EPA-650/2-73-048b
December 1973
Environmental Protection Technology Series
-------
EP A-650/2-73-048 b
EVALUATION
OF THE FLUIDIZED-BED
COMBUSTION PROCESS
VOLUME II -
FLUIDIZED-BED BOILER COMBINED-CYCLE POWER
PLANT DEVELOPMENT-VOLUME I APPENDICES
by
D.L. Keairns, D.H. Archer* J.R. Hamm,
R.A. Newby, E.P. O'Neill, J.R. Smith, and
W.C. Yang
Westinghouse Research Laboratories
Pittsburgh, Pennsylvania 15235
Contract No. 68-02-0217
ROAP No. 21ADB-09
Program Element No. 1AB013
EPA Project Officer: P.P. Turner
Control Systems Laboratory
National Environmental Research Center
Research Triangle Park, North Carolina 27711
Prepared for
OFFICE OF RESEARCH AND DEVELOPMENT
U.S. ENVIRONMENTAL PROTECTION AGENCY
WASHINGTON, D.C. 20460
December 1973
-------
This report has been reviewed by the Environmental Protection Agency and
approved for publication. Approval does not signify that the. contents
necessarily reflect the views and policies of the Agency, nor does
mention of trade names or commercial products constitute endorsement
or recommendation for use.
-------
ACKNOWLEDGMENTS
The results, conclusions, and recommendations presented in
this volume represent the combined work and thought of many persons at
Westinghouse and the Office of Research and Development (ORD). Other
ORD contractors have'freely shared with us their ideas and the results of
their research and development effort.
In particular, we want here to express our high regard for
and acknowledge the contribution of personnel at Westinghouse Research
Laboratories and the personnel at ORD 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. F. P. Turner,
Chief of the Advanced Process Section, has served as project officer on
our work. Numerous enlightening and helpful discussions have been held
with Mr. Turner; with section members D. Bruce Henschel and Sam Rakes;
and with R. P. Hangebrauck, Chief of the Demonstration Projects Branch.
Mr. E. J. Vidt participated in the evaluation. Mr. R. E. Brlnza,
Mr. W. F. Kittle and Mr. C. Spangler assisted in the collection of data.
-------
PREFACE
The Office of Research and Development (ORD) 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 utility power plants. These methods should:
• Meet environmental goals for S0_, NO , ash, smoke emissions,
£ X
and wastes
• Compete economically with alternative means for meeting
these abatement goals.
Westinghouse Research, under contract to the Office of Research
and Development (ORD) of the Environmental Protection Agency, is carrying
out a study to evaluate and develop fluidized bed combustion and oil
gasification in air pollution abatement. The goals of this work are
• To identify which fluidized processes might be economically
employed in utility power plants or industrial boilers to
reduce SO., particulate, and NO emissions
£* X
• To assist in planning and implementing a program to
develop the fluidized bed systems deemed effective in
air pollution control and economical in steam/power production.
Tasks in the evaluation of fluidized bed combustion set forth
by EPA which have been completed at Westinghouse under a previous
contract are
• To search the technical and patent literature in fluidized
bed combustion, to canvass commercial organizations with
expertise pertaining to this field, and to survey the
market for industrial boilers and utility power systems
iii
-------
• To design a fluidized bed industrial boiler and two
fluidized bed utility power systems — one 300 MW
capacity, the other 600 MW; and to provide performance
and cost projections for the equipment designed
• To provide technical consultation and assistance on the
ORD fluidized bed fuel processing program, including
both combustion and gasification
• To conceptualize a development plant which will prove
the configurational or operating features of the fluidized
bed boiler designs not confirmed by prior art or experience
• To assess the effectiveness and economics of an
atmospheric-pressure, fluidized bed oil gasification-
combustion system and to aid in planning for a demonstration
installation of such a system.
The results of these surveys, designs, evaluations, and
comparisons were published in a three-volume report, "Evaluation of
the Fluidized Bed Combustion'Process," in November 1971 under contract
No. CPA 70-9.
These results provided the basis for the work reported in this
four-volume report. Tasks in the evaluation of the fluidized bed
combustion process set forth by EPA under contract 68-02-0217 have
focused on the development of pressurized fluidized bed combustion
for power generation and fluidized bed oil gasification for power
generation. These tasks, which are presented here, have included:
• Extensive process evaluation studies of the pressurized
fluidized bed combustion system for power generation.
These investigations predict the sensitivity of operating
and design parameters selected for the base power plant
design on plant economics; provide additional experimental
data on sulfur removal and sorbent regeneration using
iv
-------
limestones and dolomites; project economics and performance
for various sulfur removal systems, both regeneration/sulfur
recovery and once-through sorbent; establish a plant operation
and control philosophy and evaluate alternative pressurized
fluid bed combustion concepts, both economics and performance.
• Preparation of preliminary plans and a cost estimate for a
30 MW (equivalent) pressurized fluidized bed combustion
boiler development plant. The design provides sufficient
detail to locate a suitable plant site and to obtain a fixed
cost bid for the preparation of detailed plans. The plant
will provide the capability for studying the remaining
technical problems effectively in order to achieve the
greatest potential for reducing emissions and for generating
economical electrical energy.
• Identification of a project team to demonstrate fluidized
bed oil gasification/desulfurization for power generation.
A cooperating utility — New England Electric System —
has been identified to carry out a 50 MW demonstration
plant program. Further process evaluation has been
carried out and experimental data obtained on sulfur
removal and spent stone disposal.
• Evaluation of pressurized oil gasification for combined
cycle power generation. Oil gasification process concepts
and options have been reviewed, material and energy balances
projected, performance projected,and capital and energy
costs estimated.
• Provision of technical consultation and assistance on the
ORD 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.
-------
Volume I contains new data on pressurized fluid bed combustion
and an evaluation of the current technology state. This volume (Volume II)
contains the appendices to Volume I. Volume III presents the preliminary
design of a 30 MW (equivalent) pressurized fluid bed boiler development
plant. Volume IV includes the work on fluid!zed bed oil gasification/
desulfurization.
vi
-------
CONVERSION FACTORS — ENGLISH TO METRIC UNITS
Length
Area
Volume
Mass
Pressure
Temperature
Energy
Power
English System
in
ft
gal
oz
Ib
ton
lb/in2
in H20
°F
°R
Btu
Btu/min
Metric Equivalent
2.54 cm
0.305 m
6.45 cm2
0.0930 m
16.39 cm3
28.32 1
3.785 1
28.35 gm
453.6 gm
907.2 kg
51.70 mm Hg
1.865 mm Hg
1.8 (°C) + 32
1.8 °K
252 cal
252 cal/min
vii
-------
TABLE OF CONTENTS
ECONOMIC SENSITIVITY
A. Boiler Tube Specifications
B. Supplementary Cost Data
C. Parametric Study of Elutriation Rate from a
Fluidized Bed Combustion Boiler
SULFUR REMOVAL SYSTEMS
D. Base Design — High-Pressure One-Step Process
E. Base Design — Low-Pressure One-Step Process
F. Base Design — Two-Step Process
G. Temperature Control and Process Turndown
H. Effect of Boiler Conditions
I. Once-through Process
J. Constant Load Concept
K. Process Comparisons
L. General Study of One-Step Process
M. Low-Sulfur Coal
N. Limestone Wet Scrubber Cost
PLANT OPERATION AND CONTROL
0. Fart Load Operation
ALTERNATIVE FLUID BED BOILER CONCEPTS
P. Recirculating Bea Boiler Studies
ix
-------
LIST OF FIGURES
ECONOMIC SENSITIVITY Page
Appendix A
1. Maximum Allowable Stress for Different Tube Materials A-3
2. Maximum Required Tube Wall Thickness A-3
3. Minimum Required Tube Wall Thickness (Bed-Tube Heat A-7
Transfer Coefficient 35 Btu/ft2-hr-°F)
4. Minimum Required Tube Wall Thickness (Bed-Tube Heat A-7
Transfer Coefficient 75 Btu/ft2-hr-°F)
5. Minimum Required Tube Wall Thickness (Bed Temperature A-8
1407°F)
6. Minimum Required Tube Wall Thickness (Bed Temperature A-8
1522°F)
•7. Minimum Required Tube Wall Thickness (Bed Temperature A-8
1636°F)
8. Minimum Required Tube Wall Thickness (Bed Temperature A-8
1750°F)
Appendix B
1. Effect of Bed-Tube Heat Transfer Coefficient on the B-2
Steam Generator Cost (by Design Bed Temperatures)
2. Effect of Bed-Tube Heat Transfer Coefficient on the B-2
Steam Generator Cost (by Maximum Allowable Bed Depth)
3. Dependence of the Steam Generator Cost (318 MW) on B-3
the Maximum Allowable Bed Depth (Bed Temperature
1750°F)
4. Dependence of the Steam Generator Cost (318 MW) on B-3
the Maximum Allowable Bed Depth (Bed Temperature
1636°F)
5. Dependence of the Steam Generator Cost (318 MW) on B-3
the Maximum Allowable Bed Depth (Bed Temperature
1522°F)
6. Dependence of the Steam Generator Cost (318 MW) on B-3
the Maximum Allowable Bed Depth (Bed Temperature
1407°F)
xi
-------
Page
7. Dependence of the Steam Generator Cost (318 MW) B-4
on the Maximum Allowable Bed Depth (Heat Transfer
Coefficient 75 Btu/ft -hr-°F)
8. Dependence of the Steam Generator Cost (318 MW) B-4
on the Maximum Allowable Bed Depth (Heat Transfer
Coefficient 50 Btu/ft2-hr-°F)
9. Dependence of the Steam Generator Cost (318 MW) B-4
on the Maximum Allowable Bed Depth (Heat Transfer
Coefficient 35 Btu/ft2-hr-°F)
10. Effect of Bed-Tube Heat Transfer Coefficient on B-5
the Steam Generator Cost (1-in. O.D. Tubes)
11. Effect of Bed-Tube Heat Transfer Coefficient on B-5
the Steam Generator Cost (l-inf O.D. Tubes)
12. Dependence of the Steam Generator Cost on the B-6
Maximum Allowable Bed Depth (1-in, O.D. Tubes,
Bed Temperature 1750°F)
13. Dependence of the Steam Generator Cost (318 MW) B-6
on the Maximum Allowable Bed Depth (1-in. O.D.
Tubes, Bed Temperature 1636°F)
14. Dependence of the Steam Generator Cost (318 MW) B-6
on the Maximum Allowable Bed Depth (1-in. O.D.
Tubes, Bed Temperature 1522°F)
15. Dependence of the Steam Generator Cost (318 MW) B-6
on the Maximum Allowable Bed Depth (1-in. O.D.
Tubes, Bed Temperature 1407°F)
16. Dependence of the Steam Generator Cost (318 MW) B-7
on the Maximum Allowable Bed Depth (1-in. O.D.
Tubes, Bed-Tube Heat Transfer Coefficient 75 Btu/
ft2-hr-°F)
17. Dependence of the Steam Generator Cost (318 MW) B-7
on the Maximum Allowable Bed Depth (1-in. O.D.
Tubes, Bed-Tube Heat Transfer Coefficient 50
Btu/ft2-hr-°F)
18. Dependence of the Steam Generator Cost (318 MW) B-7
on the Maximum Allowable Bed Depth (1-in. O.D.
Tubes, Bed-Tube Heat Transfer Coefficient 35
Btu/ft2-hr-°F)
19. Effect of Bed-Tube Heat Transfer Coefficient on B-8
the Steam Generator Cost (2-in> O.D. Tubes)
20. Dependence of the Steam Generator Cost (318 MW) B-9
on the Maximum Allowable Bed Depth (2-in. O.D.
Tubes, Bed Temperature 1750°F)
xii
-------
Page
21. Dependence of the Steam Generator Cost (318 MW) on B-9
the Maximum Allowable Bed Depth (2-in. O.D. Tubes,
Bed Temperature 1636°F)
2'2. Dependence of the Steam Generator Cost (318 MW) on 5-9
the Maximum Allowable Bed Depth (2-in. O.D. Tubes,
Bed Temperature 1522°F)
2'3. Dependence of the Steam Generator Cost (318 MW) on 5-9
the Maximum Allowable Bed Depth (2-in. O.D. Tubes,
Bed Temperature 1407°F)
24. Dependence of the Steam Generator Cost (318 MW) on B-10
the Maximum Allowable Bed Depth (2-in. O.D. Tubes,
Bed-Tube Heat Transfer Coefficient 75 Btu/ft2-hr-°F)
25. Dependence of the Steam Generator Cost (318 MW) on B-10
the Maximum Allowable Bed Depth (2-in. O.D. Tubes,
Bed-Tube Heat Transfer Coefficient 50 Btu/ft2-hr-°F)
26. Dependence of the Steam Generator Cost (318 MW) on B-10
the Maximum Allowable Bed Depth (2-in, O.D. Tubes,
Bed-Tube Heat Transfer Coefficient 35 Btu/ft -hr-"F)
Appendix C
1. Different Combinations of Elutriation Rates for a C-3
Dust Loading Equal to that of Westinghouse-Foster
Wheeler Basic Design
2. Different Combinations of Elutriation Rates for a C-3
Dust Loading Double that of Westinghouse-Foster
Wheeler Basic Design
3. Different Combinations of Elutriation Rates for a C-3
Dust Loading Triple that of Westinghouse-Foster
Wheeler Basic Design
4. Possible Dust Loading from FBC at 0% Carbon C-4
Elutriation Rate
'5. Possible Dust Loading from FBC at 6% Carbon c-4
Elutriation Rate
6. Possible Dust Loading from FBC at 10% Carbon C-4
Elutriation Rate
7. Possible Dust Loading from FBC at 15% Carbon C-5
Elutriation Rate
8. Possible Dust Loading from FBC at 20% Carbon C-5
Elutriation Rate
SULFUR REMOVAL SYSTEMS
Appendix D
1. One-Step Regeneration Process Elements D-2
2. High-Pressure One-Step Regeneration Process Flow Diagram D-5
3. High-Pressure One-Step Regeneration Schematic Flow DiagramD-5
xiii
-------
Page
Appendix E
1. Low-Pressure One-Step Regeneration Process Flow Diagram E-3
2. Low-Pressure One-Step Regeneration Process Flow Diagram E-4
Sorbent Temperature Control
3. Low-Pressure One-Step Regeneration Schematic Flow Diagram E-5
Appendix F
1. Two-Step Regeneration Process Elements F-2
2. Two-Step Regeneration Material and Energy Balances F-4
3. Two-Step Regeneration Schematic Flow Diagram F-7
Appendix G
1. Dolomite Circulation Effect on Temperature Control — G-2
2000°F Regenerator
•2. Dolomite Circulation Effect on Temperature Control — G-2
1900CF Regenerator
3. Limestone Circulation Effect on Temperature Control — G-3
2000°F Regenerator
4. Limestone Circulation Effect on Temperature Control — G-3
1900°F Regenerator
5. Enthalpy of Regenerator Product Gas G-4
6. Gas Producer Characteristics G-5
7. Gas Producer Fuel Requirement G-7
6. Reducing Gas Temperature G-8
9. One-Step Process Turndown G-ll
Appendix H
1. Carbonating Boiler Conditions Flow Diagram H-2
2. Calcining Boiler Conditions Flow Diagram H-2
Appendix I
1. Sorbent Handling and Feeding Costs 1-2
2. Once-through Energy Cost 1-2
Appendix J
1. Constant Load Concept Flow Diagram J-2
•2. Constant versus Variable Load for the Low-Pressure J-3
One-Step Process
3. Constant versus Variable Load for the High-Pressure J-5
Processes
xiv
-------
Appendix K
1. Regeneration Process Investment Comparison K-2
2. Regeneration Process Energy Cost Comparison K-2
•3. Energy Cost Comparison: Disposal before Regeneration, K-3
$2/ton Dolomite
4. Energy Cost Comparison: Disposal before Regeneration, K-3
$5/ton Dolomite
5. Energy Cost Comparison: Disposal before Regeneration, K-4
$10/ton Dolomite
6. Energy Cost Comparison: Disposal before Regeneration, K-4
$10/ton Dolomite
7. Energy Cost Comparison: Disposal after Regeneration, K-6
$2/ton Dolomite
6. Energy Cost Comparison: Disposal after Regeneration, K-6
$5/ton Dolomite
9. Energy Cost Comparison: Disposal after Regeneration, K-7
$107ton Dolomite
10. Energy Cost Comparison: Disposal after Regeneration, K-7
$15/ton Dolomite
Appendix L
1. One-Step Regeneration Process Elements L-4
2. Regeneration Process Sulfur Load L-13
3. Tail Gas Recycle to Boiler L-14
4. Gas Producer Fuel Requirement L-15
5. Total Regeneration Process Fuel Requirement L-15
6. Effect of Sulfur in Gas Producer Fuel L-18
7. Sulfur Production Rate L-19
8. One-Step Regeneration Process Flow Diagram L-23
9. Regenerator Element Capital Investment L-26
10. Sulfur Recovery Element Capital Investment L-27
11. Tail Gas Handling Element Capital Investment L-28
12. Gas Producing Element Capital Investment L-30
13. Sorbent Circulation Element Capital Investment L-32
14. Energy Cost — Methane at 10 Atmospheres L-35
15. Energy Cost — Methane at 5 Atmospheres L-35
xv
-------
Page
16. Energy Cost — Methane at 2 Atmospheres L-35
17. Energy Cost — Oil at 10 Atmospheres L-36
16. Energy Cost — Oil at 5 Atmospheres L-36
19. Energy Cost ~ Oil at 2 Atmospheres L-36
20. Energy Cost — Coal at 10 Atmospheres L-38
21. Energy Cost — Coal at 5 Atmospheres L-38
22. Energy Cost — Coal at 2 Atmospheres L-38
23. Energy Cost Comparison — High-Sulfur Load L-39
24. Energy Cost Comparison — Medium-Sulfur Load L-39
Appendix M
1. Dependence of the Shell Cost on the Fluidizing Velocity M-5
2. Effect of Bed-Tube Heat Transfer Coefficient on the M-6
Steam Generator Cost
Appendix N
1. Energy Cost of a Conventional Coal-Fired Power Plant N-2
with Limestone Wet Scrubbing
PLANT OPERATION AND CONTROL
Appendix 0
1. Temperature-Energy Diagram for Pressurized Utility 0-4
Boiler
2. Part Load Characteristics 0-7
3. Effect of Design Point Temperature and Percent 0-8
Excess Air on Plant Turndown Capability
Appendix P
1. Detailed Schematic of the Two-Dimensional Cold Model P-2
2. Two-Dimensional Cold Model with Modified Draft Tube P-5
Inlet
3. Dimensions of the Serpentine Copper Rods P-7
4. Experimental Data for Case I and Case II P-10
5. Experimental Data for Case III P-ll
6. Experimental Values of Total Air By-Passing P-13
7. Schematic Representation of Flow Characteristics P-49
in Vertical Pneumatic Transport
xvi
-------
Page
8. Comparison between the Experimental Data and the P-50
Saturation Carrying Capacity
9. Pressure Loss Due to Slugging Transport P-52
xvii
-------
LIST OF TABLES
ECONOMIC SENSITIVITY
Appendix A
1. Boiler Tube Specification A-2
•2. Boiler Tube Specification for L-in, O.D. Tubes A-4
3. Boiler Tube Specification for 2-in, O.D. Tubes A-5
4. Design Metal Temperatures and Tube Material A-6
Appendix C
1. Design Bases for Westinghouse-Foster Wheeler Fluidized C-2
Bed Combustion Boiler (318 MW/four-module design)
SULFUR REMOVAL SYSTEMS
Appendix D
1. Design Specifications D-3
2. Design Assumptions D-3
3. Material and Energy Balance — One-Step Regeneration at D-6
High Pressure (1% S02)
4. Material and Energy Balance — One-Step Regeneration at D-7
High Pressure (2% S02>
5. 1% S02 Investment Breakdown, $/kw D-10
6. 2% SO„ Investment Breakdown, $/kw D-10
7. 1% S02 Energy Cost (mills/kWh) D-ll
8. 2% S02 Energy Cost (mills/kWh) D'11
9. Cost Sensitivity of High-Pressure One-Step Regeneration
Appendix E
1. Design Specifications g_l
2. Design Assumptions g_2
3. Material and Energy Balances - Low-Pressure One-Step E-6
Regeneration
4. Low-Pressure One-Step Process Investment Breakdown, $/kw E-8
5. Energy Cost (mills/kWh) E-9
6. Cost Sensitivity of Low-Pressure One-Step Regeneration E-10
xix
-------
Page
Appendix F
1. Design Specifications F-3
2, Design Assumptions F-3
3. Two-Step Regeneration Material and Energy Balance (Process F-5
Sulfur Load = 0.03)
4. Two-Step Process Investment Breakdown, $/lw F-8
5. Energy Cost (mills/kWh) F-9
Appendix H
1. Cost Increase for Carbonating Boiler Conditions — H-3
High-Pressure and Low-Pressure Regeneration Processes
•2. Cost Increase for Calcining Boiler Conditions — H-5
Two-Step Process
Appendix J
1. Constant Load Concept .Energy Costs J-4
2. Constant Load Concept Energy Costs (mills/kWh) J-6
Appendix L
1. Process Specifications L-5
2. Operating and Design Variables L-6
3. Performance Variables L-6
4. Design Options L-7
5. Simplified Variables L-8
6. Ranges of SO. Mole Fractions L-10
7. Total Installed Regeneration Process Cost, $/kw L-34
8. Conventional Plant Energy Cost with Limestone Wet L-42
Scrubbing
9. Energy Cost Comparison L-42
Appendix M
•1. Cost Basis M-l
2. Cost Results with 1 wt% Sulfur Coal M-2
3. Cost Results with No Desulfurization M-3
4. Design Factors Resulting from No Sulfur Removal M-4
Requirements
Appendix N
1. Conventional Plant Capital Investment N-3
•2. Conventional Plant Energy Cost N-3
xx
-------
PLANT OPERATION AND CONTROL
Appendix 0
1. Alternative Turndown Procedures 0-2
2. Process Variable Limitations 0-5
ALTERNATIVE PRESSURIZED FLUIDIZED BOILER CONCEPTS
Appendix P
1. Particle Size Distribution of the Ottawa Sand Used "-8
As Bed Material
2. Experimental Data of Test Series 1 (Case I) P-14
3. Experimental Data of Test Series 2 (Case I) p_i5
4. Experimental Data of Test Series 3 (Case I) P-16
5. Experimental Data of Test Series with Tubes in P-17
Downcomers (Case II)
6. Experimental Data of Test Series with Tubes in P-18
Downcomers (Case II)
7. Experimental Data of Test Series with Tubes in P-19
Downcomers (Case II)
8. Experimental Data of Test Series with Boot (Case III) P-20
9. Comparison between Experimental and Calculated Jet P-28
Penetration Depth (Case I)
10. Comparison between Experimental and Calculated Jet P-30
Penetration Depth (Case II)
^
11. Comparison between Experimental and Calculated Jet P-32
Penetration Depth (Case III)
12. Comparison of Pressure Drop Predictions in the P-34
Draft Tube
13. Comparison of Pressure Drop Predictions in the P-36
Draft Tube
14. Comparison of Pressure Drop Predictions in the P-38
Draft Tube
15. Comparison between Theoretical Predictions and P-40
Experimental Values in the Downcomer (Case II)
16. Comparison between Theoretical Predictions and P-42
Experimental Values in the Downcomer (Case III)
17. Comparison between the Calculated and the Observed P-44
Solid Flow Rate
18. Comparison between the Predicted and the P-53
Experimental Solid Recirculation Rate (Case I)
xxi
-------
Appendix P (Continued) Page
19. Comparison between the Predicted and the P-54
Experimental Solid Recirculation Rate (Case II)
20. Comparison between the Predicted and the P-56
Experimental Solid Recirculation Rate (Case III)
xxii
-------
-------
APPENDIX A
BOILER TUBE SPECIFICATION
The minimum requirement of boiler tube wall thickness is
dependent on the design fluid bed temperature, steam temperature in
the tube, bed-tube heat transfer coefficient, and the size of the boiler
tube. For ferrous tubing up to and Including 5 in. outside diameter,
the equation used is
t = 2g^-p + 0.005 D + e , (A-l)
where
t = the minimum required wall thickness, Inches
P = the maximum allowable working pressure, psi
D = the outside diameter of tubes, inches
S = the maximum allowable stress value at the operating
temperature of the metal, psi, and
e = the thickness factor for expanded tube ends
The maximum allowable stress value, S, was obtained from the
ASME Boiler & Pressure Vessel Code and plotted as shown in Figure A-l
for the specific tube materials used in the present design.
