preliminary analysis has indicated that this specification
can probably be met. However, the beds do have a substantial thermal
inertia, and the residence time of the fuel in the bed is relatively
high. Therefore, a comprehensive analysis of the load change rate capa-
bility is needed.
Shut-down.
The normal shut-down procedure for the high-pressure
boiler is basically a reversal of the start-up procedure. Fluidizing air
to the beds will be maintained after fuel feed has stopped in order to
purge and cool the beds. If it is necessary to keep the bed on hot stand-
by, the flow of feedwater to the beds would be maintained.
Two situations would lead to emergency shut-down. One is loss
of load on the gas turbine, and the other is 108s of load on the steam
turbine.
An emergency relief valve is provided in the air by-pass line
to handle loss of load on the gas turbine (see Figure 1-122).
If such
a loss were to occur, the fuel flow to the beds would automatically be
stopped, and the relief valve would open to divert flow from the gas
turbine.
Emergency shut-down due to loss of load on the steam turbine
would be accomplished in the same manner as in a conventional steam
316
-------
TABLE 1-42
FLUIDIZED BED PARAMETERS
DESIGN LOAD
(1 MODULE)
GAS TEMPERATURE LEAVING BOILER(a)
- 1650°F
---------..---'-
FUEL FLOW AIR FLOW FLUE GAS SUPERFICIAL BED
(LB/HR) (LB/HR) (LB/HR) VELOCITY TEMPERATURE
Pre-evaporator 17,680 192,100 206,200 9.0 1,750
st . 13,100 142,700 153,700 6.8
1 Superheater 1,750
nd 10,740 116,900
2 Superheater 125,900 5.6 1,750
Reheater 12,390 134,800 145,200 6.4 1,750
CBC 6,494 42,600 44,660 6.1 2,000
(a)
Includes CBC Gas
(b) 31. 7% carbon
63.3% ash
5.0% dolomite
317
-------
TABLE 1-43
FLUIDIZED BED PARAMETERS
70% PLANT LOAD
(1 MODULE)
GAS TEMPERATURE LEAVING BOILER(a)- 1400°F
FUEL FLOW AIR FLOW FLUE GAS SUPERFICIAL BED
(LB/HR) (LB/HR) (LB /HR) VELOCITY TEMPERA TURE
Pre-evaporator 12,600 199,200 209,200 8.7 1,47Q
st 8,950 141,200 . 148,600 6.3 1,470
1 Superheater
nd 7,370 116,300 122,400 5.2 1,470
2 Superheater
Reheater 9,680 152,200 160,800 7.0 1 ,530
CBG 4,650 20,200 21,690 3.2 2,000
(a)Inc1udes GBG gas
(b) 31. 2% . carbon.
62.1% ash
6.7% dolomite
318
-------
TABLE 1-44
STEAM GENERATOR PERFORMANCE SUMMARY
(FOR TOT1\L PLANT)
l~--~;:"::-T 70~~~
Fuel flow -- M lb/hr
Air flow -- M lb/hr
Flue gas flow -- M lb/hr
215
2,506
154
2,506
2,6.50
Feedwater inlet temperature -- of
Feedwater inlet pressure -- psig
2.705
578
547
2;650
1,000
Main steam
utlet ressure -- psig
.2,824
1,000
2,500
1,727 .
650
2,444
1,140
Main steam outlet temperature -- of
Main steam flow -- M lb/hr
Reheat inlet temperature -- of
Reheat inlet pressure -- psig
Reheat outlet temperature -...; of
600
1,000
584
405
Flue gas exi t temperature -- of
Flue gas pressure drop-- psig
Fuel/air ratio
580
1,644
1.650
4.8
0.0856
.1,000
391
Reheat outlet pressure -- psig
Reheat flow --M lb/hr
1,090
1,400
4.8
0.0615
319
-------
plant.
Flow would be shut off to the turbine by an emergency trip valve,
and the steam would be diverted to the condenser.
Fuel flow to the beds
would be shut off automatically, and airflow would be maintained to purge
and cool the beds.
If possible, feedwater flow would be maintained until
the beds had cooled sufficiently so that tubes would not be overheated.
Control and Instrumentation
A composite flow diagram for the high-pressure fluidized bed
boiler system is shown in Figure 1-123. The control system for this
plant is designed to perform the following functions:(l)
.
Maintain steam header pressure to the turbine at the
desired value by regulation of the turbine governor
valve
.
Maintain the steam header temperature at the turbine
at the desired value by regulation of the fuel input
to the beds
.
The boiler feedwater flow is set by the load demand
from the dispatcher.
Many of the plant subsystems are automated, e.g., that which maintains
solid levels in feeder bins. Other functions are manually controlled,
e.g., the sorbent recycling, rate to the beds where a monitor sounds an
alarm when the recycle rate is too low or too high.
The Foster Wheeler
report (Appendix
H
) gives complete details on the controls for the
boiler and the boiler auxiliaries and lists the instrumentation required
for control and monitoring.
Plant Layout
United Engineers and Constructors prepared a plant layout for
a 635 MW high-pressure fluidized bed boiler system (see Figure 1-124).
A hypothetical site was assumed, located in the eastern United States
on a large river - the source of once-through cooling water.
provided to a state highway by an existing secondary road.
Access is
Rail access
(l)Mote details are given in the United Engineers and Constructors report
(see Appendix J ).
320
-------
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BOILER COMBINE.D CYCLE PlANT
COMPOSITE FLOW DIAGRAM
BOOSTER AIR
COMPRl$$OA
Fig.
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@ STOIlA<:;E &1115 @ DEAEIlA'TOIl
0 ADM,I'j15TRATlOl'j &LOG;. @ ASI! 'DOLO""Ti SILO
o _TlCULAlE ~MOVAL E'QIJIPT @ WATEIl S'TOIC'AG. TAI'j"S
CD CiAS TUIl5INE G1."IEIlAT\)1i'5 @ LIGIITOIL TANK.
@ $TAClC6AS COOLE'~ @ LIGIITOIL UNLOADING PUMP
6) STAC'I<: @ ClRC1JU.Tlt.IG W..1EQ INTAICE PipES
@ I~ To.. SuRGE e'N @ ""IlII(;, HOPPR @ GASES (STQClACiE)
@ rnCLA'M HOPPEI;e @ FILTEIlED6 FiU ~I/ICE \IIATEIl T4Hr.
@ SILO (11/X>O TONS) @ SEl.'I/lcE WA1BI PUMP IIOU~
@ OI2YER @ CIIZeULATII'jG WO~ INTAICE $TIZUCTU!lE
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@ WATEIl TIlEATMEIiT @ Ii"oRIAIIU; WElllS
@ DIESEL GUIEIi'ATOI? Ii.'OOM @ SWITCH VAllO
@ CONTk'OL!ZOOM @ CO,NEYOIlS
@ MACIUIIE SWOP @) DOLDMITE S'LO ('2,000 TONS)
@ TRANSFOIlMEIlS @ MIL.L
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SITE PLOT
PLAtJ
o 50 100 'ZOO
1.......01 1.......01 1.......01
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80lLER C0I"18/NEO CYCLE
,PL.IlNr
SITE PLOT PLAN
IIIl'II8J,,'--6._..
Fig. 1-124
323
-------
can be obtained by the construction of a five-mile spur.
Other drawings
prepared by United Engineers and Constructors and included in this'
report are:
Plan of Main Equipment at Grade Floor, Figure 1-125
Plan of Main Equipment at Operating Floor, Figure 1-126
Elevation of Main Equipment, Figure 1-127
Plan of Main Equipment at Mezzanine Floor, Figure 1-128
Drawings included only in the United Engineers and Constructors report
(Appendix J ) are:
Isometric of Boiler Modules and Particula~e Separator
Details of Gas Piping from Separators to Gas Turbines
Electrical Single Line Diagram
Steam Turbine and Surge Bin Foundations
Steam Turbine and Surge Bin Structures
Steam Turbine Foundations - Plans and Sections.
In order to have a basis for comparing the physical character-
istics of the high-pressure fluidized bed combined cycle plant with con-
ventional fossil-fueled plants, United Engineers and Constructors pro-
vided drawings of a 1000 MW coal-fired plant (Figures 1-129 - 1-132)
and of a 1000 MW oil-fired plant (Figures 1-133 -1-136).
Plant Performance
Figure 1-102 shows the heat rate versus the load for the 635 MW
high-pressure fluidized bed boiler system for the following conditions:
. 'Back pressure - 3 in Hg
. Combustion efficiency - 89.6%
. Gas turbine system pressure drop - 3%.
The site chosen by United Engineers and Constructors for com-
parison of energy costs has a supply of once-through cooling water which
will give a back pressure of 1-1/2 in. Hg. The estimated combustion
efficiency for the high-pressure combined cycle system is 88.6%. The
estimated pressure losses in the gas turbine combustor loop at design
load are as follows:
325
-------
RR.
OIL UNLOADING
51''''1'ION ~
~
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: TUNS/Nt: GEN.
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CHLOR-IN"TIO"-l
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INTAKE SiR.UCTURE
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-------
Campanent
Ta tal
0.75
0.20
0.10
0.05
1.00
3.10
0.10
0.48
0.77
0.20
0.75
7.50
is being can-
Air .outlet transitian
Piping between transitian and bailer
Cantral valves in .open pasitian
Internal air passages ta plenum
Distributian plate
Bed
Internal gas passage ta
stage callectar
First-stage callectar
Secand-stage callectar (4 units).
inlet .of first
Piping fram secand-stage separatar ta
transition . .
Hat gas inlet transitian
An alternative arrangement far the secand-stage callectar
sidered wherein twa tarnada callectars in parallel are used in place .of
faur.
The pressure lass far this canfiguratian is 1.38%, giving a tatal
system pressur~ drap .of 8.11%.
The full laad heat rates far these sets .of canditians are as
fallaws:
COMBUSTION PRESSURE PLANT HEAT
EFFICIENCY-% DROP-% RATE':"BTU/kwh
3 89.6 3 8974
1-1/2 88.6 7.5 8892
1-1/2 88.6 8.1 8897
3 88.6 3 9075
347
-------
Recommendations
Pressurized fluidized bed power plants show greater pollution
abatement, lower costs, and greater development potential than atmos-:-
pheric pressure fluidized bed or conventional p.f. plants.
Based on the
projected market for coal and the technical assessment, we recommend
that:
. Emphasis be placed on the development of pressurized
fluidized bed power plants
. Alternate and advanced concept evaluation be conducted
to assess potential
. A 10-30 MW pressurized fluidized bed boiler development
plant be designed and constructed
. Laboratory research and development on pressurized
fluidized bed combustion continue in key areas:
sulfur removal, NO reduction, stone regeneration,
x
sulfur recovery, materials evaluation, advanced concepts,
particulate removal, and turbine blade erosion.
348
-------
Atmospheric Pressure Fluidized Bed Boiler DeSign
Specifications
Specifications for the power plant and the fluidized bed
boiler were prepared by Westinghouse for use by Foster Wheeler in making
a preliminary design of the atmospheric boiler.
The specifications for the power plant for which the
Plant.
atmospheric pressure fluidized bed boiler was to be designed were as
follows:
. Capacity - 300 MW
. Type - intermediate load
. Turn-down requirements - 4:1
. Transient requirements - 5% per min.
. Geographical location - Eastern United States
. Pollution control
Particulate emissions. - 0.01 gr/SCF
Sulfur removal - 90%
Nitrogen oxide emissions - <0.7 lb/l06 BTU
Boiler.
The specifications for the atmospheric pressure
fluidized bed boiler were as follows:
. Fuel - Ohio Pittsburgh No.8 Seam Coal
Proximate analysis (wt
Volatile matter
Fixed carbon
%)
39.5
48.7
8.5
3.3
100.0
Ash
Moisture (as mined)
Sulfur content - 4.3%
Particle size - 1/4 in x 0
Complete data on coal is given on page 167.
349
-------
. Sotbent - 1359 Limestone
Particle size -
-1/4 in with minimum fines
Feed ratio - 6 x stoi~hiometric
Complete data on sorbent is given on page 170
. Bed Design Parameters
Primary beds
Temperature - l600°F design pt~
l700°F max:
l400°F
min.
Size distribution of material in bed
Max. 5000 Jl
Min.. 1000 ).l
Mean 2500 ).l
Superficial velocity - 10 to 15 ft/sec
Excess air - 10%
Combustion efficiency- 87% design pt.
Density - 45 lb/ft3 .
Expanded height - 2-1/2 to 3 ft.
Elutriated material
Ash - 100%
Coal - 13%
Sorbent - 3% of recirculated limestone
Gas-side heat
. .
Size distribution - see page 127
transfer coefficients
In .hed':' 50 BTU/ft2 hr - of
Above bed ~ 40 BTU/ft2 hr - of
. Carbon Burn-upCe11
Temperature - 1900°F,design
Range - lBOO°F to2000~F.
.Size distributionbf material in bed
Max. 5000 J.l
. .
Min. 1000'. Jl .
. .
. Mean 2500 ).i
350; .
-------
Superficia+ velocity - 10 to 15 ft/sec
Excess air - at least 50%
Combustion efficiency - 85% design pt.
Density - 45 Ib/ft3
Expanded height - 2~1/2 to 3 ft
Recovery of coal elutriated from primary beds - >90%
Elutriated material.
Ash - 100%
Coal - 15%
Sorbent - 100% of sorbent in with \
recovered coal plus 3% of recirculated
limestone.
Size distribution - see page 127
Gas-side heat .transfer coefficients
Same as for primary beds
Stack gas temperature - 275°F
Air preheat temperature - to be determined
Cycle Selection
Consideration was given to a subcritical cycle (2400 psig/
1000°F/lOOO°F) and a supercritical cycle (3500 psig/lOOO°F/lOOO°F).
The
subcritical cycle was selected for this intermed~ate load application
because of its advantages in reliabilitYt availabilitytand delivery
and start-up.
Alsot in practice the supercritical cycle plants have
not demonstrated a significant advantage in unit heat rate.
Steam
temperatures currently in use in pulverized coal units were selectedt
although the possibility of reduced boiler tube corrosion in fluidized
beds indicates that this is a conservative position.
A limited study was made to determine the effect of increasing
the steam temperature on turbine power and plant net heat rate.
condenser pressure of 3 in. Hg. was used in these ca1culations~
A
351
-------
STEAM SUPERHEAT REHEAT STEAM TURBINE PLANT
PRESSURE TEMP. TEMP. FLOW POWER HEAT RATE
PSIA of of % % %
2400
3300
4500
1000
1100
1200
1000 100 100 100
1100 100 110.7 95
1100 100 118.0 92.1
that the use of a fluidized bed boiler may provide
This study indicates
, a significant growth potential for conventional Rankine cycle systems.
The once-through system was selected over the natural cir-
culation system for the subcritical cycle because it permits the use
of horizontal boiler tubes in the fluidized beds.
The steam system used in this boiler design is that from the
Hammond No.4 unit of Georgia Power and Light (see following matrix).
This system has seven regenerative feedwater heaters and its turbine
heat rate versus net load is plotted in Figure 1-137.
Boiler Design
Alternative Concepts. Three basic conceDts were considered for
the pre'liminary design of the atmospheric pressure fluidized bed'
boiler:
. Vertically stacked beds with four modules in a contiguous
square array
. Vertically stacked beds in four separate modules in an in-
line arrangement
. Horizontal tandem arrangement with four modules in parallel.
An early decision to use a modular arrangement with each bed having a
~ingle function precluded modules having a single bed and a single module
boiler.
The modular arrangement is necessitated by turn-down require-
ments and facilities start-up.
352
-------
I-
i
,
STEAM CYCLE CONDlT.lONS
300 MW ATM FLUIDIZED BED BOILER
Primary steam flow
Steam pressure, superheater out
Reheater in
Reheater out
Steam temperature, superheater out
Reheater in
Reheater out
Steam flow (reheater)
Boiler feedwater temperature
353
6
1. 9xlO lb/hr
2400 psig
601 psig
581 psig
lOOO°F
650°F
lOOO°F
1. 6xl06 lb/hr
480 ° F
-------
10,000
9800
!.... 9600
.c
~
~
--
::3
-
CCI
Q,) 9400
-
m
a:::
-
C'C
Q,)
::I:
C 9200
C'C
0..,
9000
8800
o
40 ro
Plant Load - %
20
80
Fig. 1-137 - Part load performance of Hammond #4 un it
2400/1000/1000/3 In.
354
100
-------
Preliminary Design of Preferred Concept. The in-line arrange-
ment of four separate modules of vertically stacked beds was selected
as the preferred concept (see Figure 1-138).
The vertical beds were
chosen over the horizontal beds to simplify the steam piping, the
distribution of air to the beds, and the removal of products of com-
bustion from the beds.
The in-line arrangement was chosen over the
square array because there were more disadvantages than advantages in
sharing common walls, and symmetry of auxiliary equipment was lacking
so that ductwork and piping would have been complicated.
The preliminary design configuration consisted of four modules,
each containing six beds, including a carbon burn-up cell. Two of the
beds are assigned to evaporating functions, two to superheating, and
one to reheating. The carbon burn-up cell bed has no immersed tubes.
The water walls used for enclosing the beds are evaporating tubes.
Preheating of the feedwater is done in separate convection passes
located in the overhead above each bed.
The advantages of this over
an external economizer are:
. The gases are cooled within the boiler enclosure, thus
eliminating the need for high-temperature ducts leading to
an economizer
. The convection surface acts as a particle screen and helps
in the detrainment of the larger particles elutriated from
the beds
. It makes more effective use of the boiler cross sectional
area.
The dimensions of the beds in the four module boilers .are
13 ft x 12 ft x 2-1/2 ft, and the module height is about 100 ft.
Since
the overall dimensions exceed the limits for land transport, field
erection will be necessary.
However, the maximum degree of shop
fabrication of subassemblies will be employed.
Horizontal tube bundles are used in all beds and for the con-
vection passes above the beds.
All bundles are arranged in a rectangular
355
-------
_ml
STEAl .1
.
SUPERHEATED I I
SUPERHEATER' STEAM .
REMEATER
SUPERHEATER .
EVAPORATOR
EVAPORATOR
eBe
.. WATEI.. STEAM
~ US
c::> AIR
18" ,.
AIR HEATER.
~AIRIN
HALF PLAN VIEW
,
,
,
,
,
lUBULAR
. DUST
eOUECTOR
'n----------
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HOPPER
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'I 1
. IN
o 25'
----------
12'
! GRADE Elevation'
25'
I
1
ELEVATION
ATMOSPHERIC-PRESSURE FLUIDIZED BED STEAM GENERATOR
FOUR 141 REQUIRED
Fig. 1 - 1 38
356
-------
pattern, and the tubes are on a square or rectangular array.
The
submerged tubes have both vertical and horizontal pitches of 4 in.
Two-inch O.D. tubes are used in the reheater bundle and 1-1/2-inch O.D.
tubes in the superheater and evaporator.
In the preheater bundles,
the first two loops have a vertical pitch of 4 in and a horizontal
pitch of 8 in.
The remainder of the bundle is on a 4 inx 4 iri
The preheater tubes are 2 in'O.D.
square array.
The water walls which form the enclosure for the beds are
of welded fin configuration.
They are constructed of l-3/4-in tubes
. on 2-i~ centers ~ith a 1/4-in fin.
This spacing is the minimum
practical with this type of construction.
The water-steam circuitry is shown on an isometric flow
diagram in Figure 1-139.
Water enters the feedwater inlet header and
then proceeds in series through the preheater bundles which are located
in the convection passes above each bed.
From the feedwater outlet
header at the top of the module, the water goes to the first evaporation
bundle in the bed above the carbon burn-up cell. This is in series
with the second evaporation bundle in the next bed above.
Evaporation
is completed in the water wall enclosure.
in series.
The four. walls are connected
The evaporated water then goes through the two bundles of
superheater tubes in series.
A mixing header is located between the
two superheater bundles.
The superheated steam leaves the module
through a single-ended header and goes to the high-pressure steam
turbine. Reheat steam enters the module through a header at the inlet
end of the reheat bundle and leaves through a header at the exit end
of the bundle;
Both headers are single-ended.
The auxiliary equipment for the atmospheric pressure fluidized
bed steam generator includes an air preheater, a coal handling and storage
system, a combined coal drying and feeding system, a limestone handling
and feeding system, a limestone regeneration system, and fans and
particulate removal equipment.
357
-------
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FLU1 DIIED BED
,STEAM GENERATOR
CIRCUITRY
'c(
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ACTI08I MIAD
COIIT. IVP'V.
Fig. 1-139
11010~""""-"'-
~ !"OSTER WHIEIEU:R CoRPORATION
il ..8 80. ~ 1ft.. UWW8'RIfL .. J.
.... .. ....... wrTI80VY ~A""" cn'1t8It ""A. .,..
------ ~ nu.., IT -"L.L. .... - ..
....-uca. COP8aD. -.-T. -..............:n.T-
tM88II8CTLY .... U888 ... AII"f' ........ -- .,...,.
'hUto., .... WtMCIf IT .. 818C1PtCAL.L.Y ~ .....
.......TU8 -- ... ..... ....- .. C8W8I88 ..
...--.
. ,,359
~ .-.-._~ ~ .-.--
-------
Air from the air preheater passes through two external ducts
to a common internal duct which runs the full height of the boiler
module.
From this common duct, air is metered by dampers into plenum
chambers below the fluid bed distributor plates.
A separate air stream
which by-passes the preheater goes directly from the forced draft fan
to the carbon burn-up cell.
An internal common duct collects the flue gas with carbon-rich
ash from the five primary fluidized beds and carries it to the primary
cyclone separator.
A separate duct located within the main flue gas
duct carries the flue gas from the carbon burn-up cell to another dust
collector.
Coal is pneumatically transported through sixteen separate
feed lines to each bed as a dilute phase of dried coal in 250°F air.
The coal feed lines enter the gas duct from either side of the module,
pass through the flue gas exit openings in the water wall, then up
into the plenum chambers and through the grid plates.
The coal is
injected into the bed in a horizontal direction in a volume free of
tube bundles.
One coal injector is provided for each ten square feet
of fluid bed surface.
The recycle ash feed lines are brought into
the module at a point below the carbon burn-up cell.
Sorbent feeding into the fluidized beds falls into two
categories; make-up feed and recycled sorbent feed.
The make-up sorbent
is fed into the beds along with the coal.
Each bed is provided with
one injection point for the regenerated sorbent and one withdrawal
point for the spent sorbent.
. The boiler modules are not bottom supported, but hang from
structural steel.
Regeneration - Sulfur Recovery System
A "direct reduction" process was selected for regeneration
of the sulfated limestone.
The calcium sulfate is converted to CO
by reaction with a reducing gas
CaS04 + HCO ~ CO + SO + C02
2 2 H20
(1-40)
361
-------
A flow diagram and material balance are given in Figure 1-140 and the
following matrix for a 300 MW plant. Figure 1-141 is a schematic plant
layout showing the boiler-regenerator interconnections.
MATERIAL BALANCE FOR 300 MW ATMOSPHERIC PRESSURE
REGENERATION SYSTEM
Spent stone
56.6 tph = 113,000 Ib/hr
37.8 wt % CaS04
l500°F
43.5 tph
1. 33 wt % CaS04
2000°F
"-9000 Ib/hr
2800 Ib mole/hr
5.6 tph
0.48 tph
3500 lb mole/hr
2000°F
Regenerated stone
Waste stone
Air
Coal
. Ash
Sulfur gas
S02
H20
C02
N2
CO
H2
Adiabatic heat
6.3% (mole %)
13.3%
18.9%
61.1%
0.3%
0.1%
9 MMBtu/hr
Spent stone discharges from each boiler bed into .a hold vessel,
one hold vessel per boiler module. The stone is then pneumatically
transported to the regenerator vessel, using air as the carrier gas.
The regenerator vessel is a fluid bed reactor operating at
1 to 2 atm. .and2000°F. Coal is charged into the reactor bed. An air-
stream mixture fluidizes the reactor bed and combusts the coal to provide
reducing gas for the CaS04 reaction.
36.2
-------
Dwg. 2950A15
To Sulfur
Plant
Regenerator
Vessel
Coal
Distributor
Ve sse I
Spent Stone
Ash
( 4 Lines)
Steam
Regenerated
Stone (24
Lines)
Air
Waste
Spent Stone: 56.6 tph = 113,000 Ib/h r, 37.8 wt % CaSO 4' 1500°F
Regenerated Stone 43. 5 tph-= 87,000 Ib/hr, 1. 33wt% CaSO 4' 2000°F
Waste Stone: ,..., 9000 Ib/hr
Ai r: 2800 Ib mole/h r
Coal: 5.6 tph
Ash: 0.48 tps
Sulfur Gas: 3500 Ib mole/hr
S02: 6.3% (mole %)
H20: 13.3%
C02: 18.9%
Adiabatic Heat: --9 MMBTU/hr
20000F
N2: 61.1%
CO: O. 3%
H2: 0.1%
Fig. 1-140 -Atmospheric pressure regeneration flow diagram
363
-------
Dwg. 29S0A13
Atmospheric-Pressure 300 MW Boiler System
I. Boiler Module..........16' x 16' ea
2. Regenerator Vessel.....See Sketch
3. Distributor Vessel.....12' g x 12' ht.
4. Hold Orums.............S' g x S' ht.
5. Cyclone ...............
6. Waste Stone Hopper.....8' g x 121 ht.
7. Air Blower...'..........
l.-J
0\
~
-------
vessel.
Regenerated stone flows from the reactor vessel to a distributor
Twenty-four lines return regenerated stone from the distributor
to each bed in the boilers. - Some deactivated stone is discarded from
the distributor vessel to provide for the fresh limestone make-up
required to achieve the desired SOZ removal. Fresh stone is charged
to the boilers with coal through the coal feed system.
Experimental results from Esso R&E [37] and Esso England [Z9]
indicate that the SOZ concentration of gas leaving the regenerator
vessel will contain the equilibrium concentration of S02' approximately
4% to 10% by volume. The sulfur recovery process selected for this gas
is the one recommended by Allied Chemicals [2].
The hot regenerator
top gas will pass first through a gas cooler and cleaner. The SOZ
concentration will be enriched to 90% by volume by a dimethylamine (DMA)
absorber-stripper process. The concentrated SOZ stream is reduced to
sulfur by the Asarco process (described by Allied Chemicals) using
natural gas as the fuel.
The obvious points of sulfur loss in the-regenerator are by
the carrier gas for the spent and regenerated stone. The expected
sulfur recovery efficiency for the DMA process is 98% to99%.
This
process consumes soda ash and dilute sulfuric acid..
Vented gas from
the DMAscrubber should contain 450 ppm S02' The Asarco process is
expected to recover approximately 9~% of the sulfur fed to it. Adding
a third Claus reactor to the two planned, or the Claus tail gas clean-
up, could possibly improve the performance. The overall sulfur
recovery efficiency for the regenerator-sulfur recovery plant should
be between 85% and 90%.
Emissions from an acid 'plant should contain 500 ppm or less
SOZ' corresponding to a sulfur recovery efficiency in excess of 99%.
The overall plapt sulfur recovery should be approximately 95%.
The cost estimate for the regeneration system for a 300 ~v
plant is $6.3/kw.
The cost estimate scaled to 635 MW is $4.Z/kw.
The
sulfur recovery system cost estimate is $ll/kw for a 635 MW plant.
365
-------
Claus plant tail gas clean-up will add an estimated $3-4/kw. Cost
estimates for the various components of this regeneration system were
obtained from the following sources.
. Solids feed system:
scaled from Petiocarb system data
. Sulfur plant:
cost estimates in Allied Chemical reports
. . Process vessels: . scaled from Westinghouse He~t Transfer
Department estimate for high-pressure process vessels
using extrapolation procedure given by Guthrie. (1)
Auxiliary Equipment
Coal Handling and Crushing. The coal handling system includes
receiving, storage, crushing, and conveying facilities (see Figure 1~142).
Coal (1-1/2 in x 0) arrives at the plant by unit train twice a week
and is unloaded in about five hours. A 12,000-ton s110 acts as a buffer
between the batch-type receiving operation and the continuous feeding
operation.
deliveries.
It also provides live storage for slmrt-term delays in
Dead storage for long-term delays in deliveries is .pro-
vided by a 50,000-ton open pile.
Coal from the silo is crushed to 1/4:x 0 in. in a reversible
hammermill.
The size distribution of this crushed coal 'is given in'
Appendix K.The crushsd coal is conveyed through a 150-ton surge
hopper to a group of 16 bunkers or hoppers which serve the individual
feed injectors to the boiler modules.
Make-up Sorbent Handling. The limestone sorbent is purchased
. .
crushed to size (1/4 x 0 i~and is delivered in covered 'cars. A
silo is provided with sufficient capacity for both live and dead
storage.
The crushed sorbent is fed onto the coal conveyor just ahead
of the l50-ton surge hopper.
Coal Drying and Feeding. The coal drying and feeding system
. .
is shown in Figure 1-143. A direct fired system was chosen over
(l)H P .
. opper, Modern Cost Engineering Techniques, New York: McGraw Hill,
19 70 .
366
-------
4 FEEDERS
500/700 tph
EACH
DC
~ RECLAIM
- HOPPER El.270'
\400 TONW -:0~D~~hF2
~ 80'~'-~METAL DETECTOR
~1!~200 tph 12 x 0 .
. ELT 54" BELT CONV. NO.1
56' SCALE
'NO.1 DC
BY-PASS' TO DEAD
STORAGE .
12,000
TON
SILO
70' DIA.
DEAD STORAGE
50,OQO TONS
24" BELT
CONVEYOR NO.2
150 tph
. CRUSHER
~ DUST COLLECTOR
150 TON
SURGE BIN
\..o.J
0\
'-I
E1.170'
..
2 FEEDERS F4
0/75 tph EACH
COAL FEEDER SURGE SILOS BY OTHERS
-- I
GROUND El. 100 '
7Ji 7::-/7;'....' '.' .
iI- .... '=
/ILY/~"-"
'- -.
FIG. 1-142 COAL HANDLING PLANT
-------
~LJI
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1 I I
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1-143
ll. 22'..'
PlA.N
tlE.'I
40',0'
A1
C>-~~ - "-' 2
369
NOTES
1. DO MO'f 8(;ALI[ THI8 0"..""0
D.....8ION8 ONL"
.a. ._UVIATtOH8 UND ON TMI. CNlAWtttG A"
.. ACCOIltOANC:II: WfTM .aeJlIC.A.N 8TANDAIItO
-~T'IOI'oI8 "CMII U.. 0frII DIltAWIHG8-
14 F_L~' D e>l.O
"""~Oboll.(,
-r
-r
~ .,,~.'. '011.- 0
~,
I o,"'i 1..1,. b:
+--- --I .'
o
I' ..
O\J(O ....
c.,., I.. "t..t. t
2071-.4('.'/.."
I...".'l"{ 0
0
"e;
--.J
-I-i-
: !
0,
~. i
0,
-,
-----
Da:8CIt'P'TION
0.. T ~\ ::> ~ P ,", l ~ . <:
,,".".D'ltD
t> t.. D
co:....'!.
o ~'( \ N"
AI< D
."JlC"Oo<
F7~
:,'( ~'T ~ .,
~oo M oN
~\_~~"T
8
no.. o.-tII8 .. .. p,......., f1If ..
F08TlER WHEELI:" CO-.o"ATION
,t. 80. OIIANGII Aft UV1N08TON, .. ~
...-- u""~c__."'- -",....n.
_-0..- ~.._"" nu,1' ". .......... "'"" .. ..
-00...1:..- c-.. ....""". --.. 0# ~'I'-
...__C'T\.'I' - ~ .- AlOfT "".-. 0..-. TMA..
"""''' .- ...c:. ". .. _C8I'te...~'I' ........-. 'n4
:~...::;.~. .-. ... "... .......... .. 1;0_"0 8'1'
-1'.0"
APPttOV1rD 8Y,
RO-7,\-12':)
-------
several alternatives.
This feeding system includes bunkers for
receiving coal from the coal handling systems) volumetric feeders)
combination fluid bed dryers and distributors) dust collectors for
the fluidizing air) primary air fans, and dilute phase coal transport
lines to the individual coal injection points. The flow diagram for
the coal feeding system is shown in Figure 1-144. Heated air from
the air preheater goes to the primary air fans for use in the dryer
and fuel injector.