The design metal temperature, minimum tube wall thickness, and
tube material required for the basic design at bed temperature 1750°F
are presented in Table A-l. The minimum wall thickness required for
designs at bed temperatures of 1636°F, 1522°F, and 1407°F was calculated
using Equation A-l and is also presented in Table A-l and Figure A-2 for
comparison.
The boiler tube specifications for 1-in. O.D. and 2-in. O.D.
tubes are summarized in Tables A-2 and A-3. When the bed-tube heat transfer
coefficient changes, the design metal temperature and tube material
A-l
-------
TABLE A-l
BOILER TUBE SPECIFICATION
(Bed-Tube Heat Transfer Coefficient - 50 Btu/ft2-hr-°F)
Tube Location
Tube Outside
Diameter , in.
Tube Material
Design Metal
Temperature, °F
Minimum Wall
Thickness, in
Bed Temperature 1750°F
Water walls
Preevaporator
Lower superheater
Upper superheater
(lower loops)
Upper superheater
(upper loops)
Reheater
Bed Temperature 1636°F
Water walls
Preevaporator
Lower superheater
Upper superheater
(lower loops)
Upper superheater
(upper loops)
Reheater
Bed Temperature 1522°F
Water walls
Preevaporator
Lower superheater
Upper superheater
(lower loops)
Upper superheater
(upper loops)
Reheater
Bed Temperature 1407 "F
Water walls
Preevaporator
Lower superheater
Upper superheater
(lower loops)
Upper superheater
(upper loops)
Reheater
2
1-1/2
1-1/2
1-1/2
1-1/2
2
2
1-1/2
1-1/2
1-1/2
1-1/2
2
2
1-1/2
1-1/2
1-1/2
1-1/2
2
2
1-1/2
1-1/2
1-1/2
1-1/2
2
SA-213-T22
SA-210-A1
SA-213-T2
SA-213-T22
SA-213-TP304H
SA-213-T22
SA-213-T22
SA-210-A1
SA-213-T2
SA-213-T22
SA-213-TP304H
SA-213-T22
SA-213-T22
SA-210-A1
SA-213-T2
SA-213-T22
SA-213-TP304H
SA-213-T22
SA-213-T22
SA-210-A1
SA-213-T2
SA-213-T22
SA-213-TP304H
SA-213-T22
975
732
900
1058
1150
1122
953
724
884
1038
1136
1104
935
717
870
1023
1120
1085
915
705
858
1004
1105
1067
0.280
0.150
0.165
0.318
0.238
0.186
0.249
0.149
0.160
0.290
0.226
0.170
0.231
0.147
0.156
0.266
0.213
0.150
0.218
0.145
0.153
0.246
0.200
0.135
A-2
-------
18
16
~S14
x
a 12
m"
i/»
o>
KM
5
i 8
I 4
SA-213-T2
SA-210-A1
SA-210-A1 Carbon Steel S-Si C-Si
SA-213-T2 Low Alloy Steel 'Ml Cr-1/2 Mo
SA-213-T22 " 21/4Cr-lMo
SA-213-TP304H High Alloy Steel 18Cr-8Ni
i i i I L
_L
_L
0' 100 200 300 400 500 600 700 800 900 1000 1100 1200
Metal Temperature, °F
Figure A-l: Maximum Allowable Stress for Different Tube Materials
Curve 6U7IS7-A
0.4
« 0.3
i
I
e 0.2
0.1
Upper Superheater
(lower loops)
Water Walls
Upper Supperh
(upper loops)
Lower Superheater -\
Preevaporator
Reheater
1300
Bed-Tube Heat Transfer Coefficient 50 Btu/ft2-hr-°F
Tube-Pitch/Diameter Ratio ViF'FW Basic Design
1400 1500 1600
Bed Temperature, °F
1700
1800
Figure A-2: Minimum Required Tube Wall Thickness
A-3
-------
TABLE A-2
BOILER TUBE SPECIFICATION FOR 1" O.D. TUBES
(Bed-Tube Heat Transfer Coefficient = 50 Btu/ft2-hr-°F)
Tube
Location
Tube Outside
Diameter, in.
Tube Material
Design Metal
Temperature, °F
Minimum Wall
Thickness, in
Bed Temperature 1750°F
Water walls
Preevaporator
Lower superheater
Upper superheater
(lower loops)
Upper superheater
(upper loops)
Reheater
Bed Temperature 1636°F
Water walls
Preevaporator
Lower superheater
Upper superheater
(lower loops)
Upper superheater
(upper loops)
Reheater
Bed Temperature 1522°F
' Water walls
Preevaporator
Lower superheater
Upper superheater
(lower loops)
Upper superheater
(upper loops)
Reheater
Bed Temperature 1407°F
Water walls
Preevaporator
Lower superheater
Upper superheater
(lower loops)
Upper superheater
(upper loops)
Reheater
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
SA-213-T22
SA-210-A1
SA-213-T2
SA-213-T22
SA-213-TP304H
SA-213-T22
SA-213-T22
SA-210-A1
SA-213-T2
SA-213-T22
SA-213-TP304H
SA-213-T22
SA-213-T22
SA-210-A1
SA-213-T2
SA-213-T22
SA-213-TP304H
SA-213-T22
SA-213-T22
SA-210-A1
SA-213-T2
SA-213-T22
SA-213-TP304H
SA-213-T22
975
732
900
1058
1150
1122
953
724
884
1038
1136
1104
935
717
870
1023
1120
1085
915
705
858
1004
1105
1067
0.140
0.100
0.110
0.212
0.159
0.093
0.125
0.099
0.107
0.193
0.151
0.085
0.115
0.098
0.104
0.177
0.142
0.075
0.109
0.965
0.102
0.164
0.133
0.068
A-4
-------
TABLE A-3
BOILER TUBE SPECIFICATION FOR 2" O.D. TUBES
(Bed-Tube Heat Transfer Coefficient • 50 Btu/ft2-hr-cF)
Tube Tube
Outside
Location Diameter, in.
Bed Temperature 1750°F
Water walls
Preevaporator
Lower superheater
Upper superheater
(lower loops)
Upper superheater
(upper loops)
Reheater
Bed Temperature 1636°F
Water walls
Preevaporator
Lower superheater
Upper superheater
(lower loops)
Upper superheater
(upper loops)
Reheater
Bed Temperature 1522° F
Water walls
Preevaporator
Lower superheater
Upper superheater
(lower loops)
Upper superheater
(upper loops)
Reheater
Bed Temperature 1407°F
Water walls
Preevaporator
Lower superheater
Upper superheater
(lower loops)
Upper superheater
(upper loops)
Reheater
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
2
Tube Material
SA-213-T22
SA-210-A1
SA-213-T2
SA-213-T22
SA-213-TP304H
SA-213-T22
SA-213-T22
SA-210-A1
SA-213-T2
SA-213-T22
SA-213-TP304H
SA-213-T22
SA-213-T22
SA-210-A1
SA-213-T2
SA-213-T22
SA-213-TP304H
SA-213-T22
SA-213-T22
SA-210-A1
SA-213-T2
SA-213-T22
SA-213-TP304H
SA-213-T22
Design Metal
Temperature, °F
975
732
900
1058
1150
1122
953
724
884
1038
1136
1104
935
717
870
1023
1120
1085
915
705
858
1004
1105
1067
Minimum Wall
Thickness, in
0.280
0.200
0.220
0.424
0.318
0.248
0.332
0.199
0.214
0.386
0.302
0.227
0.308
0.196
0.208
0.354
0.284
0.200
0.291
0.193
0.204
0.328
0.266
0.180
A-5
-------
change as well (Table A-4). The effects of design bed temperature and
bed-tube heat transfer coefficient on the tube wall thickness requirement
for the basic design are presented graphically in Figures A-3 through
A-8.
TABLE A-4
DESIGN METAL TEMPERATURES & TUBE MATERIAL
Tube
Location
Tube Material
Design Metal Temperature, °F
1750°F
(Bed-tube Heat Transfer Coefficient =
Water walls
Preevaporator
SA-213-T22
SA-210-A1
Lower superheater SA-213-T2
Upper superheater SA-213-T22
(lower loops)
Upper superheater
(upper loops)
Reheat er
SA-213-TP304H
SA-213-T22
928
710
869
1023
1125
1088
(Bed-tube Heat Transfer Coefficient =
Water walls
Preevaporator
Lower superheater
Upper superheater
(lower loops)
Upper superheater
(upper loops)
Reheater
SA-213-T22
SA-210-A1
SA-213-T2
SA-213-T22 °r
SA-213-TP304H
SA-213-TP304H
SA-213-T22
1035
775
947
1103
1190
1167
Bed Temperature
1636°F f 1522°F
35 Btu/ft2-hr-°F)
914 900
704 695
859 849
1010 998
1114 1101
1077 1060
• 75 Btu/ft2-hr-°F)
1007 981
765 749
929 909
1083 1060
1170 1145
1139 1117
1407°F
885
687
839
988
1090
1047
954
735
889
1038
1125
1089
A-6
-------
Curve M7I93-A
0.4
E
3
0.2
0.1
Bed-Tube Heat Transfer Coefficient 35 Btu/K2 - hr - °F _
Tube Pitch/Diameter Ratio W - FW Basic Design
Upper
SuperheaterMowerjoops)
Water Walls-
"Upper Superheater I uwiei loops)
Lower Superheater
Preevaporator —
Reheater-
1300 1400
1500 1600
Bed Temperature, °F
1700
1800
Figure A-3. Minimum Required Tube Wall Thickness
Curve 6li719Ii-A
0.4
'- 0.3
0.2
0.1
Water Walls
Lower Superheater
Reheater
Preevaporator
Bed-Tube Heat Transfer Coefficient 75 Btu/ftf - hr - »F
Tube Pitch/Diameter Ratio W - FW Basic Design
1300 1400
1500 1600
Bed Temperature, °F
1703
1800
Figure A-4: Minimum Required Tube Wall Thickness
A-7
-------
Curve 6''JlbS-A
>
00
0.4
E 0.3
i
E
i
I 0.1
0.4
0.3
= 0.2
i
E
I 0.1
Upper Superheater
(lower loops)
Water Walls
Upper Superheater
(upper loops)
Lower Superheater
:eheater
Preevaporator
Bed Temperature 1407°F
Tube Pilch/Dia. Ratio W -FW Basic Design
102030405069708090
Bed-Tube Heat Transfer Coefficient. Btu/ft2-hr-°F
100 120
Figure A-5 Minimum Required Tube Wall Thickness
Curve W.7166-A
1 r
Superheater
lower loops)
Upper Superheater
(upper loops)
Reheater
Lower Superheater
Preevaporator
Bed Temperature 1636°F
Tube Pitch/DiiT Ratio W -FW Basic Design
1 J,—i
10 20 M 40 50 60 70 80 90 100 110 120
Bed-Tube Heat Transfer Coefficient, Btu/tt2-hr-°F
Figure A-7: Minimum Required Tube Wall Thickness
0.4 -
0.2
0.1
0.4
S
•3
0.1
Curve W7lBb-A
—I 1
Upper Superheater
(lower loops)
Water Walls
Upper Superheater
(upper loops)
Reheater
Lower Superheater
Preevaporator
Bed Temperature; 1522°F
Tube Pitch/Dia. Ratio; W-FW Basic Design
J I l 1 J_
10 20 30 40 SO 60 70 80 90 100 110
Bed-Tube Heat Transfer Coefficient, Btu/ftz-hr-°F
Figure A-6. Minimum Required Tube Wall Thickness
Curve M7I67-A
T
T
T T
Upper Superheater
I lower loops)
Water Walls
Upper Superheater
(upper loops)
Reheater
Lower Superheater
Preevaporator
Bed Temperature 1750°F
Tube Pitch/OlaTRatlo W -FW Basic Design
10 20 30 40 50 60 70 80 90 100 110 120
Bed-Tube Heat Transfer Coefficient, Btu/ftz-hr-T
Figure A-8. Minimum Required Tube Wall Thickness
-------
REFERENCES
1. ASME Boiler and Pressure Vessel Code, Power Boilers, Section I,
1971.
A-9
-------
B
-------
APPENDIX B
SUPPLEMENTARY COST DATA
The effects of operating conditions and boiler design on the
total boiler cost have been discussed in detail in Volume I, Section
2. To complete the evaluation, the cost information is also presented
in the following figures to show the interacting effects of the tube
pitch/diameter ratio, the bed-tube heat transfer coefficient, the maximum
allowable bed depth, and the design bed temperature:
Figure B-l shows the effect of the bed-tube heat transfer coefficient
on the steam generator cost at different design bed temperatures. The
tube pitch/diameter ratio is similar to that of the Westinghouse-Foster
Wheeler basic design.
Figure B-2 shows the effect of the bed-tube heat transfer coefficient
on the steam generator cost at two maximum allowable bed depths (20 feet
and unrestricted'). The tube pitch/diameter ratio is similar to. that of
the Westinghouse-Foster Wheeler basic design.
Figures B-3 to B-6 present the effect of the maximum allowable bed
depth on the steam generator cost at a constant bed temperature and with
the bed-tube heat transfer coefficient as a parameter. The tube pitch/
diameter ratio is similar to that of the Westinghouse-Foster Wheeler
basic design.
Figures B-7 to B-9 present the effect of the maximum allowable bed
depth on the steam generator cost at a constant bed-tube heat transfer
coefficient and with the design bed temperature as a parameter. The
tube pitch/diameter ratio is similar to that of the Westinghouse-Foster
Wheeler basic design.
The same series of figures was also prepared for 1-in. O.D. tubes
at H = 4 in. and V = 2 in. (Figures B-10 to B-18) and 2-in. O.D. tubes at
H = 8 in. and V = A in. (Figures B-19 to B-26).
B-l
-------
Curve M721I-B
T
1-0
S
X
«»
s"
. Curve Maximum Allowable Bed Perth Bed-Tube Heat Transfer Coefficient-
la
Ib
Z
3a
3b
20ft
unrestricted
20ft
20ft
unrestricted
50 Btu/ft2 - hr - °F
75 Btu/lt2 - hr - f
35Btu/it2-hr-cF
Tube Pitch/Diameter Ratio
W - FW Basic Design
'1300
1400
1500 1600
Bed Temperature °F
1700
1800
Figure B-ls Effect of Bed-Tube Heat Transfer Coefficient on the Steam Generator
Cost (not including erection)
10
x 6
10
Curve Max. Allowable Bed Perth
20ft
unrestricted
Tube Pitch/Diameter Ratio W - FW Basic Design
1407T
1522°F
2D 30 40 50 63 70 80
Bed-Tube Heat Transfer Coefficient. Btu/H2 hr - °F
90
100
Figure B-2. Effect of Bed-Tube Heat Transfer Coefficient on the Steam Generator
Cost (not Including erection)
-------
ID
Curve W7I66-A
Curo M7IW-A
Bed Temperature 1750°F
Tube Pitch/ Diameter Ratio w-fW Basic Design
Btu/fl-hr-T
Bed-Tube Heat
Transfer Coefficient.
35
50
75
10 20 30
Maximum Allowable Bed Depth, fl '
Figure B-3 Dependence of the Steam Generator Cost
(31BMWI on the Maximum Allowable
Bed Depth (not Including erection)
6
I
Bed Temperature 1636°F
Tube Pitch/Diameter Ratio W-FW Basic Design
Bed-Tube Heat
Transfer Coefficient.
Btu/ftZ-hr-°F
10 20 30
Maximum Allowable Bed Depth. It
Figure B-4: Dependence of the Steam Generator Cost
(318 MW) on the Maximum Allowable
Bed Depth (not Including erection)-
Cui«e 61./I70-A
Cirrv, W7I71-A
s
I
Bed Temperature 1522T
fube Pilch/Olameler Ratio W-FW Basic Design
Bed-Tube Heat
Transfer Coefficient
Btu/lt-hr-°F
35
50
75
10 20 30
Maximum Allowable Bed Depth, ft
Figure B-5 Dependence of the Steam Generator Cost
(318 MW) on the Maximum Allowable
Bed Depth (not Including erection)
i
s
Bed Temperature 140rf
Tute Pitch/Diameter Ratio
W-FW Basic Design
Bed-Tube Heat
Transfer Coefficient,.
Blu/ftz-hr-°F
10 20 30
Maximum Allowable Bed Depth, ft
Figure B-6: Dependence ol the Steam Generator Cost
(318 MW) on the Maximum Allowable
Bed Depth (not including erection)
B-3
-------
Steam Generator Cost * x 10
o rsi * o* 00 3
Curve 6b71F7-A
Bed-Tube Heat Transfer Coefficient 75 Btu/lt2-hr-°F
Tube Pitch/Diameter Ratio W-FW Basic Design
i Bed
~| | — • Temperature
\—T~L 1 W07°F
T_ *-> \— \K7?°r
T ' 1636°F
1750°F
i i i
) 10 20 30
Cur.,. 6I.7173-*
10
Maximum Allowable Bed Depth, ft
Figure B-7 Dependence of the Steam Generator Cost
(318 MW) on the Maximum Allowable
Bed Depth (not including erection)
: 6
3
s
S 4
E
Bed-Tube Heat Transfer Coefficient 50 Btu/lt -hr-°F
Tube Pilch/Diameter Ratio W-FW Basic Desigr
Temperature
1407T
1636T
1750°f-
10 20 30
Maximum Allowable Bed Depth, ft
Figure B-8: Dependence of the Steam Generator Cost
(318 MW) on the Maximum Allowable
Bed Depth (not including erection)
Bed-Tube Heal Transfer Coefficient 35 Btu'ft -hr-°F
Tube Pitch/Diameter Ratio W-FW Basic Desigr
10 20 30
(Maximum Allowable Bed Depth, ft
Figure B-9 Dependence of the Steam Generator Cost
(318 MW) on the Maximum Allowable
Bed Depth (not including erection)
-------
Curve f*7J>3-B
tfl
Ln
_- 4
S
1300
Curve
1
2
3
Maximum Allowable Bed Depth Heat Transfer Coefficient
75Btu/ft2-hr-f
SOBtu/lt^-hr-f
35Btu/ftz-nr-«F
20ft
20 It
20ft
Tube Pitch/Diameter Ratio
1"0. 0. Tubes at H=4'&V=2'
1400
1500 1600
Bed Temperature. °F
1700
1800
Figure B-10: Effect of Bed-Tube Heat Transfer Coefficient on the Steam Generator
Cost (not including erection)
Curve M72M-B
Curve
Max. Allowable Bed Depth
10ft
20ft
Tube Pitch/Diameter Ratio 1" 0. D. Tubes at
1636PF
1750"F
10 20 30 « 50" 60 70 " 80 90 10D 110
Bed-Tube Heat Transfer Coefficient. Btu/ft2 hr - °F
Figure B-ll: Effect of Bed-Tube Heat Transfer Coefficient on the Steam Generator
Cost (not including erection)
-------
10
x
^ 6
fe
04
E
S
Bed Temperature 1750T
Tube Pitch/Diameter Ratio 1" 0. D. Tubes at
H = 4" & V = 7'
\
Bed-Tube Heat
Transfer Coefficient.
Btu/ft^hr-T
35
50 .
10 20 30
Maximum Allowable Bed Depth, ft
Figure 8-12 Dependence of the Steam Generator Cost
(318 MW) on the Maximum Allowable
Bed Depth (not including erection)
,- *»«• W75-A
8
•°o
X
* 6
i
&
01
° d
E
3
(/I
2
n
' i i
Bed Temperature 1636°F
Tube Pitch/ Diameter Ratio 1" 0. D. Tubes at
H = 4"&v = 2"
-
Bed-Tube Heal
^ Transfer Coefficient. "
Btu/ft2-hr-°F
3;
1 ™
-
i i i
10 20 30
Maximum Allowable Bed Depth, ft
Figure B-13 Dependence of the Steam Generator Cost
(318 MW) on the Maximum Allowable
Bed Depth (not Including erection)
Curve U7196-A
10
8
•0
S
x
~. 6
nerator Cos
E *
3
2
n
i i i
Bed Temperature 1522 "F
Tube Pitch/Diameter Ratio 1" 0. D. Tubes
at H=4" & V=2'
1 Bed-Tube Heat
1 Transfer Coefficient. "
SU Btu/ftz - hr - »F
"
75
-
i i i
10
10 20 30
Maximum Allowable Bed Depth, ft
Figure 8-14- Dependence of the Steam Generator Cost
(318 MW) on the Maximum Allowable
Bed Depth (not including erection)
x
~ 6
u 4
E
2 -
Curve M7176-A
Bed Temperature 1407T
Tube Pitch/ Diameter Ratio 1" 0. D. Tubes at
Bed-Tube Heat
Transfer Coefficient.
Btu/nz-hr-°F
10 20 30
Maximum Allowable Bed Depth, ft
Figure B-13- Dependence of the Steam Generator Cost
(318MW) on the Maximum Allowable Bed
Depth (not including erection)
B-6
-------
Curve (A7IJJ-*
Cunc W7I79-A
10
8
la
X
« 6
•g
i5Z2«f
L \ /
\N-1636°F
Bed-Tube Heat Transfer Coefficient
50Btu/ftz-hr-T
Tube Pitch/ Dla. Ratio: r 0. D. Tubes
at H = 4" & V = 2'
8
"s
v
«. «
S 4
I
2
n
Bed-Tube Heat Transfer Coefficient 35Btu/flZ-hr-cF
Tute Pitch/Diameter Ratio 1" 0.0. Tubes at
' ~\
lin
l — i
— i h „„„-
iHT^ mrr
-
i i i
10 20 30
Maximum Allowable Bed Depth, ft
Figure B-1& Dependence of the Steam Generator Cost
(318 MW) on the Maximum Allowable Bed
Depth (not Including erection)
10 20 30
Maximum Allowable Bed Depth, ft
Figure B-17: Dependence of the Steam Generator Cost
(318 MW) on the Maximum Allowable Bed
Depth (not including erection)
10 20 30
Maximum Allowable Bed Depth. H
Figure B-18: Dependence ol the Steam Generator Cost
(318 MW) on the Maximum Allowable Bed
Depth (not Including erection)
-------
Curve 61*721 It-B
1/1
o
o
0>
s
ID
1300
Curve Maximum Allowable Bed Depth Bed-Tube Heat Transfer Coefficient
1
2a
2b
3a
3b
20ft
20ft
unrestricted
20ft
unrestricted
- Tube Pitch/Diameter Ratio
75Btu/ft2hr-°F
50 "
50 "
35 "
35 "
2"0. D. TubesatH = 8"&V=4" -
1400
1500 1600
Bed Temperature, °F
1700
1800
Figure B-19: Effect of Bed-Tube Heat Transfer Coefficient on the Steam Generator
Cost (not including erection)
B-B
-------
Curve U7I97-A
10
".6
s
Bed Temperature W50°F
Tube Pitch/Diameter Ratio 2' 0. D. Tubes at
H = 8"&V = 4"
Bed-Tube Heat
Transfer Coefficient.
- hr -"F
35
0 10 20 30
Maximum Allowable Bed Depth, ft
Figure B-20: Dependence of the Steam Generator Cost
1318 MW) on the Maximum Allowable Bed
Depth (not including erection)
10
Curvo 6b7l80-A
I
i
o
E
Bed Temperature 1636°F
Tube Pitch/Diameter Ratio ? 0. D. Tubes
at H = S' & V = 4"
Bed-Tube Heat
Transfer Coefficient,
Btu/ft2-hr-"F
35
50
75
°0 10 20 30
Maximum Allowable Bed Depth, ft
Figure B-21. Dependence of the Steam Generator Cost
(318 MW) on the Maximum Allowable Bed
Depth (not including erection)
Curve W7I98-A
Curve M7I8I-A
10
* 5
8
u
o
Bed Temperature 1522°F
Tube Pitch/Diameter Ratio 2"0. D. tubes at
H=8"&V=4"
Bed-Tube Heat
Transfer Coefficient,
I) 10 20 30
Maximum Allowable Bed Depth, ft
Figure B-22: Dependence of the Steam Generator Cost
(318MW) on the Maximum Allowable Bed
Depth (not including erection)
10
I
i
35
Bed-Tube Heat
Transfer Coefficient,
Btu/ft2-hr-°F
Bed Temperature 1407T
Tube Ditch/Diameter Ratio 2" O.D. Tubes
atH = 8"&V =
10 20 30
Maximum Allowable Bed Depth, ft
Figure B-23: Dependence of the Steam Generator Cost
(318 MW) on the Maximum Allowable Bed
Depth (not Including erection)
B-9
-------
w
M
o
10
Curve 6"*7i99-A
Curvi C./.62-A
Bed-Tube Heat Transfer CoeHicienl 7? Btu/ftz - hr -T
Tube Pitch/Diameter Ratio 2'0.0. tubes at
". 6
I
1«7
-------
-------
APPENDIX C
PARAMETRIC STUDY OF ELUTRIATION RATE
FROM A FLUIDIZED BED COMBUSTION BOILER
A parametric study of the elutriation rate from the fluidized
bed combustion boiler (Westinghouse-FW basic design) was performed. The
design bases for the Westinghouse-FW basic design are summarized in
Table C-l. To generate a dust loading identical to that projected for
the Westinghouse-Foster Wheeler fluidized bed boiler (7 grains/SCF from
the combustor), an infinite number of combinations is possible, as
shown in Figure C-l. The basic design specification of this boiler is
shown as a point which represents an 8.5% ash coal with 6% carbon elutria-
tion rate and 0.7% dolomite elutriation rate. (The design assumed six
times the stoichiometric Ca/S ratio.) For the same coal, a 0% carbon
elutriation rate will require a 5.3% dolomite elutriation rate to pro-
duce the same dust loading. For a coal of 13.4% ash, the elutriated ash
alone will give the same dust loading with 0% carbon and dolomite elutria-
tion rate. Similar diagrams are presented for cases which double and
triple this basic design dust loading (Figures C-2 and C-3). The data
from Figures C-l to C-3 are cross-plotted in Figures C-4 to C-8 for a
constant carbon elutriation rate to illustrate the possible dust loading
for coals of different ash content and stone loss.
C-l
-------
TABLE C-l
DESIGN BASES FOR WESTINGHOUSE-FW
FLUIDIZED BED COMBUSTION BOILER (318 MW-4 MODULE DESIGN)
Coal feed (ash content = 8.5%)
Air feed
Dolomite feed plus make-up
53,910 Ib/hr-module
586,500 Ib/hr-module
44,500 Ib/hr-module
Ca/S ratio
Dust loading from primary bed
7 grains/SCF
Composition: Carbon
Ash
Dolomite
31.9%
63.4%
4.6%
C-2
-------
Curv. 6I.903I-*
Curve M9032-A
o
I 5
Carbon Elutnatlon Rate;
-—ftof Total Carbon
Input
0 5 10 15
Ash Content In Coal, *
Figure C-l: Different Combinations of Elutriation
Rates for a Dust Loading Eflual to That
of ®-FW Basic Design
Carbon Etutrution Rate
* d Total Carbon Input
10 20
Ash Content in Coal.
30
Figu re C-2. 0 ifferent Combinations of Elutriation
Rates for a Dust Loading Double That
of i® -fW Basic Design
*
1
c
a
i
I
Carbon Elutriation Rate,
% of Total Carbon Input
10 20 30
Ash Content in Coal, %
40
Figure C-3: Different Combinations of Elutriation Rates for
a Dust Loading Triple Thai of ©-FW Basic
Design
-------
Curve M90J6-A
Curve 6b903l*-A
Curv. W90J5-A
Dolomite Elutnation Rate
% of Sulfated Dolomite
ID 20
Ash Content in Coal. %
30 40
I
Trt
5
u
I
10 20
Ash content in Coal, %
40
"Dolomite Elulnation Rate
I of Sulfated Dolomite
Figure C-4 Possible Dust Loading from FBC at 0% Carbon
Elutriatnn Rate
Figure C-5: Possible Dust Loading from FBC at» Carbon
Elutrialton Rate
10 20
Ash Content in Coal. %
Figure C-4. Possible Dust Loading from FBC at 10% Carbon
Elutnation Rate
-------
Curve M9037-A
S
V. Dolomite Elutriation Rate
I
10 20
Ash Content in Coal. %
30
Figure C-7: Possible Dust Loading from FBC at 15% Carbon
Elutriation Rate
Curve W9038-A
.E 4
I
I
10 20
Ash Content in Coal, %
30
Figure C-8: Possible Dust Loading from FBC at 20% Carbon
Elutriation Rate
-------
0
-------
APPENDIX D
BASE DESIGN-HIGH PRESSURE ONE-STEP PROCESS
PROCESS. DESCRIPTION
The one-step dolomite/limestone regeneration process consists
of a single fluidized bed reaction vessel in which utilized' dolomite/
limestone (CaSO.) from the pressurized fluid bed boiler (coal-fired)
is reacted with a H./CO reducing gas to produce CaO and SCL. The
reaction
H -> Ho°
CaS04 + (CQ} «- CaO + {^ } + S02 (D-l)
requires temperatures on the order of 2000°F to produce high levels
(1 to 15%) of S0_ at reasonable rates. The S0_ equilibrium concentration
is favored by reduced pressure, since it is inversely proportional to the
total pressure. The competing reaction
H, H,0
CaS04 + 4 (CQ} ^ CaS + 4 (CJ } (D-2)
also occurs to produce the unwanted product CaS.
The regenerated dolomite/limestone is returned to the fluid
bed boiler with the required amount of fresh dolomite/limestone to
supplement the regenerated sorbent's reduced activity. The S02 stream
from the regenerator is sent to a sulfur recovery plant in which
sulfuric acid or elemental sulfur is produced.
The high-pressure (10 atmosphere) one-step regeneration process
is broken down into five interrelated elements: a regenerator element, a
reducing-gas producer element, a sulfur recovery element, a tail gas
handling element, and a sorbent circulation element. These elements and
their relationship to one another are shown in Figure D-l. The simple
block flow diagram indicates the basic streams in the process. Fuel, air,
and/or steam are converted to a H.CO reducing gas in the reducing gas
D-l
-------
Make-Up
Sorbent
Hot Gas
to
Turbine
Fluid Bed
Boiler
Coal Air
Sorbent
Disposal
Sorbent
Circulation
Element
Regenerator
Element
Reducing
Gas
Producer
Element
Fuel
Air
Regenerator
Product
Gas
Dwg. 72^8352
Sulfur
Sulfur
Recovery
Element
Sulfur Recovery
Tail Gas
Tail Gas
Handling
Element
Regeneration
Tail Gas
Figure D-L- One-Step Regeneration Process Elements
-------
producer element. This reducing gas is combined with tail gas recycled
from the sulfur recovery element to make up the reducing gas provided
to the regenerator. Utilized sorbent from the fluid bed boiler is
transported to the regenerator element to produce a regenerated sorbent
and an SO. stream. The regenerator product gas is processed in the
sulfur recovery element, with the sulfur recovery tail gas being either
recycled to the regenerator or processed in the tail gas handling
element. The performance and design of these five regeneration process
elements are interrelated in a complex manner.
SPECIFICATIONS AND ASSUMPTIONS
The high-pressure one-step regeneration base design is founded
on the following specifications and assumptions.
TABLE D-l
DESIGN SPECIFICATIONS
Fixed coal rate
Sorbent type
Boiler sulfur removal
Variable process sulfur loads
Recycle of tail gas to boilers
Recovery of elemental sulfur
by methane reduction of S0»
Reducing-gas production from coal
(methane for sulfur recovery)
Four regenerator vessels
(one for each boiler module)
430,000 Ib/hr (*
dolomite
95%
0.01, 0.03, 0.06
635 MW)
TABLE D-2
DESIGN'ASSUMPTIONS
Regenerator pressure
Regenerator temperature
S02 mole fraction in regenerator product
CO + H2 mole fraction in reducing gas
Sulfur recovery efficiency
Dolomite utilization in boiler
Dolomite utilization after regeneration
Dolomite make-up rate
Fluidization velocity
Boiler condition
Dolomite temperature to regenerator
Calcium sulfide in regenerated dolomite
Ash in utilized dolomite from boilers
Heat losses from regenerator neglected
10 a tin
2000°F
1 and 2% cases
1.5 and 3%, respectively
95%
30%
10%
1 mole Ca/mole S
5 ft/sec
calcination
1200°F
0%
D-3
-------
MATERIAL AND ENERGY BALANCES
Material and energy balances for the process streams indicated in
Figure D-2 are shown in Tables D-3 and D-4, for the 1% and 2% S0_ cases,
respectively. The basis for Tables D-3 and D-4 is a process sulfur load
of 0.03. The energy balance indicates that the process is essentially
balanced — the steam raised from the hot regenerator gas is sufficient
to provide the auxiliary power requirements of the process. Thus, the
regeneration process has a negligible effect on the electrical output
of the power plant.
EQUIPMENT
The major equipment items considered in the base design of
the high-pressure one-step regeneration process are shown in Figure D-3.
Regenerator Element
The regenerator element consists of four regenerator vessels (V5).
Each regenerator vessel is followed by a multistage particulate clean-
up unit (V6). The regenerator vessels have been assumed to be 35 ft in
total height with operation at a fluidization velocity of 5 ft/sec. Carbon
steel shells with refractory lining and insulation and a refractory
distributor plate have been utilized in the design. The particulate
control equipment used for each module is a two-stage Ducon cyclone
followed by an Aerodyne unit. The cyclones are enclosed in refractory
lined pressure shells.
Sorbent Circulation Element
The sorbent circulation element consists of four identical
modules to circulate dolomite pneumatically between each boiler-regenerator
combination. Each circulation module collects utilized dolomite from
the boiler through four transport lines into a single hold vessel (V2).
The utilized dolomite is then injected into the circulation line using
the Petrocarb injection system (V3). Regenerated dolomite leaves the
-------
Dwg. 724B67
Regenerator Recycle Gas
J Codl
* G<
Prnri
r
as
b Element O
TS i- *" .... . n.jM
2)A,r ^
Regenerator
Element
x 10 aim
!:__ 2000°F
•. T
Al
^ ^
Utilized A.
H20 Methanes
i 1
Sulfur'
L
Tail Gas
k. Hanrilinn
(4) Element T Element
Gas (T) Sulfur ,
1 Recovery i
^ * Tail Gas 1
^£r \i^/ I
Steam Sulfur A2
(f) Process
Tail Gas
" Regenerated
Dolomite
Figure 0-2: High-Pressure One-Step Reqeneration Process Flow Diagram
-------
TABLE D-3
MATERIAL AND ENERGY BALANCES ~ ONE-STEP REGENERATION AT HIGH PRESSURE (1Z SO.)
Scream
1. Coal
2. Air
3. Reducing Gas
4. S02 Gas
Conditions
Temperature, °F/ Pressure, psia
100
600
2250
2000
5. Sulfur Recovery
Tail Gas 300
6. Process Tail
Gas
7 . Regenerator
Recycle .Gas
8. Utilized
Dolomite
O
1 9 . Regenerated
Dolomite
10. Water
11. Methane
12. Steam
13. Sulfur
A. Auxiliary
Power
A. Auxiliary
Power
400
400
1200
2000
100
100
750
350
200
Sorbent Transport) 600 H.P.