This air is then cleaned and used as part of the
combustion air for the primary beds.
Minor amounts of coal transport
and sealing air are also required ,by the coal feeding system.
Carbon Burn-up Cell Feeding.
The particulates elutriated
from the primary beds are assumed to contain about 13% of the fuel
fired.
Over 90% of this material is recovered for recycle to the
carbon burn-up cell (see Figure 1-145).
The recovered material at a
temperature of 840°F is collected in a dust hopper from which it is
periodically transferred into a fluid bed injector which meters the
flow of solids to the feed points in the carbon burn-up cell.
Draft Fans and Air Preheater.
A forced draft system was
. selected over a balanced draft system since the gas-tight fluidized
bed boiler construction is well suited to the forced draft system.
The secondary air enters the system at the forced draft fan and is
heated to 735°F in the air preheater.
It is fed from there into the
internal air distribution duct.
A separate cool stream which is taken
off ahead of the air heater goes to the carbon burn-up cell.
A tubular
type air preheater was selected over the somewhat less expensive
regenerative type because of the severe leakage problem which would
. .
result from the use of the latter in combination with the high-pressure
drop fluidized bed and the forced draft fan.
Particulate Removal.
The particulate removal system in the
fluidized bed boiler has two functions:
to recover most of the unburned
fuel in the solids elutriated from the primary beds for recycle to the
carbon burn-up cell and to control the emission of particulates to the
atmosphere (see Figure 1-146).
There are two'stages of particulate
371
-------
FIG. 1-144 AIR FLOW DIAGRAM
SEALING
FAN
I
, FEEDERS
I
.
L._--.L-.: .,-;, '.
~;.<.' ,,,",,
FUEL
INJECTOR
COAL ,I
TRANSPORT AIR !.
---@--C:r-----S
FEEDERS
SEALING
FAN
I
i FUEL
i INJECTOR
COAL
TRANSPORT AIR i
~----~
LEGEND .
COAL FLOW
AIR FLOW
F .D. FA.."l
-......-
CYCLONE
PRIMARY
AIR FANS
CYCLONE
PRIMARY
AIR FANS
-,
I
I
I
I
I
-i
t
. FBC //'/ I
///1 I
I I
-,
,
I
I
I
-..J
I
I
J
I
,
I
-1
I
,
I
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-'~ ,
I I
I I
, I
'. I I
6 ~
I I
I I
I I
T"'- ------_.J ~
I ...------~
I I -
I I
I .
,
I
,
I I
a I I
,..---- ------,+---_.L -_J
I I
! AIR
IpREHE4TER
-,
I
I
I
I
,
.'
DUST
COLLECTOR
L- ----- J. - - - !'~'-~~ -----
TO PRECIPITATOR
372
GAS
AIR AND C;OAL --~
---
-------
I
I
L--
.,.. ---
,-
/
-{!I
I
~I
,,~
i?
~
.-
't
11 t
~'I
~I I~,
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- /
h ...~I
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-
--1
t
-
~
~
t-
-q
:;:.
....
~~~
~
~
~
CO L~~";"Oi..
DLJS1 -
AIMO)?Hn\c. flUID'? t D
e,E D SHAM GUIEH10 ~
FO ~
300 MWPUdH
{'\ ~t.Q.U\t.EO
FOutZ.. ..
373
-------
FROM FBC
10.8 GRAINS/SCF
% Ash 33.75
% Coal. 59.45 .
% U.L. 6.80
(UNREACTED LIMESTONE)
W
'-J
V1
0.44 GRAINS/SCF
>-
E '" 96%
TO CBC
FROM cac
25.6 GRAINS/SCF.
% Ash
% Coal
. % U.L.
70.8
14.0
15.2
MASS BALANCE--
PARTICULATE REMOVAL
SYSTEM--300 MW
ATMOSPHERIC FLUID BED
BOILER
. FIG. 1 - 146
\
FLUE GAS
1.03 GRAINS/
SCF
E = 96%
TO ASH REMOVAL
0.58 GRAINS/SCF
ELECTROSTATIC
. PRECIPITATOR
l
TO ASH REMOVAL
PRIMARY CYCLONE
CBC DUST REMOVAL
TO STACK
>-
0.01 GRAINS/SCF
-------
iemoval for the products of combustion from both the primary beds and
the carbon burn-up cell.
Multiclone collectors were selected for the first stage of
both st reams.
The two are physically integrated, with the streams
isolated so that the carbon-rich ash from the primary beds is not
mixed with the spent ash from the carbon burn-up cell.
For the predicted
particle size distributions from the primary beds and the carbon burn-
up cell, the collection efficiencies are estimated to be 96% and 90%,
respectively.
The two gas streams are mixed after the multiclones, and an
electrostatic precipitator is used to reduce the particulate concen-
tration to less than the specified 0.01 grains per standard cubic foot.
Systems Analysis
The preliminary design for the atmospheric pressure fluidized
bed boiler was to have been made for a 300 MW plant.
Due to an error
in the steam flow specified for this plant, however, the design
capacity was 280 MW.
A flow sheet for this boiler at full load with
solids flows is shown in Figure 1-147, and an energy balance at full
load is shown in Figure 1-148.
In these mass and energy balances, the
combustible solids in the ash elutriated from the primary beds and the
carbon burn-up cell were considered to have the same composition as
the original coal.
elemental carbon.
In actuality, the unburned solids would be mostly
A temperature-enthalpy diagram for the atmospheric pressure
. boiler is shown in Figure 1-149. The theoretical gas temperature, the
fluidized bed temperature, and the cold-side fluid temperatures are
shown for the full-load conditions.
This plot graphically describes
the allocation of boiler functions and illustrates the heat transfer
situation in each boiler component.
Tables 1-45, 1-46, and 1-47 give the fluidized bed parameters
for one of four boiler modules for 100%, 75%, and 50% boiler loads.
376
-------
w COAL
" T/hJ
" 100 Tons/Hr
SORBENT-45.5
AIR (80°F)
FLUE GAS
--
PARTICULATE: 21.0 T/hr
ASH: 33.75%
COAL: 59.45%
u.S. :
6.80%
FLUID BED
COMBUSTION
.1600°F
VEL.. ~11 ft/see
X's AIR 10%
EFF. 87%
AIR
780°F)
\
, ) I USED
SORBENT
50.8 T/hr
cas04:37.8%
CaS04 11.4%
SORBENT
5.1 T/hr
AIR BOoF
PRIMARY
CYCLONE
n=96%
PARTICULATE: .84 T/hr
20.1 TONS/HR
COAL CONC. 59.4%
C.B.C.
19000F
VEL. 10 ft/see
X's AIR 70.5
EFF. 85%
FLUE GAS
--
PARTICULATE 10.B5 T hr
ASH: 70.79%
COAL 14.01%
U.S.: 15.2%
AIR (BO°F)
(U.S .)
USED SORBENT
5.65 T/hr
CaS04:37.B%
56.5 T/hr
---
50.6 T/hr
- 43.5 T/hr
7.1 T/hr
MAKE-UP SORBENT
FIG. 1-147 OVERALL MATERIAL
BALANCE--FULL LOAD
1 27 T/HR
to ELECTRO-
sunc
PRECIPITATOR
,
.43 T/HR
1)=96% SECONDARY
CYCLONE
10.42 T/HR
~ TO SORBENT
REGENERATOR
,
REGENERATED
SORBENT
-------
SUPHT. STEAM
2400 PSIG - 1000°F
---
~TER IN
430°F
--
REHEAT
STEAM
6.50°F
601 PSIG
COAL AND
SORB'ENT
,- -~
I
I
ELUTRIATED
. SOLIDS
(59.4% ',CARBON)
I
I
1- -- --
IN
'-J
00
SORBENT
REHEAT STEAM
581 PSIG - 1000°F
~
FLUID BED CELL
---
Heat. of Comb~stion
22.46 x 10
Sensible Heat of Air
4.05 x 108
Heat of Reagtion .
0.44 x 10
Sensible Hegt - CaO
0.06 x 10
Heat Losses
2.99 x 108
Heat to Steam
16.45 x 108
Boiler Performance 85%
FLUE GAS 16000F
USED SORBENT
----
l600°F
---~_.....-- --._-
CARBON BURN-UP CELL
Heat of Combustion
2.79 x 108
Sensible Heat of Air
~.
Heat of Reaction
0.05 x 108
Sensible He~t - CaO
0.01 x 10
Heat Losses8
0.73 x 10
Heat to Steam
0.58xl08 .
. .
---..---
. USED SORBENT l400°F
-.---
>
FLUE GAS
1900°F
STACK
TO GAS
340°F
CONVECTION PASS
HEAT TRANSFERRED
8
4.56xlO
HEAT TRANSFERRED
8
1. 18x10
AIR. IN 780 of
(10% X's)
840°F
840°F
AIR IN 80°F (70.5% X's)
FIG. 1-148 OVERALL ENERGY BALANCE
(FULL LOAD)
AIR 80°F
AIR HEATER
- 8
4.05 x 10
HEAT TRANSFERRED
-------
3000
LL.
o
W
-....J .
\01
..
Q,)
L..
.a 2000
ra
L..
Q,)
0-
E
Q)
I-
1000
'"'"
,.,."
""""
""
""
",""
",
",
",
",
", .
",
",
"
. ",
. ".
"
6 8 10
Energy in MAF Coal - 1000 B tu/Lb
Fig. 1-149-Temperature-energy diagram for the atmospheric utility boiler design
«xJO
P reheater
CBC 1900°
,.......,
I I FBC 16000
I
I I
. . I I
I I
I
Economizer
Superheater v)
Evaporator
4
12
Dwg. 6161A83
Adiabatic
Flame
Temp.
V)
Q)
-V)
oV)
o
Q,) ...J
U L..
c: Q,)
ra -
(6.0
OJ OJ
14
16
-------
TABLE 1-45
FLUIDIZED BED PARAMETERS
DESIGN LOAD
(ONE MODULE)
I FUEL FLOW AIRFLOW FLUE GAS SUPERFICIAL BED TEMPERA-
LB/HR LB/HR LB/HR VEL. FT/SEC TURE 0 F
I
REHEATER 1l.600(a) 6 6 10.9 1600(d)
0.126x10 0.137x10
SUPERHEATER I 11 700 (a) 0.126 0.138 11.0 1600(d)
J
SUPERHEATER II 11.700(a) 0.126 0.138 11.0 1600(d)
EVAPORATOR I 7. 000 (a) 0.076 0.083 7.85 1600(d)
EVAPORATOR II 7.000(a) 0.076 0.083 7.85 1600 (d)
CBC 10.300(b) 0.10/c) 0.108 10.7 1900(d)
(a) Fuel as fired
(b) 59.45% coal
33.75% ash
6.80% sorbent
(c) 70.5% X's air based on coal flow to CBC
(d) Specified by Westinghouse
380 .
-------
TABLE 1-46
FLUIDIZED BED PARAMETERS
75% BOILER LOAD
(ONE MODULE)
FUEL FLOW AIRFLOW FL UE GAS SUPERFICIAL BED TEMPERA-
LB/HR LB/HR LB/HR VEL. FT/SEC TURE of
8,630(a) 6 6 8.10 1425
REHEATER 0.0935xl0 0.102x10
SUPERHEATER I 8 790(a) 0.095 0.104 7.91 1390
,
SUPERHEATER II 8,790 (a) 0.095 0.104 8.29 1433
EVAPORATOR 5 250(a) 0.057 0.062 4.74 1350
'
EVAPORATOR II 5,250 0.057 0.062 4.74 1350
CBC 7 670 (b) 0.076 (c) 0.081 6.10 1900(d)
'
(a) Fuel as fired
(b) 59.45% coal
33.75% ash
6.80% sorbent
(c) 47% X's air based on coal flow to CBC
(d) Specified by Westinghouse
381
-------
TABLE 1-47
FLUIDIZED BED PARAMETERS
. 50% BOILER LOAD
(ONE MODULE)
FUEL FLOW AIRFLOW FLUE GAS SUPERFICIAL BED TEMPERA-
LB/HR LB/HR LB/HR VEL. FT/SEC TURE 0 F
5 , 800 (a) 6 6 4.80 1223
REHEATER 0.063x10 0.068x10
SUPERHEATER I 5,050(a) 0.064 0.069 4.79 1170
SUPERHEATER II 5,850(a) 0.064 0.069 4.70 1218
EVAPORATOR I 3,500(a) 0.038 0.041 2.79 1144
EVAPORATOR II 3, 500 (a) 0.038 0.041 2.79 1144
CBC 5,150(b) 0.052(c) 0.054 5.40 1900(d)
(a) Fuel as fired
(br59.45% coal
33.75% ash
6.80% sorbent
(c) 21.4% X's air based on coal flow to CBC
(d) Specified by Westinghouse
382 .
-------
This shows the variation in the temperatures of the primary beds
required at part load to match the functional duty requirements.
Boiler Performance.
The losses for the atmospheric pressure
fluidized bed boiler are as follows:
Losses
Percent
Dry gas loss (stack temp - 275°F)
Loss due to hydrogen and moisture
in coal
Loss due to moisture in air
Radiation loss
Incomplete combustion .
Unaccounted for and manufacturer's
margin
4.28
5.05
0.10
0.18
2.39
1.50
13.88
Radiation, moisture, and unaccounted for losses and manufacturer's
margin are the same as for a conventional boiler.
Dry gas loss is
slightly lower than that of most units because of the lower stack gas
temperature.
The losses due to incomplete combustion are higher
because of the elutriation of unburned carbon from the fluidized beds.
This gives a boiler efficiency of 86.1%, which is significantly lower
than that for the conventional coal-fired boilers.
Plant Performance.
has been estimated to be 86.1%.
The boiler efficiency at design capacity
Figure 1-150 shows the part-load boiler
efficiency, which was patterned after that of a converitional boiler.
Using this boiler efficiency gives the plant heat rate relationship
shown in Figure 1-151.
This is two to three percent poorer than that
for a conventional plant because of the lower boiler efficiency.
Load Control.
A maximum primary bed temperature of l600°F
was chosen on the basis of ash characteristics and sulfur recovery
considerations; a lower limit of l300°F for continuous operation was
chosen on the basis of sulfur recovery and combustion efficiency con-
siderations. At part load, the gas-side heat transfer coefficient is
nearly constant, and there is only a small reduction in steam-side
383
-------
fF..
...,
'"
....
>.
u
c:
Q)
:~ 89
-
L;:j
....
~
.0
CX)
Curve 645901-A
93
92
91
Conventional
( Hammond No.4)
90
88
\
,
"
"-
"-
"
..................
~
. "
,
,
Atmos. Fluid. BED
(Estimated)
87
86
850
100
Fig. 1-150 = Part load boiler efficiency for conventional
and fluid bed boiler
11, 000
10.500
10,000
a::
:I:
~
::.<::
-
:::>
I-
CX)
9500
$
ro
a::
-
ro
Q)
:I:
9000
8500
8000
o
Curve 6459U2-A
Atmospheric
Pressure
F.B. Boiler
(Estimated)
GA. PWR. & LT,
Hammond No.4 Unit
~
~ 00
Plant Load, "/0
100
80
Fig. 1'~ 151 .- Heat rate vs plant load for conventional and F. B. Boiler
2400/1000/1000/3 in.
-------
coefficient because of Reynolds number effect.
The heat transferred
from the beds. therefore. is primarily a function of the bed tempera-
ture.
Because of the constraints on bed temperatures discussed above.
the turn-down capabili ty of a boiler module is only 35% of boiler flow.
which is equivalent to about 40% reduction in,plant power.
To attain
a 4:1 turn-down capability for the plant it is necessary to use four
boiler modules which can be shut down individually.
plant load profile shown in Figure 1-152.
This gives the
Ignition.
Ignition of the primary beds is done in three
steps:
l.
2.
3.
Warm-up of bed and pressure parts
Heating of bed to ignition temperature
Injection of coal at ignition temperature.
The ignition temperature of the specified coal is about 850°F. Six
6 x 106 BTU/hr burners located in the air ducts between the air preheater
and the boiler module are used to warm up the beds and the pressure
parts. During step 1 the airflow is increased until the bed is fluidized
and the bed temperature approaches 650°F. Then a set of igniter burners
iocated immediately above the bed surface is turned on to heat the bed
. 6
to the ignition temperature. The igniter burners are rated at 2 x 10
BTU/hr.
The carbon burn-up cell will receive less reactive fuel. so
it will require a higher ignition temperature.
Since ignition of the
carbon burn-up cell will lag behind that of the primary beds. high-
temperature air will be available for preheating the bed. The carbon
burn-up cells will be equipped with anabove-the-bed igniter burner
with off-on control sized to provide the:heat necessary for ignition.
S t:art-up.
The start-up of the atmospheric pressure fluidized
bed boiler is the same as that of a conventional once-through boiler
(see Appendix K). but with the modular boiler only one module is
started up at a time. The start-up of the second. third. and fourth
modules will differ slightly from that of the first module since feedwater
385
-------
Curve 645614-A
100 (1)
"0 t
I'tJ
0
....J
::J 80 t
u..
-
0
~
~
V1
C1>
::J (f)
"0
0
~
I'tJ
::J
"0
:> 40
"0
c:
-
0
~
c:
+-- aJ
C'O
C:::
(1, 2): Boiler Module No. I
and 2 Operating
(1,2,3): Boiler Module No.1,
2, and 3 Operating
a
a
aJ 40 (f) 80
Total Boiler Load, % of Full Load
100
Fig. 1-152 - Boi Ie r load reduction
386
-------
will be at 500°F rather than 80°F.
are given in Appendix K.
Details of this start-up procedure
Shut-down.
the start-up procedure.
Normal shut-down. is essentially a reversal of
Fuel flow to individual modules would be ,cut
off, but fluidizing air would be maintained to purge and cool the beds.
The air would be shut off when the bed was sufficiently cooled.
If
it was desirable to keep the bed in stand-by ,condition, feedwater flow
at about 500°F would be continued through the pre-evaporator, the
evaporator walls, and the first superheater.
Emergency shut-down due to steam-side problems is accomplished
in the same manner as in conventional units.
The fluidized beds do
have a. high thermal inertia and long residence time for the fuel.
The
bed temperatures are low, however, ,and it is estimated that the tempera-
. '
, ,
ture excursion resulting from interruption of steam flow will not exceed
tube material limitations.
Pressure Losses.
The gas side pressure losses at full load
are as follows:
COMPONENT
PRESSURE LOSS IN H2~
Ducts
1.,43
10.8
27.5
0.1
3.0
Distribution plate
Bed
,Convection'bank
Dust collectors
Air hea ters'
Air-side
Gas-side
3.0
1.5
TOTAL
47.33
Since the bed pressure loss is predominant and independent of load,
the overall gas-side pressure drop will. be relatively unchanged with
load.
387
-------
The stearn-side pressure losses are as follows:
COMPONENT
PRESSURE LOSS-PSI ..
Primary Loop
Feed heating, convection bank
Evaporator bundle 1 }.'
Evaporator bundle 2
Water walls pass 1
125
96
Water walls pass 2
Water walls pass 3
Water walls pass 4
41
55
55
51
Superheater bank12}
Superheater bank
TOTAL
142
565
17
582
SUBTOTAL
Reheater Loop
388
-------
Economics
Capital and energy costs were prepared for pressurized and
atmospheric pressure fluidized bed boiler power plants burning high-
sulfur coal based on the preliminary designs~ Capital and energy costs
were also estimated for fluidized bed boiler power plants burning high-
sulfur oil. Costs are projected for plants with once-through limestone/
dolomite and '7ith stone regeneration and sulfur recovery. These costs
are compared with the cost of conventional coal-and oil-fired plants
with stack gas cleaning.
Assumptions common to all plants are:
. Northeast site
. January 1970 costs
. Begin construction July 1971
. Supplementary cooling
. Contingency of 6%
. Escalation at 7-1/2%
'. Interest during construction at 7-1/2%
not included
. 70% capacity factor
. 15% leve1ized fixed charge
. No credit taken for sulfur recovery
. 45~/106 Btu high-sulfur coal (~ 4%)
8 45~/106 Btu high-sulfur oil (~3%)
. $2/ton limestone or dolomite.
The heat rates for each coal-fired plant are:
Conventional
Atmospheric pressure
fluidized bed
boiler
9230 Btu/kwh
9550 Btu/kwh
Pressurized fluidized
bed boiler combined cycle
8967 Btu/kwh
The heat rates are for 3 in. Hg turbine back pressure to allow for a
cooling tower.
389
-------
Capital and energy costs for the plants are for 635 MW capacity.
The costs assume each concept has been developed and thus represent
commercial costs. Costs as a function of plant capacity were projected
using conventional plant characteristic cost curves.
Capital Costs
Capital costs for 635 MW coal-fired plants are summarized in
Table 1-48. Boiler plant equipment and stearn turbine plant equipment
component costs are presented' in Tables 1-49 and 1-50.
Capital cost
as a function of plant capacity is shown in Figure 1-153.
control costs are summarized in.Table 1-51.
Air pollution
Pressurized Combined Cycle Plant. The total capital cost
estimate for the plant with sulfur recovery is $276/kw. This represents
an 18% cost reduction over a conventional plant with stack gas cleaning.
The cost estimate assumes no advances in subsystem concepts
(e.g., solids handling, particulate removal) beyond current technology.
Cost reductions may be achieved by using a1te~native subsystem concepts.
The following components have been studied for potential savings:
. Particulate Removal: The projected system utilizes four
tornado collectors for final particle removal before the
gas turbine. Two alternative systems have been considered
which reduce the number and size of the units. Cost
estimates for these systems show cost reductions of up to
$5/kw(1) for the particulate removal equipment.
. High-Temperature Gas Piping:
The high-temperature gas
piping cost would be reduced if the, particulate removal
system were simplified by using two tornado units per
module for the 600 MW plant. Additional savings might be
realized if high-alloy steel is not required between the
(1) ~ased on quotation from Aerodyne.
390
-------
TABLE 1-48
BOILER PLANT EQUIPMENT COSTS
CONVENTIONAL
P.F. BOILER
$/KW
Once-Through
ATMOSPHERIC-PRESSURE
FLUID BED BOILER
$/KW
Regeneration Once-Through
PRESSURIZED
FLUID BED BOILER
$/KW
Regeneration Once-Through
.st:~,.lIn G~nerator
31. 50
14.49
Draft System
F.D. Fans
Particulate Removal
Draft Flues & Ducts
Piping (to gas turbines)
Stack and Foundation
0.98
(a)
2.63
0.47(C)
3.30(d)
Air Heater
Coal Handling & Feeding Equipment
13.58
2.69
20.0 (g)
Ash & Dust Handling Systems
Stack Gas Cleaning System
R~generation System
Sulfur Recovery System
,,0 Control
x
Instru~entation & Controls(i)
3.50
2.64
~iscellaneous Equipment
0.94
82.23
25.10
25.10
14.49
12.76
1.66
2.77
0.47
12.76
1.66
2.77
0.47
1.50
5.00
1.66(b)
1.50
5.00
1.66
10.94(e)
0.55(f)
14.94
1.55
0.47
4.63
1l.50(e)
2.69(f)
0.47
4.63
14.71
3.69
8.90
3.78(h)
3.10
~
60.36
3.10
~
52.68
4.2
11.0 (h)
3.10
~
3.10
0.94
71.79
60.80
(a) No particulate control assumed beyond wet scrubber. (see note 9)
(0) Ducts incorporated with steam generator; cost between air preheaters, precipitator, and stack obtained from
United Engineers cost estimate for ducting at the same conditions for the pressurized plant.
(c) Stack height assumed for all plants of 280 ft based on emissions projected from pressurized fluid bed boiler
combined cycle plant.
(d) Air heater cost supplied by Foster Wheeler.
(e)
The coal handling system costs for both systems were prepared by McNally-Pittsburg. The coal feeding system
cost for the pressurized system was prepared by Petrocarb and United Engineers and Constructors. The coal
feeding system equipment cost for the atmospheric-pressure system was prepared by Foster-Wheeler. The cost
for foundation and erection of the system was added. Differences in the coal feed systems include: atmos-
ph8ric system has four times the coal feed pipes and includes coal drying as part of the injector system
which requires a cyclone system for each injector; pressurized system has one lock hopper per injector,
atmospheric system has two volumetric feeders per injector; atmospheric system includes primary air fans,
pressurized system includes booster compressor. (Pressurized system coal drying included with coal handling
cost.)
(f) The atmospheric ash handling system is assumed to be the same as a conventional plant. The cost for the
pressurized ash handling system was prepared by United Engineers and Constructors. The pressurized system
is able to utilize a dry conveying and storage system and thus only includes a conveying system from the
boiler modules and regeneration systems to the ash silo. The ash silo and compressor costs are included
elsewhere. The atmospheric-pressure systems use wet transport which is more expensive due to the extra
equipment -- ash sluice pumps, ash hoppers, dewatering bins, ash settling basin, exhausters, silo, etc.
(g) Recent cost data presented at
indicate the cost for S02 and
half for S02 control and half
the International Symposium for West Lime/Limestone Scrubbing, Nov. 1971.
particulate control equipment for a new plant will be ~ $40/kW. Approximately
for particulate control.
(h) Recent information from Ford, Bacon, and Davis indicate they have
recovery process which reduces sulfur emissions from the process.
emissions and increase their reported cost estimate.
(i) Instrument&tion for stack gas cleaning, regeneration, and sulfur recovery systems is included with the
respective systems.
developed modifications to the sulfur
These modifications will reduce sulfur
391
-------
TABLE 1-49
COAL-FIRED POWER PLANT COST BREAKDOWN
635 MW Plant
(Sulfur Removal process(es) Designated Under Concept)
CONVENTIONAL
P.F. BOILER(a)
$/KW
Once-Through
ATMOSPHERIC-PRESSURE
FLUID BED BOILER(b)
$/KW
Regeneration I Once-Through
PRESSURIZED
FLUID BED BOILER
COMBINED CYCLE(c)
S/KW
Regeneration I Once-Through
Land & Land Rights 1.13 1.13 1.13 1.13 1.13
Structures & Improvements 27.28 25.80 24.80 19.57 18.52
Boiler Plant Equipment 82.23 71. 79 60.80 60.36 52.68
Gas Turbine-Generator Equipment 14.80 14.80
Steam Turbine-Generator Plant Equipment 55.76 54.96 54.96 44.14 44.14
Electric Plant Equipment 14.99 14.99 14.99 15.91 15.91
~lisc. Plant Equipment 4.21 4.21 4.21 3.56 3.56
Undistributed Costs(d) 30.01 29.20 29.20 28.35 28.35
Other Plant Costs(e) ~ ~ ~ ~ ~
Sub-Total 218.52 204.99 193.00 190.73 182.00
Normal Contingency 13.11 12.30 11.58 11.44 10.92
Sub-Total 213.63 217.29 204.58 202.17 192 .92
Escalation 52.12 48.08 45.26 37.60 36.17
Sub-Total 283.75 265.37 249.84 239.77 229.09
Interest During Construction 47.88 41.80 39.35 31. 30 30.07
General Items & Engineering ~ ~ ~ ~ ~
TOTAL CAPITAL COST 337.13 312.67 294.69 276.57 264.66
Construction time:
conventional plant 4.5 years(f)
atmospheric pressure fluid bed boiler plant 4.2 years(g)
pressurized fluid bed boiler plant 3.5 years(h)
(a) Conventional plant cost data provided by United Engineers and Constructors.
wet scrubbing system and NOx control.
(b) Cost estimate prepared by Westinghouse on component data supplied by Westinghouse, Foster Wheeler, and
United Engineers and Constructors. Plant includes sorbent regeneration and sulfur recovery system.
Plant includes throw-away
(c)
Cost estimate prepared by United Engineers and Constructors based on component
Foster IVheeler, and United Engineers and Constructors. Plant includes sorbent
recovery.
(d) Engineering and construction management and supervision, temporary facilities, construction services, etc.
data supplied by Westinghouse,
regeneration and sulfur
(e)
Operator training, spare parts, etc.
(f) Source -- United Engineers and Constructors.
(g) Construction time reduction for boiler projected by Foster Wheeler assumed for total plant.
(h) Construction schedule prepared by United Engineers and Constructors.
392
-------
TABLE 1-50
STEAM TURBINE PUU~T EQUIPMENT COSTS
CONVENTIONAL P.F.
BOILER
$/KW
ATMOSPHERIC PRESSURE
FLUID BED BOILER
$/KW .
PRESSURIZED
FLUID BED BOILER
COMBINED CYCLE(a)
$/KW
Turbine genera tor.
Circulating water system
. ,.'. (b)
Other turbine plant equlpment
Instruments and controls
22.60
5.46
7~33
13.35
.5.57
1.45
22.60
5.46
7.33
l2.96(c)
5'.16 (c)
1.45
18.69
4.11
5.46
10.01 (c)
4.63(c)
W.
\0
W
Condensing system
. .' (b)
'Feedwater system
1.24
TOTAL
55.76
54.96
44.14 .
(a) Power from steam turbine-generator plant is ~82.6% of total power.
(b)
Includes station piping -- boiler feed 'piping, main steam piping.
in pressurized. system.
Includes stack gas coolers
(c) United Engineers and Constructors developed preliminary piping schematics for the pressurized
system and estimated the cost. The small, modular pressurized boilers reduced the piping dis-
tance and allowed for smaller, more flexible, pipe which reduced the piping cost.. The
atmospheric-pressure boiler piping cost was estimated from the pressurized system cost.
-------
500
~ 400
-...
~
,
.....
VI
o
u
(..J
1.0
~
~ 300
c...
to
c...>
~
.....
o
..,..
200
100
200
Cu rve 645592-A
Fuel: Coal
Conventional Plant I ncl udes Wet Scrubbi ng
FI uidized Bed Plants I ncl ude Sulfur Recovery
Conventional P. f. Boiler
1976 Operation
Pressurized
FI uidized Bed Combustion
1975 Operation
300
400
600 700
Plant Capacity, MW
Fig. 1-153 - NtW Plant Cost vs. Capacity
500
800
Atmospheric-Pressure
FI uidized Bed Boiler
1976 Operation
900
1000
lIOO
-------
TABLE 1-51
AIR POLLUTION CONTROL EQUIPMENT COST.
ATMOSPHERIC. PRESSURE PRESSURIZED
CONVENTIONAL FLUID BED BOILER FLUID BED BOILER
P.F.BOILER $/KW $/KW
$/KW Regenerationl Once-Through Regeneration I Once-Through
S02.removal 20(a) (b) (b) (b) (b)
Sulfur recovery 15.2 12.68
l;.J. Particulate control(c) (a) 5.0 5.0 (d) (d)
\J:)
VI
NO minimization 3.50 (e) (e) --
x
23.50 20.2 5.0 12.68
(a) The wet scrubber is assumed to achieve adequate particulate removal. Recent cost data [Second
Annual Lime/Limestone Wet Scrubbing Symposium, New Orleans, La., November, 1971] indicate that
the cost for S02 and particulate control by wet scrubbing for a new plant will be ~. $40/kw
(b) Sulfur dioxide is removed in the boiler.
(c) Ash ha~dling cost is not included.
(d) Particulate removal is required for gas turbine operation.
pollution control.
This cost is not attributed to
(e) Additional NO control may be required to meet proposed regulations.
x
-------
tornado units and the gas turbine.
The present design
uses the high-alloy steel to assure protection of the
gas turbine from additional particulates.
. Stone Regeneration System:
The present design is based on
a multistep process which 'separates each function and
utilizes coal as fuel. Advanced concepts were proposed
which combine two or more process steps into a single'
operation.
If such a concept is successful, the regenera-
tion system cost could be reduced $3 to $4/k~. The use
of natural gas or oil as a fuel for the present regenera-
tion system concept could result in savings of approxi-
,mately $l/kw.
. Coal Feeding System:
The coal feeding system design is
based on systems which have,been built and operated.
The
design provides a. separate coal' feeding system for each
fluidized bed in order to assure control of the coal feed
rate to each bed. It may be possible, however, to reduce
the number of coal feed systems from 16 to 4 if independent
control of solids flow to each bed in a module can be
achieved from a single pressurized injector.
cost reduction is estimated to be $1 to $2/kw.