Power J5.600 H.P.
/ 165
/ 165
/ 150
/ 140
/ 135
/ 155
/ 155
/ 150
/ 150
/
/ 140
/ 600
/ 75
/ 135
36,400 H.P.
2.100 H.P.
Moles/Hr Enthalpy, 1
Jtu/lb. Fuel Comments
57,100 Ib/hr 1725C 85Z thermal efficiency In coal partial
combustor
16,500
139 92Z of stolchlometric combustion of coal
40,000 1594 1.5Z H2 + CO assumed requirement
AUj^ = -104 Btu/lb fuel in regenerator
40,300 1404 1Z S02
40,600 172 (+53)d AIL = -40 Btu/lb fuel in Claus plant
18,000 84 (+23)d recycled to boiler
22,600 108 (+30)d
2015 moles Ca ~~\ 50Z MgO. 15Z CaSO^, 35Z CaO
1 change in sensible heat.
f AH = 86
2015 moles Ca ___/ 50Z MgO, 5Z CaSO^, 45Z CaO
411,000 Ib/hr
310
354.000 Ib/hr (600*)
67,000 Ib/hr (750)
435
214
1270 1110 Btu/lb fuel - 600 psia, 750°F steam
160 Btu/lb fuel - 75 psia, 350° F steam
95Z Claus sulfur removal efficiency assumed
70Z compressor efficiency
70Z compressor efficiency
(boosting 10 atm transport gas) or
(compressing atm press, trans, gas)
Uusis:
Process sulfur load = 0.03
Holler coal rate - 430,000 Ib/hr.
j
BTU per lb coal fed to the fluid bed boilers.
Based on 13,000 Btu/n, coal.
Heat of combustion In parentheses.
-------
TABLE D-4
MATERIAL AND ENERGY BALANCES — ONE-STEP REGENERATION AT HIGH PRESSURE (2Z
Conditions
Stream Temperature ,°F/ Pressure, p
1. Coal 100°F / 165 psia
2. Air 600 / 165
3. Reducing
Gas 2500 / 150
4. S02 Gas 2000 / 140
5. Sulfur Recovery
Tall Gas 300 / 135
6. Process Tail
Gas 400 / 155
7 . Regenerator
Recycle Gas 400 / 155
8. Utilized
Dolomite 1200 / 150
9. Regenerated
Dolomite 2000 / 150
10. Water 100 /
11. Methane 100 / 140
12. Steam 750 / 600
(Total) 350 / 75
13. Sulfur 200 / 135
A. Auxiliary
Power 20. 400 H.P.
A. Auxiliary
Power 1, 200 H.P.
a 1
sla Moles/Hr Enthalpy. Btu/lb Fuel 1 Comments
34,800 Ib/hr 1050b 852 thermal efficiency in coal partial
combustor
9,400 80 872 of stolchiometrlc combustion of coal
19,800 892 32 H2 + CO , AHg ° -104 Btu/lb Fuel
20,150 702 2Z S02
20,440 92 (+50)c AHR = - 40 Btu/lb Fuel in Claus plant
10,740 52 (+25)c Recycled to boiler
9.700 49 (+25)c
2015 moles Ca ~' change in sensible heat, 502 MgO, 15Z CaS04> 35Z CaO
V AH ° 86
2015 moles Ca J 502 MgO. 52 CaSO^, 452 CaO
210,000 Ib/hr
291 200
179.000 Ib/hr (6000) 650 561 Btu/lb Fuel - 600 psia, 750°F steam
31,000 Ib/hr (750) 89 Btu/lb Fuel - 75 psia, 350°F steam
414 952 Claus sulfur removal efficiency
702 compressor efficiency
70Z compressor efficiency
Sorbent Transport^ 600 H.P. (boosting 10 atm trans, gas) or
Power J 5,600 H.P. (compressing atm press, trans, gas)
a
Btu per Ib coal fed to the fluid bed boilers
* Based on 13,000 Btu/lb coal
'Heat of combustion in parentheses
-------
VI
V2
V3
Vt
V5
V6
V7
BOILER
HOLD VESSEL
SORBENT INJECTION VESSEL
SURGE VESSEL.SORBENT INJECTION VESSEL
REGENERATOR VESSEL
MULTISTAGE CYCLONE
CYCLONE
V8 - COAL PARTIAL COHBUSTOR V*
V9 - REDUCTION REACTOR C5
VI0 - COAL FEEDING SYSTEM C6
VII.VI2.VI3 - CLAUS REACTOR El
VII) - COALESCER E2.E3.E6
CI.C2 - TRANSPORT AIR COMPRESSOR ft.E5
C3 - REGENERATOR RECYCLE COMP
Dug 72<*B665
AIR COMPRESSOR
AIR COMPRESSOR
BOILER RECYCLE COMPRESSOR
STEAM GENERATOR
SULFUR CONDENSER
HEAT INTERCHANGER
SULFUR
| SORBENT CIRCULATION ELEMENT
UTILIZED
I ^^f DULUHITL I
I
L
REDUCING GAS PRODUCER ELEMENT
-
PROCESS TAIL GAS |
RECYCLE TO BOILERS
I TAIL GAS •
| HANDLING ELEMENT |
Figure D-3: High-Pressure One-Step Regeneration Schematic Flow Diagram
-------
regenerator and enters a distribution vessel (V4). The dolomite is
distributed to four lines and injected into the transport lines returning
the sorbent to the boiler beds. Waste dolomite may be withdrawn from
either the hold vessels or the distribution vessels (shown in Figure D-3
for disposal. A pair of air compressors is required for transporting
the sorbent (C1.C2). Air at about 175 psia is required for transport.
Reducing-Gas Producer Element
A single gas producer module is required for generation of the
reducing gas. Coal storage, handling and preparation are required,
though some of the equipment required may be integrated with the main coal supply
system. In the high-pressure one-step regenerator the levels of hydrogen
and carbon monoxide required for regeneration are so low (1.5-3.0%)
3.0%) that it is assumed that reducing gas can be produced by operating a
slagging combustor (V8) at slightly below stoichiometric combustion
conditions. Coal is fed to the slagging partial combustor using a Petro-
carb feeding system (V10). Compression of combustion air is also needed
(C5). The reducing gas passes through a single stage cyclone (V8)
and is mixed with recycled tail gas by means of recycle compressor C3.
Sulfur Recovery Element
The sulfur recovery element is a three-stage Claus process
(V11,V12,V13) with a portion of the entering SO- stream reduced
to H.S by means of a catalytic reaction with methane (V9). The sulfur
* 1
recovery element is described in a previous report by Shell.
Other possibilities exist for sulfur recovery such as an Allied Chemical
process which is nearer commercial development than the process assumed
here.2
Tail Gas Handling Element
In the case of the high pressure one-step regeneration process
the tail gas handling element consists of a compressor (C6) which
recycles the tail gas to the fluid bed boilers. It has been assumed that
D-9
-------
no modification to the boilers is necessary to handle the recycled gas,
though this may be a large stream relative to the boiler air rate.
CAPITAL INVESTMENT BREAKDOWN
The breakdown by element of the capital investments for the
1 and 2% SO. cases is shown in Tables D-5 and D-6. Investments are given
in $/kw based on 635 MW of plant capacity.
TABLE D-5
1% S02 INVESTMENT BREAKDOWN, $/kw
Element
0
.06
Process
1
0
Sulfur
.03
Load
1
0
.01
Regenerator Element
Sulfur Recovery Element
Tail Gas Handling Element
Dolomite Circulation Element
Gas Producer Element
TOTAL PROCESS
11.2
6.7
2.8
2.5
37.4
60.6
7.6
A.I
1.6
1.5
22.0
36.8
3.4
1.9
0.9
0.6
11.3
18.1
TABLE D-6
2% SO2 INVESTMENT BREAKDOWN, $/kw
Element
0
.06
Process
1
0
Sulfur
.03
Load
1
0
.01
Regenerator Element
Sulfur Recovery Element
Tail Gas Handling Element
Dolomite Circulation Element
Gas Producer Element
TOTAL PROCESS
7.6
5.0
1.8
2.5
24.0
40.9
4.1
3.2
1.0
1.5
14.8
24.6
2.3
1.5
0.5
0.6
6.8
11.7
The costs given in Tables D-5 and D-6 are based on the following:
• Mid-1973 costs
• Interest during construction and general items and engineering
not included. These items are added to obtain the total power
plant cost.
D-10
-------
ENERGY COSTS
The energy cost contribution of the high-pressure one-step
regeneration process to the total power plant cost is shown in Tables
D-7 and D-8. These costs were obtained on the'following basis:
0 Interest during construction at 7-1/2% per year
e Capital charges at 15%
e O&M at 5%
e 70% capacity factor
o No credit for recovered sulfur
e Coal at 45
-------
SENSITIVITY OF RESULTS
The sensitivity of the. cost results for the high-pressure one-
step regeneration process with respect to the various parameters which
apply to the process are summarized in Table D-9. The variables which
have been investigated by the base case study are the process sulfur
load, the S0_ mole fraction, and the boiler condition (calcining or
carbonating). The effect of the dolomite cost and make-up rate on the
plant energy cost is shown in Appendix H. Some of the design and oper-
ating variables considered in Table D-9 are also investigated in the
general study of Appendix L.
The capital cost of each of the five process elements are
considered in Table D-9, in addition to the total process capital
investment. The sensitivity of the regeneration process fuel cost to
the variables is also indicated in the table. The quantitative state-
ments in Table D-9 refer to the.case of 0.03 process sulfur load and
indicate a comparison with the base case results with all parameters
fixed at the base value but the parameter in question.
D-12
-------
PAGE NOT
AVAILABLE
DIGITALLY
-------
REFERENCES
1. Shpall, R. T. Shell Development Company, "Claus Technical and
Economic Study," National Air Pollution Control Administration,
GAP Contract EHSD-71-45 (1972).
2. "Evaluation of the Fluidized Bed Combustion Process," Westinghouse
Research Laboratories, Office of Research and Monitoring, EPA,
Contract No. 68-02-0217 and 68-02-0605, Thirty-Seventh Monthly
Progress Report, p. 23, April 1973.
D-15-
-------
E
-------
APPENDIX E
BASE DESIGN — LOW-PRESSURE ONE-STEP REGENERATION PROCESS
PROCESS DESCRIPTION
The same chemistry and basic flow diagram applies for the
low-pressure one-step regeneration process as for the high-pressure
one-step process. The same five process elements are Identified as
in the high-pressure process: the regenerator element, the sorbent
circulation element, the reducing-gas producer element, the sulfur
recovery element, and the process tail gas handling element. The
characteristics and performance of the five elements are different from
the high-pressure case, however, and the motive for examining the
low-pressure one-step regeneration process is Indicated by the results
of Appendix L as having, the potential for a much higher concentration
of SO£ in the regenerator product.
PROCESS SPECIFICATIONS AND ASSUMPTIONS
The low-pressure one-step regeneration base design is founded
on the following set of specifications and assumptions. The sensitivity
of the cost results to these assumptions and specifications are considered
in this appendix and in Appendix L.
TABLE E-l
DESIGN SPECIFICATIONS
Fixed coal rate 430,000 Ib/hr
Sorbent type dolomite
Boiler sulfur removal 95%
Variable process sulfur loads 0.01, 0.03, 0.06
Incineration of process tail gas
Recovery of elemental sulfur
by methane reduction of S02
Reducing gas produced from coal
(methane for sulfur recovery)
Four regenerator vessels
(one for each boiler module)
E-l
-------
TABLE E-2
DESIGN ASSUMPTIONS
Regenerator pressure 1 atm
Regenerator temperature 1900°F
S02 mole fraction in regenerator product 10%
CO + H£ mole fraction in reducing gas 16.5%
Sulfur recovery efficiency 95%
Dolomite utilization in boiler 30%
Dolomite utilization after regeneration 10%
Dolomite make-up rate 1 mole Ca/moles S
Fluidization velocity 5 ft/sec
Boiler condition calcination
Dolomite temperature to regenerator 1900°F
Calcium sulfide in regenerated dolomite 0%
Ash in utilized dolomite from boilers 0%
Heat losses from regenerator neglected
MATERIAL AND ENERGY BALANCES
Material and energy balances for the low pressure one-step
regeneration process streams in Figures E-l and E-2, are shown in Table E-3.
Temperature control and process turndown of the low-pressure process
are discussed in Appendix G. Figure E-2 represents the process streams
required to preheat the sbrbent to 1900°F, and Figure E-l deals with
the remainder of the regeneration process. Somewhat more steam is
produced in the low-pressure one-step process than is required to
generate the process auxiliary power. The effect of the regeneration
process on the power plant electrical output is neglected.
EQUIPMENT DESCRIPTION
The major equipment items considered in the base design of the
low pressure one-step regeneration process are shown in Figure E-3.
Regenerator Element
The regenerator element follows the same basic design used
in the high^pressure case: four regenerator vessels (V13), refractory
lined, are each in series with a multistage cyclone (V14) to
remove fines from the SO. gas.
E-2
-------
Dwg. 721*6670
(7) Regenerator Recycle Gas
P\ r«-»i
ij ooai ^
[)Air "
i
G
Prod
Elen
H20 Methanes
r
Regenerator
uccr » Element
-"» ^... iff.
°"
|
(D
A, Utilized A
Dolomite *&
i
k.
Gas
^
1 I
Sulfur
Recovery
Element
.
r i
(
® ®
Steam Sulfur
i
^
1
Tail Gas
' ^ Ha nil linn
Element
^
J) Sulfur i
Recovery i
Tail Gas 1
1
A2
(6) Process
Tail Gas
Regenerated
Dolomite
Figure E-l: Low-Pressure One-Step Regeneration Process Flow Diagram
-------
Dug. 6207A29
8
UtilFzed
Dolomite
to
Regenerator
Sorbent
Heater
^
Utilized
Dolomite
Waste
Heat
Boiler
Coal
Combustor
Coal
Air
Regenerated
Dolomite
Sorbent
Cooler
1
-------
Owg 721(6666
VI -
V2 -
V3 -
Vl) -
V5 -
V6 -
V7 -
V8 -
BOILER
HOLD VESSEL
DISTRIBUTION AND INJECTION
SORBENT COOLER
SORBENT HEATER
LOCK HOPPER SYSTEM
SORBENT INJECTOR
LOCK HOPPER SYSTEM
V9 - SORBENT HEATER Cl
VI0 - SORBENT COOLER C2
VII - WASTE HEAT BOILER C3
VI 2 - COHBUSTOR C<»
VI3 - REGENERATOR VESSEL C5
Vl
-------
TABLE E-3
MATERIAL AND ENERGY BALANCES — LOW-PRESSURE ONE-STEP REGENERATION
M
Stream
1.
2.
3.
4.
5.
Coal
Air
Reducing Caa
S02 Gaa
•F
200*
200
3150
1900
Sulfur Recovery 300
Tall Can
Conditions
/ DBla
F /
/
/
/
/
20 pala
20
17
16
15
Moles/Hr
(unless Indicated otherwlaa) Enthalpy. R
12,750 Ib/hr 385
2,500 2
3.630 210
4.030 126
tu/lh F.,.1 Comments
B5Z thermal efficiency In coal gaalfler
63Z of stolchlometrlc combustion of
16. 5Z H, + CO assumed requirement
aH_ • -104 Btu/lb fuel In generator
10Z S02
coal
* .300 12 (+47) H.V. of unuaed CH, In parenthesis
6. Process Tall 700 / 15
Caa
7. Regenerator 300 / 15
Recycle Caa
8-
Doloolce
190t> / 20
3,640 + 5.700
660
2015 moles Ca
9.
10.
11
12.
13.
14.
15
16.
17.
18.
19.
20.
21.
11.
23
24
25
Regenerated
Tall CUB
Water
Methane
Steam
Sullur
Utilized
Dolomite
Kuiiuncruted
Dolnnltt!
Cod)
Air
Air
Air
Heating
Heater Flue
Las
Hot Tall Cab
Cold Tall Caa
Water
St. •am
1900
100
100
750
350
200
1200
800
200
100
100
1HOO
2JOO
MOO
1400
300
800
.'50
/
/
/
/
/
/
/
/
/
/
/
/
/
/
/
/
/
15
—
16
600
75
15
20
15
15
30
25
20
20
17
17
15
200
600
2015 moles Ca
51,000 Ib/hr
270
43.000 Ib/hr
8,000 Ib/hr
395
2015 moles Ca
2015 moles Ca
3410 Ib/hr
1083
4360
4360
4750
4750
5700
S700
46.000 Ib/hr
46.000 Ib/hr
3 (+7)
change In
sensible heat,
AH - -23
186
154
change In sensible
heat AH - 99
change In sensible
heat AH - 157
104
157
221
121
156
27
10
119
AH - -40 Btu/lb fuel In Claus plant
combustion of sulfur plant tall gas with no CH4 addi-
tion and combination with stone heating tall gas
50Z MgO, 15Z CaSO, . 35Z CaO-
50* MgO, 5Z CaS04, 45Z CaO
135 Btu/lb fuel - 600 pala, 750*F steam
19 Btu/lb fuel - 75 psla, 350°F steam
95Z efficient Claus plant
-------
Sorbent Circulation Element
The sorbent circulation element used in the base design re-
quires a system of lock hoppers to transport the dolomite pneumatically
between the 10 atmosphere boilers and the 1 atmosphere regenerators.
Cooling of the hot sorbent is required in order to meet the temperature
limitations of the system valves (V4, V10). The Petrocarb design for
pressurized feeding of solids was applied (V6, V7, V3, V10). Heating
of the dolomite to 1900°F before it enters the regenerator is carried
out by using the hot gas from the sorbent cooler (V10) combined with
the hot combustion products supplied by a coal combustor (V12). Energy
from the sorbent heater is recovered in a waste heat boiler (Vll).
Stone for disposal is withdrawn from the hold vessels (V2). High-
temperature valves and/or a dense phase transport scheme could reduce
the complexity of this element.
Reducing-Gas Producer Element
A single fixed-toed type coal gasifier (V17) is used to supply
reducing gas to the regenerator. The gasifier product is combusted with
sufficient air in a combustor (V15) to provide the proper reducing- gas
composition and temperature when it is combined with the tail gas
recycle stream. Compressor (C3) provides the required air for gasifica-
tion and partial combustion. Compressor (C4) handles the tail gas recycle
stream. Coal storage, handling, and preparation are also accounted for.
Sulfur Recovery Element
The same sulfur recovery concept has been applied to the low-
pressure one-step regeneration process as was used for the high-pressure
case. The SO- product from the. regenerator is partially reduced to ItS
by the catalytic reaction of methane. This step is followed by a
three-stage GLaus system.
E-7
-------
Tail Gas Handling Element
The tail gas from the sulfur recovery plant is split into a
regenerator recycle stream and an Incineration stream. Blower (C5)
provides the air required to incinerate the tail gas in the incinerator (V18).
CAPITAL INVESTMENT BREAKDOWN
The capital cost breakdown by elements is given in Table E-4
The costs are in 1973 dollars and do not include interest during
construction or general items and engineering.
TABLE E-4
LOW-PRESSURE ONE-STEP PROCESS INVESTMENT BREAKDOWN, $/kw
Element 1
Regenerator Element
Sulfur Recovery Element
Tail Gas Handling
Element
Dolomite Circulation
Element
Gas Producer Element
TOTAL PROCESS *
Process Sulfur
0.06 ! 0.03
8.2 5.6
5.4 3.1
0.6 0.4
25.3 14.7
8.1 4.1
i7.6 27.9
Load
1 0.01
2.2
1.5
0.2
6.8
2.7
13.4
About 60% of the dolomite circulation element cost is for heating the
dolomite prior to regeneration, waste heat recovery, and coal combustion
to supply hot gas to the dolomite heater.
ENERGY COST
The energy cost contribution of the low-pressure one-step
regeneration process is shown in Table E-5. Costs have been obtained
on the basis:
• Interest during construction at 7~l/2%
• Capital charges at 15%
• Operating and maintenance at 5%
E-8
-------
• 70% capacity factor
• No credit for recovered sulfur
• Coal at 4SC/MM Btu
• Methane at 80C/MM Btu
• Dolomite at $2/ton.
The total fluid bed combustion power plant would have an energy cost
of 11.3 mills/kWh, plus the regeneration process energy cost shown in
Table E-5.
TABLE E-5
ENERGY COST (mllls/kWh)
Capital Charges
plus O&M
Fuel Cost
Dolomite Cost
TOTAL
Process Sulfur Load
0.06 | 0.03 1 0
1.79 1.05 0
0.46 0.23 0
0.24 0.12 Q_
2.49 1.40 0
.01
.50
.08
^04
.62
SENSITIVITY OF RESULTS
The sensitivity of the cost results for the low-pressure
one-step regeneration process with respect to the various parameters
which apply to the process are summarized in Table E-6. Numerical
results in the table refer to a process sulfur load of 0.03.
E-9
-------
PAGE NOT
AVAILABLE
DIGITALLY
-------
F
-------
APPENDIX F
BASE DESIGN — TWO-STEP REGENERATION PROCESS
PROCESS DESCRIPTION
The two-step sorbent regeneration requires two fluidized
bed reactors to carry out the conversion of calcium sulfate to calcium
carbonate. In the first step calcium sulfate is reduced to calcium
sulfide by the reaction
at about 1500° F and 10 atmospheres. The second step converts calcium
sulfide to calcium carbonate by the reaction
CaS + H20+C02 + CaCO_ + H2S
at about 1250°F and 10 atmospheres. The regenerated sorbent is recycled
to the boilers while the H«S stream is sent to a sulfur recovery plant.
The two-step regeneration process provides a low-temperature route to
regenerated sorbent.
Figure F-l illustrates the elements of the two-step process. The
process consists of seven elements — the calcium sulfate reducer element,
the HjS generator element, the reducing- gas producer element, the CO-
recovery element, the sulfur recovery element, the tail gas handling
element, and the sorbent circulation element. Reducing gas for the first
step sulfate reduction is provided by the reducing-gas producer element,
while CO. for the second step calcium carbonate production is provided by
recovering C02 from stack gas and recycled sulfur recovery tail gas.
Tail gases from the calcium sulfate reducer element and the CO-
recovery element are incinerated or recycled to the fluid bed boilers
in the tail gas handling element. The sorbent circulation element carries
out the function of circulating the sorbent between the fluidized boiler
F-l
-------
Dwg 6202A67
Hot Gas
to
Turbine
A/lake-up
Sorbent
71
NJ
Sorbent
Disposal
Boilers
Sorbent
Circulation
Element
Coal Air
1st Step
Calcium
Sulfate
Reducer
Element
2nd Step
KLS Generator
Element
Reducing
Gas
Stream
Reducing
Gas
Producer
Element
Fuel
Sulfur
Recovery
Element
Sulfur
Recovery
'Tail Gas
CO 2
Recovery
Element
T TStack
Steam Gas
Tail Gas
Handling
Element
Regeneration
Tail Gas
Figure F-l: Two-Step Reqeneration Process Elements
-------
and the regeneration process. Elemental sulfur is recovered in the sulfur
recovery elements.
SPECIFICATIONS AND ASSUMPTIONS
The base design for the two-step regeneration process follows
from the set of specifications and assumptions shown in Tables F-l and F-2.
TABLE F-l
DESIGN SPECIFICATIONS
Fixed coal rate 430,000 Ih/hr (^ 635 MW)
Sorbent type dolomite
Boiler sulfur removal 95%
Variable process sulfur loads 0.01, 0.03, 0.06
Recycle of tail gas to boilers
Recovery of elemental sulfur
by Claus reaction
Reducing-gas production from coal
Four regenerator modules
TABLE F-2
DESIGN SPECIFICATIONS
CaSO^ reducer pressure 9 atm
CaS04 reducer temperature 1500°F
t^S generator pressure 11 atm
H£S generator temperature 1250°F
H2S mole fraction 3%
CO + H2 mole fraction in reducing gas 30%
Sulfur recovery efficiency 95%
Dolomite utilization in boiler 30%
CaSO/ reduced in reducer vessel 100%
Dolomite utilization after H^S generator 10%
Ash in utilized dolomite 0%
MATERIAL AND ENERGY BALANCES
Material and energy balances for the two-step process streams
shown in Figure F-2 are listed in Table F-3. The process is nearly self-
balanced: the steam generated by the process is sufficient to provide
the required auxiliary power. Temperature control and process turndown
is discussed in Appendix G.
F-3
-------
Dwg 6197A62
Utilized
Dolomite
Sulfided
Dolomite
Regenerated
Dolomite
Steam
CaS04
Reducer
Vessel
1500°F
9atm
2)Reduc-
ing Gas
Waste
Heat
Boiler
Coal
Gasifier
[4 ) Coal
Air
CIO) Stack Gas
!4 H.,0
Heat
Generator
Vessel
1250°F
llatm
Reducer Exchanger Steam
Tail Gas
HO
Condens-
er
Tl
Cooling
Water
Stack
t '
Make-up
co2
Recovery
srt\
Rec
C
Air (22) H00
1 I2
Recovery
Heat
Exchanger
Sulfur
Recovery
rj
C2D Sulfur(23)Steam
Steam
Tail
Gas
Handling
Regeneration
Tail Gas
Figure F-2: Two-Step Regeneration Material and Energy Balances
-------
TABLE F-3
TWO-STEP REGENERATION MATERIAL AND ENERGY BALANCE (Process Sulfur Load
0.03)
Stream
Temperature,
'f
Pressure ,
osia
Moles/Hr
Comments
T
Ul
1. Utilized Dolomite 1200
2. Sulflded Dolomite 1500
3. Regenerated Dolomite 1250
4. Coal 200
S. Air 600
6. Reducing Gaa 1900
7. Reducer Tall Gas 1SOO
8. Reducer Tail Gas 1100
9. Reducer Tail Gas 300
10. Stack Gas 300
11. C02 Make-Up 600
12. CO. Reactant Gas 600
13. CO. Reactant Gas 710
14. HjS Gas 1250
15. H2S Gas 110
16. HjS Gas 850
17. Sulfur Recovery 300
Tall Gas
18. CO, Recovery 400
Tall Gas
19. Regenerator Tail 600
Gas
20. Air 600
21. Sulfur 100
22. Water 100
23. Steam saturated
24. Water 100
25. Steam 1000
150 4,030 moles Ca
165 4;030 moles Ca
160 4,030 moles Ca
150 38,600 Ib/hr
150 6,550
135 7,750
132 7,750
128 7,750
124 7,750
15 7,500
170 2,150
170 13,800
165 13,800
160 13,400
157 8,000
153 8,000
147 8.100
140 1,600
155 9,350 (to boiler)
155 650
150 400
210 5.5xl04 Ib/hr
200 5.3x10* Ib/hr
A
2400 6x10 Ib/hr
2400 6x10 Ib/hr
SOX MgO, 15.0ZCaS04, 35. OZ CaC03
501 MgO, 15.0ZCaS, 35.OZ CaC03
50Z MgO, 5.0ZCaS, 45.OZ CaCOj
30Z H
CO
16X C02
502 C02, 50Z
50Z C02, 50Z
50Z C02, 50Z
3Z H2S
5Z H2S
5Z
-------
EQUIPMENT DESCRIPTION
Figure F-3 shows the major equipment items considered in the
two-step process base design.
CaSO, Reducer Element
Four refractory lined pressure vessels (V5), each in series with
a multistage cyclone (V7), make up the CaSO, reducer element. The reactor
and cyclone design is similar to the design used in the one-step process.
H-S Generator Element
As in the CaSO, reducer element, four pressure vessels (V6)
3 \
element.
in series with multistage cyclones (V8) comprise the H_S generator
Sorbent Circulation Element
The sorbent circulation element is identical to the pneumatic
transport scheme used in the high-pressure one-step regeneration process.
Reducing-Gas Producer Element
The reducing-gas producer element consists of a fixed-bed
coal gasifier (V10) with a cyclone particulate control unit (V9)
followed by a waste heat boiler (E4) for temperature control of the CaSO,
reducer. Compressed air for partial combustion of the coal is supplied
by compressor Cl. The Petrocarb coal feeding system (Vll) is used.
Coal storage, handling, and preparation are Included.
C00 Recovery Element
«.
The C0_ recpvery element consists of a pair of hot carbonate
CO. recovery processes. One process recovers CO. from stack gas to make
up for C02 consumed in the H0S generation, and the other recovers CO., from
F-6
-------
VI - BOILER
V2 - HOLD VESSEL
V3 - SORBENT INJECTION VESSEL
Vf - SJRGE VESSEL, SORBENT INJECTION VESSEL
US - CaSOl, REDUCER VESSEL
V6 - H,S GENERATOR VESSEL
V7.V8 - HOLT I STAGE CYCLONE
El - CONDENSER
E2.E3 - HEAT IMTERCHANGERS
E4 - WASTE HEAT BOILER
V9 - CYCLONE
V10 - COAL GAS1FIER
VII - COAL FEEDING SYSTEM
Cl - AIR COMPRESSOR
C2 - TAIL GAS RECYCLE COMPRESSOR
Dwg 724B667
1 I AIR
J_T_ r
LU'i A
ELEMENTS
SORBENT CIRCULATION
CaSOj, REDUCER & H2S GENERATOR
SULFUR RECOVERY
C02 RECOVERY
REDUCING GAS PRODUCER
TAIL GAS HANDLING
L__Oj__L©_
C02/H20 REACTANT GAS
Figure F-3: Two-Step Regeneration Schematic Flow Diagram
-------
the sulfur recovery element tail gas. The combined C0_/H_0 stream from
the two C02 recovery processes is heated in exchange (E2) before entering
the H_S generators.
Sulfur Recovery Element
The sulfur recovery element recovers elemental sulfur from
the H2S stream in a Glaus plant. The H2S stream from the H2S generator
element is cooled to condense out excess water in exchange (El) and
preheated before entering the Glaus plant in exchanger (E3).
Tail Gas Handling Element
The process tail gas is recycled to the fluid bed boilers by
means of compressor (C2). No boiler modifications have been assumed.
CAPITAL INVESTMENT BREAKDOWN
Table F-4 gives the capital investment for the two-step
regeneration process broken down by element. The investments are given
in $/kw based on 635 MW of total plant capacity.
TABLE F-4
TWO-STEP PROCESS INVESTMENT BREAKDOWN
Element
Reducing Gas Producer
CaSO, Reducer
H_S Generator
Dolomite Circulation
Sulfur Recovery
Tail Gas Handling
CO, Recovery
TOTAL PROCESS
Process Sulfur
0.06 | 0.03
39.2 22.8
2.4 1.6
2.9 2.0
4.3 2.6
9.0 5.8
0.9 0.6
18.0 10.6
76.7 46.0
Load
I 0.01
9.7
0.9
1.1
1.0
2.8
0.3
5.2
21.0
F-B
-------
ENERGY COST
The energy cost for the two-step process has been computed
on the same basis as the other regeneration processes examined. The
breakdown is shown in Table F-5.
TABLE F-5
ENERGY COST (mills/kWh)
Cost
Capital Charges
plus O&M
Fuel Cost
Dolomite Cost
TOTAL
Process Sulfur Load
0.06 I 0.03 1 0
2.88 1.72 0
0.73 0.37 0
0.24 0.12 0
3.85 2.21 0
.01
.80
.12
.04
.96
SENSITIVITY OF RESULTS
The sensitivity of the cost results for the two-step regenera-
tion process is similar to the high-pressure one-step process behavior.
The most important variables in the two-step process are theH?S concentra-
tion in the H-S generator product stream and the change in dolomite
B R
utilization, Xc - X_. Other variables have a secondary effect.
o o
F-9
-------
-------
APPENDIX G
TEMPERATURE CONTROL AND PROCESS TURNDOWN
TEMPERATURE CONTROL — ONE-STEP PROCESS
Temperature control of the one-step regeneration process is
accomplished by adjustment of the air/fuel ratio in the gas producer element
and adjustment of the rate of sulfur recovery element tail gas recycle to
the regenerator. In addition to providing the proper regenerator temper-
ature, the gas producer air/fuel ratio and the tail gas recycle stream
must provide the proper CO + H. concentration in the regenerator reactant
gas stream. The regenerator energy balance is sensitive to the dolomite
circulation rate, the dolomite inlet temperature, the SO. mole fraction
in the regenerator product stream, and the regenerator temperature.
Figures G-l and G-2 represent the effects of the dolomite circulation
rate, the dolomite utilization, and the dolomite inlet temperature on the
regenerator heat load. The change in the sorbent enthalpy across the
regenerator is shown as a function of the change in the sorbent utiliza-
tion, X - XV, for dolomite utilizations in the boiler of 0.5, 0.3, and
0.15. Curves for inlet dolomite temperatures of 200 to 1900°F are shown.
Figure G-l represents a bed temperature of 2000°F and G-2 represents a
1900°F regenerator temperature. Figures G-3 and G-4 show the case of
limestone as a sorbent rather than dolomite.
Figure G-5 represents the enthalpy of the regenerator product
gas as a function of the SO. mole fraction. Curves for a regenerator
temperature of 1900°F and 2000°F are shown. The total heat to
be supplied by the gas producer element is then the sum of the regenerator
product enthalpy, the change in the enthalpy of the sorbent stream, and
the heat of the regenerator reaction (3.47 M Btu/lb sulfur converted
if no CaS is formed);
Figures G-6, G-7 and G-8 indicate some gas producer operating
characteristics as a function of the regenerator heat requirement.
G-l
-------
o
36
34
32
30
28
26
24
22
20
18
16
12
10
8
6
4
2
0
Curve 652706-A
Dolomite
Regenerator (T8nipcrature = 2
X B =0.5
s 0.3
0.15
200.°F = Temperature of
Sorbent to Regenerator
1700°F
1500°F
0.1 0.2 0.3 0.4 0.5 0.6
Change in Sorbent Utilization, x$B -XSR
Figure G-l: Dolomite Circulation Effect on
Temperature Control - 2000°F
Regenerator
36
32
30
28
26
24
22
20
18
16
14
12
10
8
6
4
2
0
Curve 652704^
Dolomite
Regenerator Temperature = 1900°F
XSB=0.5
0.15
200°F = Temperature of
Sorbent to Regenerator
0.1 0.2 0.3 0.4 0.5
Change in Sorbent Utilization, XSB - XjR
Figure G-2: Dolomite Circulation Effect on
Temperature Control - 1900°F
Regenerator
-------
Curve 65Z703-A
Limestone
Regenerator Temperature = 2000°F
- lemperaiure 01
Sorbent to Regenerator
0.1 0.2 0.3 0.4 0.5 0.6
Change in Sorbent Utilization, XSB - XSR
Figure G-3: Limestone Circulation Effect on
Temperature Control - 2000°F
Regenerator
Curve 65270;-r.
Regenerator Temperature = 1900°F
0.1 0.2 0.3 0.4
Change in Sorberrt Utilization, Xs
0.5
B-X.B
Figure G-4: Limestone Circulation Effect on
Temperature Control - 1900°F
Regenerator
-------
Curve 652709-B
100
f 70
o
CJ
60
50
03
1 40
o
ol
I 30
01
o>
S1
« 20
rtJ
£ 10
Regenerator heat requirement (to be supplied by
the gas producer element) = regenerator
product enthalpy (Fig. G-5) ^Sorbent stream
enthalpy change (Fig. G-l - G-4) *• heat of reaction
(3.47 M Btu/lb sulfur converted)
Regenerator Temperature = 2050°F
8
10
12
14
16
18
20
22
Mole Percent SO, in Regenerator Product
Figure G-5: Enthalpy of Regenerator Product Gas
-------
I I I I I
Vertical lines at end points of air/fuel
ratio curves indicate maximum possible
regenerator heat input.
s
in
re
o
en
O)
c
o
tJ
ra
o
u
1.0
0.9
0.8
0.7
0.6
SO,
"5 0.5
o
tS
re
o
'•s
on
"35
8
•o
o
Q_
ra
CD
0.4
0.3
0.2
0.1
/ / Atmospheric I Gas Producer Air/Fuel'
/ / — - — 1 Ratio
lOatm
__ Fraction of, Total Reducing Gas
Atmospheric] Stream Generated by the Gas
"" Producer (the remainder by
tail gas recycle)
0 10 20 30 40' 50 60 70'
Regenerator Heat Requirement, M Btu/lb Sulfur Converted
Figure G-6: Gas Producer Characteristics
90
G-5
-------
Figure G-6 shows the gas producer air/fuel ratio (fraction of stoichio-
metric) and the fraction of the total reducing gas stream which is
produced by the gas producer (the remainder being recycled from the
sulfur recovery element) as a function of the regenerator heat require-
ment. Curves of constant SO. mole fractions of 1, 2, 5, 7, 10, and 15%
are shown with pressures of 1 and 10 atmospheres. The maximum feasible
heat requirement for each of the SO mole fractions is also indicated.
A general assessment of the figures (G-l to G-6) indicates that SO. mole
fractions greater than about 15% will not be feasible under most operating
conditions. The only alternatives in order to achieve SO, mole fractions
greater than 15% would be to heat the incoming sorbent to a high temp-
erature (> 2000CF) or to provide heat transfer surface for heating the
vessel, but these alternatives are probably not feasible.
Figure G-7 shows the fuel required to maintain the regenerator
temperature and provide the proper reducing gas as a function of the
regenerator heat requirement. The amount of fuel (coal) is expressed as
the fraction of the boiler fuel input divided by the process sulfur
load. Curves of constant SO. mole fraction (1, 2, 5, 10, and 15%) are
shown with regenerator pressures of 1 and 10 atmospheres.
Figure G-8 shows the temperature of the regenerator reducing
gas as a function of the regenerator heat requirement. Curves of
constant SO. mole fraction are shown.
Figures G-7 and G-8 with Figures G-l to G-6 indicate that lower
SO- mole fractions require higher fuel requirements and thus lower plant
efficiencies, while higher SO- mole fractions require higher reducing-
gas temperatures and thus more severe design conditions. For this
reason the 1 atmosphere one-step process is operated with a regenerator
temperature of 1900°F and with preheat of the inlet dolomite to 1900°F.
Temperature control of the high-pressure (low S02 mole fraction) process
is simpler than that of the low-pressure (high S0_ mole fraction) process.
G-6
-------
•o
S
'
Curve 652702-A
CO
I/)
8 8
o
-x 6
5 5
00 J
c
o
T3 4
to H
I 3
QJ
'5
£ 2
"eu
ra
c
S i
CO
cu
c
O)
S1
I I I I
Does not Include Fuel for Sulfur
Recovery or Tail Gas Incineration
Atmospheric Pressure
1% SO;
10 Atmospheres Pressure
0 10 20 30 40 50 60 70
Regenerator Heat Requirement, M Btu/lb Sulfur Converted
Figure G-7: Gas Producer Fuel Requirement
G-7
-------
Curve 652707-B
4000
o
"HJ
OS
O
CJ
00
3000
3
"S
a:
2000
1400
10 20 30 40 50 60 70 80 90
Regenerator Heat Requirement, M Btu/lb Sulfur Converted
100
110
Figure G-8: Reduclng-Gas Temperature
-------
TEMPERATURE CONTROL — TWO-STEP PROCESS
The two-step regeneration process is a low-temperature process
(when compared to the one-step process). As such, the complications in
dealing with extremely high-temperature gas streams, which exist in the
one-step regeneration process, .are not important •considerations in the
two-step process. The temperatures of the CaSO, reducer and the H«S
generator should be controllable -by .using waste heat 'boilers and heat
exchangers >to regulate the temperature .of the reducing gas .and the COL/ti-0
reactant 'gas. Because process temperature control appears feasible with
the two-step process, detailed parametric studies, as 'were performed for
the one-step process, have not been carried out.