The potential
. Stack Gas Cooler Design:. Cost'estimates were obtained for
the stack gas coolers, but no attempt was made to optimize
the design or c6nsider nonconventional designs, .such as those
using fluidized beds. Preliminary conceptual evaluation
indicates that the cost might be reduced.~ $l/kw from the
present $4.40/kw.
Capital cost savings may also be realized by an increase in
plant performance over that projected., The gas turbine temperature of
l600°F is based on a temperature drop of 150°F of the flue gas between
the fluidized bed and gas turbine.
Two factors indicate this may
. 396
-------
be well below the actual gas turbine temperature: burning above the bed
is expected to increase the gas temperature above the l750°F bed
temperature, and the heat losses are not expected to be as large as
projected. If the gas turbine temperature were increased to l700°F,
the total capital cost reduction would be equivalent to $6 to $7/kw.
The total capital cost would be reduced ~ $20/kw if all the
potential savings in equipment cost and improvement in plant performance
were achieved. This would give a plant cost of ~ $255/kw -- about 25%
less than the conventional plant cost.
Fora plant using a cooling tower instead of a once-through
cooling system, an additional saving in the cooling tower cost can be
credited to the pressurized system over a conventional plant because
of the reduced heat rejection from the condenser in a-combined cycle
plant and the higher plant efficiency. The cost estimate for cooling
water equipment for a conventional plant is ~ $8/kw.
will depend on the combined cycle plant efficiency.
The cost savings
These design alternatives do not represent cost savings
which might be achieved by advanced concepts in boiler design, sulfur
recovery processes, or power cycles.
The pressurized plant could also be operated by using a once-
through dolomite system instead of regenerating stone and recovering
sulfur or sulfuric acid. The capital cost for this system is estimated
to be $264/kw.
Atmospheric-Pressure Fluidized Bed Boiler Plant. The preliminary
design of the atmospheric pressure fluidized bed boiler was made by
Foster Wheeler for a nominally 300 MW plant. The boiler cost estimates
were extended to 635 MW capacity. The total capital cost estimate for
the plant with sulfur recovery is $3l2/kw. This represents a 7% cost
reduction over a conventional plant with stack gas cleaning. The cost
of a once-through limestone system is estimated to be $295/kw.
397
-------
Cost reductions are not projected for the atmospheric pressure
system design. The auxiliary equipment to the boiler incorporates an
advanced coal feeding system and a single-stage regeneration system.
Thus, additional possible cost reductions in system components have not
been identified. Advanced concepts in boiler design and advanced steam
conditions may offer the potential for additional savings.
Capital costs were projected for oil-fired plants burning
high-sulfur residual oil. Capital costs for 635 MW oil-fired power
plants are summarized in Tables 1-52 and 1-53. Capital cost as a
function of capacity is shown in Figure 1-154 Capital cost for
pressurized fluidized bed combustion combined cycle oil-fired plants
with sulfur recovery may be ~ 14% less than conventional plants with wet
scrubbing without sulfur recovery. The atmospheric pressure fluidized
bed boiler plant is only 2% less than a conventional plant.
These capital costs are higher than many reported in other
studies. However, since the costs for the various types of plant were
developed on the same basis, it is thought that they provide valid
comparisons.
Energy Costs
Energy costs for coal-"and oil-fired plants are summarized in
Tables 1-54 and 1-55. Energy cost reductions for pressurized fluidized
bed combustion combined cycle coal-fired plants with sulfur recovery
may be ~ 10% or 1-1/4 to 1-1/2 mills/kwh beiow cdnventional "plant
energy costs. The reduction in energy costs for atmospheric pressure
fluidized bed combustion with sulfur recovery is estimated to be ~ 2%.
Energy cost reductions for pressurized fluidized bed combustion combined
cycle oil-fired plants with sulfur recovery may be ~ 7% below conventional
plant energy costs. Atmospheric pressure fluidized bed systems do not
show significant energy cost reductions over new conventional plants.
398
-------
TABLE 1-52
BOILER PLANT EQUIPMENT COSTS
CONVENTIONAL (a)
Steam generator(b) 29.40
Draft system
FD fans 0.90
Precipitators
Draft flues/supports 2.63
Stack/foundation 0.47
Fuel oil equipment 8.60
Ash handling
NOx control 1. 70(f)
Instrumentation 2.75
w
\0 Miscellaneous 0.86
\0
TOTAL 47.31
ATMOSPHERIC PRESSURE
FLUID BED COMBUSTION
PRESSURIZED
FLUID BED COMBUSTION
26.80
12.90
1.50(c)
5.00(c) 7.50(d)
1. 66 (c) 4.43(c)
0.47(c) 0.47(c)
9.60(e) 1l.00(e)
2.69(c) 0.55(c)
3.l0(c) 3.1O(c)
0.94(c) 0 . 94 (c)
51.76 40.89
(a) Based on cost data from United Engineers and Constructors.
(b) Includes air heater cost for atmospheric pressure systems. Boiler cost for oil gasification assumed to
be the same as conventional boiler--this will depend on burner cost, flame characteristics, etc. Fluid
bed boilers assumed to be 10% less than coal-fired boiler designs since no carbon burn-up cell is
required and the gas flow is reduced. No credit taken for reduced fouling from vanadium and sodium
in fluidized bed systems.
(c)Costs same as coal-fired plant.
(d) Primary collectors in coal-fired plant not included since there is no carbon burn-up cell.
(e)Cost will depend on oil piping and pump cost to gasifier or combustor relative to conventional cost to
burners. This cost is 20% of the fuel oil equipment cost in a conventional system. The cost for oil
gasification is not expected to increase since the pressure and number of feed points are similar. The
cost for fluid bed combustion boilers is projected to increase due to more feed points and a higher
pressure requirement for the pressurized system.
(f) Flue gas recirculation.
-------
TABLE i-53
OIL-FIRED POWER PLANT COST BREAKDOWN
635 MW Plant
ATMOSPHERIC PRESSURE PRESSURIZED
FLUID BED COMBUSTION FLUID BED COMBUSTION
CONVENTIONAL (a) $/KW $/KW
$/KW Regenerative Regenerative
Land & land .rights 1.13 1.13 1.13
Structures & improvements 24.95 25.80(b) 19.57(b)
Boiler plant ~quipment 47.31 51. 76 40.89
Sulfur removal equipment 20.00(c) 14.10 12.68
Gas turbine-generator equipment 14.80
Steam turbine-generator equipment 55.76 54.96 .44.14
+=""' Electric plant ~quipment 14.40 14.40 l5.30(d)
o Miscellaneous plant equipment 4.08 4.08 3.45(d)
o 27.80(d) 27.00(d)
Undistributed costs 28.55
Other plant costs 2.91 2.91 2.91
. Subtotal 199.09 196.94 181.87
Normal continge~cy 11.95 11.82 10.91
Subtotal 211.04 208.76 192.78
Escalation 47.48 46.19 36.15
Subtotal 258.52 254.95 228.93
Interest during construction 43.63 40.14 30.05
General items and engineering 5.5 5.5 5.5
TOTAL CAPITAL COST 307.65 300.60 264.48
(a) Assumes stack gas scrubbing system, no particulate control beyond scrubber, no additional
stack height over alternates. No by-pass around the scrubber system is provided.
(b) Same as coal-fired plant.
(c) Once-through throw-away system.
(d) Estimated by assuming same percent reduction for oil-fired plant as was assumed with coal-
fired plant over conventional coal-fired plant.
-------
Curve 645275-A
500
Fuel: Oil
Conventional Plant Includes Wet Scrubbing
Flu idized Bed Plants Include Su Ifu r Recovery
:i: 400
.::&.
-
~
+:-
o
I-'
.......
~ At mospher ic-pressu re
u
co Flu idized Bed Gas ification
~ 300 1976 Operation
co
u
Conventional
1976 Operation
co
15
I-
Pressurized Fluidized
Bed Combustion
1975 Operation
Atmospheric - Pressure
Flu idized Bed Boiler
1976 Operation
200
100
200
400
700
800
900
1000
300
500 600
Plant Capacity, mw
Fig. 1-154- New plant cost vs capacity
-------
TABLE 1-54
ENERGY GENERATION COSTS FOR COAL-FIRED POWER PLANTS
635 MW PLANT
CONVENTIONAL ATMOSPHERIC PRESSURE PRESSURIZED
PLANT FLUIDIZED BED BOILER FLUIDIZED BED BOILER
Regenerative Once-through Regenerative Once-through
-
Fixed charges 8.10 7.61 7.20 6.75 6.44
Fuel 4.11 4.44 4.15 4.31 4.00
.r::-
a
I\) Dolomite or limestone 0.05 0.29 0.12 0.52
Operating and maintenance 0.99 0.86 0.67 0.90 0.71
Total, mi11s/kw h 13.20 12.96 12.31 12.08 11.67
-------
TABLE 1-55
ENERGY GENERATION COSTS FOR OIL-FIRED POWER PLANTS
635 MW PLANT
Fixed charges
Fuel
g ,
w
Dolomite or limestone
Operating and maintenance
TOTAL, mills/kwh
CONVENTIONAL
BOILER
(ONCE-THROUGH)
7.52
4.11
0.05
0.91
12.59
ATMOSPHERIC PRESSURE
FLUIDIZED BED BOILER
(REGENERATIVE)
7.35
4.60
0.05
0.78
12.78
PRESSURIZED
FLUIDIZED BED BOILER
COMBINED CYCLE
(REGENERATIVE)
6.47
4.35
0.12
0.82
11.76
-------
Evaluation
The fluidized bed combustion systems are compared with each
other and with conventional systems with stack gas cleaning. Potential
advances and uses for fluidized bed combustion boilers are reviewed.
Comparison
A comparison of atmospheric and pressurized fluidized bed
combustion power systems must consider several factors:
.
Design characteristics
Effectiveness in pollution abatement
Effectiveness in fuel utilization
.
.
.
Economy of power generation
Status of technology
.
.
Development requirements.
Design Characteristics.
A summary of the design characteris-
tics for each system is presented in the next matrix.
The heat transfer
characteristics -- heat and
release rates and heat transfer surface --
are compared in Table 1-56. The height of the fluidized bed boilers is
approximately 100 ft for both 300 MW and 600 MW capacities. Conventional
coal-fired boiler height is approximately 200 ft. The pressurized con-
cept is modular and can~ for the most part~ be shop-fabricated.
This
greater degree of standardization and shop fabrication will allow the
power industry to reduce costs and to capitalize the experience obtained
during initial development.
Heat Transfer Surface.
The salient advantage of the fluidized
bed boiler over the conventional boiler is the reduced heat transfer
surface.
A comparison of the surface requirement for the atmospheric
and pressurized fluidized bed boilers with a conventional p.f. boiler is
as follows:.
404
-------
COMPARISON OF HEAT TRANSFER SURFACE AREAS
Conventional Boiler
Function Boiler(a) ressurized c
Evaporator 4 4 (FB) 4 (FB)
2.97 x 10 5.28 x 10 1.21 x 10
Superheater I 8.16 1. 75 (FB) 1.27 (FB)
Superheater II 1. 75 (FB) 1.67 (FB)
Reheater 11.3 1.65 (FB) 1.67 (FB)
Economizer 15.36 11.44 (CONV) 1. 906 (CONV)
Total 37.79 21. 86 x 104 9.918
Regenerative Tubular
Air Preheater 53.0 x 104 208 x 104
(a)
Georgia Power & Light Ft. Hammond No.4, 530 MW, 2486/1000/1000,
3.6 x 106 lb. steam/hr.
(b)
600 MW, 2400/1000/1000,
6
3.8 x 10 lb. steam/hr.
6
3.26 x 10 lb. steam/hr.
(c)
600 MW, 2400/1000/1000,
Pollution Abatement Potential. The projected emissions to
the atmosphere from the fluidized bed boiler plants are compared with a
conventional plant with stack gas scrubbing in Table 1-57. These emis-
sions are consistent with the plant designs and costs. Further reduction
in emissions is possible for each system but would represent additional
costs. The solids disposal problem is significantly reduced with the
fluidized bed combustion system where the waste solids are dry and
large -- 1/16 inch. . Thus, the solids are easier to handle and have a
greater potential for market va1ue[25] than the slurries produced from
405
-------
FLUIDIZED BED BOILER DESIGN CHARACTERISTICS
DESIGN 'CHARACTERISTICS
PRESSURIZED
UTILITY
ATM. PRESSURE
Capacity
Power
Steam, lb/hr.
Conditions
Design Features
Water circulation
Modules
Number
Beds/module (inclCCBCJ:
Bed orientation
Heat transfer surface
Arrangement
Tube size, in.
Bed depth, it.
Gas velocity, fps
Bed temperature, of
Design point
Part load
CBC
Particle size, in.
Excess air, %
Primary beds
CBC
Total
Control method
Sulfur removal
Stone
Ca/S ratio
Make-up, %
Coal feeding .
Primary bed area, f~2.
Total bed area, ft2
Feed points to each
primary bed
600 MW
6
3.8 x 10
2400/1000/1000
once through
forced circulation
4
6
stacked
horizontal
1-1/2 & 2
2.5
6-10
serpen tine
1600
> 1300
1900-2000
-1/8 or -1/4
10
50
23.6
bed temperature and
module shut-down
limestone
6
'V10
312
7500
'V 30
406
635 MW
(17.4% gas turbine
'V 3.4 x 106
2400/1000/1000
power)
once through
forced circulation
4
5
stacked
horizontal
1-1/2 & 2
10-16
5.6-9.1
serpentine
1750
> 1300
1900-2000
-1/8 or -1/4
used 10%
80
15.8
bed temperature and
module shut-down
dolomite
6
'V10
70
1200
4-8
-------
TABLE 1-56
HEAT TRANSFER CHARACTERISTICS
Conventional
.f. Boiler
Heat release rate
Boiler cross-section
Btu/hr - ft2
6
2.2 x 10
Bed volume
Btu/hr - ft3
01 1 (d)
Bo~ er vo ume
3
Btu/hr - ft
4
1.7 x 10
(e) 2
Heat transfer surface, ft /MW
740
160
(a)Based on ::0 -Foster Wheeler 280 MW boiler design.
(b) Based on @-Foster Wheeler 317 MW boiler design.
(c) Average which includes carbon burn-up cell.
(d) Boiler volume within water walls.
(e) Excludes air preheater or stack gas coolers.
407
POWER PLANT CONCEPT
Atmospheric Pressure(a)
Fluidized Bed Boiler
6 (c)
o . 7 x 10
6
0.3 x 10
4
5.1 x 10
410
Pressurized(b)
Fluidized Bed Boiler
6 (c)
4.2 x 10
6
0.3 x 10
20.
x 104
-------
TABLE 1-57
AIR POLLUTION ABATEMENT
- .~.,. -- -- +" '; ~.
.-- ... .. --POWER PLANT CONCEPT
Conventional (a) Fluidized Bed Combustion
, Atmospheric Pressure T Pressurized
Sulfur dioxide
(lb S02/106 Btu)
1.5
0.7
0.7
Nitrogen oxides
(lb N02/106 Btu)
1.2-1.4
0.4-0.8
0.1-0.2
Particulates
(lb/106 Btu)
0.05-0.15
0.05(b)
. 0 .15 (c)
B!AS :
4.3% Sulfur, 8.5% Ash Coal
(a)With stack gas scrubber for sulfur and particulate control; no NO
x
(b) Electrostatic precipitator.
(c) Does not include particulate removal after the gas turbine.
control.
408
-------
wet scrubbing systems. The atmospheric pressure system also offers the
advantage of large solids relative to the conventional plant. A con-
ventional steam or water transport system is projected for solids
handling, which increases the solids handling problems relative to the
pressurized system.
Fuel Utilization.
The plant heat rates and alternate fuel
capabilities are summarized in the next matrix.
The atmospheric pressure
fluidized bed boiler plant has a lower heat rate than the conventional
plant because of the higher carbon loss. The combined cycle pressurized
fluid bed boiler plant has a higher efficiency than the conventional
plant. No attempt was made in the present design to exploit the poten-
tial of the combined cycle plant to achieve higher plant efficiencies.
The fluidized bed plants have the potential for burning low-grade fuels
which cannot be burned in conventional plants or, if they are burned,
result in high plant cost or low plant efficiency and plant availability.
The low bed temperatures permit high-alkali fuels to be burned without
corrosion or fouling of the boiler tubes .(1) Since many of the western
coals have high alkali and high ash content, the fluidized bed concept
offers an economical means of utilizing these fuels. The Igni-fluid
process illustrates the potential of fluidized beds to burn low-grade
fuels for power generation. [35] The fluidized bed concept also offers
potential for burning solid wastes, either as the only fuel or as a
supplement~ry fuel.
(1)
See pp. 132-156.
409
-------
TABLE 1-58
FUEL UTILIZATION
Conventional
POWER PLANT CONCEPT
Fluidized Bed Combustion
Atmospheric Pressure" Pressurized
9230(c) 8892
(37%) (38.4%)
Plant heat rate(a)
,-;Ji tu /:kwh .
, 9l22(b)
(3'7.4%)
Alternate fuel
capability
limited to low-
alkali, low-ash
coals; low-vana-
dium oil; or gas
potential for burning high- and
low-grade coals, soiid wastes
tailings from fuel processing
facilities, oils, or gas.
(a)Based on 1-1/2 in. Hg steam turbine back pressure.
(b)Hammond #4, Georgia Power & Light.
(c)This assumes a boiler efficiency of 88.6%. If the boiler efficiency
is 86.1%, as projected in the design, the heat rate would be 9500 Btu/kwh.
hT
410
-------
Development Status of Fluidized Bed Combustion.
The amount of
time and effort devoted to atmospheric pressure fluidized bed combustion
has thus far been one order of magnitude greater than that devoted to the
pressurized concep't. The technical accomplishments represented by the
capabilities and performance of fluidized bed boiler pilot plants, how-
ever, appear comparable. The next matrix presents technical and operat-
ing data on the Pope, Evans and Robbins (PER) atmospheric system and the
National Coal Board (NCB) pressurized boiler. These two pilot plants
represent the most advanced experimental work on the respective fluidized
bed combustion concepts.
.
The size and capacity of these tWo systems do not
differ greatly
.
The pressurized boiier carried out longer runs with
fewer operating difficulties at an earlier date
(January 1971) than the atmospheric system (July 1971)
although it was commissioned a year later
.
Results on sorption of S02 and reduction of NOx were
more favorable in the pressurized system (see
Table 1-57)
.
Adequate particle removal was achieved in the pressur-
ized system so that the gases did not erode or deposit
on turbine stator blades.
Several additional factors must be considered in comparing the accomplish-
ments of the atmospheric and pressurized systems. The PER atmospheric
boiler did operate at design conditions of pressure, temperature, and
velocity; and it did illustrate operation with an auxiliary unit for
burning carbon carried by combustion gases from the fluidized bed and
recovered by dust removal equipment. The NCB pressurized boiler operated
at conditions selected for British designs but not at those for the
Westinghouse-Foster Wheeler design; it contained neither an auxiliary
carbon combustion system nor a sorbent regenerator.
4n
-------
PRESENT EXPERIMENTAL APPARATUS (a)
ATMOSPHERIC PRESSURE
PER SYSTEM
6 ATMOSPHERIC-PRE&SURE
NCB-BCURA SYSTEM
Primary System
Componen ts
NCB
FBM, CBC, vertical welded wall
heat transfer surface in bed,
coal feed systems.
Primary bed, particulate remo-
val, pressurized coal feed sys-
tem, turbine blade cascade,
horizontal heat transfer sur-
face in bed as tube bundles.
2
Bed area, ft
Max. coal feed rate reported,
lb/hr
8.75
"'900
8
"'600
+"
I-'
I\)
Projected equivalent capa-
city, kw
Longest runs, hrs
"'1300
Years of operation
Hours of operation
OAP sponsored
3-1/2
"'900
350 (b)
2-1/2 (Feb.
1969)
435 (Aug. 1970-July 1971)
Integrated System
Components
Primary system with
regenerator
156 (c)
460
sorbent
Longest run, hrs
Coal feed rate, lb/hr
Projected equivalent
capaci ty, kw
700
\a)Includes only PER and BCURA Pressurized Apparatus.
(b)The apparatus was designed for "'100 hrs continuous operation. The long-term tests were thus
interrupted to refill the coal feeding system and were not interrupted due to mechanical or
operating problems.
(c) Because of operating problems, there were three interruptions, two of them due to coal feeder
plugging.
-------
Both atmospheric and pressurized fluidized bed power systems
require further development effort. The technical problems on which
further information is required are listed in the next matrix. This
information can be most rapidly and economically acquired on develop-
ment scale equipment.
Economy of Power Generation. Capital and energy generation
costs for coal-:andoil-fired fluidized bed combustion power plants and
..
conventional power plants were presented on pages 389-403. A comparison
of these cost data shows
Potential
.
Pressurized fluidized bed combustion combined cycle power
plants using a high-sulfur coal fuel, with sulfur recovery,
may reduce energy costs 10% below conventional plants with
stack gas cleaning.
.
Pressurized fluidized bed combustion combined cycle plants
burning high-sulfur oil, with sulfur recovery, may reduce
energy costs 7% below conventional plant burning high-
sulfur oil.
.
Atmospheric pressure fluidized bed combustion power plant
systems burning high-sulfur c9alor oil may reduce energy
costs less than 2%.
.
Environmental control costs for 802' NO , and particu-
lates represent ll%tto 20%(1) of convent~onal plant equi.p-
ment costs and 5%,to7%(2) of fluidized bed plant equipment
costs. The cost for control of heat rejection to water
sy&tems adds ~ 4% to the conventional plant equipment cost.
The cost for the fluidized bed systems is less since only
50%:.to 80% of the energy is produced by steam turbine gen-
erators.
Potential energy cost reductions for fluidized bed power
generation systems are projected as the result of:
(1)20% assumes $40/kw for 802 and particulate control.
(2)
Does not include particulate removal required for gas turbine opera-
tion.
413
-------
PROBLEMS IN FLUIDIZED BED POWER SYSTEMS
FOR INVESTIGATION IN DEVELOPMENT PLANTS
ATMOSPHERIC PRESSURE
PRESSURIZED
control
Operation of Deep Beds
NO
--X
Is two-stage combustion necessary?
Alternative techniques for control
Solids handling
Sorbent activity and utilization
Feasibility of producing high S content gas
Sulfur r 3covery s ys tern
Temperature gradients?
Particle circulation characteristics
Combustion efficiency
S02 removal
NOx minimization
Evaluate particulate carry-over and control
e':J.uipment
Evaluate particulate carryover and
control equipment
Re~eneration system
Turbine blade t es ts
Boiler tube materials checked
Re~eneration system
Solids handling
Sorbent activity
Feasibility of producing high S content gas
./='
f-'
./='
Heat transfer surface studies
Boiler tube spacing and heat transfer coeffici-
ents above bed
Sulfur recovery system
Boiler tube materials checked
Operation:
s tart-up. shutdown. load follow
Advanced concepts:
higher steam temperatures
Hea t trans fer surface studies
Boiler tube arrangement and heat transfer
coefficients
Operation:
start-up. shutdown. load follow
Advanced cvncepts: h.igher g.as turbine tempera-
tures. circulating beds. higher steam tem-
peratures
NOTE:
Very little information on the development plants is transferrable from one concept to the
other. For example. work on large cross-section shallow beds provides no insights into the
potential problems with deep beds; particulate control equipment will operate at different
temperature and pressure levels; regeneration systems have to process stones with entirely
different histories; the pressure drop across the distributor can be much greater in the
pressurized design. but the design may be more sensitive for the large area shallow. beds;
and boiler tube materials are subjected to different temperatures and pressures. Work on
a sulfur recovery system might be common. although even this is not yet clear.
-------
.
Modification of the present design concepts
Advanced boi1er-regeneration-su1fur recovery designs
and power cycles
.
.
Utilization of low-grade fuels.
Modification of Present Concepts.
Capital cost reductions
from simplifications and improvements in the present design concepts
were discussed on pages 390 and 396.
The potential capital cost reduction for the pressurized
system is estimated to be ~$20/kw. This corresponds to an energy cost
reduction of ~0.50 mills/kwh, a 40% reduc~ion in the present design
energy cost. If the gas turbine temperature can also be increased as
the result of burning above the bed or lower heat losses than projected,
the plant heat rate would increase 1% to 2%, resulting in an additional
energy cost reduction of ~0.1 mills/kwh. The potential energy cost
reduction over the conventional plant with wet scrubbing using the
modifications in the present design with sulfur recovery is 13%. Mod-
ification in the present atmospheric pressure boiler plant design con-
cept does not appear to result in significant cost reductions.
Advanced Concepts. Advanced designs have the potential for
'reducing capital costs and reducing the plant heat rate over the present
design concepts. Three areas have been considered: boiler design,
regeneration-sulfur recovery processes, and power cycle conditions.
.
Advanced boiler design concepts which have potential
advantages in maintaining uniform bed temperature, in
feeding reactants to the bed, in separating ash, and
in power plant control were considered. One of these is
a deep, recirculating, fluidized bed boiler concept.
.
Regeneration-sulfur recovery processes have also been
assessed in order to develop a process with reduced
capital cost and improved operability. Particular
emphasis was placed on processes to eliminate a
separate C02 generation system or for an H2S enrich-
ment system.
415
-------
.
The power cycle operating conditions can be advanced to
improve the plant heat rate -- increased gas turbine tem-
perature and increased system temperature and pressure.
In the high-pressure combined cycle system the bed temperature
is limiting, rather than the gas turbine i,nlet temperature. If a way is
found to decrease the temperature difference between the bed and the gas
turbine, or if a supply of gaseous fuel can be obtained for heating the
exhaust products from the beds to the current limiting level for the gas
turbine, Le., l800ciF-l900°F, the plant heat rate would decrease 2% to 3%,
and the plant capacity would increase 5% to 8%. The effect on specific
plant cost would be somewhat less than the capacity change. Gas turbine
temperature of 2500°F is projected for 10 to 15 years from now, so the
potential for growth in this area is substantial. No significant
improvements are expected in gas turbine aerodynamic efficiencies in the
foreseeable future.
There is, however, a definite trend in gas turbine
design toward higher mass flows on an area basis. This will bring about
further reduction in specific plant cost. Reduced corrosion in fluidized
bed boilers could result in increasing steam temperatures up to l200°F
or higher. Both the plant capacity and the plant heat rate would change
about 4% per 100°F increase. Near-term improvements in gas turbine and
steam turbine utilization could result in heat rate improvements on the
order of 10% and in specific plant cost reduction on the order of 10% to
l5%~ Longer term advances in gas turbine temperatures would bring about
further improvements.
The potential energy cost reductions for the atmospheric
pressure and pressurized fluidized bed boiler plants based on successful
development of the advanced concepts are summarized below. The energy
costs for the atmospheric pressure and pressurized systems represent
reductions of 11% and 23%, respectively, over a conventional plant with
stack gas cleaning:
416
-------
TABLE 1-59
ADVANCED FLUIDIZED BED COMBUSTION POWER PLANT COSTS
WITH SULFUR RECOVERY
BASIS: 600 MW
Capital cost, $/kw
Heat rate, Btu/kwh'
Energy cost, mills/kwh"
ATMOSPHERIC PRESSURE
PLANT
290
8300
11.90
PRESSURIZED
PLANT
240
7600
10.30
Fuel 'Utilization.
Fluidized bed combustion boilers offer a
means for burning low-grade fuels which cannot be effectively burned in
conventional boilers. Two advantages can be realized: effective
utilization of fuel resources and reduction in energy cost as the result
of low-cost fuel supplies.
Fluidized bed combustion boilers may also be attractive for
burning char produced from coal gasification for power generation. A
possible combination of gasification and combustion is shown in Figures
1-155 and 1-156. The gasifier produces a clean fu~l gas by driving off
the volatile components from the coal and removing sulfur. The residual
devolatilized char from the gasifier is burned in a fluidized bed
combustion boiler which also acts as a heat recovery boiler for the gas
turbine exhaust. The system may have improved economics and provide
improved turn-down capability by storing char.
Assessment. In comparing technology, the state of development
and the economics of the problem must be considered and a balanced assess-
ment rendered. Demonstration installations of wet scrubbers are well
underway; 70%xo 90% removal of S02 can be achieved; some operating
problems -- corrosion and deposition -- must be solved. Costs of new
scrubber installations -- about $20 to $40/kw' -- have been esti-
mated with fair accuracy.
Combustion modification techniques on large
conventional boilers have been under investigation and it appears that
NO
x
emissions can be reduced.
417
-------
CaO
Coal
Dwg. 2920A65
Combustor
Turbi ne Compressor
I\)
Air
Generator
CaS
Gasifier
CaD
Ash
Char Storage
Combustor
Generator
Steam Turbi ne
Fig. 1~I55 -FI uidized bed gasification-combustion power plant
418
-------
'--I
. I
I
I
I
I
I
-- -~----------~
- - -~- - - - - - - - - -. :
-- -~------i I I
I I I
-~--A I I I
I I I I I
....... F. B. II I HPT LPT G
:: Gasifier I
-:" . ... . .. '. ,." .. . . - . I I ~ + + ~ I + I
~: 1 1 10-
~ L_--~-~--L-_l_-----r--~ c qOut
:.J. I I I' I
Ash """"r~-'0-----j-___~tIF~~ B~P
II . I
C T G I i. -~ J
~-~ -= =-- - -+To Stack
.
.
.
I
Air
F-
Sorbent Out
F. B. Boiler
Sorbent In
Coal
+:""
I-'
\D
Dvlg. 6171A66
-----Air
- - - - Comb. Prod.
.....-... Fuel Gas
- - - Steam
- --Water
Solids
Figu re 1-156
-------
Atmospheric fluidized bed combustion is still in a develop-
mental stage. Demonstrated S02 removal is 80% .to 90%; NOx reduction may
not be adequate to meet emission standards. The technology of sorbent
regeneration and sulfur recovery is not sufficiently developed for
accurate syste~s design and economic evaluation. With reasonably
optimistic assumptions, energy costs from atmospheric fluidized bed
boiler power plants are estimated to be 2% ,to 3% less than those from
conventional plants equipped with wet scrubbers. This cost advantage
might well disappear as further technical data are developed on the
atmospheric system.
Pressurized fluidized bed combustion is also in the develop-
ment stage. Demonstrated S02 and NOx reductions are adequate to meet
emission standards. Energy costs are projected to be 10% and poten-
tially 24% less than conventional power plants. There is thus a much
larger economic margin for solving technological problems.
System evaluation and experimental test results indicate that
the pressurized system is more attractive for pollution abatement and
for economy of operation. A summa.ry of the plant comparisons and
potential is presented in Table 1-60. The commercialization of either
the atmospheric or the pressurized system would require 9 to 11 years
6 .
and $65-75 x 10 , primarily because the construction of a demonstration
plant dominates the time and cost estimates.
The number of common technical problems remaining inatmo-
spheric and pressurized systems appears very small (see page 414). The
order of magnitude differences in pressure, in mass flow per unit of bed
area, and in bed depth make it difficult to transfer technical informa-
tion and design procedures.
In conclusion, the conventional power plant with combustion
modification for NOx control and wet scrubbers for S02 and particulate
.control is the most highly developed technology. Pressurized fluidized
bed combustion technology is the most promising of its kind from the
point of view of effective pollution abatement, efficient fuel
420
-------
TABLE 1-60
COAL-FIRED POWER PLANTS
BASIS: 600 CAPACITY
4.3% S COAL
FLUIDIZED BED
CONVENTIONAL ATM. PRESS. PRESSURIZED
WITH SCRUBBING ONCE-
SYSTEM REGENERATION REGENERATION THROUGH
337 313 295 277 264
13.40 13.12 12.46 12.12 11. 71
Cost
Installed capital cost, $ikw
Power cost, mills/kwh
w/o sulfur credit
Environmental Factors
+=-
f\)
~
Plant heat rate (exclusive of
sulfur recovery or removal
systems)6Btu/kwh
S02' (lb/10 Btu)
NOx' (lb N02/106 Btu)
Particulate, (lb/106
Potential for Advanced Concepts
Cost - capital
Cost - energy
Efficiency
Fluidized Bed Boiler Systems
Relative advantages (technical)
Technical problem areas
9230
1.5
1.2-1.4 w/o
0.05-0.15
9550
0.70
0.4-0.8
0.05
8967
0.70
0.1-0.2
0.05-0.15
recir.