PROCESS TURNDOWN — ONE-STEP PROCESS
Various possibilities exist for changing the regeneration
process load in response to a change in the boiler load. It is assumed
that the S02 mole fraction in the regenerator product should remain nearly
constant at.all loads in order to maintain the performance of the sulfur
recovery element. In the interest of simplifying the control system and
the design of the sorbent circulation element, it is assumed that the
sorbent circulation rate should be constant for all loads. Then, the
only way to keep the SO- mole fraction constant during turndown is to
reduce the fluidization velocity in proportion to the load, or to
maintain a constant fluidization velocity and adjust the regenerator
pressure and temperature with the load change in order to achieve a
constant SO mole fraction. The latter option is much too complex, so
the option of fluidization velocity reduction has .been chosen. In
order to maintain fluidization >of the regenerator during turndown»
multiple vessels are required. Four regenerator vessels have been
specified in the base designs, though two may-be sufficient.
Since the sorbent circulation rate is constant during turndown
and the reducing-gas volume flow rate is being reduced with reduced load,
the temperature of the reducing gas must increase with reduced load to
G-9
-------
maintain the regenerator temperature. This behavior is demonstrated in
Figure G-9. The reducing-gas temperature is shown as a function of the
percent turndown. Curves of constant SO. mole fraction are shown. The
extent of temperature increase is seen to become a greater problem with
the case of high-SO. concentrations than it is for the low-concentration
cases. With four regenerator modules the percent turndown of each vessel
need not exceed 50 percent, so the high temperature problem is reduced.
The four-module concept allows each of the four boiler modules
to be coupled with a single regenerator module so that the regenerator
module can follow the load of the individual boiler module.
PROCESS TURNDOWN — TWO-STEP PROCESS
Due to the low temperature of the two-step regeneration
process the power plant load follow using a fixed sorbent circulation
rate should be accomplished without the high-temperature complications
found in the one-step process. Using four modules, the CaSO, generator
and H.S generator performance should easily be maintained during turndown.
Fewer modules could be considered; The two-step process turndown has not
been studied in detail.
G-10
-------
420D
4000
3800
3600
Q
B 3400
ro
L_
Qi
d.
E
Curve 652701 -A
o
CD
32°°
3000
2800
k_
o>
c
o>
CT>
« 2600
2400
2200
200P
Conditions:
Fixed S(L mole fraction
Fixed regenerator pressure and temperature
Fixed dolomite circulation rate
at full load XSB= 0.3, XSR=0.1
dolomite inlet temperature = 1200°F
) 10% S02; 1 atm , 1900°F regenerator:
dolomite preheat to 1900°F
D 2%S02; 10 atm
Regenerator
3) 1%S02: 10 atm
2000°r
_L
25 50
Percent Regenerator Turndown
Figure G-9: One-Step Process Turndown
75
G-ll
-------
H
-------
APPENDIX H
EFFECT OF BOILER .CONDITIONS ON REGENERATION PROCESS COST
CARBONAT-ING 'BOILER CONDITIONS ~ ONE-STEP 'PROCESS
If the conditions in the fluid bed 'boilers (temperature,
pressure, 'CO™ concentration) are.such that the dolomite in the bed is
in the 'Carbonated form (MgO • CaCO_), then the base design and cost for
the 'high-pressure -and low-pressure one-step regeneration .processes
must -be reevaluated. Because .the 'one-step regenerator -qp'erating 'conditions
will always lead to calcination of the dolomite, and because the
endothermic nature of the calcination reaction would not permit
simultaneous 'maintenance of the regenerator temperature and the SO.
mole fraction, the .process must be altered
-------
Dwg 6207A27
Owg 6207A28
N>
Fresh
Dolomite
Coal
Air
Compressor
Booster
Compressor
Fluid Bed
Boiler
(1600°F)
T
Cyclone
Regenerated
„ Sorbent ICaO)
Utilized
Sorbent (1500°F)
Fines (CaCOj)
I Combustion
_GasJ1900°F)
Calcining
Vessel
(1900°F)
t
Cyclone
Calcined Sorbent ICaO)
to High- or Low-Pressure
Regenerator (1900°F)
Fines
Combustion Gas (2500°F)
t>~
->
^
Coal
Slagging
Combustor
I Sorbent
1500°F from Boiler
Fines
Carbonating
~1200°F
Utilized Sorbent (CaCCty
1200°F to Two-Step
Regenerator
Stack Gas Compressor
Slag
Stack Gas
300°F
Figure H-2: Calcining Boiler Conditions Flow Diagram
Figure H~l: Carbonating Boiler Combination Flow Diaqram
-------
TABLE H-l
COST INCREASE FOR CARBONATING BOILER CONDITIONS - HIGH-PRESSURE
AND LOW-PRESSURE REGENERATION PROCESSES
Process Sulfur
Load
Increase in Capital
Investment, $/]**
100% Recarbonation
50% Recarb'onation
Increase in Energy
Cost, mills/kWh
100% Recarbonation
50% Recarbonation
0.01
0.03
0.06
13.5
29.0
46.0
9.0
19.5
30.0
0.45
0.86
1.35
0.30
0.57
0.90
-------
consider in the process design. In addition to the large cost increase
in the low pressure and high-pressure one-step regeneration processes
if carbonation takes place in the boiler, the drastic increase in the
process complexity must also be acknowledged.
The costs in Table H-l are extremely sensitive to the dolomite
circulation rate, in addition to the process sulfur load and the
percent recarbonation. The process cost increase should change with
the circulation rate in exactly the same manner as it changes with
the process sulfur load. Low circulating rates, or identically high
B U
X -X , are advantageous in all cases.
3 b
CALCINING BOILER CONDITIONS — TWO-STEP PROCESS
If the conditions in the fluid bed boilers are such that the
dolomite in the bed is in the calcined form (MgO * CsO) then the base
design and cost of the two-step regeneration process must be re-
evaluated. The conditions in the two-step regeneration process will always
lead to carbonation of the dolomite, and, because of the highly exothermic
nature of the carbonation reaction and the large quantity of CO-
consumed, temperature control of the H2S generator would not be feasible.
Carbonation of the dolomite must be carried out before the regeneration
process in a separate vessel.
Figure H-2 shows the major equiument required to carrv out
the carbonation step using stack gas as the CO- source. Stack gas
compressed to 10 atmospheres is reacted directly with the utilized
dolomite in a fluidized bed at about 1200°F. The 1200°F gas produced
is cleaned in a cyclone and then cooled in a waste heat boiler to recover
some useful energy. Table H-2 estimates the increase in the orocess
capital investment and energy cost above the base case cost for 100%
and 50% recalcination of the dolomite in the fluid bed boilers. The
costs in Table H-2 are extremely sensitive to the dolomite circulation
rate, which has been assumed to be the base circulation rate.
H-4
-------
TABLE H-2
COST INCREASE FOR CALCINING BOILER CONDITIONS
TWO-STEP PROCESS
Process Sulfur
Load
Increase in Capital
Investment, $/kw
100% Recalcination
50% Recalcination
Increase in Energy
Cost, mills /kWh
100% Recalcination
50% Recalcination
0.01
0.03
0.06
9.5
20.4
32.9
6.3
13.6
22.0
0.36
0.79
1.33
0.24
0.53
0.89
-------
-------
APPENDIX I
ONCE-THROUGH PROCESS
The capital investment and energy cost of pressurized fluid
bed combustion with once-through utilization of the sulfur sorbent (dolo-
mite or limestone) has been estimated. The parameters involved in the
once-through estimate are the weight fraction of sulfur in the coal, Us,
the sulfur removal efficiency, in the boiler, e, the sorbent feed rate
to the boiler, moles Ca/mole sulfur fed, m, the sorbent purchase plus
disposal cost, S/ton, and the boiler conditions — calcination or
carbonation.
The cost of the regenerative boiler system has been previously
estimated. The once-through boiler system consists of the same
components — sorbent storage and handling, sorbent feeding, and sorbent
disposal — designed for higher sorbent rates. The once-through process
excludes the cost of sorbent regeneration and circulation. The cost in
$/kw for sorbent storage and handling, sorbent feeding, and sorbent
disposal are estimated in Figure 1-1, based on 635 MW of plant output.
The costs .shown do not include interest during construction or general
items and engineering.
The true investment expressed as $/kw depends on the energy
lost in heating the sorbent and calcining the sorbent. With a base
cost of $226.62/kw for a once-through process excluding sorbent storage,
handling, feeding, and disposal the total plant investment is, in $/kw,
[226..6Z + Cost of Sorbent System] 1.13 + 5.5
based on 635 MW plant output. This cost is corrected'by an efficiency
loss term which for dolomite is
[0.9938 - 0.286 W m]'1
s
with calcination of the dolomite in a boiler operated at 1750°F, or
[0.9959 - 0.0955 W m]"1
1-1
-------
Curve 651667-A
e>
o<
I
andjtora
^•—•
Sorbent Disposal
1000 2000 3000 4000
Dolomite/Limestone Feed Rate, Tons/Day
Fiqure I~l- Sorbent Handling and Feeding Costs
Curve 651666-A
19
5T17
'e
13
12-
11
Fixed Coal Rate to Boiler =430,000 Ib/hr,
—^— Calcination in Boiler
— — No Calcination in Boiler
W =wt Fraction Sulfur
5 in Coal
m = Dolomite Feed Rate,
Moles Ca/Mole
Sulfur
0.10
0.20
0.30
0.40 0.50
0.60
Figure 1-2: Once-through Energy Cost
1-2
-------
with no calcination of the dolomite in a boiler operated at 1750°F.
The correction can be substantial for large dolomite rates.
The plant energy cost is computed using the basis previously
applied. The purchase plus disposal cost of dolomite is
used as a parameter having values of 2, 5, 10, and 15 $/ton. Again, the
energy cost, in mills/kWh, has been corrected to reflect energy losses
associated with the sensible heat of the discarded dolomite and the
calcination of the sorbent. Figure 1-2 shows the energy cost results for
a fixed coal feed rate of 430,000 Ib/hr. Note that the single variable
W m is sufficient to account for all once-through conditions, independent
of sulfur removal efficiency. The variable W m is shown for values as
large as 0.60 which would correspond to, for instance, a coal containing
6 wt %•sulfur requiring a dolomite feed rate of 10 mole Ca/mole sulfur.
This extreme case appears to be an upper limit which would never be
approached. A more realistic value of the variable would be 0.08
corresponding to a sulfur content of 4 wt % and m = 2 (calcium utilization
of 45% for 90% sulfur removal). Even lower values seem probable.
The energy cost results in Figure 1-2 indicate that once-through
operation with high-sulfur coals and with calcium utilizations of 50%
or better will yield cost advantages over regenerative operation or a
conventional coal-fired power plant using limestone wet scrubbing even
for dolomite costs of $15/ton.
1-3
-------
J
-------
APPENDIX J
CONSTANT LOAD CONCEPT
The major equipment required for the constant load concept
is shown in Figure J-l for the cases of the 1 atmosphere one-step
regeneration process and the two 10 atmosphere regeneration processes.
The equipment consists of circulation elements, storage vessels,
sorbent heaters and coolers, and single regenerator modules rather
than the multiple modules used in the variable load concept. It has
been assumed that the normal load variation of the power plant is over
a short time period so that a two-week storage capacity is sufficient
to allow constant operation of the regeneration process. On-off
control can also be used.
Table J-l gives the capital investment for the constant load
concept applied to the three regeneration processes. Figure J-2 compares
the capital investment for the constant load concept and the variable
load concept with the low-pressure one-step regeneration process.
Because the variable load concept of the low-pressure one-step regen-
eration process required sorbent heaters and coolers and much of the
circulation equipment shown in Figure J-l, the cost of the constant
load concept is only slightly greater than that of the variable load concent,
Thus, it may be advantageous from the process control standpoint
to use the low-pressure'one-step process as a constant load process
rather than as a variable load concept.
Figure J-3 shows the comparison between the capital investment
for the constant load and variable load concepts with the two high-
pressure regeneration processes. The constant load concept is
considerably more expensive than the variable load with these processes.
Table J-2 presents the energy cost breakdown for the three
regeneration processes using the constant load concept. The energy
cost for the constant load 1 atmosphere one-step regeneration process
is nearly identical to the cost of the variable load concept. The energy
cost of the high-pressure constant load concepts is considerably greater
than that of the variable load concepts.
J-l
-------
Owg 72IiB668
LOW PRESSURE REGENERATION CONCEPT
REGENERATED
DOLOMITE
TO STACK
AIR
HIGH PRESSURE REGENERATION CONCEPT
TO STACK
r JCYCLONE
TH— INJECTOR
AIR
COOLER
HEATER
KTO STACK
REGENERATION
PROCESS -
HIGH-
PRESSURE
Figure J-l: Constant Load Concept Flow Diagram
-------
Curve 65279UA
50
40
30
20
10
I \ \ I
One-Step at Low-Pressure
tf*
^r
&%?.
K&
-------
TABLE J-l
CAPITAL INVESTMENT FOR CONSTANT LOAD CONCEPT
Process
Sulfur Load
Investment, $/kwa
One-Step-1 Atm
10% S0->b
One-Steo-10 Atm
1% SO?0 | 2% SO?0
Two- Step0
0.01
0.03
0.06
15.1
29.9
48.9
22.1
45.8
73.4
17.6
37.5
59.0
24.2
52.9
86.8
Does not include interest during construct! and fee - 635 MW(e)
plant capacity, mid-1973 costs, 70% capacity factor.
Incineration of process tail gas.
'Recycle of process tail gas to fluid bed boilers.
-------
Curve 652792-A
100
90
80
I I I I I
Two-Step Process
High-Pressure One-Step Process -1% S02
High-Pressure One-Step Process - 2% S02
i
Variable Load Concept
Constant Load Concept
•W-
o>
60
r so
CO
40
30
20
10 -
1
1
0 0.01 0.02 0.03 0.04 0.05 0.06
Process Sulfur Load
Figure J-3: Constant vs Variable Load for the High-
Pressure Processes
J-5
-------
TABLE J-2
CONSTANT LOAD CONCEPT ENERGY COSTS (mills/kWh)
Process
One-Step at 1
W ( -MX )
Atm
0.06
0.03
0.01
Capital and
O&M
1.84
1.13
0.56
Fuel
0.46
0.23
0.08
Dolomite
($2/Ton)
0.24
0.12
0.04
Total
2.54 + 11
1.48
0.68
.30
One-Step at 10 Atm (1% S02>
0.06
0.03
0.01
2.75
1.72
0.83
1.30
0.65
0.22
0.24
0.12
0.04
4.29
2.49
1.09
+ 11.30
One-Step at 10 Atm (2%
0.06
0.03
0.01
2.21
1.40
0.66
0.88 0.24 3.33 + 11.30
0.44 0.12 1.96
0.15 0.04 0.85
Two-Step
0.06
0.03
0.01
3.26
1.98
0.92
0.73 0.24 4.23 + 11.30
0.37 0.12 2.47
0.12 0.04 1.08
J-6
-------
K
-------
APPENDIX K
PROCESS COMPARISONS
The costs associated with the base designs of the three regen-
eration processes examined are compared as a function of the process
sulfur load in Figures K-l and K-2. Figure K-l gives the process capital
investments, for both boiler carbonation and calcination conditions,for
each of the processes. Figure K-2 shows the process energy costs, for a
dolomite cost of $2/ton.
Figures K-2 through K-10 show the total power plant energy cost
for the three regeneration processes and for the once-through dolomite
operation. The costs are expressed as a function of the dolomite make-up
rate, the weight fraction of sulfur in the coal, and the dolomite purchase
plus disposal cost. A fixed sulfur removal efficiency of 95% in the
boiler and a calcium utilization of 30% in the boiler is assumed.
Figures K-3 to K-6 represent the case in which utilized dolomite is removed for
disposal before the regeneration process while Figures K-7 to K-10 indicate
the cost behavior for the case in which the dolomite is removed for disposal after
the regeneration process. The optimum boiler condition (calcination or
carbonation) is assumed for each regeneration process. The cost behavior
with both calcination and carbonation conditions are shown for the once-
through process.
For the case of disposal before the regeneration process
(Figures K-3 to K-6) it is seen that the energy cost is reduced as the dolomite
make-up rate is increased for both the $2/ton and the $5/ton dolomite.
This is because the process sulfur load is reduced as the dolomite make-
is
up rate is increased [being equal to Wc (£- mx_)], which reduces the
O 9
process energy cost to a greater extent than the cost of dolomite
increases the energy cost. This behavior is sensitive to the calcium
utilization in the boiler, X_, but is not present for the case of disposal
K-l
-------
ro
100
90
4 80
70
3 60
Q.
» 50
O
30
20
10
Curve 6S27<»-A
Curve 652796-B
High-Pressure One-Step -
Low-Pressure One-Step
High-Pressure One-Step
Two-Step Process
-»- J7>
-2*SO, /> ' J3)
Calcination
Conditions
Carbonation
Conditions
*'/
_L
I
0.1 0.2
0.3 0.4 0.5
Process Sulfur Load
0.6
Figure K-l. Regeneration Process Investment Comparison
5.0
4.0
S
o
:3.0
-s 2.0
re
I
1.0
High-Pressure One-step - I/ SOj
Low-Pressure One-step
High-Pressure One-step - Z'/ SO
Two-step Process z
Calcination
Conditions
Carbonation
Conditions
I
I
I
I
0.01 0.02
0.03 0.04 0.05
Process Sulfur Load
0.06
Fiqure K-2: Regeneration Process Energy Cost Comparison
-------
CO
12
11
Curve 652797-B
Disposal before Regeneration
Oolonite Utilization In Boiler = 0.30
Dolomite Purchase f Disposal at 52/Ton
95V Boiler Sulfur Removal
Fixed Coal Rote to Boilers
——^^— One-step, Low-Pressure
— — One-step, High-Pressure, S02 s I/
_._ One-step, High-Pressure, S02 • 2/
««• Two-step (Carbonation Conditions)
...._• Once-through Process-Top Curve for
Calcination, Lower for Carbonation
\^J /
15
I14
t 13
12
11
Curve 6S2798-B
Disposal before Regeneration
Doloni te'Uti lizaiion In Boiler • 0.30
Dolomi te Purchase * Disposal at 55/Ton
95/ Boiler Sulfur Removal
Fixed Coal Hate to Boilers
— . -One-step. Low-pressure
— -—One-step, High-Pressure, S02 • I/
— •—One-step. High-Pressure, SO, 3 2/
• »— -Two-step, (Carbonation Conditions)
.......Once-through Process-Top Curve for
Calcination. Lower for Carbonation
^
N>N
\\
6wt%
Sulfur
Zj Coal /
f*** J
4*
I
0 1 2 3 4 5 6
Dolomite Make-up Rate, moles Ca/mole Sulfur
Figure K-3. Energy Cost Comparison Disposal before Regeneration
$2/Ton Dolomite
1123456
Dolomite Make-up Rate, moles Ca/mole Sulfur
Figure K-4: Energy Cost Comparison: Disposal before
Regeneration, $5/Ton Dolomite
-------
Curve 652799-B
.c
in
E
s'
u
>!
Plant Enert
t—i
UJ
1_
i
Q.