>cost projected
above
285-290
11. 90
240-250
10.10-10.50
No increase pro-
jected
40%
47%
Less stringent particu-
late removal
More expo data on opera-
tion
Coal feeding
NOx removal
Control
Higher efficiency
Reduced size
Reduced heat transfer
surface
Greater potential
Reduced construction time
Deep beds
Gas turbine blade
erosion
S02 removal
Solids handling
Coal feeding
NOx emission
Control
SOL removal
-------
utilization, and economicai power generation. Atmospheric fluidized
bed combustion may be better than conventional systems, but further
development is required.
422
-------
2.
ASSESSMENT OF A FLUIDIZED BED OIL
GASIFICATION-COMBUSTION POWER
SYSTEM
INTRODUCTION
Atmospheric pressure residual oil gasification/desulfurization
concepts are being studied under contract with OAP by Westinghouse and
Esso (England) to make possible the on-site production of a low-sulfur
fuel gas suitable for power plant utilization.
A number of factors
indicate a promising market potential for such a process: stricter S02
emissions standards, unstable prices for low-sulfur fuels (both coal
and oil in the 0.3 to 0.5 percent sulfur range), and a growth in the
use of residual oil for power generation on the East coast of the United
States, in both new plants and conversions, (expected annual growth
rate between 1969 and 1975,13.5 percent).
Fundamentally, oil gasification/desulfurization consists of a
complex ~rack~ng-partial combustion phenomenon in a fluidized bed (the
I
gasifier) of lime. H2S evolved during this ~omplex operation is
simultaneously absorbed by the lime bed to yield a clean fuel gas and a
sulfided lime at a control temperature of 1600°F. The hot, low-sulfur,
fuel gas from the gasifier is transported directly to the boiler, where
combustion is completed.
Two operational modes for this process are
considered in this preliminary study: the regenerative mode and the
once-through mode. In the regenerative mode, the utilized gasifier
lime is continuously circulated to an air-fluidized bed regenerator in
which an S02-rich gas stream is produced. This S02 stream is converted
to elemental sulfur, or sulfuric acid, in a sulfur recovery section of
the system, while the regenerated lime is returned to the gasifier along
with a nearly stoichiometric amount of fresh make-up limestone. On the
other hand, the once-through mode utilizes the gasifier lime to a much
423
-------
greater extent than does the regenerative mode, and the sulfided lime is
disposed of directly, leading to a stone disposal rate three to four
times that of the regenerative mode; but the overall operation is
simpler.
This study investigates both modes to evaluate quantitatively
the merits and disadvantages of each.
Esso (England) has provided experimental information on the
oil gasification/desulfurization process based on small-scale batch
fluidized units (seven-inch diameter beds). Concurrently with the batch
experimental program, Esso (England) has constructed an approximately
1 MW continuous gasification/desulfurization system (operated regenera-
tively) which is now in its commissioning tun. This batch experimental
I I I
work has provided the basis for preliminary design considerations.
The scope of the preliminary investigation includes:
.
The evaluation and scale-up of the Esso (England) batch
data
.
The generation of conceptual system designs
.
The evaluation of both new boiler concepts and retrofit
boiler concepts in terms of feasibility and performance
.
The development of preliminary cost estimates and overall
system evaluations
.
Recommendations for future design considerations and
further experimental studies to be carried out by Esso
(England) (both continuous and batch)
.
Recommendations on the development of a utility
demonstration plant.
In general, the investigation considers boiler capacities greater than
100 MW, with interest centered on high sulfur removals (90% to 95%).
The basic operating variables are identified from the Esso
(England) nonsteady batch studies for both modes of operation, and this
information is interpreted in terms of the large-scale continuous
424
-------
operations of a commercial system (> 100 MW). The data indicate a
potential for 90% to 95% sulfur removal in a commercial unit under a
variety of operating conditions for both regenerative and once-through
operations.
Information to be obtained from the Esso (England) con-
tinuous unit will reveal more scale-up information needed in a detailed
design phase.
Material and energy balances are computed for the conceptual
system designs proposed with once-through and regenerative operations.
The methods used to maintain temperature control in the gasifier,
regenerator, and sulfate generator are design variables of major impor-
tance in those computations, and the results provide both definite
conclusions and an indication of the need for further experimental
information in the area of temperature control.
The feasibility and performance of retrofit design concepts
are considered on the basis of the boiler modifications needed for
conversion and the performance of the retrofit boiler.
Retrofit
concepts are divided into two groups:
"internal designs," in which the
gasifier system is placed directly beneath the boiler and the burners
are placed in the upshot position at the base of the boiler to provide
intimate contact between the burners and gasifier; and "external
designs," in which the burners are left in their normal boiler locations
and the gasifier system is placed external to the boiler. Preliminary
quantitative evaluations of these retrofit design concepts are made for
the retrofit cases of coal- and oil-fired boilers in terms of necessary
modifications and boiler performance.
New boiler designs are also
considered in terms of boiler-gasifier configuration. Burner design
and system operation are discussed with reference to both new and
retrofit boilers.
Capital and operating costs are estimated for regenerative
and once-through operations, both new and retrofit boilers, for capacities
of 600 MW with 80% and 40% load factors and 90% to 95% sulfur removal.
The estimates point out the most important system cost contributions.
425
-------
Reg,enerativ,eand :once-,through operations are ,co~pared w.itheach :o,ther"
and both .are ,comparedwiththeS02 control schemes .of low-sulfur oil
and s;tack gas cleaning. Other ,environmental factors, .such as NOx"
particulates, .and solid wast.eare included in the comparison,.
The preliminary analysis leads to lo,g:i:cal ,design conclusions
and future design recormnendations,and poi.nts 'out .some specificar,eas
needing further ,experimental analysis to assist in ,a de,tail.eddesign
phase. The most advantageous characteristic.s -of a gasification/
desulfurizationdemons.tration plant and the timing needed in the next
phases of ,development are also stated.
426
-------
GASIFICATION/DESULFURIZATION CONCEPTS
Two possible modes for gasification/desulfurization operation
are the regenerative mode and the once-through mode.
Figure 2-1
illustrates the major process streams and identifies the basic elements
of the two operational modes, without reference to the specific system
configuration or specific retrofit concepts. The operating variables
and temperature control schemes differ slightly for the two modes.
The basic elements of the regenerative operation are the
gasifier vessel and the regenerator vessel, as shown in Figure 2-1.
In
the regenerative operation, residual oil is injected into the gasifier
vessel, an air-fluidized bed of lime at l600°F operated with substoichio-
metric air (~ 20% of stoichiometric), to yield by cracking, partial
combustion, and H2S absorption by the lime, a hot, low-sulfur, fuel gas
and sulfided lime. The fuel gas is transported to the boiler burners,
where combustion is completed, and the sulfided lime is sent to the
regenerator. The regenerator is an air-fluidized vessel operated with
a slight excess of air at 1900°F. Regeneration takes place by reaction
of oxygen with the utilized lime to give an S02-rich stream (of about
10 mole % S02) and a regenerated lime with a decreased activity compared
to that of fresh lime. The S02 stream is transported to a sulfur
recovery system, and the regenerated lime is returned to the gasifier.
along with a nearly stoichiometric amount of fresh make-up limestone.
A variety of temperature control schemes may be useful in
controlling the temperatures of the gasifier and the regenerator.
tempe~ature control scheme used in the gasifier must conserve the
The
energy of the fuel in a form usable in the boiler.
For this reason
water injection, steam injection, and stack gas recycle are considered
for temperature control of the gasifier by the storage of sensible heat;
while the use of heat transfer surface in the gasifier to preheat
427 .
-------
Combustion Air
-\--- -Ciean ~;lGas
I. r----
I
I
I
I Fuel
I Oil
I I
J-- - ..,
I I
I
I
L~
To
+ Stack
I
. Combustion Air
--------T--
li mestone
N\ake-Up
Gasifier
Fuel Oil
Boiler
---
Clean Fuel Gas
"" Temperature
"" Control
Regenerated Strea m
Lime
+"
I\)
. ex>
---
Air
Regenerator
S02 Rich
Stream
Sulfur
Recovery
Sulfated
Lime
Disposal
Regenerative Mode
---Gas
liquid
Sol id Feed
- Solid Circulation
Fig. 2-I-Modes of operation
Dwg. 6167A48
+ To
I Stack
Boiler
Gas ifi er
li mestone Feed
Temperature Control
----Stream
Sulfate
Generator
Sulfated
Lime
Disposal
Once-Through Mode
-------
boiler water is also an attractive possibility. A fifth possibility is
to reduce the air/fuel ratio to a value low enough (~ 14% of stoichiometric)
that the gasifier is thermally balanced and no further temperature
control need be considered.
These five possible control schemes are
discussed in greater detail below.
Because of the high cost of sulfur recovery, temperature
control of the regenerator is best accomplished in such a way as to
avoid diluting the S02 stream produced by regeneration. For this
reason temperature control by heat transfer surface in the regenerator,
by the addition of fresh make-up stone to the regenerator, or by con-
trolling the rate of lime circulation between the gasifier and regenerator
are investigated.
The important operating variables in the regenerative operation
are the air/fuel ratio, the gasifier and regenerator temperatures, the
limestone particle sizes, the gasifier and regenerator fluidization
velocities and bed depths, the stone residence times in both the gasifier
and regenerator, and the limestone make-up rate. The effects of these
variables are interrelated with the temperature control schemes used
and the fuel and limestone properties.
An important phenomenon which takes place in the gasifier
during the cracking-partial oxidation is the deposition of carbon, or
coking, on the gasifier lime. Carbon deposition affects the sulfur
removal efficiency, the thermal efficiency of the system, and the overall
operation of the gasifier and regenerator when in the regenerative mode.
Minimizing the rate of carbon deposition is, therefore, an important
design consideration.
The basic elements of the once-through operation, shoWn in
Figure 2-1 are a gasifier vessel and a sulfate generator, or predisposal
vessel.
The operation of the once-through gasifier is the same as that
of the regenerative operation gasifier, while the sulfate generator
operates similarly to the regenerator but at a lower temperature (~ l500°F).
Thus, the sulfated lime from the gasifier is converted to calcium sulfate
429
-------
rather than calcium oxide.
The calcium sulfate may be disposed of while
the gas stream from the sulfate generator is sent to the gasifier.
The once-through operation requires a limestone addition rate three
to four times that used in the regenerative operation.
The same five temperature control methods proposed for the
regeneraiive gasifier are considered for the once-through gasifier,
while the sulfate generator temperature is controlled by one of two
proposed schemes:
heat transfer surface or excess air circulation.
The same operating variables are important in the once-through
operation as in the regenerative operation, and carbon deposition is
again a major factor in determining the sulfur removal efficiency and
the overall system operation.
However, in the once-through operation
the carbon deposition rate does not affect the system thermal efficiency
because the thermal energy of the deposited carbon is recycled as
sensible heat from the sulfate generator back to the gasifier.
The two conceptual modes of operation described illustrate
the basic characteristics of the systems considered in preliminary design
investigations.
430
-------
DATA EVALUATION
Results of the Esso (England) experimental program have
identified the important operating variables and the critical phenomena
associated with atmospheric pressure gasification/desulfurization.
Mechanisms involved in gasification, desulfurization, and regeneration
are discussed briefly, along with the essential experimental findings.
The application of the Esso (England) results to the design of a utility
gasification/desulfurization system is considered, and the areas of
particle carryover, metal retention, burner design, and the Esso (England)
continuous unit are included.
431
-------
The Esso (England) Program
Under contract with OAP, Esso (England) is conducting
laboratory tests on atmospheric pressure batch fluidized bed equipment
to investigate lime sulfur absorption and lime regeneration operating
variables [20, 66]. The batch work is being carried out in conjunction
with the construction and operation of a 750 kw continuous pilot unit.
The unit will be operated for 200 hours under the present contract.
The information obtained from the batch tests and the continuous unit
provides the basis for the design of a gasification/desulfurization '
retrofit system to be operated as a demonstration unit on a utility
boiler.
Investigations were carried out in four different batch units,
all very similar in design. All four beds were seven inches in diameter;
three of the units were 5-1/4 feet high from distributor to lid, and a
fourth unit was 12-1/2 feet high. Other design differences in the four
units involved location of cyclones and differences in insulation on
the units. Thus, results from the units are comparable and have been
combined.
Two types of experimental operations have been carried out:
. Fresh bed tests, in which the bed is loaded with a given
amount of fresh limestone, heated and calcined, and then
operated as a gasifier-qesulfurizer to note the effects
of the operating variables on the sulfur removal efficiency
as a function of time.
After two to three hours of
operation the fuel oil is no longer fed to the bed, and
regeneration of the sulfided stone is recorded as a
function of time.
432
-------
.
Cyclic tests have been carried out to investigate the
continuous behavior to be expected from a regenerative
process.
In these tests, the procedure carried out in
the fresh bed tests is repeated in cycles of gasification
steps and regeneration steps with a set amount of fresh
limestone added to the bed at the end of each cycle to
make up for losses due to attrition and sampling, and
to maintain a level of bed activity.
Two residual fuel oils have been involved in the batch tests -
a U.K. fuel oil with 3.03 wt % sulfur and a U.S. fuel oil containing
2.22 wt % sulfur.
Some of the fuel properties are listed in Table 2-1:
TABLE 2-1
FUEL OIL PROPERTIES
PROPERTY U.K. FUEL U.S. FUEL
Specific gravity' 0.95 0.95
Kinematic viscosity 98.2 (140°F) 54.7 (180°F)
Carbon (wt %) 84.8 85.9
H2 (wt %) 11.6 11.5
Sulfur (wt %) 3.03 2.22
Vanadium (ppm) 140 270
Nickel (ppm) 20
Sodium (ppm) 37 13
. Iron (ppm) 1 4
Asphaltenes (wt %) 3.72 6.3
Conradson carbon (wt %) 9.32 7.5
From these properties, the heating values of the two fuel oils were
estimated to be about 17,700 Btu/lb (low heating value). Batch operations
with these two fuels were found to be nearly identical.
433
-------
Three limestones have been utilized in the batch tests, and
two of them, a U.K. and a U.S. limestone, have been
well in the gasification/desulfurization operation.
two limestones are indicated in Table 2-2:
found to function
Compositions of the
TABLE 2-2
LIMESTONE COMPOSITIONS
COMPONENT
U.K. LIMESTONE
(wt %)
U.S. LIMESTONE
. (1961)
(wt %)
CaO
55.2
0.27
0.32
0.07
0.06
0.03
43.4
45.0
2.02
14.2
0.53
1.14
0.17
35.4
1.0
MgO
Si02
Fe203
A1203
Na20
C02 released
S as 803
Esso (England) examined gasification, desulfurization, and
regeneration performance as functions of the following operating variables:
gasifier and regenerator temperature, air/fuel ratio, fluidiza~ion
velocities, bed heights, limestone particle size, and limestone make~up
rate.
In applying the results of the Esso (England) batch tests to the
desig~ of a continuous commercial power plant, the following factors
must be considered:
.
The data are nonsteady state, batch data.
.
Experiments were limited by heat losses, because temperature
control could not be maintained under all conditions due
to the heat losses from the small batch units.
were also limited by this factor.
Bed heights
434
-------
.
The quality of fluidization may differ in the batch units
from that expected in a commercial unit; the batch units
were often operated in a state of slugging, or vigorous
bubbling.
435
-------
Gasification
The gasification phenomenon involves the cracking operation,
which leads to carbon deposition and the evolution of low: molecular
weight hydrocarbons, combined with the combustion of both deposited
carbon and fuel to provide the heat needed to maintain the system
temperature.
Since the bed is operated with excess fuel, the region
of the bed above the distributor and below the point of fuel injection
is likely' to be essentially an oxidizing region, while the zone above
the fuel injection point iS,essentially reducing. Cracking takes
place substantially in the ar:ea of the fuel injection point, while
deposited carbon combustion and fuel combustion are dominant in the
oxidizing zone. Gas'ifier product equilibrium' compositions, if
equilibrium does exist, are established in the reducing zone' of the bed.
Carbon deposition and product compositions have been studied in, the
Esso (England) program to provide information about the gasification
operation.
Carbon Deposition
Since carbon deposition affects the overall gasification!
desulfurization operation, some understanding of the. phenomenon is
necessary.
The operating and design variables affec.ting the carbon
deposition rate in the Esso (England) batch tests are:
.
Gasifier temperature.
.
Air/fuel ratio
.
Residual oil composition
.
Fuel distribution
.
Particle' sizes or surface area in bed.
436'
-------
Lowering the bed temperature increases the carbon deposition on the
lime; a lower air/fuel ratio has the same effect. Since temperature
control of the gasifier was not maintained, lower air/fuel ratios led
to lower bed temperatures, making the two variables interdependent.
The carbon deposition should be proportional to the amount of Conradson
carbon in the fuel (Conradson carbon is a measure of the amount of
residue left after fractionation of the fuel).
The location and quality
of fuel distribution in the bed are important and are highly dependent
on the quality of fluidization and solids mixing. The particle size or
lime surface area in the bed is probably a factor in the deposition of
carbon, but it has not been investigated directly.
If the rate at which carbon is laid down on the lime bed is
assumed to be a constant, Kl' for a given fuel, stone, and air/fuel
ratio, and if the rate at which carbon is burned off in the oxidizing
region of the bed is assumed to be first order in the fraction of carbon
on the stone, X , and first order in the oxygen concentration in the
c
oxidizing region, then for perfect mixing of particles in the bed
dX
c
WB ~ = Kl - K2Xc
(2-1)
where WB is the weight of the bed and K2 is the product of a kinetic
constant, the volume of the oxidizing zone, and the oxygen concentration.
Then, for a batch test, the carbon fraction and carbon deposition rate
are given by
Xc = Kl/K2
[1- exp
K2 J
(- - t) .
WB
and
dXc K2
WB ~ = Kl exp (- WB t)
.,
where t is the time variable.
On the other hand, for a continuous
operation in which lime is being withdrawn from the gasifier at a mass
rate NC' pound per pound of fuel, the carbon deposition is represented by
437
-------
KING
XcNG = NG + K2 .
(2-2)
Therefore, in a continuous operation, the carbon deposition rate will be
largely dependent on the lime residence time in the bed. For sho~t
residence times (N »K) the carbon deposition rate will be a
G 2
constant equal to Kl (Kl equals the initial carbon deposition rate in
a .batch operation), while for large residence times (NG « K2) Xc
becomes a constant equal to the equilibrium value found in a batch
operation. It is felt that the latter will generally be the case with
the conditions considered in this investigation.
Figure 2-2 summarizes the carbon deposition rates obtained in
batch tests as a function of the air/fuel ratio and indicates the
experimental conditions. Although lower air/fuel ratios necessarily
meant lower bed temperatures, the data do indicate a reasonable range
in which the carbon deposition should £a11 in a continuous operation.
Also, because o£ the superior temperature control expected in a
commercial unit, lower carbon deposition rates are expected.
Gasifier Product Analysis
Table 2-3 shows the results of analysis of the gasifier
product in three of the batch runs. The chemical analysis is incomplete
in that it excludes condensible tars, water, and H2S, all of which are
present in the product. The complete analysis can be estimated by using
the fact that the inert nitrogen should all be present in the analyzed
.product gas. The material balance is completed for run 53 to yield
the results shown in Table 2-4. This approximate product analysis
corresponds to a hot gas thermal efficienc~ of 92% and a cold gas thermal
efficiency of about 80%. The hot gas heating value of the product is
299 Btu/S ft3, 12% of which is sensible heat. The total heat generated
in this run by vaporization, cracking, deposition, oxidation, and sulfur
removal was about 3600 Btu/hr/1b fuel/hr, 40% of which was lost to the
surroundings to yield the observed gasifier temperature. Material
losses in the form of tars and soot have been neglected in the product
analysis, so the results are probably optimistic.
438
-------
Q)
:::s
L.&-
a 0.05
-... '
c
o
..c
50...
ro
U 0.04
..c
I
Q)
.....
ro
c:: 0.03
c
o
.-
.....
-!:='"
W
\0
In
8-
Q) 0.02
o
c
o
..c
50...
~ 0.01
, ..... ' ,
Q)
z
Curve 643909-A
. 'UK Fuel, UK Stone, Tall Un it
~'4 fVsec, 790-865°C
o UK Fuel, 9.3 wt % Con radson Carbon
UK Limestone, 600-1«>0 ~
'" 4 fVsec, Bed Temp. 8L1O-870oC
. UK Fuel, UK Stone
~ 4 fVsec, 780-850°C
.
.
.
o
A
~ 0
A
o
17
A US Fuel, '7.5 wt % Conradson Carbon
US Stone, 300-3175 ~
'" 6 ft/sec, '" 850°C
.
. 0
oM:) 0
.
.
o
A
.
.
.
19
20 21 22 23 24
Air/Fuel Ratio, % of Stoichiometric Air
27
18
25
26
Fig. 2-2-Carbon Deposition Data
-------
TABLE 2-3
PRODUCT ANALYSIS
RUN
CONDITIONS 53 65 86
Air/fuel (% of stoichiometric) 23.3 23.3 24.1
Fuel-limestone UK-US UK-US (1691) UK-US (1690)
% Sulfur removed 72.5 70.3 79.1
Temperature 880°C 885°C 865°C
Component (mole %)
N. 64.0 57.0 62.0
2
CO 10.4 10.3 8.3
C02 10.1 9.7 10.9
H2 4.2 10.0 7.4
CH4 4.4 6.5 5.1
C2 5.4 5.8 5.4
C3 0.3 0.5 0.8
C4 0.2 0.2 0.1
4'40
-------
TABLE 2-4
.-....... ----
ESTIMATED CaEMICAL ANALYSIS
COMPONENT
Wt. %
N2
CO
C02
H2
CH4
C2H4
C3H6
C4H8
Condensable tars, soot
58.70
9.55
14.60
0.28
2.31
4.96
0.41
0.37
'5.61 carbon + 0.92 H2
H2S
H20
0.20
1.99
441
-------
The total heat production in the gasifier as a function of the
. .
air/fuel ratio can be established by carrying out the same analysis for
runs 65 and 86, in addition to other data supplied by Esso giving CO
and C02 concentrations from the gasifier. The results of this procedure,
shown in Figure 2-3, are needed to co~pute the temperature control
requirements in the gasifier.
The data also indicate a mean molecular weight of the gasifier
product of 29, and a mean heat capacity of 0..35 Btu/l6°F, both nearly
independent of the air/fuel ratio.
The product analysis shown in Table 2-4 may be a reasonable
analysis to expect from a continuous gasifier-desulfurizer operating
with a similar air/fuel ratio and sulfur removal, and with no special
provisions for temperature control. At other air/fuel ratios, with
higher sulfur removals, and with temperature control of the gasifier by
the addition of some component such as stack gas or steam, the results
,
in Table 2-4 are not of quantitative importance. However, the design
work can be carried out without complete knowledge of the product
composition, since Figure 2-3 provides sufficient information on which
to base the preliminary design.
442
-------
:J: 3800
-
Q)
::J
.....
..0
-=:::::.
~ 3600
:::>
I-
ca
+0-
+0-
W
c::
o
~ 3400
~
Q)
c::
Q)
<..::>
.-.
ra
Q)
:J: 3200
4000
3000
Curve 643988-B
/
/
-0
~~
.~
~~
-~~/
-~~
~'l"
'?
~oo/"
~" -y
r
/
o
o
19
20
21
22 23 24 25
Ai r/Fuel Ratio, %of Stoichiometric Air
26
27
28
Fig. 2-3- Gasification -Desulfurization Heat Generation Rate
-------
Desulfurization
,
H2S evolved during fuel cracking reacts with the CaO in the
bed by the reaction
+
CaO + H2S + CaS + H20.
(2-3)
For both fresh bed tests and cyclic tests, the extent to which this
reaction proceeds depends on the gasifier temperature, the air/fuel
ratio, the gas residence time, the lime particle sizes, the bed lime
content relative to the rate of sulfur fed to the bed, the amount of
lime utilized, the limestone make-up rate, and the quality ,of fluidization.
Regenerated lime loses activity compared to that of fresh lime. This
may be due to a continuous deterioration in the pore structure of the
lime particles brought about by the high temperatures in regeneration.
Sulfate formation taking place in the regenerator may become a barrier
to further sulfur removal; and carbon deposited on the lime is also
believed to be detrimental to sulfur removal.
The next matrix summarizes the experimental results for sulfur
removal found in atmospheric pressure fresh bed and cyclic batch tests.
For fresh bed tests operated at a temperature and air/fuel
ratio at which carbon deposition does not reduce sulfur removal
(temperature greater than 830°C and air/fuel ratio greater than 20% of
stoichiometric), the important variables are the gas residence time,
bed CaO content per pound of sulfur fed per hour to the bed, the average
particle diameter, and the CaO utilization (the fraction of calcium
oxide converted to calcium sulfide, or sulfate).
Figure 2-4 illustrates the results of three fresh bed tests
under identical conditions of gas residence time .(~ 0.22 sec), bedCaO
content per pound of sulfur fed per hour (~ 24 lb CaO/lb S/hr), bed
444
-------
VARIABLE
Gasifier
Temperature
Air/Fuel
Gas Residence
Time
+:-
+:-
Vl
Average Particle
Diameter
Bed CaO Content
Sulfur Feed Rate
Limestone Makeup
Rate
Quali ty of
Fluidization
SULFUR REMOVAL - EXPERIMENTAL OBSERVATIONS
RANGE
790-950°C
15-30% or
Stoichiometric
Bed depth
10-20 inches
Velocity
3-8 ft/sec
4.96}.1,
682}.1,
l700}.1
Up to 1 mole CaO
per mole S fed
COMMENTS
For best sulfur removal 830-880°C
< 830°C carbon deposition increases
> 880°C release of 802 in bed increases
Temperature not independent of air/fuel
in batch tests
For best sulfur removal found for> 20%
Air/fuel controls temperature and carbon depo-
sition rate in batch tests carried out
rated
Increased residence time increases sulfur
removal
Behavior dependent on particle size
See Figure 2-5
Finer particles give greater sulfur removal
. for a given lime utilization, bed depth,
etc.
See Figures 2-4 and 2-5
Increase gives greater sulfur removal
Not varied independent of air/fuel ratio,
velocity, bed height, etc., in batch tests
See rigure 2-5
Nearly 100% sulfur removal obtained in cyclic
tests with 1 mole CaO/mo1e S fed make-up
rate
See Figure 2-6
No observable effects due to increased pressure
. drop across distributor of higher velocities
Batch tests are in a state of vigorous bubbling
Good mixing must be maintained to prevent car-
bon bUild-up on bed line
-------
100
V)
.W
..
+:-
+:-
~
~ 60
E
Q).
a::
~
:J 40
-
:J
V')
~
Curve 643911-A
80
.........-.
1 ............ .
"
""4
20
- --
...........
o "'"""
0'0,
~
"-
A dp = 496 ~ "
. dp = 682 ~ " "
o dp = 1700 ~ 3
Bas is: dp = 1700 ~ . . 2 ..
For small u: ulul = (dPl/dPi at Equal Es
curves 1 and 2 .
3/2
. For large u: uiul~ (dPl/dPi at EQualEs
curves 3 and 4 .
2
10
o
10
20 30
% CaG Utilization, u
40
Fig. 2-4- Effect of particle size on sulfur removal
-------
temperatures (~ 860°C), and air/fuel ratios (~ 24%). Each run represents
a different average particle diameter, d , and shows the sulfur removal
p
efficiency, £ , plotted against the CaO utilization,~. The finer lime
s
particles give greater sulfur removal for a given CaO utilization
because of easier accessibility to the finer particle surfaces. The
figure indicates that for small utilizations -- small enough that external
particle surface is dominant in the reaction -- the utilization is
inversely proportional to the average particle diameter.
the performance of two particle sizes, 1 and 2,
Or, comparing
U2/Ul = (d /d )2
PI P2
at equal £ .
S
(2-4)
For larger lime utilizations, where the internal particle surface enters
the reaction, the proportionality is observed to be the diameter to the
-3/2 power, or
U2/Ul = (d /d )3/2 at equal £ .
PI P2 s
(2-5)
Equations (2-4) and (2-5) are demonstrated in Figure 2-4~ with the
experimental curve for d = 1700 ~ taken as the basis of comparison.
p .
Note that the point of transition between equations (2-4) and (2~5) is
also a function of the average particle diameter.
The behavior of fresh bed tests in the temperature range of
. r--' ',"
830°C to 880°C may be summarized in terms of the operating variables 'G'
the gas residence time, and S, the lb CaO per lb of sulfur fed/hr. This
is shown in Figure 2-5 with the sulfur removal efficiency, £ , plotted
. s
against the product S 'G for average particle diameters 496, ,682, and
1700 ~ and a calcium oxide utilization equivalent to 5 wt % sulfur on
the bed lime. Curves for 4 and 6 wt % sulfur on the bed lime are
included for the case of d = l700~. The behavior of Figure 2-5 is
p
similar to that of Figure 2-4. The curves for diameters of 496 ~ and
682 ~ and 5 wt % sulfur on the stone are in the region where external
surface dominates in the reaction, and the data appear to be consistent
with the relationship
447
-------
100
90
80
~
..
+:""
+:""
CP
cu
~ 70
E
Q,)
0::
~
.260
::]
V)
50
40
"".
.,
/
I
I
I
I
I
I
I
I
I
I
I
I
. I
o
Curve 643910-A
2
. .
. ".. -- -4% - -- - -5%
.",. -- --
if' ",""" --- -- (ffo
/ / ~
/'
I jJ " .
./ I /
/ / /
/ I f
I I /
I I I
I /
o I /
I
I ' I
Il'
P I I
I 01 0'
I I ,
4
18
5 wt % Sulfur on CaO
Temperature =---- 850°C
6 dp =4961J
. dp = 682 IJ
o dp = 1700 IJ (Includes curves for
4 and 6 % Sulfur)
6 8 . . 10 12 14
S TG (Ib CaO/lb Sulfur/Hr) (Sec) .
16
Fig. 2-5-5ulfur removal correlation
-------
(S 'G) 1
.(S 'G)2
d 2
PI
= (-)
d
P2
at equal sulfur removal.
(2-6)
The curve identified with the particles of 1700 ~ average diameter is
in the region of internal diffusion, as is seen in Figure 2-4.
I t may
be shown that, by looking at the 682 ~ particles at higher utilizations,
3/2
S 'G a dp
at a given sulfur removal,
( 2-7)
in the region of internal surface reaction. Figure 2-5 permits the
prediction of ~he fresh bed operating conditions, bed depth, and velocity
necessary for a given sulfur removal for a given particle size and lime
utilization.
Cyclic tests differ from fresh bed tests in that the limestone
make-up rate and the conditions of the stone regeneration cycle area of
major importance to the sulfur removal efficiency. Except for these
two factors, the behavior of cycled beds should be similar to the
behavior found in fresh bed tests and summarized in Figures 2-4 and 2-5.
Results of three cyclic tests are shown in Figure 2-6 as the sulfur
removal versus the cycle number. The experimental conditions of the
tests and the duration of the gasification and regeneration cycles are
indicated on the figure. Esso (England) has observed in similar tests
that the sulfur removal efficiency reaches a steady-state value of
nearly 100% at a make-up rate of.l mole of CaO per mole of sulfur fed
with the U.K. limestone and the U.S. limestone 1691.
--_._-- .--
-- - .-.. . .--- - - "_. .---. .
-. - .. - ._,- -. --.- . .
_..- ---- - ~...u. ._h
The batch, nonsteady, tests carried out by Esso (England)
. . .'
do not apply directly to the steady operation of a commercial size
continuous unit.
Fortunately, the tests indicate that extremely fast'
reactions are involved in the desulfurization phenomena, so that scale-
up to cases with much different fluidization characteristics should be
possible. For the design of a once-through operation, the fresh bed
test results may be applied directly to the design, with Figure 2~5
providing a great deal of design information.
The regenerative operation
449
-------
~80
~
Vl
a
,
ro
e;
E
C1>
Q:: 70
~
::J
-
::J
V"I
Curve 643912-8
100
90
o
0
. 0
6 .
. .
.
A 0 A 0
! A 0
. A
A
A Q 0 0
. 8
. . e .
A
A . . 0
o . 0 .
.
A
6
o UK Stone, dp = 850-1400~, 4 fVsec
00 24% Air/Fuel, Make-Up = 0.74 wt. CaO/Wt. S
-. US Stone 1691, 000-1400~, 4 fVsec
24% Air/Fuel, Make-Up = 0.75 Wt. CaO/Wt. S
50 AUS Stone 1690, OOO-I400~, 4 fVsec
240/0 Air/Fuel, Make-Up = 1. 00 wt. CaO/Wt. S
Each Cycle Consists of
I Hr Gasification
< 1/2 Hr Regeneration
After Gasification Bed Contains
3-5 wt. % Sulfur
I
3
7
5
9 II
Cycle Number
13
15
17
19
Fig. 2-6- Batch Cyclic Bed Tests
-------
design may be somewhat less directly related to the cyclic bed tests,
but the make-up rates necessary for a given sulfur removal should be
valid for the continuous regenerative case, even though the regenerator
performance in the batch and continuous units may differ somewhat.
Temperature levels found optimum for desulfurization in batch
tests should apply to continuous operation, but the effects of the
air/fuel ratio on sulfur removal observed in batch tests may differ
greatly in a continuous unit in which temperature control is maintained.
In addition, the gasifier temperature control scheme used in a continuous.
unit may itself reduce or promote sulfur removal. For example, steam or
water injection in the gasifier may greatly affect the kinetics of the
reaction in equation (2-3) in such a way as to reduce sulfur removal, but
carbon deposition may be simultaneously reduced by the reaction
+
C + H20 + CO + H2
(2-8) .
to increase the sulfur removal.
.Factors such as these call for further experimental investigation,
but sufficient information is available from the Esso desulfurization
batch data to permit preliminary design.
451
-------
Regeneration
The regenerator is operated to give a slight oxygen break-
through in the S02-rich product stream, thus assuring predominantly an
oxidizing atmosphere in the bed. Besides converting sulfided lime back
to CaO, a portion of the calcium sulfide is converted into calcium
sulfate, and the carbon deposited on the lime is combusted.
dominant reactions are
Thus, the
+
CaS + 3/2 02 + CaO + S02
(2-9)
+
CaS + 2 02 + caS04
(2-10)
.+
C + 02 + C02'
(2~11)
The main factors influencing the observed loss in lime activity
may be taking place in the regeneration step. These factors are probably
sintering of the stone (loss of porosity), due to the high generator
temperatures, and sulfate build-up in the stone.
The mechanisms involved in batch regeneration differ somewhat
from those expected in a continuous regeneration operation mainly
because of differences in the temperature history of the stone in the
two cases. In the batch case the whole bed heats up relatively slowly,
with different reactions prevalent at different stages in the regeneration.
The continuous, steady-state operation will lead to a quick heat-up of
the sulfided lime particles entering the hot regenerator.
From the standpoint of regeneration of the sulfided lime and
the maintenance of stone activity, about 1050°C has been found to be a
suitable regenerator temperature. 802 concentrations of about 10 mole %
have been obtained from batch regeneration runs. The S02 levels
452
-------
expected in a continuous operation will depend strongly on the rate of
carbon deposition and the amount of sulfate production in the regenerator.
With sulfided lime containing 3% to 5 wt % sulfur, Esso (England) has
nonlally produced a final regenerated lime containing 1 to 2 wt % sulfur
in the form of calcium sulfate.
Lime residence times in the regenerator
have varied from five to fifteen minutes.
Esso (England) has recorded
CO, C02' 502' 02 concentrations from the regenerator and the regenerator
temperature as a function of time.
The performance of a cqntinuous regenerator is expected to be
superior to that found in batch tests in terms of the amount of sulfate
production in the regenerator.
The operation of the Esso (England)
continuous unit will supply additional information on regeneration which
will be needed for more detailed design work.
453
-------
Particle Carry-Over
Bed losses during the Esso (England) batch tests were caused
by attrition of the particles in the calcination, gasification, and
regeneration operations.
The factors controlling bed losses are:
.
Superficial velocity
.
Quality of fluidization
.
Limestone physical characteristics
.
Air distributor design
o
Gasifier vessel design - bed depth, freeboard height,
internals, etc.
.
Bed operation - recirculation of fines, turn-down.
Lime particle size distributions were. measured as a function
of time during the calcination, gasification, and regeneration stages of
operation.
Average rates of bed losses varied from 1 to 5 wt % of the
bed/hr for U.K. and U.S. 1691 limestones at velocities of from 3 to 8
ft/sec, though the rate of.elutriation varied greatly with time.
Since. the factors listed above probably differ between the
Esso (England) batch operations and a commercial power plant design, the
bed losses recorded by Esso are difficult to relate to design consider-
ations.
Again, the operation of the Esso 750 kw pilot unit will provide
useful information in this area.
454
-------
Metal Retention
Batch tests have indicated that the bed lime removes nearly
100% of the fuel vanadium and about 20% of the fuel sodium during
gasification. Tests have also indicated that this build-up of vanadium
and sodium on the lime has no effect on sulfur removal, though this is
not conclusive. This high metal retention behavior of the atmospheric
pressure gasification/desulfurization operation should also be observed
in continuous operation and is an important factor in material corrosion.
Burner Design
Esso (England) has constructed and operated a burner utilizing
the clean, hot, fuel gas from batch gasification/desulfurization
operations. A luminous flame, free of smoke, has been obtained with
proper burner operation, indicating that the low Btu gas is easily
combustible.
Preliminary Results of Continuous Operation
Operation of the Esso (England) 750 kw, regenerative, gasifi-
cation/desulfurization pilot unit is now in the preliminary st~ges
leading to the completion of the 200 hr commissioning run. Preliminary
indications are that the continuous unit behaves much as expected
from batch experience, with about 90% sulfur removal obtained at the
expected limestone make-up rate of 1 mole CaO per mole of sulfur, fed.
Also, apparently because of the maintenance of the gasifier temperature,
the continuous unit has been operated at much lower air/fuel ratios
than were possible in the batch studies (presently down to 17% of
stoichiometric). Overall behavior of the continuous unit is promising.
455
-------
MATERIAL AND ENERGY BALANCES
Batch information provided by Esso (England) is taken as a
basis for the preliminary design of the atmospheric pressure gasification/
desulfurization concept.
Both regenerative and once-through modes of
operation are examined, and the temperature control schemes proposed
for the systems are of prime concern. Overall material balances are
computed, and the power requirements for the two modes are estimated.
The preliminary design is limited to a specific set of operating
conditions. Further data, however, will permit a more complete
examination of the relative merits of the two modes of operation and
will make possible the determination of optimum operating conditions.
456
-------
Modes of Operation
Flow diagrams for both regenerative and once-through modes of'
operation were shown in Figure 2-1. In the regenerative operation, fuel,
air, and fresh make-up limestone are introduced into the gasifier vessel
to produce a hot, low-sulfur, fuel gas. The gasifier product is sent
directly to the boiler, or possibly to cyclone clean-up before being
transported to the boiler. The hot sulfided lime produced in the
gasifier is sent to the regenerator, converted to a less active form of
lime, and returned to the gasifier. The 802 released from the
regenerator is transported to a sulfur recovery system, possibly
yielding a useful by-product to reduce the system operating .cost. In
the regenerative system, the control of the gasifier and regenerator
temperatures is interdependent with the regenerator temperature control
scheme determining the temperature control requirements in the gasifier;
and the carbon deposition rate is an important factor in the overall
operation of the system.
The once-through mode of operation differs from the regenerative
operation in that the lime is utilized to a greater extent in the
gasifier and is disposed of after being converted to the sulfate form
in the sulfate generator.
Temperature control of the gasifier and
sulfate generator is an important consideratiori, as in the regenerative
mode, but the carbon deposition rate is less important to the overall
once-through system operation..
457
-------
Design Parameters
Design parameters of importance in oil gasification/desulfuri-
zation are listed in the following matrix. The four factors listed as
process specifications (geographic location, fuel oil, limestone, and
sulfur removal) are determined by the market outlook, by residual oil
and limestone availability, and by the expected 802 emission standards,
respectively.
The design variables listed refer to the broad design factors
which will eventually determine the feasibility of preliminary design
concepts. The design considered may be that of a new boiler or a
boiler retrofit, while three factors -- boiler size, boiler load factor,
and boiler turn-down ratio -- are important in determining the design
for either a new system or a retrofit system. In the retrofit case the
type of boiler being retrofit is important because coal- and oil-fired
boilers differ in retrofit characteristics. The location of the
gasification system relative to the boiler structure (internal or
external to the boiler) and the configuration of the gasification system
(singular or modular) are basic considerations in both new and retrofit
designs. The other important design variable considered in the preliminary
design is the mode of operation (regenerative or once-through).
The operating variables listed are those found to be important
in the Esso (England) batch studies.
In addition, the schemes proposed
to control the temperatures of the gasifier, regenerator, and sulfate
generator are listed. Possible methods of controlling gasifier
temperature are the recycle of stack gas (presently being used with the
Esso continuous pilot unit), water or steam injection, the use of heat
transfer surface in the gasifier to preheat boiler water or, possibly,
reducing the air/fuel ratio to the point of thermal balance. In the
regenerative operation, the regenerator temperature might be controlled
. 458
-------
PROCESS
SPECIFICATIONS
DESIGN
VARIABLES
OPERATING
VARIABLES
TEMPERATURE
CONTROL
DESIGN PARAMETERS
Geographic location
Fuel oil
'Limestone
Sulfur r,cmoval
New boiler design
Retrofit design - (coal- or
oil-fired)
Gasification system location
Internal to boiler
External to boiler
Gasification system operational mode
Regenerative or once-through
Gasification system configuration
Unit design or modular, etc.
Fluidization velocity
Particle sizes '
Air/Fuel ratio
Limestone make-up rate
Limestone utilization
Bed depth
Bed temperature - gasifier and
regenerator
Gasifier temperature control
Stack gas recycle
Water injection
Steam injection
Heat 'transfer surface
Air/fuel ratio
Regenerator temperature control
Stone circulation rat~
Addition of make-up stone
Heat transfer surface
Sulfate g=nerator temperature control
Excess air circulation
Heat transfer surface
459
Boiler
Size
Load
Factor
Turn-Down
Ratio
-------
by regulating the lime circulation rate (as in the Esso continuous
pilot unit), by adding fresh make-up limestone to the regenerator, or
by including heat transfer surface in the regenerator vessel. Methods
proposed for control of the sulfate generator in the once-through
operation are the circulation of excess air through the sulfate generator
or the inclusion of heat transfer surface in the vessel.
The specifications assumed for the conceptual design ate
presented in the next matrix. The process specifications include a
geographic location for the commercial development of the atmospheric
pressure residual oil gasification/desu1furization process in the
Eastern United States according to market predictions. Preliminary
designs are based on a residual oil containing 3 wt% sulfur with low
heating value of 17,700 Btu/lb, a limestone having the composition of
the U.K. limestone investigated by Esso (England) (Table 2-2), and a
sulfur removal efficiency of 90% to 95% to meet present and future
pollution standards.
For cost estimates, a boiler capacity of 600 MW and load
factors of 40% and 80% were considered. In general, capacities of 100 MW
and greater are of interest in the preliminary designs. A turn-down
ratio of 4/1 was of interest in both new and retrofit designs.
The operating variables are specified according to the Esso
(England) batch data. These data indicate that the gasifier temperature
of 1600°F is optimum for both regenerative and once-through operations
from the standpoint of sulfur removal.
From the standpoint of lime
regeneration, the data indicate 1900°F for the regenerator temperature.
A temperature of l500°F is chosen for the sulfate generator, based on
chemical equilibrium to give an S02 partial pressure of less than 0.01 atm
coming off the generator, although the kinetics of the reaction have not
been investigated. Based on the Esso. (England) cyclic batch tests a
limestone make-up rate of 1 mole CaO per mole of sulfur fed is specified
for 90% to 95% sulfur removal in the regenerative operation, with 5 wt %
sulfur on the gasifier lime, an average stone diameter of 2000 ~, a
460
-------
PROCESS SPECIFICATIONS
Fuel oil:
Geographic location:
Eastern United States
Limestone:
3 wt. % sulfur; LHV = 17,700 Btu/1b.
U.K. limestone
Sulfur removal:
90-95%
Gasifier temperature
Regenerator (sulfate
generator) temperature
Stone make-up rate
Air/fuel ratio
Limestone utilization
Fluidization v~locity
Particle sizes (avg.)
Gasifier bed depth
(static)
DESIGN VARIABLES
Boiler size:
Load factor:
600 MW
40%, 80%
ONCE-THROUGH
1600°F
1500°F
3 moles CaO/mole sulfur
20% of stoichiometric 20% of stoichiometric
~ 5 wt % sulfur in bed ~ 19 wt % sulfur .in bed
PROJECTED PERFORMANCE
REGENERATIVE
90-95%
90-98
~ 2-5%
~ 100%
of bed/hr
Turn-down:
4/1
8 ft /sec
~ 1000 .~
3.5-4.0 ft
Sulfur r.emova1
Thermal efficiency
E1utriation
Vanadium retention
Sodium retention
OPERATING VARIABLES
REGENERATIVE
1600°F
1900°:F
.1 mole CaO/mo1e sulfur
8 ft /sec
~ 2000 ~
2.5-3.5 ft
~ 20%
[
ONCE-THROUGH
90-95%
90-98
~ 2-5% of bed/hr
~ 100%
~ 20%
461
-------
velocity of 8 ft/sec in the gasifier, and a gasifier bed depth of
3.5 to 4.0 ft. An air/fuel ratio of 20% of stoichiometric is specified
from the Esso data in order to assure reasonable carbon deposition rates
in the regenerative and once-through operations.
The projected performance for regenerative and once-through
operations is identical and is based on the batch experimental results
provided by Esso (England).
462
-------
Temperature Control
Energy and material balances are carried out to determine the
requirements and feasibility of the methods proposed for controlling the
temperatures of the regenerative and once-through systems.
Factors such
as the system thermal efficiency, gasifier and regenerator product
characteristics, apparatus dimensions, and temperature control require-
ments are investigated with respect to the proposed temperature control
methods as a function of the carbon deposition rate.
The advantages and
disadvantages of the proposed temperature control schemes. are compared.
Regenerative Operation
Control of the gasifier temperature at l600°F and the
regenerator temperature at 1900°F is interdependent,with the temperature
control scheme used in the regenerator affecting the temperature
control.requirements in the gasifier.
The performance of the regenerative
operation is highly dependent on the temperature control schemes
applied to the system and the rate of carbon deposition occurring in
the gasifier.
Regenerator
The methods proposed for controlling the regenerator temperature
at 1900°F are regulation of the rate of lime circulation between the
gasifier and regenerator, addition of fresh make-up limestone to the
regenerator, or inclusion of heat transfer surface in the regenerator
vessel to preheat boiler water.
Heat is liberated in the regenerator by the three reactions
-+
CaS + 3/2 02 + CaO + S02' ~Hl
(2-9)
-+
CaS + 2 02 + CaS04' ~H2
(2-10)
-+
C + 02 + C02' ~H3'
(2-11)
463
-------
where ~Hl' ~H2' and ~H3 are the heats of reactions of equations (2~9),
(2~lO), and (2-11), respectively, in Btu/lb-mole. The following
assumptions are made in evaluating the regenerator temperature control
requirements:
.
Per pound of fuel entering the gasifier, YE lbs of
s
sulfur enter the regenerator (where Y is the wt fraction
of sulfur iri the fuel and E: is the gasifier sulfur
s
removal efficiency), and 90% of this entering sulfur is
converted to S02 by reaction (2~9) while the remaining
10% is converted to calcium sulfate by reaction (2-10).
.
Deposited carbon enters the regenerator at a rate of XcNG
(lb carbon/lb fuel) where X is the weight fraction of
c
carbon of the gasifier lime and NG is the rate of lime
circulation to the regenerator (lb lime/hr per lb of
fuel/hr).
All of the deposited carbon is combusted by
reaction (2-11).
.
Fresh make-up limestone is added to the regenerator at
a rate of NMR (lb limestone/hr per lb fuel/hr) and is
calcined completely by the reaction
-+
CaC03 + CaO + C02 ' ~H4
(2-12)
having a heat of reaction ~H4.
.
Sulfided lime enters the regenerator with a temperature of
l600°F and is withdrawn at a temperature of 1900°F, the
lime having a heat capacity, Cp , of 0.3 Btu/lb of. Air
s .
and fresh limestone enter the regenerator at 77°F and
exit at 1900°F. A stoichiometric amount of air is fed
to the regenerator.
.
Heat losses to the surroundings are neglected.
464
-------
An energy balance written around the regenerator indicates
that for adiabatic operation
(Y) (£s)
0= ~Hl (.9) 32
(Y) (£8)
+ ~H2 (.1) 32
X N
+ ~H ...s....Q
3 . 12
NMR
+ ~H4 (.434) ~ + NG Cp
s
(300) + NMR Cp
s
(.566) (1823)
+ FN Cp (1823) + FCO C (1823) + FSO C (1823)
2 N2 2 PC02 2 PS02
(2-13)
where FNZ' FCOZ' and FSOZ are the molar flow rates of nitrogen, carbon
dioxide, and sulfur dioxide per lb of fuel, respectively, and Cp .' Cp .
. . . N2' C02 '
and CPS02 are the heat capacities of N2' C02' and S02' respectively,
FN ' FCO ' and FSOZ are given by
2 Z
X N Y£ Ye:
F = -£....Q. + 1. 5 ( . 9) ---!!. + O. 1 ---!!. ,
N 2 12 3Z 32
.434 N X N
MR+~
44 12
(2-14)
FCO =
2
(2-15)
Ye:s
FSO = (.9) ~ .
2
(2-16 )
Placing the proper values in equation (2-13) and solving for NG' the lime
circulation rate per lb of fuel/hr, gives
NG = 88 XcNG + 41.3 Y£~ - 14.5 NMR'
(2':'17)
Equation (Z-l~ relates the flow rate of sulfided lime necessary for tem-
perature control of the regenerator to the carbon deposition rate, XcNG'
the rate of sulfur removal per Ib of fuel/hr, Ye: , and the rate of
s
addition of make-up stone to the regenerator per lb of fuel/hr, NMR'
465
-------
(2~17) is illustrated in Figure 2-7.
On the basis of 95% sulfur removal, or E = 0.95, equation
s
Equation (2-17) becomes
NG = 88 XcNG + 39.3 (Y - 0.369 NMR)
(2-18)
and Figure 2-7 is a plot of NG versus XcNG with Y-0.369 NMR as a
parameter. Curves of constant X are also shown on Figure 2-7 (as
c
broken lines) to reflect the behavior expressed by equation (2-2).
That is, Xc is a constant independent of NG for NG «K2' Two schemes
of regenerator temperature control are incorporated in the figure:
control by sulfided lime circulation (represented by NMR = 0), and
control by addition of fresh make-up limestone to the regenerator
(represented by NG = N~IN' the minimum circulation rate necessary to.
maintain the maximum allowable gasifier lime utilization). The minimum
circulation rate is given by
YE
N '" s
GMIN XSG - XSR
(2,-19)
.where XSG is the maximum weight fraction of sulfur on the gasifier lime
.in ord~r to achieve a sulfur removal of ES' and XSR is the weight fraction
of sulfur on the regenerator lime, in the form of calcium sulfate.
. For Y = 0.03, E
S
data, and XSR is
lb lime/lb fuel.
= 0.95, XSG must be 0.05 according to the Esso (England)
assumed to be about 0.005, giving a N of 0.633
GMIN
Figure 2-7 indicates that the control value of NG can easily
rise to three or four times the minimum value, NGMIN' while relatively
small amounts of fresh make-up limestone can be used to accomplish the
temperature control (NMR = 0.05 lb/lb fuel with Xc = 0.003, Y = 0.03,
and NG = NGMIN)' In fact, equation (2-17) indicates .that the critical
value of X above which temperature control of the regenerator can no
c
longer be accomplished with NMR = 0 is 0.0114. That is, for X > 0.0114.,
c.
an increase in NG will have no effect on the net rate at which heat is
generated in the regenerator. This conclusion holds only if the air is
466
-------
(.,:)
z
2- 2.0
'"
a:::
c:
o
......
ro
~
u
~
u
Curve 643987-B
3.0
Temperature = 1900° F
Sulfur Removal = 95%
~\).
Xc = O. 002 0 003 ~~.
I .j O. 004 ~~~
/ ,,~
I / / ~,~.
I /
I I
I
1.0
o
0.01 0.02
Net Carbon Deposition Rate, XCNG
Fig. 2-7 -Regenerator Temperature Control
467
-------
kept at the stoichiometric feed rate to. give complete combustion of the
deposited carbon at all values of X. It is necessary to operate in
c
this manner, though, to ensure regeneration of the lime.
The third temperature control scheme, the use of heat transfer
surface in the regenerator with NG = N~IN' is limited by the wall-
surface to volume ratio in the regenerator. Based on the air rate to
the regenerator at'1900°F, a superficial velocity of 8 tt/sec, 95%
sulfur removal, a gasifier thermal efficiency of 95%, and a plant heat
rate of 104 Btukwh, the cross-sectional area of the regenerator per MW
is shown in Figure 2-8 versus the carbon deposition rate, with the weight
% of sulfur in the fuel as a parameter.
The regenerator cross-sectional
area is strongly dependent on the carbon deposition rate, and assuming
an expanded bed height of about three feet, the possibility of using
a water-jacketed regenerator for temperature control appears limited
to small units. A 20 MW unit would require a reasonable heat transfer
coefficient of about 25 Btu/hr-ftZ-OF for temperature control based
on the wall-surface area of the vessel, while a 200 MW unit would
require a heat transfer coefficient of about 90 Btu/hr~ftZ-oF. The
additional surface available with a water-cooled air distributor in
the regenerator does not appear sufficient to give reasonable heat
. transfer coefficients with capacities greater than 200 MW. Thus, the
inclusion of heat transfer surface internal to the fluid bed, or the
restriction to modular regenerator designs, must be considered in order
to use this temperature control scheme.
Another important factor in regenerator temperature control is
the concentration of SOZ leaving the regenerator. As the carbon
deposition rate, XcNG' increases, the molar concentration of S02
leaving the regenerator will decrease, making sulfur recovery more
expensive.
From equations (2-14) through (2-16) the S02 concentration
is given by
YE
s
XS02 = 12.7 XcNG + 6.64 YES + 0.316 NMR
(2-20)
468
-------
0.8
0.7 Fluidization Velocity = 8 ft/sec
Sulfur Removal = 95%
s Gasifier Thermal Efficiency = 95%
N~
0.6
t-
L.L..
I
S
~ 0.5
~
Q)
a..
c:
0
:.;::
u 0.4
Q)
V')
I
In
.+=- In
0\ 0
\0 ~
(J 0.3
~
.9
~
~
Q)
c:
Q)
C'I
Q,)
~
0.2
0.1
o
0.01
0.02
Net Carbon Deposition Rate
0.03
Fig. 2-8-Regenerator Design
n
t:
-.
<
(1)
(7'\
~
w
U)
co
U)
I
~
-------
where XS02 is the mole fraction of S02 in the regenerator product.
Equation (2-20) is shown in Figure 2-9. Since the temperature control
scheme "NMR = 0" operates with a high carbon deposition rate, the S02
concentration will be reduced below that of the other two schemes.
For instance, if X = 0.004 and Y = 0.03, then Figure 2-7 indicates that
c ,
for temperature control with NMR = 0 the carbon deposition rate will be
almost 0.01 lb carbon/lb fuel, while with NG = N~IN the deposition rate
will be 0.0025 lb carbon/lb fuel. Then from Figure 2-9 the mole
fraction of S02 from the regenerator will be 0.090 for NMR = 0, 0.130
-- .
for NG = NGMIN' and 0.130 for control with heat transfer surface.
The advantages and disadvantages of the three regenerator
temperature control schemes are compared in the next matrix. The points
summarized there indicate that:
.
The use of stone circulation for regenerator temperature
control is attractive only if the amount of carbon deposited
on the gasifier lime does not exceed the value of X '" 0.004
c
and if the carbon deposition phenomenon is well-behaved
and predictable (that is, the carbon deposition must not
fluctuate during operation and must be a predictable
quantity for design purposes).
.
The other two temperature control schemes become more
'" '
attractive than stone circulation for X > 0.004.
c
.
The use of make-up limestone in the regenerator yields a
high S02 concentration in the regenerator product and
greatly reduces the stone circulation rate from the value
which would be needed with the stone circulation temperature
control method.
If the carbon deposition phenomenon is
not well-behaved~ then the specified make-up rate of
1 mole CaO/mole sulfur fed must be sent totally to the
gasifier to achieve the prescribed sulfur removal and
minimize the control problem involved in solids flow
while the make-up necessary for regenerator temperature
470
-------
0.16
0.14
L.
o
-
co
L.
~ 0.12
Q)
g
a:::: .
E
e O. 10
-
N
o
V1
'0 0.08
c:
o
U
co
L.
~ 0.06
(5
=E
0.04
0.02
Curvf' ~46232-A
Sulfur Removal = 95%
NMR = 0
---- N =N
G GMIN
o
0.01 0.02
Net Carbon Deposition Rate, lb. Carbon/lb. Fuel
0.03
Fig. 2-9-Regenerator Operation
471
-------
CONTROL SCHEME
Stone circulation
(NMR = 0)
+:-
-.1
.1\)
Make-Up $ tone
(NG = NG )
MIN
Heat transfer surface
(H.T.S.) .
COMPARISON OF REGENERATOR TEMPERATURE CONTROL SCHEMES
ADVANTAGES
DISADVANTAGES
Circulation of stone is already a
necessary element of operation
A simple scheme if the carbon depo-
sition behavior is predictable
Best of the three schemes for
X < .0.002
c
Minimizes stone circulation and
carbon deposition rate (NGMIN =
0.633 with Y = 0.03 and
ES = 0.95) .
Higher S02 concentration than with
. stone circulation scheme. (XS02
!:O 0.12)
Minimizes stone circulation and
carbon deposition rate
(NG =NGmN) .
Highest S02concentration and
higpest thermal efficiency of the
three schemes. (XS02!:O 0.12,
thermal efficiency !:O 0.97% depend-
irig on gasifier control scheme)
May necessitate extremely large
circulation rates for which capi-
tal and operating expenses are
high. (At Xc ~ 0.005, Y = 0.03,
and ES = 0.95, NG = 2 lb stone/
lb fuel)
Gives higher carbon carry-over to
the regenerator than other schemes,
decreasing the S02 concentration
from the regenerator (a probable
decrease in XS02 of about 25% as
compared to other schemes)
Makes gasifier and regenerator
performance interdependent
Calls for increase in the amount of
make-up stone fed to system (make-
up rate to gasifier !:O 0.09 lb/lb
fuel, while NMR !:O 0.05 lb/lb fuel)
Slight decrease in thermal effici-
ency compared to other schemes
(~ 1% less efficient)
Control is complex
High regenerator capital cost
May be feasible only with modular
regenerator design of H.T.S. inter-
nal to the bed in systems> 100 MW
-------
control is added to the regenerator.
This requires a
greater consumption of limestone and, thus, a higher
operating cost. If the carbon deposition is a well-behaved
phenomenon then the total make-up rate of 1 mole CaOper
mole of sulfur fed may be distributed between the gasifier
and regenerator to maintain regenerator temperature control
and remove the specified amount of fuel sulfur.
Thus, the
rate of limestone consumption might be minimized in this
case.
This indicates the need for further experimental
investigation of carbon deposition during continuous
operation.
.
The use of heat transfer surface maximizes the 802 .
concentration relative to the other two control methods,
but it may be associated with high capital costs and may
be limited in feasibility.
Gasifier-Regenerative Operation
The gasifier temperature may be maintained at1600°F by one
of five methods:
stack gas recycle, steam injection, water injection,
heat transfer surface, or operating at the minimum air/fuel ratio.
The first three depend on the addition of a heat storage component,
stack gas; steam, or water, to the gasifier to carry the excess heat
generated in the gasifier as sensible heat. In this way no heat is
lost, but the system equipment is enlarged. A penalty must be paid in
the steam and water injection cases, since the heat of vaporization of
water is nonrecoverable. Also, stack gas, .steam, and water are not
inert to the gasification/desulfurization phenomena. In particular,
steam and water injection may lead to a reduction in the rate of carbon
deposition by the reaction
+
C + H20 + CO + H2'
(2-21)
473
-------
On the other hand, the presence of large excesses of H20 in the gasifie~ .
may lead to a reduction in sulfur removal by reaction (2-3). The sudden
evaporation of injected water droplets may also contribute to the
. distribution of injected residual oil in the fluidized bed upon hitting
the hot gasifier lime. Further experimental ,studies must be carried out
to investigate the effects of stack gas, steam, and water on the gasifica-
tion/desulfurization operation. For the purposes of preliminary
design it is assumed that ,these temperature control.components behave
as inerts.
The inclusion of heat transfer surface in the gasifier will
allow the system to operate with high therm~l efficiency if boiler water
is preheated in the gasifier, and will lead to a reduction in equipment
sizes over those of the first three proposed methods.
The cost of the
gasifier will increase greatly, and a detailed cost analysis of the
trade-off between heat transfer surface and stack gas recycle, steam
injection, or water injection must be performed to evaluate the relative
merits of these proposed temperature control methods.
If the gasification/desulfurization operation can be carried
out at a low enough air/fuel ratio (about 14% of the stoichiometric
air rate) then the gasifier may be thermally balanced at l600°F so that
'further temperature control consideration is unnecessary.
Operating at
this air/fuel ratio would minimize equipment sizes, control problems,
and overall system complexity. It has not been demonstrated that
operation with the minimum air/fuel ratio can be realized, and at such
a low air/fuel ratio the desulfurization may be disrupted by high carbon
deposition rates. In any case, the lowest air/fuel ratio which avoids
excess carbon deposition is the most reasonable choice for this operating
variable. For preliminary design it is assumed that the minimUm air/
fuel ratio may be applied to the gasification/desulfurization process
without increasing the carbon deposition rate above that expected at 20%
of stoichiometric air.
474
-------
In the regenerative operation, the choice of regenerator
temperature control scheme affects the temperature requirements in the
gasifier and the overall system performance. This is demonstrated by
examining the gasifier thermal efficiency and the gasifier product hot
gas heating value as a function of the carbon deposition rate and the
gasifier and regenerator temperature control methods.
Energy and material balances are carried out around the
gasifier in terms of the following quantities: ~HG = the heat generation
rate at 77°F in the gasifier due to cracking, combustion, and desul-
furization, in Btu/lb fuel (as in Figure 2-3); F = mass flow rate of
p
gasifier product, excluding the temperature control stream, in lb/hr/lb
fuel/hr; ~Hp = the enthalpy change in the gasifier product going from
77°F to l600°F, in Btu/lb, = -Cp l523°F (from the Esso data Cp ~ 0.35
Btu/lbOF); FS = mass flow rate of the temperature control stream (stack
gas, steam, or water) in lb/hr/lb fuel/hr; ~HS = the enthalpy change in
the temperature control stream going from its initial conditions to.
l600°F (~HS ~ 1800 Btu/lb water based on water at 77°F, 710 Btu/1b steam
based on saturated steam at 1 atm, and 358 Btu/lb stack gas based on
stack gas at 300°F).
is given by
The limestone make-up rate to the gasifier, NMG'
Y 56
NMG = 32 .552.= 3.17 Y lb limestone/hr/lb fuel/hr
(2-22)
based on 1 mole of CaO per mole of sulfur fed.
Then, the limestone
contributions to the gasifier energy balance are
(1)
Energy to calcinate the make-up limestone in the gasifier
= 0.434 (3.17 Y) 1740 = 2397 Y Btu/lb fuel.