12
11
i i i i l i
Disposalbefere Regeneration
Dolomite Utilization In Boiler = 0 30
Dolomite Purchase I Disposal at $IO/Ton M
95/ Boiler Sulfur Revival / ^ 1
1 ' /
! 1 /
- — 1 / / / /
/ ' / /
/ / / / /
_ ^r >f /
// / / / /
r*S f / / '
^f^^^^^^ / / * /
~'~ ' , / '
H W/// /
L-^1^ I f / ' ^'
r-*4fJ/ //''
*^JJ7 ^'"
// S- • One- step, Low-pressure
Sulfur-
Coal
_
>» _
/ '* — — One-itep, High-Pressure, SO, = 1%
^«* __.._ One-step, High-Pressure, SI
35 - 21
JP- —— - ~ Two-step (Carbonatlon Conditions)
9^ ---^Once-throu^i Process-Top Curve for
Calcination, Lower for Carbonatlon
111111
1123456
Dolomite Make-up Rate, moles Ca/mole Sulfur
Figure K-5 Energy Cost Comparison: Disposal before
Regeneration. $10/Ton Dolomite
I14
13
12
11
Curve bS2800-B
Sulfur
Coal
_L
I
Disposal bafere Regeneration -
Dolomite Utilization in Boiler o 0.30
dolomite Purchase 4 Disposal at $IS/Ton
One-step, Low-Pressure
•^— ^•One-step, High-Pressure. S02 s 1%
^—• ^One-step, High-Pressure. SOj B 1%
• ___T«o-step (Carbonatlon Conditions)
-••• ••Once-through Process-Top Curve for
Calcination, Lower for Carbonatlon
I I I I
0123456
Dolomite Make-up Rate, moles Ca/mole Sulfur
Figure K-6 Energy Cost Comparison. Disposal before Regeneration,
$15/Ton Dolomite
-------
after the regeneration process (Figures K-7 to K-12) unless the calcium
utilization after
R
being Wg(e - mxg)]
i»
utilization after regeneration, X_, is large [the process sulfur load
K-5
-------
Curve 652801-B
Curve 652802-6
0\
15
14
!
I 13
12
Disposal fler Regeneration
ion after Regen = 0 10
+ Disposal at $5/Ton
Removal
o Boilen
^^— One-step, Low-Pressure
— —One-step. High-Pressure. S02 = IX
^.—One-step, High-Pressure, SOj = 2%
.__.Two-step (Carbonation Conditions)
.,. ^.Once-through Process-Top Curve for
Calcination, Lower for Carbonation
I I I
123456
Dolomite Make-up Rate, Moles Ca/mole Sulfur
Figure K-7 Energy Cost Comparison: Disposal after Reqeneration.
$2/Ton Dolomite
123456
Dolomite Make-up Rate, moles Ca/mole Sulfur
Figure K-8: Energy Cost Comparison
$5/Ton Dolomite
Disposal after Regeneration.
-------
Curve 6S2803-8
15
14
£ 1J
12
11
/ft
_L
_L
Disposal after Regeneration
'Dolomite Utilization after Hegen = 0 10_
Dolomite Purchase + Disposal at$IO/Ton
95% Boiler Sulfur Ronoval
FTxed Coal Rate to Boilers
-One-step. lov-Pres:,... „
— — One-step. High-Pressure. S02 = IX
— .——One-step, High-Pressure, S02 - 2X
— — — — Two-step (Carbonation Conditions)
• Once-through Process-Top Curve For
• Calcination. Lower for Carbonation
1 2 3 4 5 6
Dolomite Make-up Rate, moles Ca/mole Sulfur
Figure K-9: Energy Cost Comparison Disposal after Regeneration,
tlO/Ton Dolomite
Curve
16
15
13
12
Disposal after Regeneration
Dolomite Utilization after Regen. = 0.10
Dolomite Purchase 4 Disposal at $15/Ton
95/ Boiler Sulfur Removal
Fixed Coal Aaie to Boiler»
//
X /•
X X
y x
. f— -One-step,
y X — ^On"-"*?.
//V // X _ _..,_0ne-step, High-Pressure, SflJ - 2?
Xf / * /x* --_.Two-step (Carbonation Conditions)
// SS ^—..•——Once-Ehraugh Process-Top Curve far
Ji \fS 1 Calcination, Lowrr for Carbonation
LoM-Pressure
— — One-stei*. High-Pressure, S02 = I1/
9123456
Dolomite Make-up Rate, moles Ca/mole Sulfur
Figure K-10 Energy Cost Comparison- Disposal after Regeneration.
*15/Ton Dolomite
-------
L
-------
APPENDIX L
A GENERAL STUDY OF THE
ONE-STEP REGENERATION PROCESS
INTRODUCTION
The regeneration of dolomite or limestone sulfur sorbents is
perhaps the area of greatest uncertainty in the design of the pres-
surized fluid bed combustion power plant. Various regeneration schemes
have been proposed, but only bench-scale investigation has been carried
out at this time. A two-step regeneration process has been considered
in the design of a 635 MW pressurized fluid bed combustion power plant,
with the regeneration process design of a preliminary nature and
based on little experimental information. A one-step regeneration
process has also been proposed in which the CaSO, from the fluid bed
boiler is reacted with a reducing gas of H_ and CO to produce CaO and
SO.. The one-step process has been investigated by Esso with bench-
scale equipment and will be utilized in a continuous pressurized fluid
2
bed combustion miniplant which they will operate.
Sensitivity analyses are being carried out to assess both
the one-stage and the two-stage pressurized regeneration systems and
selected alternatives — e.g., once-through, atmospheric pressure re-
generation — and the corresponding sulfur recovery system. An analysis
of the one-step dolomite regeneration process as it applies to pres-
surized fluidized bed combustion is being carried out first. The goals
of this sensitivity analysis are to:
• Identify design parameters, operating variables, and
major problem areas involved in the one-step process
o Estimate regeneration system costs and identify the
critical cost factors
• Provide direction for experimental studies to be
carried out by Esso
L-l
-------
• Establish feasibility bounds on the regeneration system
performance variables which can be used to evaluate the
experimental results provided by Esso and Westinghouse.
ONE-STEP REGENERATION CONCEPT
The one-step dolomite/limestone regeneration process con-
sists of a single fluidized bed reaction vessel in which utilized
dolomite/limestone (CaSO,) from the pressurized fluid bed boiler (coal-
fired) is reacted with a H./CO reducing gas to produce CaO and SO-.
The reaction
H2 + H2?
CaS04 + (CQ } + CaO + {£ } + S02 (1)
requires temperatures on the order of 2000°F to produce high levels
(1 to 15%) of SO- at reasonable rates. The SO. equilibrium concentration
is favored by reduced pressure since it is inversely proportional to the
total pressure. The competing reaction
H -> H2°
CaS04 + 4(CQ} * CaS + U{^ } (2)
also occurs to produce the unwanted product CaS.
The regenerated dolomite/limestone is returned to the fluid
bed boiler with the required amount of fresh dolomite/limestone to
supplement the regenerated sorbent's reduced activity. The SO- stream
from the regenerator is sent to a sulfur recovery plant in which
sulfuric acid or elemental sulfur is produced.
L-2
-------
'STUDY BASIS
The elements of the -one-step regeneration process and .the
various design -options --which .are possible tshould be understood .before
.analysis
-------
Make-Up
Sorbent
Hot Gas
to
Turbine
Sorbent
Disposal
Flu id Bed
Boiler
Sorbent
Circulation
Element
Regenerator
Element
Coal Air
Reducing
Gas
Producer
Element
Fuel
Air
Regenerator
Product
Gas
Dwg. 724835^
Sulfur
Sulfur
Recovery
Element
Sulfur Recovery
Tail Gas
Tail Gas
Handling
Element
Regeneration
Tail Gas
Figure L-l: One-Step Regeneration Process Elements
-------
Process Specifications
Process specifications are presented in Table L~l« The one-step
regeneration process design study is based on a coal-fired pressurized
fluid bed combus.tion power plant capacity of 635 MW,to correspond to the
preliminary plant study carried out earlier. This capacity requires a
coal feed rate of 430,000 Ib/hr. Dolomite is specified as the SO. sorbent
for this study, though limestone could easily be considered in addition
to the dolomite. A boiler pressure of 150 psia is specified, along with
the specification that elemental sulfur rather than sulfuric acid is to
be recovered and that disposal of the spent dolomite should take place
before the regeneration step rather than after. The advantages of
recovering elemental sulfur rather than sulfuric acid are well known.
The advantages in disposing of the spent dolomite before regeneration
rather than after are that the total sulfur load on the regeneration
process is reduced, and the dolomite is in a more suitable form for
disposal before regeneration (CaSO.) than it is after (CaS).
TABLE L-l
PROCESS SPECIFICATIONS
Power Plant
Capacity
Fluid Bed Boiler Pressure
Fluid Bed Boiler Temperature
Boiler Sorbent Condition
Coal Rate
Sulfur Sorbent
Disposal of Dolomite
Sulfur Recovery
Plant Capacity Factor
Turndown Ratio
coal-fired, pressurized fluid
bed combustion power plant
635 MW
150 psia
* 1750°F
calcined
430,000 Ib/hr
dolomite (50% CaC03> 50% MgC03>
before regeneration
elemental sulfur
70%
4:1
L-5
-------
An effort has been made to specify a minimum number of items
in order to study the general cost behavior of the regeneration process.
Variables and Design Options
Operating and design variables considered in the study are
listed in Table L-2. In addition to these variables, which the designer
or operator can control or set, there are a number of performance
variables which are dependent on the design and operating variables.
These performance variables are listed in Table L-3 in very general terms.
TABLE L-2
OPERATING AND DESIGN VARIABLES
Boiler Fuel Sulfur Content
Boiler Sulfur Removal Efficiency
Regenerator Vessel Pressure
Regenerator Vessel Temperature
Regenerator Fluidization Velocity
Regenerator Bed Depth
Number of Regenerator Modules
Regenerator Vessel Design (internals, freeboard, etc.)
Gas Producer Design and Fuel Utilized
Air/fuel Ratio to Gas Produced or t^+CO Concentration in Reducing Gas
Fraction of Claus Tail Gas Recycled to Regenerator
Regenerator Vessel Temperature Control Scheme
TABLE L-3
PERFORMANCE VARIABLES
Composition of the Regenerator Gas Product Stream
Composition of Reducing-Gas Stream
Composition of Claus Tail Gas
Composition of Stone from Boiler
Composition of Regenerated Stone
Dolomite Make-up Rate
L-6
-------
•Design options which are .considered in the study are shown in
Table L-4. 'The reducing-gas producer element may consist of a coal
gasifier, an oil -gasifier, a methane .partial oxidation burner, or a
methane reformer, all of which are conventional gas producing methods.
Other reducing-gas producer options which carry out both reducing-gas
production and regeneration of dolomite in the same vessel have not
been considered because they depend largely on unexplored technology
and .may lead to many complications.
The tail gas handling element may consist of an incinerator
and a turbine lexpander to recover some of the flue gas energy. The
option of recycling the tail 'gas to -either the fluid bed boiler or
the gas-turbine
-------
TABLE L-5
SIMPLIFIED VARIABLES
Prime Variables
• Regenerator pressure - 10, 5, 2 atmospheres
• Reducing-gas producer scheme - coal gasification, oil
gasification, methane partial oxidation,
methane reforming
• Process sulfur load - Wg[e-mX ] = 0.01, 0.03, 0.06
• Mole fraction SO, in regenerator product - Y..
YS02 = 0 - 0.05 at 10 atmospheres
YgQ. = 0 - 0.10 at 5 atmospheres
YSO? =0-0.15 at 2 atmospheres
Secondary Variables
• Regenerator temperature - 1800-2200°F
• Air/fuel ratio to gas producer - 40-80% of stoichiometric
• Regenerator fluidization velocity - 5-15 ft/sec.
• Regenerator vessel height - 15-65 ft
• Fraction of Glaus tail gas recycled to regenerator - 0-0.50
• Tail gas recycle to boiler or incineration of tail gas
• Change in dolomite sulfur utilization across regenerator,
XB - XR = 0.1 - 0.3
s s
• Regeneration process thermal efficiency - 50-75%
The philosophy of the investigation is then to estimate
capital costs and energy costs for the one-step regeneration process
as a function of the prime variables. Values of the prime variables
considered are included in the table. Since the secondary variables
have a small effect on the capital and energy costs these variables
are set to span a reasonable range of values which are designed to
cover the probable situations and which are included in the table.
L-8
-------
The final result, then, is probable ranges of capital and energy costs
for the one-step regeneration process as a function of the prime
variables. These ranges are narrow enough so that concise conclusions
may be drawn as to the process feasibility and behavior.
The four prime variables listed are clearly the most signi-
ficant, based on simple considerations. The process sulfur
load, W [e-mX], is a combination of the fuel sulfur content, W ,
s ° s
(Ib sulfur/lb fuel);the boiler sulfur removal efficiency, c; the
dolomite make-up rate, m, (moles 'Ca/mole S fed);and the fractional
utilization of the dolomite in the fluid bed boiler, X_. This variable
9
is representative of the amount of sulfur which the regeneration pro-
cess must handle and, therefore, sets the scale of the process.
The mole fraction of SO. in the regenerator product gas is
representative of the size of the regeneration process equipment and
the gas 'producer fuel requirements, since the amount of sulfur handled
is fixed by the variable W [e-mX ]. Since the equipment sizes and fuel
S S
requirements are inversely proportional to Yrt , this prime variable
50-
has an extremely large effect.
Since the equilibrium SO- concentration is inversely pro-
portional to the regenerator pressure and the equipment sizes are
directly proportional to the pressure, the possible advantages or
disadvantages to be gained by operating the regeneration process at
pressures lower than the boiler pressure of 10 atmospheres should be
investigated. This also leads to the possibility of lowering the
regenerator temperature by reducing the pressure and thereby possibly
improving the regenerated dolomite activity.
The type of fuel used to produce reducing gas has a great
effect on the capital cost, system complexity, and energy cost
associated with the one-step regeneration process.
The most important of the secondary variables is the change
:e sulfur utilization, X^ - 3T. The rate
s s
lation is inversely proportional to this variable.
B _R
in dolomite sulfur utilization, Xo - x . The rate of dolomite circu-
s s
L-9
-------
MATERIAL AND ENERGY BALANCES
Thermodynamic studies indicate some of the behavior to be
expected as a function of the regenerator pressure and the operating
conditions. Energy balances around the regenerator indicate the re-
quirements for regenerator temperature control. Using the insight gained
from thermodynamics, overall material and energy balances have been
/
carried out which allow the sizing of equipment as a function of the
prime and secondary variables outlined in Table L-5.
Thermodynamics
Thermodynamics indicates that the SO. concentration resulting
from the equilibrium of reaction 1 and 2 is inversely proportional to
the regenerator pressure. This behavior provides the incentive to study
the possibility of operating the regeneration process at pressures below
the boiler pressure. Thermodynamics also predicts that nearly all of the
H_ and CO in the reducing gas will be utilized to produce S02 and CaS by
reaction 1 and 2. This leads to the expectation of low CaS contents at
pressures as low as 2 atmospheres and high CaS contents in the regenerated
dolomite at pressures of 10 atmospheres.
The trends indicated by thermodynamics can be expected in an
actual operation. For this reason the ranges of SO. mole fractions as noted
in Table L-6 are considered at each of the pressures to be investigated.
TABLE L-6
RANGES OF S02 MOLE FRACTIONS
Regenerator
Pressure, atm.
10
5
2
Range of S02
Mole Fractions
0.01-0.05
0.01-0.10
0.01-0.15
L-10
-------
These ranges are chosen on the basis of a regenerator temperature of
about 2000°F. Data collected by Esso from batch operations lie within
2
these ranges.
Regenerator Temperature Control
The temperature of the dolomite regenerator vessel may be
maintained at about 2000°F by controlling the temperature of the
streams entering the regenerator. This may require heating or cooling
of the gas producer product, the tail gas recycle stream, or the
utilized dolomite stream. The variables involved in the regenerator
energy balance are numerous and, lacking direct experimental evidence,
it appears that either cooling or heating may be required, depending on
the regenerator pressure, the reducing-gas producer fuel, the dolomite
utilization in the boiler, the temperature of the dolomite stream
entering the regenerator, and the kinetic behavior of the regenerator.
In order to account for the cost of temperature control it
was assumed that the energy loss or input required is not extensive,and
a gas-fired heater necessary to reheat the Glaus plant tail gas recycle
stream would be sufficient to control the regenerator temperature at
about 2000°F. It appears that the temperature control energy requirements
will decrease with decreasing regenerator pressure.
Material and Energy Balances
All major streams have been characterized to the extent necessary
to allow equipment capacities to be estimated as a function of the
primary and secondary variables. Fuel inputs into the reducing-gas
product, the tail gas recycle heater, the Claus plant, and the tail gas
incinerator have been estimated. Auxiliary power requirements for the
regeneration process have also been calculated. Some of the factors
which limit the one-step regeneration process feasibility are brought
out by simple material balances.
L-ll
-------
Material and energy balances have been computed for process
sulfur loads, W [e - mX0], of 0.06, 0.03, and 0.01. A sulfur load
s o
of 0.03 Is assumed to be a reasonable value to be expected with coals
having normal sulfur contents 0\<4 wt %). Values of 0.06 and 0.01 are
probably extremely high and extremely !OH respectively, and have been
chosen to allow extrapolation of the cost results to sulfur loads
within the range of 0.01 to 0.06. Figure L-2 illustrates how the process
sulfur load varies as a function of the fractional utilization of the
dolomite in the boiler for coals having weight fractions of sulfur of
0.06, 0.04, 0.03, and 0.02. A constant sulfur removal efficiency of
0.95 is assumed and dolomite make-up rates of 1.0 and 2.0 moles Ca/Mole S
are examined.
An important factor in determining the feasibility of recycling
Claus plant tail gas to the fluid bed boilers or the gas-turbine
combustor is the extent of modification required to the boiler or
combustor in order to handle the increased load from the regeneration
process. Figure L-3 is a plot of the volume of tail gas being recycled
to the boiler, in percent of the boiler combustion air volume, as a
function of Y_n . Zero recycle of tail gas to the regenerator is
bU2
assumed in the figure, so it represents the most drastic case. If it
is assumed that a 10% volumetric load increase is permissible without
modification, a mole fraction of S02 greater than 0.02 is required
for a process sulfur load of 0.03, and 0.04 for a snlfur load of 0.06.
Since no tail gas recycle to the regenerator has been assumed and the
10% limit is conservative, no boiler modification cost has been charged
for recycling tail gas to the boiler. This assumption may not be valid
for process sulfur loads as high as 0.06, with low SCL concentrations.
Fuel requirements are simply computed as a function of the
process variables. The fuel imput to the gas producer is shown in
Figure L-4 as the fraction of the Btu input of the boiler. Curves are
shown for process sulfur loads of 0.06, 0.03, and 0.01, and for reducing-
L-12
-------
Curve 6505^8-B
0.07
0.06
Boiler Sulfur Removal Efficiency = 95%.
- Sorbent Make-up Rate = 1.0 mole Ca/mole S
---- Sorbent Make-up Rate = 2.0 moles Ca/mole S
Process Sulfur Load = W
*•*
3
CO
0.05
0.04
0.03
0.02
0.01
wt Fraction Sulfur
in Boiler Fuel
l£ \ fc- • ¥ I/* £ f
W$ = wt fraction sulfur
in boiler fuel
E = boiler sulfur removal
efficiency
M = sorbent make-up
B rate
X,. = dolomite sulfur
utilization in
boiler
0.1
0.2 0.3 0.4 0.5 0.6
Dolomite Sulfur Utilization in Boiler
Figure L-2: Regeneration Process Sulfur Load
0.7 0.8
L-13
-------
Curve 650533-B
40
36
» 32
<
i_
Q>
"5
o>
E
28
24
o 20
o
f 16
'=. 12
•s
^ 8
No Recycle of Tail Gas to
Regenerator
0.06 = Process Sulfur Load
10% Increase in Boiler or
Gas-Turbine Input
0 0.02 0.04 0.06 0.08 0.10 0.12
Mole Fraction SCL in Regenerator Product
Figure L-3: Tail Gas Recycle to Boiler
0.14 0.16
L-14
-------
ve 650552-B
T
H
Lr
No Recycle o' Tail Gas to Regenerator
• No Sulfur in Gasifier Fuel
•Gas Producer Air/Fuel Ratio = 40% of
Stoicliiomeiric
Process Sulfur Load = 0.06
---- Process Sulfur Load = 0.03
--- Process Sulfur Load = 0.01
Thermodynamic Composition
at
0.1 -
0.02
0.04 O.Ot 0.08 0.10 0.12
Mole Fraction S02 in Regenerator Product
0.14 0.16
Does not include Methane required for Sulfur
recovery or Tail Gas Incineration
Process Sulfur Load =0.03
Gas Producer Air/Fuel Ratio = 40 - 60% of
Stoichiometric
Coal Gasification
_"LjT_J_ CH. Partial Oxidation
% Recycle to Regenerator
= 0
0.02 0.04 0.06 0.08 0.10 0.12 0.14 0.16
Mole Fraction S02 in Regenerator Product
Figure L-4: Gas Producer Fuel Requirement
Figure L-5: Total Regeneration Process Fuel Requirement
-------
gas production by coal gasification, oil gasification, methane partial
oxidation,and methane reforming. The following assumptions have been
made in generating the figure:
• An air/fuel ratio in the gas producer of 40% of the
stolchiometric value
• No sulfur in the fuel used for reducing-gas production.