(2)
Energy to heat the make-up lime from 77°F to l600°F
= 0.3 (3.17 Y) 0.566 (1523) = 820 Y Btu/lb fuel.
475
-------
(3)
Energy to heat the C02 calcination product from 77°F
to 1600°F
= 0.434 (N ) 1424.0 (1523) = 573 YBtu/lb fuel.
MG .. .
(4)
Energy returned from the regenerator to the gasifier as
1900°F lime = 0.3 [NG - 0.56~ NMG] 300 Btu/1b fuel.
(5)
The energy lost through the gasifier boundaries is
assumed to be 1% of the fuel oil heating value per lb
of fuel fed
= 177 Btu/1b fuel.
The combustion value of material taken from the gasifier to the regenerator
is given by
(6)
(7)
Sulfur lost = £ Y (10,630) Btu/lb fuel;
s
Carbon lost = XcNG (14,093) Btu/lb fuel.
An energy balance around the gasifier results in
FS~HS = ~HG + 90 {NG - 0.566 NMG} + Fp~Hp - 3790 Y - 177
where
(2- 23)
3790 Y = [1] + [2] + [3].
The temperature control requirements, FS' for stack gas recycle, steam
injection, or water injection can be calculated from equation (2-23) for
a given air/fuel ratio, sulfur content in the residual oil fuel, sulfur
removal efficiency, carbon deposition rate, regenerator temperature
control method, and gasifier control method.
The total loss in gasifier product heating value is given by
the sum of items (1), (2), - (4), (5), (6), (7) and FS A, where the last
item applies to the case of steam injection and water injection with A
being the heat of vaporization of water = 980 Btu/lb. Then for all five
476
-------
proposed gasifier temperature control schemes the thermal efficiency,
£T' is given by
(1) + (2) - (4) + (5) + (6) + (7)+ FSA
1 - £ ...
T
17,700
1 - £T = [£SY (10,630) + XcNG (14,093) - 90 ,[NG - 0.566 NMG]
+ 3217 Y + 177 + FSA]/17,700
(2-24)
with regenerator temperature control by lime circulation or by addition
of fresh stone to the regenerator. In the case of heat transfer surface
in the regenerator, the thermal efficiency is identical to that of the
lime circulation method. Equation (2-24) may be used to evaluate the
thermal efficiency for a given air/fuel ratio, sulfur
residual oil fuel, sulfur removal efficiency, carbon
regenerator temperature control method, and gasifier
method.
content in the
deposition rate,
temperature control
Figures 2-10 and 2-11 illustrate equation (2-24) as a function
of the carbon deposition rate for air/fuel ratios of 20% and 25% of
stoichiometr~c with 95% sulfur removal of a 3 wt % sulfur fuel oil.
Curves for 2 and 4 wt % sulfur fuel oils are very similar to Figures 2-10
and 2-11.
Note that the thermal efficiency is the same for the three
gasifier temperature control schemes in which a heat of vaporization
penalty is not paid: stack gas recycle, heat transfer surface, and,
minimum air/fuel ratios (~ 14% of the stoichiometric air rate). The
thermal efficiency is also independent of the air/fuel ratio for,stack
gas recycle and heat transfer surface temperature control schemes, while
increasing the air/fuel ratio lowers the thermal efficiency in the
steam and water injection temperature control methods. Increasing the
carbori deposition rate lowers the thermal efficiency with all five of
the proposed temperature control schemes unless the cases of steam and
water injection act to remove deposited carbon in the gasifier by
477
-------
~
I
>-
U
C
Q)
'u
5 92
ro
E
....
Q)
.c:
I-
....
Q)
;;::
'Vi
~ 90
Curve 643985-B
96
""'-
~
""'-"",,-
--.
-- Stack Gas Recycle and Heat
--. --. Transfer Surface and Min.
----..Air/Fuel Ratio
- - -. - - --..::::: - - - -
--..
~
100%Removal of Carbon
94
..............
..............
~
..............
~
..............
..............
............
""
............
. ............
100%RemOv~f Carbon- ---- - - -- - - ............~-
88
86
"
"-
"-
"
. "-
"-
"'"
"-
%Stoichiometric Air = 20.0" ' ,
%Sulfur in Fuel = 3.0 "Steam Injection
Sulfur Removal = 95% '-.....
NMR =0, Heat Transfer Surface in Regenerator
- - NG =NGMin
o
0.01 0.02 0.03
Net Carbon Deposition Rate - Ib Carbon/Hr/lb Fuel/Hr
Fig. 2-1O-Gasifier Temperature Control
418
-------
~
I
>-
u
C
Q)
u
~ 92
LLJ
i6
E
.....
Q)
..r:::.
I-
.....
Q)
:;
'V:;
(5 90
Curve 643986-B
96
-.........-.........
-.........,,-
-- ~ Stack Gas Recycle, Heat
of St . h. t. A. - 25 0 -......... -- Transfer Surface, Min.
{a OIC lo.me rlC I r - . ~ -......... Ai rIFuel
% Sulfur In Fuel = 3.0 "-
Sulfur Removal = 95% -.........-.........
NMR = 0, Heat Trans --
- - NG = NG .
Mln
94
100%Removal of Carbon
--............
--.........................
""'-..
........... .
""'-..
"
"""
""-
...........
""-
"
........
........
100% Removal of Carbon
88
--------- ---------
86
............
"
""-
............
""-
"-
""-
"
o
0.01 0.02 0.03
Net Carbon Deposition Rate - Ib Carbon/Hr/lb Fuel/Hr
Fig. 2-11-Gasifier Temperature Control
479
-------
reaction (2-21). The situation in which 100% of the deposited carbon
is removed by reaction (2-21) is indicated in Figures 2-10 and 2-11 by
the broken lines labeled "100% removal of carbon."
The method of regenerator temperature control also affects
the gasifier thermal efficiency» with the method of feeding make-up
limestone to the regenerator (NG = NGMIN) giving lower thermal
efficiencies than the stone circulation (NMR = 0) or heat transfer
surface methods for a given carbon deposition rate. The methods using
make-up limestone and heat transfer surface will also give lower carbon
deposition rates» XcNG» than the stone circulation method» because
NG ? N~IN in these two schemes.
Thus» the gasifier thermal efficiency may vary between 97%
and 85% depending on the carbon deposition rate and the choice of
gasifier and regenerator temperature control schemes.
As the air/fuel
ratio approaches the minimum value for gasifier temperature maintenanc~ '
all of the proposed gasifier temperature control schemes will tend to
approach the behavior o'f the curve for the minimum air/fuel ratio in
Figures 2-10 and 2-l~ while as the air/fuel ratio is increase~ the
proposed temperature control schemes» steam and water injection» will
become less attractive.
The projected gasifier product heating value is shown in
Figure2-12» for the regenerative operation» as a function of the carbon
deposition rate.
Curves are shown for gasifier temperature control by
use of heat transfer surface» water injection» steam injection» and
stack gas recycle at 20% (solid curves) and 25% of stoichiometric air
(broken curves). The product heating value for the case of operation at
the minimum air fuel ratio is about 500 Btu/S ft3. The method of
regenerator temperature control is also indicated on the figure» with
the use of make-up limestone (NC = NGMIN). giving» in genera], slightly
higher heating values than the use of stone circulation (NMR = 0).. The
use of regenerator heat transfer surface gives heating values identical
to the case of regenerator make-up limestone.
480
-------
Curve 643984-B
350
NG=NGMin' NMR =0
Heat Transfer
Surface
V')
-
:::>
I--
co
--------------
NG =NG NMR =0
Min'
Heat Transfer
Surface
C'f"I
-
-
300
NG =NG "
Q)
::J
I'tJ
>
en
c::
~ 250
::x:
NMR =0
Water In "ection
VI
I'tJ
<.:)
NG = NG .
M'n
--------------
-- --- -----
- -~ -;0 - - waier"Iii'ection)
N M
G GMin
NMR ::::0
Steam I njection and
Stack Gas Recycle
-
o
::x:
200
3% Sulfur in Fuel
95% Sulfur Removal
20% Ai r/Fuel
- - - 25o/q Air/Fuel
Steam I njection and
NG =NGMin Stack Gas Recycle
-------------
---
-----
N ---
. MR=O --
0.01 0.02
Net Carbon Deposition Rate
Fig. 2-12-{;asifier Product Heating Value
0.03
481
-------
The overall behavior indicated in Figure 2-12 shows that the
gasifier product hot gas heating value is relatively insensitive to the
carbon deposition rate and the method of regenerator temperature control,
but is largely dependent on the air/fuel ratio and the gasifier
temperature control method.
The product heating value is an indication
of the relative equipment sizes necessary with the various temperature
control schemes.
The next matrix summarizes the advantages and disadvantages of
the five proposed gasifier temperature control schemes based on the
considerations made in this section. The low thermal efficiency of
steam injection at air/fuel ratios of 20% to 25% of stoichiometric makes
control by steam injection unattractive, while the large equipment sizes
required with stack gas recycle at air/fuel ratios of 20% to 25% o~
stoichiometric make stack gas recycle unattractive. As the air/fuel
ratio is reduced, all of the proposed gasifier temperature control
schemes become more attractive; with operation of the gasifier at the
minimum air/fuel ratio of about 14% of stoichiometric the most attractive.
condition,
. values (X
. c
operation adds some quantitative measures
if the carbon deposition rate does not increase to unreasonable
approaches 0.01). Economic analysis of the regenerative
as a basis to compare the
proposed temperature control schemes.
Once-Through Operation
In the once-through gasification/desu1furization operation,
the same five possible gasifier temperature control methods are examined.
The sulfate generator temperature of 1500°F is maintained by one of two
proposed methods: heat transfer surface in the sulfate generator or
I circulation of excess air through the sulfate generator. The overall
behavior of the once-through operation differs from that of the regenera-
tive operation because of the higher rate of limestone addition to the
gasifier in the once-through case and the nature of the sulfate generator.
482
-------
EVALUATION OF GASIFIER TEMPERATURE CONTROL METHODS
CONTROL METHOD
DISADVANTAGES
ADVANTAGES
Heat Transfer
Surface
Minimum Air/Fuel
Ratio
~
co
UJ
Stack Gas Recycle
Steam Injection
Water Injection
Control is independent of the process
operation
Limits equipment sizes, HV ~ 325
Btu/S ft3(a)
Maximizes thermal efficiency ~ 97%
Permits operation without further
temperature control provisions in
gasifier
Minimizes equipment sizes
HV ~ 500 Btu/S ft3
Maximizes thermal efficiency
~ 97%
Simple procedure
High thermal efficiency ~ 97%
Simple gasifier design
May lower carbon deposition rate
Simple gasifier design
May lower carbon deposition rate
Fuel distribution may be improved
High capital costs
Complex design and control
Probably necessitates heat transfer
surface in the bed, or modular design
Possible high carbon deposition rate
Necessitates large equipment sizes,
HV ~ 200 Btu/S ft2(a)
Highest power requirements, particle
clean-up costs
Low thermal efficiency ~ 93%(a)
Necessitates large equipment sizes,
HV ~ 200 Btu/S ft2(a)
May be detrimental to sulfur removal
Fairly low thermal efficiency ~ 93%(a)
Water distribution may be a problem
. HVa:e~50 Btu/S ft3(a)
May be detrimental to sulfur removal
(a) At an air/fuel ratio of 20% to 25% of stoichiometric.
-------
Sulfate Generator
Heat is produced in the sulfate generator by the same reactions
operable in the regenerator, equations (2,9),. (2-10), and (2-11), at
the temperature of 1500°F. reaction (2-10) should dominate while complete
combustion of the carbon deposited on the gasifier lime is assumed.
The rate of heat generation by chemical reaction, ~HS' in the sulfate
,
generator is assumed to be
YE
~HS = 14,093 XcNG + 3.96 x 105 3~
(2-25)
in terms of the carbon deposition rate, XcNG' and the rate of sulfur
removal, YES' Utilized lime enters the sulfate generator at a rate NG
and a temperature of 1600°F, leaving at a temperature of 1500°F. If E
is the fraction of excess air circulated through the sulfate generator
then an energy balance around the sulfate generator gives, for temperature
control,
[XcNG YES] (0.79 ]
° = ~HS + 30 NG (1 + E) ~ + ~ 0.21 7.4 (1423)
X NG X NG E
c C s
-11. 6 (12)1423 - 7.8 (1"2 + 2 ~) E (1423) ,
(2- 26)
assuming no heat losses.
This reduces to
X NG + 1.05 YE
c s
E = 0.448 X NG + 0.337 YE
c s
(2-27)
Note that in the once-through operation NG is much smaller than in the
regenerative case. Where NG may be as large as 2 to 3 1b/hr/1b fue1/hr
in the regenerative operation, NG will be about 0.2 1b/hr/1b fuel/hr in
the once-through case, making X NG « YE. Thus, carbon deposition may
c s
be neglected in equation (2-27) to give
E ~ 3.1 ,
484
-------
or 310% excess air must be circulated through the sulfate generator to
maintain temperature control, independent of the rate of sulfur removal
from the residual fuel oil.
It appears that operating with excess air in the sulfate
generator should promote a high production of calcium sulfate, so that
excess air should be used even in the case of temperature control of the
sulfate generator with heat transfer surface.
At a superficial velocity
of 8 ft/sec, with 95% removal from a 3 wt % sulfur fuel oil, and a 95%
thermal
efficiency of the once-through operation, the cross-sectional
2
the sulfate generator would be 0.285 ft /MW with temperature
by heat transfer surface (based on 50% excess air). With
area of
control
temperature control by excess air circulation the cross-sectional area
would be about 1.0 ft2/MW. The cross-section of the sulfate generator
with control by heat transfer surface is nearly the same as that of
the regenerator in the regenerative operation (Figure 2-8). However,
with the rate of heat generation about .three times as great in the
sulfate generator as in the regenerator, much more heat transfer surface
will be required for temperature control.
Design will probably be
li~ited to placing heat transfer surface internal to the sulfate
generator if this method of temperature control i~ to be used.
The temperature control scheme using excess air circulation
through the sulfate generator is clearly superior to heat transfer surface:
capital costs should be lower, and the operation and control simpler.
Gasifier - Once-Through Operation
Stack gas recycle, stearn injection, water injection, heat
transfer surface, and operation at the minimum air/fuel ratio are
proposed for temperature control of the once-through gasifier. An
energy balance around the gasifier is made up of the following factors:
with a limestone make-up rate to the gasifier of 3 moles CaO/mole sulfur
fed, the heat taken up in the calcination of the make-up limestone is
three times that assumed in the regenerative case, or
485
-------
Energy to calcination = 1.137 x 104 Y Btu/1b fuel
Energy lost through boundaries = 177 Btu/1b fuel
Energy returned from sulfate generator = ~HS'
An energy balance gives
4 .
FS~HS = ~HG + ~HR - 1.137 x 10 Y - 177 + Fp~Hp
(2-28)
in terms of the nomenclature introduced earlier.
Equation (2-28)
indicates the ,requirements for temperature control with stack gas
recycle, steam injection, and heat transfer surface. Comparing
equation (2-28) with the results for the regenerative operation,
equation (2-23), FS is nearly identical in the two operations for
air/fuel ratio and carbon deposition rate.
equal
Since the heat generated in the sulfate generator, ~HS' is
recycled to the gasifier, higher thermal efficiencies are expected in
the once-through operation than were found for the regenerative operation.
In terms of the energy taken from the gasifier product, the thermal
efficiency becomes
10,651 Y + 177 - 1750 £sY + FSA
1 - £ =
T 17,700
(2-29)
For the gasifier temperature control methods of stack gas recycle,
heat transfer surface, and minimum air/fuel ratio, the thermal efficiency
is independent of the carbon deposition rate in contrast to the
regenerative case. Nor does the air/fuel ratio affect the thermal
efficiency with stack gas recycle and heat transfer surface.
For 95%
removal of sulfur in a 3 wt % sulfur fuel, the thermal efficiency is
97.5% with stack gas recycle, heat transfer surface, and minimum air/
fuel ratio. For steam and water injection the thermal efficiency is
reduced by the factor FSA much as in the regenerative case. .
486
-------
The heating value of the gasifier product from the once-through
operation is nearly the same as the heating value obtained for the
regenerative operation. As in the regenerative operation, the heating
value of the once-through gasifier product is sensitive to the air/fuel
ratio and the gasifier temperature control method, but not to the rate
of carbon deposition or the method of sulfate generator temperature
control.
In the once-through case, 'operation of the gasifier at the
minimum air/fuel ratio appears superior to the gasifier temperature
control schemes proposed for use at higher air/fuel ratios. Large
fractions of deposited carbon on the gasifier lime, X , will be more
c
tolerable in the once-through operation than in the regenerative operation
unless the rate of sulfur removal is reduced by the existence of carbon
in large amounts.
This behavior may make operation at the minimum
air/fuel ratio more easily realized in the once-through operation than
in the regenerative operation.
487
-------
Overall Material and Energy Balances
Based on the specifications presented on page 461, material
and energy balances have been carried out for regenerative and once-
through operations. Results are listed in Tables 2-5 and 2-6 for the
two schemes whose flow diagrams are shown in Figures 2-13 and 2-14.
For the regenerative operation, Table 2-5 indicates the
stream characteristics for the five proposed gasifier temperature control
schemes and the three regenerator temperature control schemes.
In
addition to mass flow rates and temperatures of the streams shown in
Figure 2-13, Table 2-5 presents other pertinent characteristics of the
gasifier product, the cross-sectional area of the gasifier and regenera-
tor per 1b of fue1/hr, and the estimated bed depths of the gasifier
and regenerator.
For the once-through operation pictured in Figure 2-14,
Table 2-6 indicates the stream characteristics and vessel dimensions for
the gasifier control schemes of stack gas recycle, minimum air/fuel
ratio, and heat transfer surface, and for sulfate generator temperature
control by excess air circulation. There is general similarity between
the regenerative and once-through operations, with the greatest difference
in the two operations in the stone flow rates.
The importance of the gasifier temperature control scheme is
evident from Tables 2-5 and 2-6. The equipment sizes, thermal efficiency,
and stone circulation rates vary greatly depending on the choice of
temperature control.
488
-------
Dwg. 295&.35
@ (j)
9 ([)
Gasifier Regenerator
Q) @
@
G) CV @
Fig. 2-13-Regenerative Operation Flow Diagram
cv
(j)
10 Su Ifate @ @
Generator Gasifier
@ 1
(j) ~ @
Fig. 2-14-0nce-Through Operation Flow Diagram
489 .
-------
REGENERATIVE OPERATION MATERIAL AND ENERGY BALANCES
TABLE 2-5
STEAM
Fuel oil, 1b/hr
Air, 1b/hr
Temperature control stream, 1b/hr(c)
+:-
\0
o
Gasifier product, 1b/hr
ft3/hr
:t(~~ ft3
T
Gasifier make-up, 1b/hr
Gasifier area, ft2/1b fue1/hr(f)
Static d2pth, ft
TEMP.
200°F
77°F
1600°F
77°F
1600°F
GASIFIER TEMPERATURE CONTROL METHOD
Heat Transfer.
Surface
1 1 1 1 1
2.76(a) 1.94 (b) 2.76 2.76 2.76
.1.70 (d) 0 'V 850 Btu/hr 0.93 0.37
[1. 50) Heat transfer rate [0.78) [0.30)
5.46 2.94 3.76 4.69 4.13
320 138 195 299 236
215 498 352 212 283
97.1 97.1 97.1 92.2 95.3
0.095 0.095 0.095 0.095 0.095
0.0093 0.003 0.0053 0.009 0.0064
2.5-3.5
STEAM
Regenerator air, lb/hr
Regenerator product, lb/hr
XS02
Regenerator make-up, lb/hr
Stone circulation, 1b/hr
XcNG
Stone circulation, 1b/hr
Sulfated stone disposal, lb/hr
Regenerator a.rea, ft2/lb fuel/hr (f)
Static depth, ft
TEMP.
REGENERATOR TEMPERATURE CONTROL METHOD
Stone Circulation I Make-Uo Stone' Heat Transfer Surface'
77°F 0.302 0.228 0.228
1900°F 0.34 0.268 0.268
0.09 0.12 0.12
77°F 0 0.049 0
1600°F 1.80 0.713 0.713
0.01 0.003 0.003
19000F 1. 75 0.68 0.66
19000F 0.051 0.077 0.051
19000F 6.23 x 10-14 -4 -4
4. 7 x 10 4.7 x 10
2.0
(a) 20% of stoichiometric
(b) 14% of stoichiometric
(c) Based on 300°F stack gas, 212°F steam
(d) Figures in brackets refer to regenerator temp.
control by make-up stone or H.T.S.
(e) Regenerator temperature control by stone circulation
(f) Based on superficial velocity of 8 ft/sec
-------
.;::- .
\0
I-'
TABLE 2-6
ONCE-THROUGH OPERATION MATERIAL AND ENERGY BALANCES
STREAM
Fuel oil. 1b/hr
Sulfate generator air, 1b/hr
Sulfate generator product, 1b/hr
1b 02/hr
Air, 1b/hr
Gasifier a.ir stream, 1b/hr
Temperature controls tream, 1b/hr
Gasifier product, lb/hr
ft3/hr
Btu/S ft3
£T.
Limestone feed, 1b/hr
Sulfided 1 ime, 1b/hr
Sulfated 1 ime, 1b/hr
Gasifier area, ft2/1b fue1/hr
Static depth, ft
2 .
Sulfate generator area, ft /lb fuel/hi
Static depth, ft
Temp.
200°F
77°F
1500°F
77°F
1600°F
77°F
1600°F
1500°F
1600°F
GASIFIER TEMPERATURE CONTROL METHOD
1
0.983(a)
0.924
O.l71(b)
2.02
2.94. ['~90°F]
1. 60 [300°F]
5.54
296
231
0.975
0.265
0.180
0.186
0.0085
3.5-4.0
0.0017
3-4
Minimum
Ai
1
0.983
0.924
0.171
1.15
2.07
o
[790°F]
Heat Transfer
1
0.983
0.924
0.171
2.02
2.94
[490°F]
2.85
148
465
0.975
0.265
0.180
0.186
0.0033
3.5-4.0
0.0017
3-4
~ 650 Btu/1b fuel
Heat transfer rate
3.94
204
336
0.975
0.265
0.180
0.186
0.0053
3.5-4.0
0.0017
3-4
(a)
Based on 310% excess air
(b) <
0.1% S02 in product
-------
Power Requirements
Pressure drops have been estimated in Table 2-7 for regenerative
and once-through operations in which low-pressure drop burners requiring
no particulate clean-up prior to combustion have been u~ed.
Bed pressure
drops are estimated to be about one inch of H20 per inch of bed height;
distributor pressure drops are assumed to be 50% of the bed bP; and
burners are assumed to contribute two to four inches of H20 pressure
drop. In the case of the sulfate generator, particulate removal has
been assumed necessary in the recycle stream going from the sulfate
generator to the gasifier to protect the air distributor.
In addition,
the air-side pressure drop in the burners is assumed to be 7 to 20
inches of H20.
Power requirements, exclusive of power for solids handling and
circulation, are listed in Table 2--8 and are based on the maximum
pressure drops shown in Table 2-7.
The table indicates the power
requirements per pound of fuel over and above that already required in
the power generating system.
This power requirement is combined in
Table 2-8 with the thermal efficiency of the gasifier to give the overall
efficiency of the gasification/desulfurization system. Though the power
requirements of the once-through operation are somewhat larger than the
power requirements of the regenerative system, the once-through operation
has a higher overall efficiency. Auxiliary power costs do not appear
to be an extremely important operating cost factor with gasification
systems, totalling at most a loss of about 1% in overall efficiency.
492
-------
TABLE 2-7
PRESSURE DROPS
REGENERATIVE OPERATION ONCE-THROUGH OPERATION
8 ft/see 8 ft/ see
Gasifier Regenerator Gasifier Sulfate Generator
(in. H20)
Bed 30-42 24 42-48 36-48
Distributor 15-21 12 21-24 18-24
~
\0
w Piping 5-7 4 7-8 7-10
Burners 2-4 2-4
Cyclone 20-30
TOTAL 52-74 40 72-84 81-112
-------
TABLE 2-8
POWER REQUIREMENTS
REGENERATIVE OPERATION ONCE-THROUGH OPERATION
Stack Gas Minimum Heat Transfer Steam Yater Stack Gas Minimum lieat Transfer
Recycle Air/Fuel Surface I nj ec tion Injection Recycle Air/Fuel Surface
Kwh/+b fuel (a) 0.021 0.008 0.013
Thermal efficiency 0.968 0.968 0.968
Overall efficiency 0.957 0.963 0.960
0.021 0.015 0.022 0.012 0.016
0.902 0.943 0.975 0.975 0.975
0.889 0.935 0.960 0.967 0.963
+-
\()
+-
(a) Based on a plant heat rate of 104 Btu/kwh, a fuel oil he~ting value of 17,700 Btu/1b, and a compressor efficiency
of 70%.
-------
DESIGN CONCEPTS
Energy and material balances provide basic information with
which to examine "the feasibility of applying the gasification/desul-
furization concept as an add.-on to an exist1ng boiler, or as a new
plant design feature.
The feasibility of the retrofit concept is
investigated in terms of the availability of space in an existing
power plant, the modifications necessary to retrofit an existing
boiler, and the performance of a modified boiler.
The performance and
design concepts projected for new boiler installations are examined.
495
-------
Retrofit Concepts
The feasibility of converting an existing boiler to one which
utilizes the fuel from a gasification/desulfurization system depends on
a number of factors) many of which will differ from one boiler to the
next. The space available for a gasification/desulfurization system
in an existing power plant) the modifications necessary to retrofit an
existing boiler) and the performance of a retrofit boiler will depend
on the specific gasification/desulfurization system design and choice
of operating conditions) the location of the gasification/desulfurization
system in the plant) the type (coal-- or oil-fired) of boiler ) size of
boiler) the turn-down needed) the boiler load factor) and the specific
design features of the boiler. These factors are discussed with respect
to the location of the gasification/desulfurization system relative to
the boiler) and the performance of the retrofit boiler.
Retrofit Location
The gasification system may be placed internal to the boiler
(directly beneath) or extern~l to the boiler (as close as possible
with9ut carrying out major boiler modifications in order to locate the
system).
These concepts are illustrated in Figure 2-15.
The
feasibility of the internal location is examined in Figures 2-16 and
2-17 with respect to the amount of space available beneath the boiler.
The two figures compare the cross-sectional area of coal- and oil-fired
boilers with the MW rating of the boiler.
Superimposed on the figure
are the gasifier cross-sectional areas for the various temperature
control schemes plotted against the gasifier MWrating. Figure 2-16
is for an air/fuel ratio of 20% of stoichiometric) while Figure 2-17
is at 25% of stoichiometric; both figures are at a superficial velocity
of 8 ft/sec.
In interpreting the figures) it should be noted that the
regenerator) or sulfate generator cross-section is :much smaller than
496
-------
Clean
Fuel
Gas
Boiler
Burners
in Normal
Positions
Gasifier
12- 25 ft
Dwg. 6167A47
Boiler
Burners
in Up-Shot
Position
Internal Gasifier Location
Clean Fuel
Gas
Gasifier
Boiler
Ash
Pit
External Gasifier Location
Fig. 2-15-Boiler retrofit concepts
497
-------
s:
:2
>-
-
'u
ra
c..
ra
U
....
.,.. Q)
\D :;::
0:> 'Vi
ra
c.:> 400
-
....
Q)
'0
CO
300
Curve
CD steam Injection - NMR = 0
CV Stack Gas Recycle - NMR = 0
(j) Steam Injection - NG = NGMin
@ Stack Gas Recycle - NG = NGM'
. In
(j) Steam Injection - 100%Removal of Carbon
@ Water Injection - NiV\R = 0
(j) Water Injection - NG =NGMin
CID Water Injection - 100% Removal
of Carbon
(2) Heat Transfer Surface
800
700
600
500
200
100
o
400
Curve 643982-6
800
1200 1600
Boiler/Gasifier Cross-Section - ft2
2400
2000
Fig. 2-16-Space Requirements for Internal Design - 20% Stoichiometric Air: 8ft/see - Carbon
Deposition Rate =0.015Ib Carbon/lb Fuel
-------
s:
2:
Z' 500
u
ra
c..
ra
U
L. 400
Q)
<="
\0
\0
. ~
'Vi
ra
~
~ 300
'0
co
Curve 643983-A
700
600
200
100
o
400
800
1200 1600
Boiler/Gasifier Cross-Section - ft2
2000
2400
Fig. 2-17- Space Requi rements for I nternal Design, 25% Stoichiometric Ai r: 8ft/see - Carbon Deposition
Rate = O. 004 Ib Carbon/lb Fuel
-------
the gasifier cross-section and was not included in the figures.
With
respect to vertical space, 12 to 25 ft of headroom is available under
boilers, which is sufficient for a gasifier 10 ft in total height.
The criterion that the internal design is feasible .if the
boiler cross-section is larger than the gasifier. cross-section suggests
the following retrofit design factors:
Critical Factors in Determining Feasibility of Internal Retrofit Design--
Regenerative or Once-Through:
.
Superficial velocity
.
Air/fuel ratio
.
Temperature control method
.
Boiler type (coal or oil) and size
Feasibility Limits on the Internal Retrofit (Based on Boiler Capacity
> 100 MW):
.
Superficial velocity ~ 8 ft/sec
.
Air/fuel ratio ~ 20% of stoichiometric
.
Temperature control by minimum air/fuel ratio and heat
transfer surface feasible for boilers> 100 MW
.
Temperature control by stack gas recycle, steam and water
injection feasible for boilers < 200 MW
.
Internal design more probable with coal-than with oil-fired
boilers.
The curve showing the cross-sectional area of the gasifier with tempera-
ture control by operation at the minimum air/fuel ratio represents the
minimum cross-sectional area possible at a given superficial velocity
and is the most promising operating condition from the standpoint of the
internal design concept. The gasifier cross-section curves in Figures
2-16 and 2-17 refer specifically to the regenerative operation, but they
are a150 representative of the once-through operation gasifier cross-
sections.
500
-------
Another factor which may limi.t the feasibility of the internal
design is the need for particulate removal before the gasifier product
reaches the burners. Though preliminary investigation indicates that
burners may be. designed which can be operated at high dust loading
levels, the situation where high-efficiency cyclones are called for is
illustrated in Figures 2-18 and 2-19. Figure 2-18 shows a 20 MW coal-
or oil-fired boiler with. a regenerative oil gasifier add-on operated
at 6ft/sec, 20% of the stoichiometric air/fuel ratio, with temperature
control by water injection, and burners modified to be in the upshot
position.
Figure 2~19 shows a 150 MW coal-fired boiler retrofit with a
regenerative gasification system with the same operating conditions as
\
in Figure 2-18.
In this case the burners have been left in the normal
boiler positions, and high-efficiency cyclones have been placed next
to the boiler due to their excessive height.
External location of the gasification/desulfurization system
would probably minimize boiler modifications, though it might necessitate
the removal of auxiliary equipment in the proximity of the boiler. The
greatest concern with the external design is the necessity of trans-
porting the l600°F reducing gas from the gasifier to the boiler. Great
lengths of piping could be expensive, and heat losses and condensations
of tars in the piping must be minimized.
The internal design concept
may exhibit the same problem. Figure 2-20 illustrates a modular external
design for a 600 MW coal-fired boiler. High~efficiency cyclones are again
included.
A more detailed look at specific boiler designs, and the modi-
fications needed for internal or external retrofit designs, is needed in
order to evaluate these concepts.
The cost involved in boiler modifi-
cation and the down-time of the boiler during modification may be the
limiting factor in retrofit feasibility.
Boiler Performance and Boiler Modifications
Degraded performance of a retrofit boiler due to a decrease in.
the temperature and luminosity of the flame produced at the burner may
501
-------
Dwg. 2950A72
8 - 2.5 x 10 Btu/hr Burners .
2 High Efficiency Cyclones per Burner - 3 x 103 ft3/ min per Cyclone
- o.
G = Gasifier
R = Regenerator
,
-"
Scale: 8 ft per inch
Make-Up
Stone
Air
S02
Fuel
H20
Stone
Disposal
Air
Air
. Stone
Pu rge
Fig. 2-18-20 MW gasifier add-on
.502
-------
Dwg. 2950A71
6 - 2.5 x 108 Btu I hr Burners
1 High Efficiency Cyclone per Burner
Scale: 16 ft per inch
50 .-
2
Stone
Fig. 2-19-150 MW gasifier add-on to a coal fired boiler
503
-------
Dwg. 2950A73
4 - 150 MW Modules, 6 high efficiency cyclones per Module
8
24 - 2.5 x 10 Btu/hr Burners
5°2
Scale 40 ft per inch
5°2
Fig. 2-a}-roO MW gasifier external to a coal fired boiler
504
-------
be a problem to be considered with all retrofit boilers.
The main fac-
tors controlling the degradation of boiler performance are the gasifier
temperature control method and the boiler type.
If no inert material
is added to the gasification product to control temperature, and if the
thermal efficiency of the gasifier is high, then the boiler degradation
should be mainly a function of the flame luminosity. Esso (England)
has obtained a luminous flame upon combustion of the gasification
product, indicating that the temperature control scheme may be the most
important single factor. Note that a decrease in the flame temperature
due to the addition of a noncombustible to the fuel will be accompanied
by an increase in gas velocity through the convective section of the
boiler, yielding higher heat transfer in that section.
The time and capital involved in the modification of a utility
boiler are important factors affecting the feasibility of boiler retro-
fits.
Any retrofit, of external or internal location, calls for the
replacement of the boiler burners.
Burner modifications will include
modifications to the burner manifolding. alterations to the boiler
water walls, and possibly changes in the structural foundation of the
boiler walls.
Minimizing the new burner size will be ~dvantageous from
this standpoint. since the size is mainly a function of the gasifier
temperature control scheme and the air/fuel ratio.
The internal add-on concept calls for the removal of the
boiler ash pit and other auxiliary equipment found beneath the boiler.
Alterations to the lower sloping water walls of the boiler will be
necessary if the burners are placed in the upshot position at the base
of the boiler.
Further modifications result from the need for addi-
tional fans and compressors, and particulate clean-up equipment.
The following two matrices summarize the points affecting the
feasibility of boiler retrofits. The first compares 'the internal and
external location of the gasification/desulfurization system, while the
second compares the modifications and ,boiler performance probable with
coal- and oil-fired retrofits.
sos
-------
The factors presented in these matrices favor the external
retrofit design over the internal retrofit'design, and' the feasibility
of retrofitting a coal-fired boiler over the feasibility of retrofitting
an oil-fired boiler. However, these preliminary investigations do not
. eliminate the possibility of the
retrofit of oil-fired boilers.
internal retrofit design or the
50G
-------
GASIFIER RETROFIT LOCATION FACTORS
Internal Location
Space Requirements
Under suitable operating conditions space available
beneath the boiler may be utilized (page 500)
Cross-sectional area needed to place burners in
the upshot position at base of boiler may limit
concept (new burners may be 1.5-3 times the size
of normal oil or coal burners)
Design limited by available overhead space if
particulate control is needed before burners
Probably not feasible with modular design in large
boilers (> 500 MW) except at velocities » 8 ft/sec
Modification Factor
~
o
~
More extensive modifications required, more down
time involved in removal of ash pit, auxiliary
equipment, alterations to water walls and found-
ations
Construction more expensive in confined quarters
beneath boiler
Piping and manifolding may be minimized with
burners placed in base of boiler
Each boiler retrofit design may need to be treated
as a unique case
Performance Factor
Need for operating conditions giving maximum
gasifier compactness may be detrimental to system
performance
External Location
Space must be found or provided as near to boiler as
possible -
More flexibility in the external design may allow
operating with maximum fuel heating value to minimize
burner sizes
Feasibility not as limited by space required for
modular design, overhead space, or other spatial
considerations
Modifications consist of removal of auxiliary equipment,
such as pulverizers, and addition of new burners
Prefabrication of vessels possible
Relatively great amount of high-temperature,
piping and manifoldi~g necessary, though the
true with internal design
More uniformity in design from one retrofit to another
noncorrosive
same maybe
Greater flexibility in operating conditions with
external design may allow operation at optimum conditions
-------
RETROFIT OF COAL- AND OIL-FIRED BOILERS
Space Requirements
Coal-fired boilers have larger cross-sections than oil-fired boilers at
large capacity (> 100 MW)
Cross-sections are comparable for smaller capacity boilers « 100 MW)
Headroom is nearly the same in coal- and oil-fired boilers at 12 to 25
feet \
More
potential space exists near a coal-fired b01lerthannear an oil-
fired boiler - pulverizers and other auxiliary equipment may be
removed - especially advantageous if the coal boiler has been
previously converted to operate with low-sulfur oil.
Modification Factors
Normal coal and oil burners are nearly the same size (200 MW Btu/hr
burner is 10 to 11 ft2 in port cross-section)
Much solid handling equipment and particulate clean-up equipment may
already exist with a coal-fired boiler. Increased particulate
control would probably be needed with an oil-fired boiler.
Performance Factors
Coal-fired boilers are designed to operate with slag on the boiler heat
transfer surface, which may make degradation of boiler performance
negligible with coal-fired boilers.. Boiler performance degradation
may be important with oil-fired boilers depending on the gasifier
temperature control method. . .
508
-------
New Boilers
In contrast to boiler retrofit considerations, the feasibility
of incorporating a gasification/desulfurization system into a new boiler
design will be limited only by the overall economics of the system and
the market potential for new oil-fired boilers.
The total space occupied
by the gasification/desulfurization system will be a small percentage of
the total plant volume.
Also, because of the flexibility in boiler de-
sign, the boiler performance will not be affected by the presence of the
gasification/desulfurization system.
Market studies indicate a decrease in the number of new oil-
fired boilers which will be constructed, with practically no new oil-
fired boilers being constructed after 1980.(1) Based ~n this projection,
the gasification/desulfurization of residual oil will be limited to the
retrofit of existing boilers, regardless of the system economics. Fur-
ther investigation of the market potential of the new boiler system is
needed.
(l)L. G. Hauser, private communication.
509
-------
Burners
Discussions with the Bloom Engineering Company and Process
Combustion Corporation suggest that with little or no development work,
burners can be constructed which can handle the hot gasifier product,
combust with low excess air to give a stable flame, and simultaneously
operate with high dust loading to eliminate the need for cyclones prior
to combustion. Bloom Engineering has had considerable experience with
boiler burners handling hot, dirty, lo~Btu fuels such as lean coke
oven gas and blast furnace gas, and they' have made the following points:
HZS Corrosion. HZS corrosion can be prevented by proper
choice of materials - stainless steel or possibly refractory.
Dust Loading.
Though Bloom has had no experience with dust
loading as high as might be expected with a gasification-desulfurization
system, and no direct experience with limestone, they felt erosion would
not be a problem if the gas fuel were carried by the large central burner
port.
The burners could be fired in the upshot position with few
problems due to limestone deposition and minimum maintenance.
The
3
effect of the limestone on the flame temperature, even at 5grains/S ft ,
will be negligible.
Pressure Drop. About Z inches ofHZO pressure drop is expected
across the central gas inlet, and 7 to ZO inches HZO across the air inlet,
though it may be desirable to increase the pressure drop across the gas
inlet for control purposes.
Control-Turn-down.
Turn-down with flame stability could be
maintained at greater than a 5:1 turn-down ratio. Control of burner
intakes could be handled with a distribution orifice at each gas inlet
to eliminate manifolding in the case of the internal gasifier location.
510
-------
Cost. An estimated cost of '" $12tOOO was placed on a
100 x 106 Btu/hr burner and port block with stainless steel construction
once a commercial design has been established.
Burner Size and Design.
6
capacity up to 100 x 10 Btu/hrt and they indicated that larger capacity
Bloom has constructed burners of
burners could easily be designed and constructed.
Since burner
cross-sectional area is approximately inversely proportional to the
heating value of the fuelt the heating value should be maximized in order
to minimize the boiler modifications required.
A preliminary design of
a burner is shown in Figure 2-21 to indicate the burner layout and
approximate dimensions.
511
-------
\J1
I-'
f\)
Di stribution
Orifice
Stai nless
Steel
i
Gas
'" 275 BTU/SFT3
Air
40 - 80 fUsee.. 30 - 46 in.
'" 2 in. H 20 ~ P
Air
..
Wi nd Box
Fig. 2-21-Design of'" 2.5 x 108 BTU/Hr Burner
Dwg. 295BA34
Refractory
Combustion
Chamber
50 - 66 in.
-------
Turn-Down
It is expected that gasification/desulfurization systems will
be operated at turn-down ratios of up to 4:1. The performance of the
gasification/desulfurization system, in terms of fuel properties,
sulfur removal, and temperature control, must be considered over a
large range of fuel feed rates.
With the system design based on the total boiler capacity,
the gasifier performance must continue at a superficial velocity which
is 1/4 of the design velocity when operating at maximum turn-down.
critical factor seems to be that the quality of fluidization remain
The
high over this wide range of superficial gasifier velocities.
Thus,. the
. . .
maximum operating velocity must be about eight times the minimum fluidi-
zation velocity of the lime particle size distribution in the gasifier.
The same restrictions hold for the velocities used in the regenerator and
sulfate generator.
Then, if fluidization with good mixing is maintained
at maximum turn-down, the sulfur removal efficiency should increase as
the system is turned down due to increased gas residence times in the
gasifier; the fuel properties and the temperature control of the system
should be unaffected by the turn-down; and the overall operation,
including particle carry-over and regeneration or sulfate generation,
should improve or remain unaffected as turn-down is achieved.
The possibility of using modular gasifier designs may also aid
in the maintenance of the system performance during turn-down.
In the
case of modular design, individual gasification/desulfurization units
may be shut down completely to achieve the required turn-down while the
remaining units would continue to operate at or near full capacity.
The details of turn-down. start-up, and other system operation
procedures will be considered in the detailed design phase.
513
-------
ECONOMICS
,Capital costs and operating costs have been estimated for both
r'egenerative and once-through operations .
Also, both new and retrofit
designs ofa 600 MW boiler are considered in the economic study for the
operating c'onditi~ns and performance factors Ij:sted on page 461.
514
-------
Cost Factors
The following matrix summarizes the capital and operating cost
factors considered:
ECONOMIC STUDY BASIS
CAPITAL COSTS
Equipment includes - retrofit case
Regenerative
Gasifier vessel
Regenerator vessel
Limestone handling equip.
Sulfur r3covery plant
Fans, compressors
Storage
Particulate clean-up equip.
Burners
Boiler modification costs
Once-Through
Gasifier vessel .
Sulfate g~nerator vessel
Limestone handling equip.
Fans, compressors
Storage
Particulate clean-up equip.
Burners
Boiler modification costs
Installation charged at 50% of equipment costs - includes materials and
labor for concrete structures, piping and ductwork, electrical
installation, instrumentation, etc.
Engineering, contingency, and construction charged at 30% of the direct
cost.
OPERATING COSTS
Capital Charges - 14% of investment
Maintenance - 5% of investment
Overhead - 4% of investment
Materials -
Limestone - $2.50/ton
Fuel oil - $3/bbl
Utilities - Charged as
Disposal - $1.30/ton of lime
fuel adder in overall system efficiency
Labor - neglected
Credits - $15/ton of sulfur when considered
Load factor - 40% and 80% examined
515
-------
Equipment costs were generated from the following estimates:
Vessel Costs [30]
(installed materials costs)
Stainless steel shell - $1.80/lb
Insulation - $20/ft2
25~/gas inlet button - 750 buttons/ft2 of grid
$27/ft2 grid surface - included bubble break-up by means of
vertical rods above the bed
Assuming the units are 15 ft tal~ the vessel cost is given by
2220 D + 167 D2,
where D is the vessel diameter.
Limestone handling costs consist of two contributions, receiving
and disposal equipment costs and recirculation costs.
are estimated from the report by Esso R&D [30].
These factors
Sulfur Recovery Plant.
Direct capital cost and operating cost
for the recovery of elemental sulfur were taken from a report by
Allied Chemical Corporation [2].
Fans.
Costs were estimated from information provided by the
Sturtevant Division of 'Westinghouse.
Particulate Clean-Up Equipment.
High-temperature cyclone
costs, with stainless steel construction, are taken as twice the cost of
low-temperature cyclones.
The costs of cyclone clean-up prior to the
boiler are included in the analysis.
Electrostatic precipitator costs
are estimated as $1.5/kw, not including foundations, to meet present.
particulate standards.
Burners.
A 100 MM Btu/hr burner designed to handle the hot
reducing gas is assumed to cost $12,000.
Burner costs are scaled
according to their cfm capacity, with $12,000 taken as a base for a
fuel with a h~ating value of ~ 300 Btu/S ft3. Burner costs are based on
516
-------
100 MM Btu/hr burners, although larger capacity burners would probably
be used - a 250 MM Btu/hr burner is estimated to cost $20,000-$25,000.
Boiler Modifications.
The cost involved in boiler modifications
is assumed to result mainly from the power production lost during
down time.
This is taken at a rate of $5000/day per 100 MW, and a
period of 30 days down time is assumed reasonable, though this will
vary from boiler to boiler and from one gasifier design to another.
517
-------
Regenerative Operation Economics - New and Retrofit
Table 2-9 lists the preliminary cost breakdown for the major
equipment included in a 600 MW regenerative retrofit and a new system.
I .
Total investment is indicated for each of the five proposed tempet"ature
control methods in $/kw.
The cases of boiler retrofit with and without
the cost of particulate control equipment (cyclones prior to combustion
and electrostatic precipitator final clean-up) are included, along with
the estimated cost for the regenerative gasifier/desulfurization system
when combined with a new boiler design.
The table indicates that operation with the minimum air/fuel
ratio necessary to achieve temperature control is the most promising
control method, if such a low air/fuel ratio can be realized in practice.
The other temperature control schemes are of comparable capital cost,
with steam injection the most expensive.
Table 2-10 summarizes operating costs for a 600 MW regenerative
operation (new and retrofit) with an 80% load factor.
The case of
retrofit with a 40% load factor is included to show the effect of this
important variable on the system operating cost.
The same relative
behavior is observed with the operating costs as with capital costs --
temperature control by minimum air/fuel gives the lowest operating cost,
while steam injection gives the highest.
Steam injection is inferior
because of the low overall efficiency inherent in the scheme at an
air/fuel ratio of 20% of stoichiometric. The analysis points out
that operating at the lowest possible air/fuel ratio, at which sulfur
removal remains high and carbon deposition does not become excessive,
gives the lowest operating and capital costs.
All of the proposed
temperature control schemes will, of course, approach the costs for
518
-------
TABLE 2-9
REGENERATIVE OPERATION CAPITAL COSTS - 600 MW
V1
f-'
\0
TEMPERATURE CONTROL METHOD
EQUIPMENT X 10-3 $ Heat Transfer Minimum
Surface Air/Fuel
Gasifier b 1 020(a) 310 850 870
,
Regenerator ( ) 135 135 138 144
Limestone handling 1,225 1.225 1,235 1,315
Fans 300 200 400 400
Burners 720 610 900 890
Particulate clean-up(c) 1,620 1,510 2,030 1,980
Boiler modifications 900 900 900 900
Equipment 5,920 4,890 6,453 6,499
Installation 2,960 2,445 3,227 3,250
Direct cost 8,880 7,335 9,680 9,749
Eng., con t., const. 2,660 2,200 2,910 2,930
Sulfur recovery 3,300 3,300 3,340 3,540
Total investment 14,840 12,835 15,930 16,219
Retrofit $/k1.v 24.7 21.4 26.6 27.0
Retrofit $/kIV
(without particulate control) 19.5 16.5 20.0 20.6
New $ /k,v
(without particulate control) 14.3 11.6 14.2 14.8
560
140
1,260
300
780
1,760
900
5,700
2,850
8,550
2,570
3,440
14,560
24.2
18.6
13.1
(a)Gasifier cost doubled to include heat transfer requirement.
(b)Regenerator temperature control by heat transfer surface doubles vessel cost.
(c) Particulate clean-up includes cyclones prior to burners and electrostatic precipitator
final clean-up.
-------
TABLE 2-10
OPERATING COSTS FOR REGENERATIVE OPERATION - 600 MW - 80% LOAD FACTOR
TEMPERATURE CONTROL METHOD
Heat Transfer Minimum Stack Gas Steam Water
Surface Air/Fuel Recycle Injection Injection
.-
Capital Charges 103 $/yr(a) 1620 1340 1770 1780 1560
Maintenance (a) 580 480 630 640 560
Overhead (a) 460 380 510 510 450
Limestone (2. 5/ton) 278 278 280 296 284
Vl Disposal ($1. 30/ton) 145 145 146 154 148
f\)
o Sulfur recovery 1210 1210 1210 1270 1240
Fuel adder (R3/bb1) (b) 928 855 1010 2940 1580
5221 4688 5556 7590 5822
Retrofit fuel adder,
-------
the minimum air/fuel .ratio as the air/fuel ratio is reduced from the
ratio of 20% of stoichiometric on which Tables 2-9 and 2-10 are based.
The three most important factors in determining the economics
of the regenerative operation are (1) the carbon deposition rate, (2) the
air/fuel ratio, and (3) the overall system efficiency. These three
factors are interrelated, with the air/fuel ratio probably the single
most important factor.
The carbon deposition rate controls the cost
of sulfur recovery and strongly affects the thermal efficiency of
the system. The air/fuel ratio determines the equipment sizes and
particulate clean-up costs, and also affects the thermal efficiency of
the system.
The overall efficiency controls the direct fuel adder
cost and is itself strongly dependent on the carbon deposition rate,
the air/fuel ratio, and the temperature control method.
521
-------
Once-Through Operation Economics - New and Retrofit
Table 2-11 summarizes the capital cost breakdown for 600 MW
once-through operation, new and retrofit.
Temperature control by
. heat transfer surface, minimum air/fuel ratio. and stack gas recycle
are considered, and the total capital investment in $/kw is shown for
retrofits with and without particulate control, and for a new boiler
excluding the cost of particulate control. Again, operation with the
. . .
minimum air/fuel ratio has the lowest capital cost, and the capital
investment is lower for the once-through operation than for a regenera-
tive operation.
Table 2-12 shows operating costs for the three temperature
control schemes in ~/106 Btu. Both 80% and 40% load factors are
considered in the table, and the results for retrofits and new designs
are shown.
As expected, operating costs are lowest with temperature
control by operation at the minimum air/fuel ratio, but operating costs
for the once-through operation and the regenerative operation are
extremely close.
This indicates that the cost incurred by increasing
.
the limestone feed rate by a factor of three, in going from regenerative
to once-through operation, does not counteract the reductioniri cost
obtained by ceasing to recover sulfur.
With once-through operation, the air/fuel ratio controls the
equipment sizes and the particulate clean-up costs and affects the
system efficiency, as in the regenerative operation. As in the regenera-
tive operation, the overall efficiency determines the direct fuel adder
cost, but with once-through operation the rate of carbon deposition
plays a much less important part in the system economics.
This is
because sulfur recovery is excluded, and the thermal efficiency is
independent of the carbon deposition rate.
522
-------
TABLE 2-11
ONCE-THROUGH OPERATION CAPITAL COSTS - 600 MW
TEMPERATURE CONTROL METHOD
. -3
Equipment x 10 $
Heat Transfer
Surface
Minimum
Air/Fuel
Stack Gas
Recycle
Gasifier
Sulfate ~nerator
Limestone handling
Fans
Burners
Particulate clean-up
Boiler modifications
Equipment
Ins talla tion
Direct cost
Eng. ,
mnt. ,
cons t.
Total investment
Retrofit, $/kw
Retrofit, $/kw
(without particulate
control)
New, $/kw
(without particulate control)
1,020
188
1,320
300
720
1,820
900
6,268
3 ,134
9,402
2,820
12,222
20.4
15.2
10.0
310
188
1,320
300
610
1,710
900
5.338
2,669
8,007
2,400
10,407
17.3
12.4
7.5
850
188
1,330
400
900
2,230
900
6,798.
3,399
10,197
3,060
13,257
22.1
15.5
9.7
523
-------
TABLE 2-12
OPERATING COSTS FOR ONCE-THROUGH OPERATION - 600 MW - 80% LOAD FACTOR.
TEMPERATURE CONTROL METHOD
Heat Transfer Minimum Stack Gas
Surface Air/Fuel. Rec de
Capital charges 103 $/yr(a)
Maintenance (a)
Overhead (a)
Limestone ($2.50/ton)
Disposal ($1.30/ton)
Fuel adder ($3/bbl)(b)
1710
610
490
835
435
830
1460 1855
.520 665
420 530
835 840
435 435
740 928
4410 5253
4910
Retrofit fuel adder
~/106 Btu (80% load factor)
(40% load factor)
lL7
18.4
New qesign fuel adder
~/106 Btu (80% load factor)
9.9
10.5 12.5
16.2 19.7
8.9 10.4
(a)
Includes particulate control.
(b)Based on overall efficiency given in Table 2-8.
524
-------
TECHNICAL EVALUATION
The gasification/desu1furization concepts can.be evaluated
on the basis of the preliminary economic studies and technical feasibility
investigations.
The regenerative and once-through modes of operation
are compared, and the merits and disadvantages of each mode are pointed
out. The costs and pollution control potential of residual oil gasifi-
cation/desu1furization is compared with the low-sulfur oil and stack gas
cleaning alternatives.
In addition, preliminary design conclusions are
listed to summarize the results of the technical feasibility studies.
525
-------
Comparison Between Regenerative and Once-Through Operations
The relative merits of the two modes ofoperatio~ are
summarized in the following matrix.
Preliminary investigations indicate
that, based on all of the points in this matrix,. the once-through opera-
tions may be somewhat more attractive to a utility customer than the re-
generative operation.
Further research into the areas. of sulfur recovery
and system dynamics may alter this outlook.
COMPARISON BETWEEN MODES OF OPERATION
Economics
. Once-through operation requires lowest capital investment
. Operating costs are nearly identical for once-through and
regenerative operations, mainly because of the high cost
associated with sulfur recovery.
. Sulfur credit does not greatly reduce the operating cost of a
regenerative operation
Space Requirements
. Both modes have the. same potential equipment sizes, though the
regenerative operation requires a large space for sulfur recovery,
making the once-through more compact overall
. Once - through may require deeper gasifier bed depth and larger
limestone handling equipment
Performance
. Both modes have same potential for sulfur removal
. Once-through may be slightly more efficient than regenerative
. Once-through has higher power requirements
. Once-through thermal efficiency is independent of carbon deposition
rate
. Regenerative operation minimizes stone disposal rate
. Once-through at present has fewer technical problems
Utility Reaction Factors
. Regenerative mode puts utility in the chemical l;>usiness
. Once-through is the simplest mode of operation, but may increase
solid waste problem unless high stone utilization can be achieved
526
-------
Comparison with Other S02 Control Methods
Table 2-13 compares atmospheric pressure oil gasification/
desulfurization with the alternative schemes of low-sulfur oil and
stack gas cleaning.
Capital costs and operating costs are compared
for new and retrofit systems.
Oil gasification/desulfurization compares
favorably with low-sulfur oil and stack gas cleaning, based on the
preliminary cost estimates presented in Tables 2-9 through 2-12.
A
reduction of about 40% in. the capital costs involved in stack gas
cleaning is estimated for new and retrofit gasification/desulfurization
systems. Operating costs appear to be about the same for stack gas
cleaning and gasification/desulfurization, while a reduction of about
20% in the operating cost for the low-sulfur fuel case by use of the
gasification/desulfurization system is estimated. These conclusions
are based on the desulfurization of high~sulfur residual oil (3 wt %
sulfur) and may be altered when a lower sulfur oil is considered (1 to
1.5 wt % sulfur). Environmental factors are also compared in Tab:Ie 2-13.
The low-sulfur oil alternative is advantageous in that capital costs
are limited to possible boiler modifications necessary when changing
from gas or coal to low-sulfur oil. On the other hand, operating costs
are higher than for stack gas cleaning or oil gasification/desulfurization.
Capital costs are extremely high with stack gas cleaning, especially in
the retrofit case, while operating costs are very near those estimated
for oil gasification/desulfurization.
527
-------
TABLE 2-13
ASSESSMENT OF FLUIDIZED BED OIL GASIFICATION-DESULFURIZATION(a)
LOW-SULFUR
OIL
STACK GAS
CLEANING
OIL GASIFICATION
Regenerative
ODeration
[
Once-Through
ODeration
COST
Capital, $/kw
New
Retrofi t
Fuel adder, C/106 Btu
New
Retrofi t
(b)
25 12-15(c). 8-10(c)
40-75 22~27 (c) 18-22(c)
11 9.5-16.0.(c) 9-10.5(c)
14-20 11-18 10.5-12.5 (c)
0.95-0.98 0.S9-0.96(d) 0.96-0.97(d)
100-110 SO-100 SO-100
(0.45) (0.35) . (0.35)
400-700 100-150 100-150
(0. S) (0.16) (e) (0.16) (e)
0.03 0.01-.0.10 0.01-0.10
(0.05) (0.02-0:2). (0.02-0.2)
25 15 45
Limestone Limestone
'V 15(f) 3.0
1.0 NA
17-23 (b)
17-23
Vl
I\)
ex:>
EFFICIENCY
ENVIRONMENTAL FACTORS
6
S02' ppm (lb/10 Btu)
6
NOx' ppm (lb N02/10 Btu)
Particulates, gr/SCF
(lb/106 Btu)
3
Solid waste, ft /MW-day
0.03
(0.06)
S removal
NA
NA
NA
Basis: 3% sulfur,- 90% sulfur removal, 600 MW capacity~
Equipment modifications are required when converting from gas or coal to low-sulfur oil.
See Tables 2-9 through 2-12 for details. . .
Overall efficiehcy is dependent on mode of temperature control; see Table 2-8..
0.01 figure based on installing electrostatic precipitator (ESP). .
f 0.1 figure based on installing high-efficiency cyclone before burners and no ESP.
( ) CalS ratio dependent on regeherator temperature control scheme.
Stone
CalS
Make-up Ca/S
(a)
(b)
(c)
(d)
(e)
-------
Preliminary Design Conclusions
Preliminary design conclusions listed in the following matrix
are based on the technical investigations conducted.
Though the analysis
has raised many technical questions, the preliminary investigation
coupled with the experimental work of Esso (England) points out the
general feasibility of oil gasification/desulfurization as a retrofit
S02 control system for utility boilers.
529
-------
PRELIMINARY DESIGN CONCLUSIONS
. ..
TECHNICAL PROBLEM AREL\S
Sulfur recovery
Degradation of bo~ler performance and boiler modifications
Calcium sulfate generation
Temperature control
Lime circulation
CURRENT OUTLOOK
Overall, once-through operation appears superior to regenerative if high
stone utilization can be achieved, though this is not definite.
(See page 526).
Gasifier temperature control by means of ,the minimum air/fuel ratio
appears superior to other methods, if it can be achieved in
practice. Temperature control with heat transfer surface also
may be attractive if high air/fuel ratios (~20%) must be used
in operation.
The regenerator temperature is best controlled by stone circulation
if carbon deposition is small. If carbon deposition is large,
then H.T.S. is the most efficient method.
Sulfate generator temperature is most easily controlled by circulating
excess air through the generator.
The situation concerning the internal and external design concept is
not clear (page 507 ) with the number of variable factors
involved. It seems reasonable to strive to minimize the
expense and time involved in boiler modifications.
The greatest present market for oil gasification/desulfurization
appears to be in retrofits. The problems involved with
retrofits should be examined in more detail.
Oil gasification/desulfurization concepts appear highly competitive with
other retrofit or new S02 control methods based on preliminary
designs (Table 2-13) and should be investigated in greater
detail. (See recommendations).
530
-------
RECOMMENDATIONS
The performance and cost for a utility boiler add-on system
are competitive with stack gas cleaning and with the purchase of clean
oil.
Based on this technical assessment, we recommend that
.
The market for oil for power generation beyond 1975-1980
continue to be assessed in order to evaluate. the market
potential for atmospheric and pressurized oil gasification.
.
Utilities be contacted and a utility partner selected for
a demonstration installation of an add-on oil gasification
unit.
.
A preliminary design of the demonstration plant be prepared.
.
Research and . ~velopment work be continued in key areas --
e.g., carbon deposition, temperature control, load range,
sulfur recovery, pressurized operation.
531
-------
3.
FLUIDIZED BED COMBUSTION DEVELOPMENT PLANTS
DEVELOPMENT PLANT ALTERNATIVES
Three advanced fluidized bed power generation systems and an
industrial fluidized bed steam generation system have been studied;
these system concepts are summarized in the following matrix. Experi-
mental development and systems evaluation have been carried out for all
three.
FLUIDIZED BED FUEL PROCESSES FOR STEAM-POWER GENERATION
FLUIDIZED BED COMBUSTION
FLUIDIZED BED GASIFICATION
Atmospheric
Pressure
.Replacement for conventional
utility boiler with S02
stack gas cleaning system
and NOx control burning
high-sulfur coal or oil.
.Substitute for S02 stack
gas cleaning system and
NOx control on utility
plant bUfnfng high-sul-
fur oil. a
10-30 Atm
Pressure
.Industrial boiler burning
high-sulfur coal.
~Pressurized fluid bed
boiler operated with
combined gas and steam
turbine-generators
burning high-sulfur coal
or oil.
(a) New plant or retrofit on Existing plant.
All of these systems offer potential benefits, and each system
might be developed independently. However, the selection of the system(s)
to be developed requires the assessment of both technical and nontechnical
factors. The potential for pollution abatement, ability to use readily
available fuels, promise for economical steam/power generation, present
state of development, technical problem areas, advance design potential,
533
-------
and projected development time and cost have been assessed for each
system and compared with conventional fossil fuel-fired steam/power
generating plants utilizing present pollution abatement technology.
, ,
Other factors, such as the level of funding, alternative technological
advancements, future energy usage, future fuel prices, and future reg-
u1ations, e.g., emission standards and resource allocations, must also
be considered.
In order to recommend a fluidized bed fuel processing deve10p-
,ment plant, the respective system evaluations and experimental deve10p-
ment and market projections were studied with respect to application --
utility/industrial, new/retrofit, and technology -- combustion/gasifica-
tion and pressurized/atmospheric pressure. Recommendations are based
on the following assumptions:
.
Available funds will be limited
Alternative technological advancements will not displace
.
.
the use of fossil fuels as a major electric utility fuel
in the next 50 to 100 years
Utility concepts should be demonstrated
to 15 years
Future fuel
within the next 10
.
combustion emissions standards for 802' NO ,
, x
and particulates will require emissions less than or equal
to the presently proposed ,national standards.
On the basis of the evaluations and the assumptions, it is recommended
that
.
Pressurized fluidized bed combustion be developed for new
electric utility plants
.
A pressurized fluidized bed combustion development plant
be constructed and operated
A retrofit atmospheric pressure oil gasification demonstra-
.
tion plant be constructed and operated by a utility.
534
-------
PRESSURIZED FLUIDIZED BED COMBUSTION BOILER DEVELOPMENT PLANT
Objectives
The pressurized fluid bed combustion boiler program at BCURA
demonstrated the feasibility of pressurized fluid bed combustion in an
8 ftZ combustor. Specifically, it demonstrated:
.
SOZ emissions of ~ 0.7 lb/l06 Btu
6
NOx emissions of 0.07-0.Z lb NOZ/IO Btu
Coal feeding at 3-l/Z and 5 atm
Continuous operation - runs up to 350 hours terminated by choice
No erosion or corrosion of gas turbine blade test passages
.
.
.
.
.
after ZOO-hour tests
Particulate removal at 'V l500°F and 5 atm by cyclones proved
.
adequate for turbine blade tests
Boiler tube materials
.
Bed operated with horizontal tube bundle.
The Westinghouse-Foster Wheeler boiler design operating con-
ditions require that the bed be operated at higher gas velocity, higher
bed temperature, higher pressure, and with a deeper bed than the BCURA
unit. Each bed area for a four--module 300 MW plant is 35 ftZ.
The development plant is required to investigate the performance of the
proposed boiler plant equipment system design at the proposed operating
conditions for the following features:
.
SOZ and NOx emissions
Operation of deep beds (10 to ZO ft) with horizontal tube
bundles and headers
.
.
Particulate control equipment
Multiple coal feeding andsorbent distribution in a large
bed ('V 35 ftZ)
.
.
Materials and component life, including boiler tubing and
gas turbine blades
535
-------
.
Sorbent regeneration processes and compatibility of su1fur-
rich gas with sulfur recovery processes
Operational techniques - start-up, shut-down, load follow
.
.
Long-term operability
Sorbent circulation.
.
Other objectives which should be considered when planning the facility
include:
.
Need to test sulfur recovery system
Feasibility of studying advanced concepts - higher steam
.
.
temperature and pressure, higher gas turbine temperatures,
circulating beds, deeper beds (30 ft), higher pressures
Feasibility of studying components of a pressurized gasifica~
tion system
.
Feasibility of expanding to multiple bed operation.
The development plant should be designed'so that sufficient information
can be obtained to design, build, and operate a demonstration plant of
150 to 300 MW.
536
-------
Plant C
-------
in the heat transfer surface.
The combustor could operate as either an
evaporatort superheatert or reheatert but could not perform all functions
together. The plant requires ~ 10 tons of coal per hour and would be
designed to handle up to ~ 500tOOO lb/hr of water or steam.
The location of the plant is an important consideration, since
large quantities of coal and water are required. Several advantages
could be achieved by locating the development plant on an existing power
plant site:
.
Availability of coal handlingt water preparationt and
solids disposal facilities
Existing plant would dispose of steam
.
.
Superheat and reheat steam generation can be studied by
using bleed stream from existing plant
Utility partnership would minimize site development and
development time, and provide plant utilities and main-
.
tenance facilities.
A partnership with a utility is therefore recommended.
538
-------
Canceptual Design
Flaw Diagram and .Material Balance
Figure 3-1 i~ a flaw diagram af the pressurized fluid bed cam-
bustian bailer develapment plant. The plant is adjacent to. a large
pawer plant to. which it has access to. pravide interfaces with the caal,
water, steam, waste stane, and utilities. The facility is designed to.
aperate aver the range af aperating canditians presented an page 537
The material balance and aperating
ment plant at the prajected 100% 10. ad design
Table 3-1. The material balance is based an
canditians far the develap-
canditians are presented in
the caal and dalamite cam-
pasitians presented in the specificatian list an page 542 and Table 3-2.
The bailer is designed to. handle up to. 500,000 lb/hr af steam flaw. The
unit wauld functian as either a pre-evaparatar, superheater, ar reheater.
Water evaparatian in the walls cauld also. be studied.
Equipment
Bailer. The bailer is a single fluidized bed with a carban burn-
up cell. The design is based an the ane prapased by Westinghause and
Faster Wheeler. The water-walled, fluidized bed bailer wauld be placed
in a pressure shell larger than required far cammercial aperatian. A
vessel 15 to. 20 feet in. diameter and 50 to. 70 feet high is prajected to.
permit easy access, flexibility far madifications, and aperatian aver a
wide range af aperating canditians.
Caal and Dalamite Handling System. Caal-receiving equipment and
starage is assumed to. be available fram the pawer plant.. Equipment
specified far the develapment plant includes .a,.-dalamite receiving happer t
dalamite starage silo., dalamite crusher, caal dryert and caal crusher.
The maisture cantent af the salids must be maintained at less than 3% to.
539
-------
Dolomite Receiving
H~pper
CD
Coal from
Power-PI ant
Storage
VI
-I='"
o
Crusher
Storage
Bin
Storage
I nj ector
Coal
I nj ector
- ,Transport
Ai r from
, Compressor
To Stack or
Scrubber
Dwg. 2949,11.35
Storage
Silo
Q)
Pressurized
Fluid Bed
- Bo i 1 er
Primary
CoIl ector
Secondary
Collector
-Turbine
Blade Test
Cascade
Waste
So 1 ids
ranspor
Ai r
Lock
Hopper
@ -
~mpressor
Steam
to Plant
Q)
@
Spent
Stone
Air
Regenerated
Stone
Natural
Gas
@
Regenerator Lock
System Hopper
Waste
Stone
Gas
Generator
Location: Existing Powe,r Plant
Fig. 3-1-Flow diagram for development plant facility'
@
To Plant
Stack
-------
TABLE 3-1
MATERIAL BALANCE - DEVELOPMENT PLANT
STREAN DESCRIPTION TEMPERATURE PRESSURE FLOW RATE COMPOSITION PARTICLE
of psia 1b/hr wt % SIZE
Coal from ambient 15 66,OOO(a) Page 1-1/2" x 0
power plan t
Sized coal "'100 15 66,000 (a) Page 1/4" x 0
Dolomite ambient 15 Table 3-2 1-1/2" x 0
Dolomite ambien t 15 6,OOO(a) Table 3-2 1/4" x 0
Coal/dolomite ~ 500 200 24,000 Streans 2 & 4 1/4" x 0
to boiler
Spent dolomite 1750 150 20,000 CaO-MgO 75 -1/4"
CaS04-MgO 25
Regenerated 1100 150 24,000 CaO-MgO 3 -1/4"
dolomite CaC03 -MgO 94
CaS04-MgO 3
8 CBC feed 1600 142 2,900 Ash 60 -1/32"
Carbon 30
Dolomite 10
Was te solids 1600 139 2,000 Ash 80 -1/32"
Ca rbon 6
Stone 14
10 Waste solids 1100 180 2,300 same as 7 -1/4"
FLOW RATE COMPOSITION SOLIDS LEADING
1b/sec ACFM mole % gr/SCF
Air ambient 15 75 56,600 N2 77.5 nil
11 02 20.9
"20 2.1
Air 600 150 75 11,100 same as 11 nil
11a
Flue gas 1700 145 73 21,800 C02 15 7-15
12 N2 74
from FBB CO 0.2
"20 9
02 2
NOx "'200 ppm
S02 "'150 ppm
Flue gas 1650 142 73 21,800 same as 12 "'1
13
from primary
collec tor
Flue gas 1750 142 8 2,380 same as 12 4-7
14
from CBC
Flue gas to "'1650 142 81 24,200 same as 12 5-8
15
secondary
collector
Flue gas to 1600 139 81 24,200 same as 12 <0.15
16
test passage
Flue gas to 81 same as 12 <0.15
17
stack
(a)Assumes crushers and dryer operate one shift per day.
541
-------
SPECIFICATION LIST
OHIO PITTSBURGH NO.8 SEAM COAL
(Source of data:
USBM, Pittsburgh, Pa.)
SAMPLE:
Run of mine - as received
PROXIMATE ANALYSIS (wt %):
ULTIMATE ANALYSIS (wt %):
(includes moisture)
GROSS HEATING VALUE:
NET HEATING VALUE:
ASH ANALYSIS (wt %):
FUSIBILITY OF ASH:
PARTICLE DENSITY:
GRINDABILITY (~ardgrove):
FREE SWELLING INDEX:
Moisture
Volatile matter
Fixed carbon
Ash
3.3
39.5
48.7
8.5
100.0
C
H
°
N
S
Ash
71.2
5.4
9.3
1.3
4.3
8.5
100.0
(~60% organic; ~40% pyritic)
13000 Btu/lb
12500 Btu/lb
Si02 45.3
A1203 21.2
Fe203 27.3
Ti02 1.0
P20S 0.11
CaO 1.9
MgO 0.6
Na20 0.2
K20 1.8
S03 0.7
100.1
Initial deformation temperature
Softening temperature
Fluid temperature
2080°F
2230°F
2420°F
Coal
Ash
-- ~1. 4 gm/cc
-- ~2.8 gm/cc
50-60
5-5.5
542
-------
TABLE 3-Z
SOz SORBENT ANALYSIS USED FOR
PRELIMINARY FLUIDIZED BED BOILER ANALYSES
Component
wt % As Received(a)
Dolomite
1337
SiOZ
A1Z03
FeZ03
MgO
CaO
0.78
0.15
0.Z5
45.0
53.0
TiOZ
SrO
NaZO
KZO
MnOz
O.OZ
< 0.03
< O.OZ
< 0.1
< 0.03
(a) .
R. W. Coutant, J. S. McNultr., R. E. Barrett, J. J. Carson, R.
Fischer, and E.H. Lougher Investigation of the Reactivity of
Limestone and Dolomite for Capturing SOZ from Flue Gas,"
August 1968 (Battelle Memorial Institute).
543
-------
prevent agglomeration in the pressurized coal injection equipment. A
fluidized bed dryer is therefore included in the coal handling system to
reduce the moisture picked up during handling and storage to ~ 3%.
Since the dolomite is received by covered railroad car and is assumed to
contain less than 3% moisturet no drying of the dolomite is required.
Re-
versible hammermill crushers are used to reduce the coal and dolomite from
l-l/Z in x 0 to 1/4 in x O.
Coal Feed System.' The Petrocarb pressurized coal feed system is
used to feed the coal and make-up dolomite to the fluid bed boiler. The
system includes a storage bin, storage injector, and coal injector which
supplies coal and dolomite to the bed at four locations.
points can be installed if required.
More feed
Particulate Removal. The primary particle collector is a cyclon~.
designed to achieve 95% particle collection efficiency by weight. The
secondary collector is a tornado unit designed to achieve 97% to 99% col-
lection efficiency.
Turbine Blade Test System.
Initial tests will be carried out
with a turbine blade test cascade.
The second phase of operation will
include a rotating gas turbine test unit.
Air Compressor.
An air compressor with motor drive, independent
of the gas turbine test system, has been selected for the development
plant.
Regeneration System.. The present conceptual design employs the
two-step process specified in the pressurized fluid bed boiler power
. '. I '. . .'
plant design.
However, additional work on regeneration is recommended
prior to construction of the regeneration system.
Auxiliary equipment cost estimates are based on the two-step
process. Equipment includes a reducer vessel, H2S generator vessel,
solids handling system, gas generator for producing a reducing gas, and
a source of COZ/HZO gas. The reducer vessel and HZS generator vessel
would each be approximately two feet in diameter. The reducing gas re-
q~irements would be approximately 2500 SCFM of gas, composition 16% CO,
8% HZ' 7% COZ' 8% HZO, 59% NZ. This gas would be supplied by a gas gene-
544
-------
rator burning natural gas. 'The C02/H20 gas requirement would be approxi-
mately 2000 SCFM of gas, composition 36% C02 and 64% H20. This gas is
supplied by purchasing merchant C02 and blending it with stearn. Absorp-
tion of C02 from the power plant stack gases should also be considered
if this regeneration process is selected for further development.
Space Requirement
The development plant will require approximately 90 ft x 55 ft,
or ~ 5000 ft2. This does not include space for coal storage, stack, or
waste solid storage.
Power Requirement
The estimated power requirement for pumps, compressors, and
crushers is 23,000 hp.
Experimental Program
The development plant program is broken down into three phases
of operation:
Phase I.
Operation of fluid bed boiler with turbine blade test
cascade and without sorbent regeneration procesa
Phase II.
Operation of integrated system, . including gas turbine
and regeneration process
Phase III.
Operation to evaluate alternative concepts, e.g., re-
generation processes, bed design, stearn conditions.
The first phase would provide information to assess the pro-
posed design concept--boiler and carbon burn-up cell. Design and operating
characteristics studied during Phase I operation include:
.
Operation of deep beds with proposed heat transfer sur-
face design
.
Temperature gradients, particle circulation, particle
carry-over, heat transfer coefficients, combustion
efficiency, S02 removal, and NOx minimization at the
projected operating bed height, temperature, pressure,
particle size, and excess air
545
-------
o
Coal feeding system
Particulate control equipment
o
.
Grade efficiencies, pressure drop, reliability
Gas turbine blade erosion/corrosion/deposition
G
.
Boiler tube material corrosion/erosion/deposition
Boiler start-up, shut-down, and load follow.
.
The second-phase operation would incorporate design modifica-
tions recommended from the analysis of Phase I data, and the'parallel
research and development efforts include a regeneration process and a
small gas turbine unit.
A sulfur recovery process should also be con-
Design and operating charac-
sidered if development effort is required.
teristics studied during Phase II include:
.
Regeneration process
Operation of the integrated system
8
.
Solids circulation, sorbent stability and activity,
and feasibility of sulfur recovery
.
Gas turbine blades and boiler tubes
Integrated system operation--start-up, shut-down,
.
load follow.
Phase III would assess alternative concepts.
Development of
the third-phase program will rely on the results of the analysis of data
from Phases I and II, results of pilot and bench-scale experimental pro-
, studies.
grams., and recommendations from systems evaluation and advanced concept
This phase might include the evaluation of higher gas turbine
temperatures, higher steam temperatures and pressures, alternative re-
generation/sulfur recovery processes, and circulating bed boiler concepts.
Cost Estimate
The estimate for the development plant equipment, installed
out-of-doors adjacent to a large central station power plant, is
$7,000,000.
presented in Table 3-3.
A cost breakdown with the sources of the respective costs is
546
-------
TABLE 3-3
DEVELOPMENT PLANT COST ESTIMATE
EQUIPMENT
Coal and dolomite
handling system
Coal Feed System
Storage bin
Storage inj ector
Coal injector
Pressurized FBB
Boiler internals
Shell
Particulate removal
Primary
Secondary
Turbine blade test
cascade
Air compressor and
motor
Regeneration system
Sub total
Engineering, piping,
erection, installation,
instrumentation, controls
TOTAL
DESCRIPTION
Hopper(s)
Storage silo(s)
Crusher(s)
Dryer
Conveyors
10' x 10' x 35' ht
6' cp, 15' h t
8'
-------
Development Plant Schedule
A development plant schedule is presented in Figure ]-2.
projected time required to assess the combustor design, develop a
The
regenerator process, and obtain data on turbine blade performance is
four years.
The plant could then be used to study alternative boiler
and regeneration system concepts.
The schedule includes time for selecting a site and design con-
tractor. A utility site is preferred. The design is broken down into a
preliminary design phase, followed by the detail design. Construction
is estimated to take one year. The plant would be operated for two years
to obtain the necessary information for the design of a demonstration
plant.
The operation would be divided into two phases:
the first would
comprise operation of the boiler, including the carbon burn-up cell,and
a turbine blade test cascade; the second would study operation of a re- .
generation process and the integrated boiler-regeneration system.
A gas.
turbine would also be installed to obtain data on a rotating unit. .
A third phase of operation is recommended to evaluate alterna~
tive concepts for the boiler, regeneration process, and sulfur recovery.
This phase is not considered necessary for design of a demonstration
plant, however.
548
-------
1972
Jan
I I r I
Preliminary Design
t----4
Detail Design
1973
Jan
I J I I I I
Construction
Start-Up
Phase I Operation-Combustor/Turbine Blade Cascade
V1
"""
'"
Modifications
Modification s
Phase II Operation-Combustor/Regeneration/Gas Turbine
Phase III Operation-Alternate Concepts
1974
Jan
I I I I I r
1975
Jan
I I I I
I
I----i
I
Fig. 3-2-10-30 MW pressurized fI uid bed boi ler development plant schedule
D"9. 7236399
I I I I I I
J---f
1976
]in
1977
Jan
I I I I I
~-----i
-------
ATMOSPHERIC PRESSURE FLUIDIZED BED OIL GASIFICATION-COMBUSTION
DEMONSTRATION PLANT
The preliminary design of the gasification/desulfuriz8tion
retrofit demonstration plant is based on the conceptual designs and
the material and energy balances presented in Section 2.
Further
experimental information for a preliminary design will be provided by
the operation of the Esso (England) continuous pilot plant unit. The
specific design of the utility boiler to be retrofit and the preliminary
design will determine the retrofit concept and operational concepts to
be applied to the demonstration plant. The cost of the development
plant and the experimental program to be followed will depend largely
on the specific design concepts applied (internal or external design,
once-through or regenerative operation, temperature control schemes)
and the characteristics of the utility boiler (size, fuel normally
utilized, location, specifics of boiler design).
Boiler capacities in
the 50 to 200 MW range should be considered for the demonstration plant.
550
-------
Objectives
The objective of the atmospheric pressure fluidized bed oil
gasification-combustion demonstration plant is to demonstrate the
technical and economic feasibility of the oil gasification/desulfuriza-
tion concept in a utility boiler.
The preliminary and detailed design'
phases will provide information about the costs and problems involved
in boiler modification for the specific boiler from which general mod-
ification information can be implied. The experimental information
gained from the operation of the plant will indicate the performance
of the system (sulfur removal, NO , particulates, boiler performance,
x
operating cost, etc.) and will permit the design of commercial retrofit
units for boilers with characteristics other than those of the develop-
ment boiler.
551
-------
Plant Concept
The plant concept
to be retrofit, the outlook
applied will depend-on the specific boiler
of the utility partner, and the findings
of the preliminary design phase.
At present, most of the concepts
proposed in Section 2 are possibilities for the demonstration
operation. The alternative of retrofitting with the internal
external design concept may be quickly decided once a utility
plant
or ,the
partner
is located, while the temperature control schemes to be applied to
the demonstration plant may be chosen on the basis of the results from
the Esso (England) pilot plant unit. The mode of operation (regener-
ative or once-through) to be applied may, be a more difficult choice.
The two alternatives will be evaluated with the utility and during
the preliminary design phase.
552
-------
Conceptual Design
Both regenerative and once-through operations are possibilities
for the demonstration plant.
The factors involved in the two modes of
operation with respect to the demonstration plant will differ; and the
equipment involved in the system, the specific experimental program
carried out, the development plant cost, and the estimated time schedule
for the development plant program will depend on which mode of operation
is selected.
Flow Diagrams
Flow diagrams for regenerative and once~through operations
were shown in Figure 3-1. The components in the flow diagrams are
those discussed in the conceptual designs of Section 2. Conceptual
plant layouts, based on these flow diagrams, are shown in Figures 3-3
through 3-6 for a 200 MW regenerative operation and a 200 MW once-
through operation. The figures are drawn roughly to scale, with each
gasification system consisting of two modules, though, a demonstration
plant would use only one such unit. The figures show the essential
vessels, fans, and streams, and the relative configuration of these
components in the regenerative and once-through operations with the
gasification/desulfurization system located external to the boiler.
Equipment
Equipment and boiler modifications are listed in the next matrix.
Equipment normally present in an existing coal-fired boiler and an
existing oil-fired boiler is. indicated to show how. capital investment
might differ in the two cases. Boiler modifications required only with
the internal design concept or the external design concept are also
noted. Finally, the equipment needed only in the regenerative operation
or only in the once-through operation is differentiated. Equipment
present in both modes of operation, such as fans and lime transport.
line~ may differ in each operation, while some equipment, such as burners,
will be identical in the two modes.
553
-------
V1
V1
-I='
200 MW Boiler Retrofit-
External Regenerative Design
1. Gasifier-Desulfurizer
2. Regenerator
3. Sulfated-Lime Bunker
4. Limestone Bunker, Feeder
5. Forced Draft Fan
6. Gasifier Fan
7. Regenerator Fan
8. Oil Feed Li ne
9. Fuel Gas Li ne
10. Regenerator Product Li ne
It Li me Ci rculation Li ne
12. Sulfated-Lime Disposal Line
13. Burners
- Solids Transport
lZZ& Gas Transport
10
4
Furnace
/
/
8
6
5
12
7
~
..
20 Ft
. Fig. 3-3-Elevation of regenerative design-200 MW retrofit
Dwg. 2960A96
-------
~
'"
200 MW Boiler Retrofit-
External Regenerative Design
1. Gasifier-Desulfurizer
2. Regenerator
3. Sulfated-Lime Bunker
4. Limestone Bunker, Feeder
5. Forced Draft Fan
6. Gasifier Fan
7. Regenerator Fan
8. Oil Feed Line
9. Fuel Gas Line
10. Regenerator Product Li ne
11. Lime Circulation Line
12. Sulfated-Lime Disposal Line
13. Burners
- Solids Transport
~ Gas Transport
4
10
8
12
3
20 Ft
Fig. 3-4-Pla n of regenerative design - 200 MW retrofit
D'''J. 723B1131
Furnace
6
-------
V1
V1
0'.
200 MW Boiler Retrofit-
External Once-Through Design
1. Gasifier-Desulfurizer
2. Sulfate Generator
3. Sulfated-Lime Bunker
4. Limestone Bunker, Feeder
5. High-Efficiency Cyclone
6. Forced Draft Fan
7. Gasifier Fan
8. Sulfate Generator Fan
9. Oil Feed Li ne
10. Fuel Gas Li ne
11. Lime Ci rculation Li ne
12. Sulfated-Lime Di sposal Li ne
13. Burners
- Solids Transport
?ZZZZ Gas Transport
9
Furnace
./
./
,
,
,
-----
4
- - -_/
/
/
/
12
...
.
20 Ft
Fig. 3-5-Elevation of once-through design-200 MW retrofit
Dwg. 2960A97
6
-------
VI
VI
~
200 MW Boiler Retrofit-
External Once-Through Design
1. Gasifier-Desulfurizer
2. Sulfate Generator
3. Sulfated-Lime Bunker
4. Limestone Bunker, Feeder
5. High-Efficiency Cyclone
6. Forced Draft Fan
7. Gasifier Fan
8. Sulfate Generator Fan
9. Oil Feed Li ne
10. Fuel Gas Li ne
11. Lime Circulation Line
12. Sulfated-Lime Disposal Line
13. Burners
- Solids Transport
m:zz Gas Transport
9
12
D"'9. 2960A98
I 1
I I
I ~II
10
I~I
I gl
lul
4
Furnace
13
. .-
20 Ft
Fig. 3-6-0nce-through design-200 MW retrofit
-------
DEVELOPMENT PLANT EQUIPMENT
Gasifier vessel
Regeneratort (1) sulfate generator(2)
Limestone handling
Limestone(~ynker(3)
Conveyors
Feeders
Sulfided lime bunker
Lime transport lines
Lime disposal line
Storage
(4)
Fuel oil 3
Limes tone ( )
Sulfur(l)
Sulfided lime
disposal storage(3)
Particulated clean-up
Cyclones
Electrostatic precipitator(3)
Fans, pumps
Fue 1 0 il ( 4 )
Gasifier
Regenerator,(l) sulfate
Pneumatic transport
Burners
Sulfur recovery. plant(l)
generator (2)
Piping and ducts
Foundations, structures, electricalt instrumentation, control, experimental
equipment
Boiler modifications (removal or alteration of)
Ash pit(S) .
Pulverizers, pumpst external equipment to be relocated(lt 6)
Boiler water walls, base of boiler(S)
Boiler burner ports, structures
(1) Regenerative operation only
(2)once-through operation only
(3)Equipment already present in coal-fired boiler
(4)Equipment already present in oil-fired boiler
(S)Modification with internal design only
(6)Modification with external design only
558
-------
Boiler Capacity
A development plant with a capacity of 50
to 200 MW may be the
best size in that it is representative of the boiler retrofit market
and minimizes the scale-up factor.
For boiler sizes less than about
50 MW the overall boiler design differs greatly from that of boilers in
the 50 to 800 MW range. The 50 to 200 MW capacity boiler may also
be more available than a smaller capacity unit. Although the scale-
up from the ~ 1 MW pilot plant unit of Esso (England) to a 50 to 200 MW
system is considerable, this represents the minimum scale-up factor
w~ich will give meaningful demonstration results within the next. five
years.
Experimental Program
The demonstration plant should be equipped to monitor the con-
ditions of major streams and vessels to demonstrate the overall perform-
ance of the atmospheric pressure fluidized bed gasification-combustion
system.
The following areas should be emphasized in investigations
with regenerative or once-through operation:
.
Boiler Performance:
Steam conditions
Combustion Characteristics of flame
Long term erosion/corrosion of boiler and. burners
.
Gasification/Desulfurization System Performance:
Performance of temperature control schemes
Gasifier thermal efficiency
Fuel and solids distribution in gasifier
Sulfur recovery-regeneration performance (regenerative
mode)
Sulfate generator performance (once-through mode)
Solids circulation and control
Long term erosion/corrosion, material build-up in
system
559
-------
~
System Performance:
Overall system efficiency
Optimum operating conditions for a given sulfur removal
(in terms of limestone make-up rate, bed depth,
particle size)
Pollution control - NO
x'
S02' particulates, metal
retention
Effects of turn-down on system performance
Overall system operation.
The experimental program will be determined on the basis of the concept
and the availability of the plant for experimental work.
Cost Estimate
Major equipment for a 100 MW demonstration plant (regeneratiye.
operation) are estimated in Table 3-4:
the cost of instrumentation,
control, and piping for the demonstration plant has been taken as a
factor twice that expected with a commercial plant, or 100% of the
equipment cost rather than 50%, because of the need for flexibility in
the demonstr~tion plant functions and the additional costs associated
with the experimental nature of the system. Engineering costs are also
assumed to be a larger factor with the demonstration plant than with a
commercial unit. The regenerative demonstration plant is estimated to
cost about $80/kw, while a once-through development plant of 100 MW
capacity would cost about $60/kw, or $6,000,000. The major difference
in capital cost for the regenerative and once-through operations is
associated with the cost of the sulfur recovery plant.
Demonstration Plant Schedule
The schedule for the atmospheric pressure oil gasification/
. .
desu1furization development plant program is shown in Figure 3-7.
560
-------
TABLE 3-4
DEVELOPMENT PLANT COST ESTIMATE
ITEM
ESTIMATED COST, $
,Gasifier
Regenerator
Limestone handling equipment
$
250,000
35,000
Particulate control
Fans
800,000
250,000
300,000
100,000
300,000
$2,035,000
Burners
Boiler modification
Equipment cost
Structures, piping, electrical, instrumentation,
Control z 100% of equipment
Direct cost
2,035,000
$4,070,000
Engineering, construction z 50% of direct cost
Sulfur recovery plant
Total cost of demonstration plant
2,035,000
2,000,000
$8,105,000
561
-------
O>lY. 7236675
FIG. 3-7-DEMONSTRATION PLANT SCHEDULE
Atmospheric Oil Gasification/Desulfurization Process for Existing Steam Power Plants
1912 1973 1974 1975
Jan. April July Oct. Jan. April July Oct. Jan. April July Oct. Jan.
I I I I I I I I I I I I I
Cost
Contact Utilities
Select Engineering Firm
Preliminary Design
..3
~
$ 30:000 a
$ 6,000,000 b
$ 24, OOO/day c
6. 5~/106 Btu fuel adder d
I\¥
Detailed Design and Installation
Developmental Operation
Conti nuous Operation
w
'r
'"
'"
f\)
Mi I estones
W Cooperating utility identified
W Engineering firm contract signed for preliminary design
W Decision on construction; agreement between OAP - utility - @ - AE
W shake down complete
Vi Engineering and economic data for future operations complete
Foot not es
a Cost estimate for preliminary design .
b Assumes complete 100 MW gasifier module, once-through limestone system
c Maxi mu m cost per day assu mes conti nuous oil feed, 0 & M for gasifier development, and no power generation
d Based on conceptual design
-------
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CONTRACT PUBLICATIONS
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for Fluidized Combustion Boilers," paper presented at Second Inter-
national Conference on Fluidized Bed Combustion. Hueston Woods,
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571
-------
ACKNOWLEDGMENTS
The results, conclusions, and recommendations presented in
this report represent the combined work and thought of many persons at
Westinghouse, at the Office of Air Programs (OAP), and elsewhere.
Westinghouse personnel in many divisions throughout the Corporation have
contributed.
Our subcontractors -- Erie City, Foster Wheeler, and
United Engineers and Constructors -- have added their expertise in
boiler and power plant designs. Other OAP contractors have freely
shared with us their ideas and the results of their research and
development effort. Suppliers of boiler and power plant auxiliary
equipment have discussed with us their products and provided cost
estimates. Other commercial firms with background in fluidized bed
technology have received us, shown us their work, and commented on ours.
We have attempted to recollect all these contributions and what is
undoubtedly (and regrettably) a partial list of organizations and
persons is presented on the following pages.
In particular, however, we want here to express our high
regard for and acknowledge the contribution of the personnel at OAP who
conceived the overall fluidized bed combustion boiler effort 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
D. Bruce Henschel and Sam Rakes; and with R. P. Hangebrauck, Chief of
the Demonstration Projects Branch.
573
-------
The following organizations and groups acted as subcontractors:
Foster Wheeler
Corporation
United Engineers
and Constructors
Inc.
Erie City Energy,
Division of Zurn
Industries
Westinghouse Divisions
Heat Transfer
Small Steam and
Gas Turbine
Power Systems
Generation Planning
Computer and
Instrumentation
Research Laboratories
. Assisted in evaluating
utility fluidized bed boiler
design concepts
. Prepared preliminary designs
of preferred concepts
. Prepared cost estimates
. Prepared plant layout,
operating procedure, and
plant cost for a
pressurized fluidized
bed combined cycle plant
. Compiled industrial boiler
market survey
. Assisted in evaluating
industrial fluidized bed
boiler design concepts
. Prepared preliminary
design of preferred concept
. Prepared cost estimate
. Prepared cost estimates for
regeneration system
vessels and stack gas coolers
. Prepared design
and cost estimate
for external combustor
gas turbine
. Prepared utility
market survey
. Provided power
system cycle
analyses
. Prepared control system
design and cost for
industrial boiler
. Gas Turbine Erosion
and Corrosion
574
R. W. Bryers
J. D. Shenker
R. Zoschak
M. Casapis
E. Berman
W. Craig
J. Crowley.
R. V. Seibel
W. Schwinden
R. Giardina
S. Jamison
R. E. Strong
B. Hugoson
S. J. Jack
R. R. Boyle.
R. J. Budenholzer
H. L. Smith
N. Weeks
R. Gessford
W. E. Young
-------
Sturtevant
Large Turbine
. Supplied design and cost
data for forced and
induced draft fans
. Supplied cost information
on steam turbines
R. Mercer
W. F. Courtney
The following vendors supplied information and cost estimates:
Aerodyne Development
Corporation
Petrocarb, Inc.
McNally~Pittsburg
Manufacturing
Corporation
Struthers Nuclear
and Process Company
Ducoh Company, Inc.
Ford, Bacon, and
Davis, Inc.
Ralph M. Parsons
Company
Wheelabrator
Corporation
Buell Engineering
Company, Inc.
Koppers Company, Inc.
Process Combustion
Corp./ Bloom
Engineering Co.
. Supplied design and cost
information for high
efficiency particulate
removal equipment
. Supplied design and cost
information for pressurized
solids feeding systems
. Supplied design and cost
information for solids
handling systems
. Submitted proposal for
stack gas coolers
. Supplied cost information
for particulate removal
. Supplied cost estimates for
sulfur recovery processes
. Supplied cost estimates
for sulfur recovery
processes
. Supplied electrostatic
precipitator cost
. Supplied electrostatic
precipitator cost
. Supplied design information
on solids feeding systems
. Supplied design and cost
information on low Btu gas
burner for boiler
575
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The following facilities were visited by Westinghouse
personnel:
Argonne National Laboratory
Chicago, Illinois
The Badger Company, Inc.
Boston, Massachusetts
Battelle Memorial Institute
Columbus, Ohio.
Bituminous Coal Research
Monroeville, Pennsylvania
Chicago Bridge & Iron Company
Research Center
Plainfield, Illinois
Consolidation Coal Company
Pittsburgh, Pennsylvania
Dorr-Oliver Fluidized Bed
Sludge Incinerator
Liberty, New York
Esso Research and Engineering
Linden, New Jersey
Esso Petroleum Company
United Kingdom
Fluidized Bed Course
University of Birmingham, England
Fuller Company
Catasauqua, Pennsylvania
Kansas Power & Light
T9peka, Kansas
NAPCA/NCB Fluidized Bed Combustion Program
Second Information Exchange Meeting
Hobart House, London
BCURA, Leatherhead
CRE, Cheltenham
National Coal Board
United Kingdom.
Pope, Evans & Robbins
Alexandria, Virginia
St. Joseph Lead Company
Monaca, Pennsylvania
Tennessee Valley Authority
Chattanooga, Tennessee
ueI Corporation
Kingston, Pennsylvania
Union Electric
St. Louis, Missouri
U.S. Bureau of Mines
Bruceton, Pennsylvania
U.S. Bureau of Mines
Morgantown, West Virginia
West Virginia University
576
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