The fuel rate to the gasifier is directly proportional to the fraction
of the tail gas which is recycled to the regenerator, so again the figure
represents the most extreme case. The curves rise rapidly at lower SCL
mole fractions and indicate that the scale of the regeneration process
may approach the scale of the fluid bed combustion process at low mole
fractions of SO™. Air/fuel ratios much larger than 40% of stoichiometric
may be suitable, which will reduce the fuel requirement (see Figure L-6).
Additional fuel, CH,, is required for the reduction of S0_ to
H S in the sulfur recovery element, for incineration of Glaus plant tail
gas if that option is chosen, and for heating the tail gas being recycled
to the regenerator. Figure L-5 shows the combined fuel requirements for
gas production and tail gas recycle heating in percent of Btu input to
the boiler. Curves are shown for coal gasification and for methane
partial oxidation with fractional recycles of the tail gas to the
regenerator of 0 and 0.50. The following assumptions contribute to the
figure:
• The process sulfur load is 0.03 and the fuel requirement
is linear in this factor.
• The recycled tail gas is heated from the 3008F Claus plant
exhaust temperature to 2000°F with an 85% efficient gas-
fired heater.
• The air/fuel ratio in the gas producer is taken over a
range of 40% to 60% of stoichiometric.
• Sulfur in the coal used for reducing-gas production
is neglected.
L-16
-------
The total fuel required for the one-step regeneration process
would be about 10% greater than that indicated in Figure L.-5 if Glaus
plant requirements and incinerating requirements were included.
The effect of utilizing a fuel for reducing-gas production
which contains a significant proportion of sulfur is shown in Figures L-6
and L-7. Figure L-6 shows the multiplication factor to be used with
Figure L-4 to account for air/fuel ratios other than 40% of stoichiometric
and fuels containing sulfur. Figure L-7 shows the plant sulfur production
rate as a function of the gas producer fuel sulfur content. The fuel
sulfur content effect can be important at small S0_ mole fractions (< 3%).
Higher air/fuel ratios will reduce the effect of sulfur in the reducing-
gas fuel shown in Figure L-7. In Figures L-6 and L-7 the SO. mole fraction
includes the sulfur contribution of the reducing-gas producer fuel.
Auxiliary power requirements for the one-step regeneration
process consist of the following contributions: compression of air for
reducing-gas- production, compression of tail gas recycled to the
regenerator, compression of tail gas recycled to the,10 atmosphere
boiler, or compression of air for tail gas incineration, power for
dolomite circulation, and power for cooling water,pumping,and fuel feeding.
The total auxiliary power requirement shows the same characteristic
increase with decreasing SO- mole fraction and may vary from 2 to 5 MW at
high S02 mole fractions (^ 10 to 15%) up to 20 to 30 MW at low S0_ mole
fractions (^ 1%).
Not all of the fuel energy which is used for reducing-gas
production, recycled tail gas heating, Glaus plant S0? reduction, and
tail gas incineration is lost. Through dolomite recycle to the boiler,
steam generation in the Glaus plant, recycle of the tail gas to the boiler,
or energy recovery from the tail gas incineration product, it is estimated
that 50 to 75% of the energy entering the regenerator vessel can be
utilized as useful energy for power generation at an air/fuel ratio of
i< 40% of stoichiometric. 'The necessity of insuring the recovery of the
useful energy for power generation is clear from the huge energy require-
ments for regeneration at low S02 mole fractions. Including the expected
L-17
-------
Curve 6505^9-b
2.3
' 2-
<• t
CD '-4^
E «
-,22.0
00
o
-------
Curve 650550-B
700
CO
600
500
c
o
£400
O-
200
100
I
I
Gas Producer Fuel
Oihwt. Fraction Sulfur =0.03
Coal-" " " =0.04
— CH., Oil, or Coal with No Sulfur Content
Air/Fuel Ratio = 40% of Stoichiometric
Claus efficiency = 0.95
Process -
Sulfur Load
0.06 -
- \
0.01
0.02 0.04 0.06 0.08 0.10 0.12 0.14
Mole Fraction SC^ in Regenerator Product
Figure L-7: Sulfur Production Rate
0.16
L-19
-------
thermal efficiency of the reducing-gas production element, the recycle
gas heater efficiency, and the process auxiliary power requirements, the
regeneration process operates with an overall efficiency of 30 to 50% at
an air/fuel ratio of * 40% of stoichiometric. Higher air/fuel .ratios will
reduce both the fuel requirement and the process overall efficiency.
The factors briefly discussed here indicated some clear advantages
in maximizing the SO- mole fraction in the regenerator product. A
quantitative picture will be generated by estimating total energy costs for
the one-step dolomite regeneration process.
EQUIPMENT DESIGN
The individual pieces of equipment making up the five elements
of the one-step dolomite regeneration process have been designed as
a function of the primary and secondary variables outlined in Table 1-5.
The considerations investigated and the assumptions applied to the
equipment design are discussed.
Figure L-8 is a schematic flow diagram showing the equipment
included in capital cost estimates. The equipment and pertinent
assumptions made are as follows:
Regenerator Element
• Regenerator vessel 5-15 ft/sec fluidization velocity
and internal's: 15-65 ft overall height
design temperature 1800-2200°F
purchased equipment cost based on
standard literature references and
price quotations
Sulfur Recovery Element
• Claus plant with SO, does not include tail gas disposal
3
reduction by CH,: cost based on Shell report
L-20
-------
Tail Gas Handling Element
• Boiler recycle option: tall gas recycle compressor
no modification of boiler
cost based on recent price
quotation data
• Incineration option: tail gas incinerator
combustion air compressor
turbine expander (except at 2 atm)
waste heat boiler at 2 atm pressure
cost based on Shell report
Reducing Gas Producer Element
• Tail gas recycle compressor: 0-50% recycle of tail gas
cost based on recent price quotations
o Tail gas recycle heater: gas-fired
0-50% recycle of tail gas
85% efficient
includes combustion air system
cost taken from published information cost
9 Gas producer: 0-50% recycle of tail gas
40-80% of stoichiometric air/fuel ratio
Option; Coal Gasification
« Array of fixed-bed type gaslfiers (Lurgior McDowell-Wellman)
• Capacity: ^320 ton coal/day per gasifier (12 ft. maximum diameter)
• Includes coal lock hoppers and ash lock hoppers
• Does not include: coal preparation or storage
sulfur removal
tar removal
• Thermal efficiency: ^80%
• Gas product: 1200-1500°F
10-25% hydrogen and carbon monoxide
4
• Cost taken from recent price quotations, recent references
• Particle collectors
• Air compressor
L-21
-------
Option; Oil Gasification (Gasification system cost would be essentially
the same for either a residual oil or a
distillate fuel.)
• Array of Shell or Texaco type heavy oil gasifiers
• Capacity: ^33 x 106 scf/hr per gasifier (12 ft diameter, 20 ft.
high)
• Does not include: sulfur removal
soot recovery and recycle
• Thermal efficiency: ^85%
• Gas product: 2000-2500°F
10-30% hydrogen and carbon monoxide
(4)
• Cost estimated from information provided by Shell Oil
• Particle collectors
• Air compressor
• Oil feeding and storage (day tank and 30 day storage tank)
Option; Methane Partial Oxidation
• Refractory lined burners
• Thermal efficiency: ^95%
• Gas product: 2000-2500°F
10-40% hydrogen and carbon monoxide
• Air compressor
Option; Methane Reforming
• Methane reformer vessels
• Thermal efficiency: ^85%
• Gas product: ~1500°F
30-50% hydrogen and carbon monoxide
• Cost taken from quote from the Selas Corporation
L-22
-------
Fit
Be
Boi
id r - —
d
ers ^^s
^^ Hold
Vessel
Sorben
Distribut
/and'
[Coolinc
Vyessel
Sorbent
FS
C
1
on
D .
) i
Circulation t
Transport / \
Air / \
Booster '
Compressor 4:
Sorbent
Disposal
Petrocarb
Feeders
(Injectors!)
r
-
Regenerator
Product Gas
Regenerator
Vessel
Dolomite
Disposal
--M-—i
Owg 721.B35I
• Sulfur
Recovery
Plant
Sulfur Steam
To Stack
I
Recycle Tail
Gas Compressor^.
j
Cyclone
Gas
Producer
Vessel
^
Fuel
Feeder
Air
Compressor
Fuel
From
Storage
Boiler
Recycle
Compressor
Air
Compressor
Sulfur
Recovery
Tail Gas
Options
To Boiler
or
Turbine
Combuster
Air
~I
Tail
Gas
\7 Air
\/Compressor
F
. CH4
Incinerator
Turbine
Expander
Stack
Figure L-8: One-Step Regeneration Process Flow Diagram
-------
Sorbent Circulation Element - X - X = 0.1-0.3
^• o o
• 10 atmosphere regenerator pressure
- Four hold vessels (one for each boiler)
- Sorbent distribution vessel
- Fetrocarb feeder assemblies
- Transport air booster compressor (175 psla)
- 2500 linear feet of solids piping
- Costs supplied by Fetrocarb
• 5 and 2 atmosphere regenerator pressure
- Four hold vessels
- Sorbent distribution vessel
- Dolomite cooling vessel: counter-current flow moving bed with
cooling by steam or water injection. Solids cooled to 800°F
to meet valve temperature restriction.
- Fetrocarb injection and feeder assemblies
- Transport air booster compressor (175 psia)
- 2500 linear feet of solids piping
The capacities and dimensions of these equipment items have
been estimated as a function of the four prime variables for the
ranges of secondary variables indicated.
CAPITAL INVESTMENT
Equipment costs have been estimated from standard cost references
appearing in the literature and from vendor quotes. Capital costs shown
are total installed costs which reflect both field and shop fabrication
costs. Each estimate was made based on factoring f.o.b. equipment cost
to total installed cost. Each item of major equipment was priced
from cost curves derived from published equipment price correlations,
modified in the case of pressure vessels and compressors to reflect
recent firm price quotations for these two categories of equipment.
L-24
-------
The f.o.b. prices of major equipment were totaled, and the total
was broken down into percentages for cost of vessels, compressors,
mechanical equipment, heat exchange equipment, tanks, and pumps.
Appropriate factors were then applied to each category of major equipment
to obtain field material costs (pipe, concrete, structural steel, etc.),
and field labor costs. Escalation was applied as required to material
and labor costs to bring them to mid-1972 levels. The subtotal of all
direct costs (equipment, material, and labor) was then factored to
provide for field indirect costs, home office engineering, field super-
vision, freight, and contractor's contingency and fee. The total
investment so arrived at is over four times the total f.o.b. equipment
cost. Costs do not include interest during construction. A contractor's
contingency and fee of 15% was applied.
Examples of installed costs, in $/kw, of the five process
elements of the one-step regeneration process are shown in the
following discussion. The importance of each element and its cost
behavior is discussed. The total one-step regeneration process cost is
discussed by means of a table of representative values.
Regenerator Element
The cost of the regenerator element for a 5 atmosphere regenera-
tor pressure is shown in Figure L-9. The cost ranges refer to a single
regenerator module with the secondary variables having the values shown
in Figure L-5. The contribution of this element to the total process cost
is small but not negligible. The cost penalty for going to two modules
rather than one will be small, especially if the size of the single
module is reduced from a field-fabricated size to two shop-fabricated
vessels. Turndown requirements may necessitate duel regenerator vessels.
Attempts should be made to maximize the SO. mole fraction in
the regenerator product by proper regenerator design. The overall cost
penalty in using low fluidization velocities and deep beds is slight.
L-25
-------
Curve 650551-B
11
10
*8
"c
I
u 7
_o
o
o
o
CO
o
3
2
1 -
Conditions:
635 MW Plant Capacity
Regenerator Pressure = 5 atm
Single Regenerator Module
Process
Sulfur Load
=0.06 -
0.03
0.01
0.02 0.04 0.06 0.08 0.10 0.12
Mole Fraction S02m Regenerator Product
Figure L-9: Regenerator Element Capital Investment
L-26
-------
Curve 6505^7-8
30.0
I 25.0
"o
•T 20.0
o>
J 15.0
CO
42
1-10.0
TO
Q_
I/)
3
iq
5.0
1
t
I
T
I I
635 MW Power Plant
Gasifier Air/Fuel Ratio = 40% of Stoichiometric
0.7 Exponential Scale Factor of Shell Cost Estimates
Gasifier Fuel
Oil - wt Fraction Sulfur = 0.03
Coal-" " " =0.04
CHL, Oil, or Coal with No Sulfur Content
10 Atmosphere Claus Operation
I
Process
Sulfur Load
0.06
0.03
0.01
0.02 0.04 0.06 0.08 0.10 0.12
Mole Fraction S09 in Regenerator Product
0.14
0.16
Figure L-10: Sulfur Recovery Element Caoital Investment
-------
Curve 6505M-B
to
to
O
o
•s
4->
s
u
-------
Sulfur Recovery Element
The capital investment for the sulfur recovery element is based
3
on a study by Shell. Figure L-10 shows the cost for a 10 atmosphere
Glaus operation with the cost of tail.gas incineration included. In-
cineration was assumed to amount to about 20% of the Claus plant invest^
ment. The effect of sulfur content in the gas producer fuel is
included in the figure. The contribution of the sulfur recovery element
to the total (process cost) is about the same as the regenerator element.
Tail Gas Handling Element
The tall gas handling element cost is shown in Figure L-11 for the
two options listed in Table L-4 with a pressure of 10 atmospheres. The
graph represents a sulfur load of 0.06 with scale factors for sulfur
loads of 0.03 and 0.01 given in the figure. Curves are shown for
percent recycles of tail gas to the regenerator of 0 and 50.
The contribution of the tail gas handling component to the
total one-step regeneration process cost is the smallest of the five
elements. Recycle of the Claus plant tail gas to the boiler is cheaper
than incineration of the tail gas at a regenerator pressure of 10
atmospheres. The costs of the two options are similar at 5 atmospheres,
and the cost of boiler recycle greatly exceeds the cost of incineration
at 2 atmospheres pressure. It should be recalled that no modification
to the boiler system was assumed In the case of the boiler recycle
option (see Figure L-3).
The advantages of recycling the Claus flue gas to the boiler
should be considered even for the low-pressure regeneration where incin-
eration is much cheaper than recycling. These advantages are
e Improved plant sulfur removal efficiency if the tail gas is
recycled to the boiler
• Possibly higher gas-turbine inlet temperatures if the tall gas
is recycled to the gas-turbine combustor
• Decreased fuel requirements for the regeneration process.
These considerations should be examined in much greater detail.
L-29
-------
Curve 6505^-8
40
30
'u
I/I
ru
O
•5 20
-#-•
8
o
I
1C
ra
•s
10
0
\
w
\
\
\ \
\
\
I I I I
Conditions
Process Sulfur Load = 0.06
Regenerator Pressure = 10 atm
Air/Fuel Ratio = 40%
Coal Gasification '
Oil Gasification
CH4 Partial Oxid.
V CH4 Reforming
\ Scale Factor for Sulfur Load of 0.03
v =0.574
\Scale Factor for Sulfur Load of 0.01-
V =0.238
\
\
\
Fraction of Tail
> Gas Recycle to
\ Regenerator
^ jk
\l
\
0.01 0.02 0.03 0.04 0.05 0.06
Mole Fraction S09 in Regenerator Product
0.07
Figure L-12: Gas-Producing Element Capital Investment
L-30
-------
Reducing-Gas Producer Elements
Reducing-gas production represents a large contribution to
the total process cost. Figure L-12 gives costs for the gas producer
element with the four fuel options of coal gasification, oil gasifica-
tion, methane partial oxidation, and methane reforming, and with sulfur
loads of 0.06, 0.03, and 0.01. Percent tail gas recycles to the
regenerator of 0 and 50 are shown. Higher air/fuel ratios will reduce the
cost of reducing-gas production,
The fuel options have the same order of relative cost, with
coal gasification being the most expensive, at all regenerator pressures.
Oil gasification is the second most expensive option at all pressures.
Methane reforming is the least expensive option at 10 atmospheres
pressure,and methane partial oxidation is the cheapest at 2 atmospheres
pressure. The cost of coal and oil gasification is reduced by increasing
the amount of tail gas recycle to the regenerator, while the cost of
methane partial oxidation and methane reforming remains nearly the same
as the amount of recycle is increased.
The merits of the various gas producing options can only be
compared for a given regenerator performance (mole fraction S0_ in the
regenerator product). The specific performance of regeneration with the
reducing gases produced by the four gas producing options is not known and
must be investigated. Investigation of one-step1 regeneration with
synthetic mixtures of hydrogen and carbon monoxide may have little
relation to the performance obtained with a commercial reducing-gas
source.
Sorbent Circulation Element
The capital investment for dolomite circulation at 5 and 2
atmosphere regenerator pressures is about four times the investment at 10
atmospheres. This is shown in Figure L-13 with capital investment plotted
against the dolomite sulfur utilization change X. - Xg. Curves for
L-31
-------
Curve 650546-B
20
18
16
o
o
4->
112
0>
o
'•is 10
o
.« 8
o
o
5 6
o
635 MW Plant Capacity
5 and 2 atm Regeneration
10 atm Regeneration
0.06= Process Sulfur Load
-\
0.1 0.2 0.3 0.4 0.5 0.6 0.7
Change in Dolomite Sulfur Utilization Across Regenerator
Figure L-13.- Sorbent Circulation Element Caoital Investment
0.8
L-32
-------
process sulfur loads- of 0.06, 0.03, and 0.01 are shown with regenera-
tor pressures of 10 atmospheres, and 5 and 2 atmospheres.
Capital investment for dolomite circulation has been- assumed
identical for the 2 and 5 atmosphere pressure cases because essentially
the same equipment is- required with both pressures.. While dolomite
circulation is- relatively cheap at a regenerator pressure of 10 atmospheres,
the. capital investment ±s- very significant at 2 and 5 atmospheres. Part
of this additional cost is due to the need to cool the dolomite down to
800°F so that existing values may be used. The introduction of high-
temperature values would1 reduce the investment.
Total One-Step. Regeneration Process Capital Investment
Samples of total process capital investments for a process
sulfur load of 0.03 at pressures of. 2, 5, and: 10 atmospheres are shown
in Table L-7. S02 mole fractions which might be expected at each pressure
have beea indicated in the table along with the corresponding, capital
investment ranges. For the most part the total capital investment is
less.than that expected for a stack gas cleaning, device on a conven-
tional coal-fired power plant (30 to 50 $/kw). Depending on the actual
regenerator performance, operation at regenerator pressures lower than
the boiler pressure may be justified, even with the large cost penalty
paid for dolomite circulation at low pressures. -The capital investment
for other process sulfur loads can be estimated by extrapolation from
the results for sulfur loads of 0.06 and 0.01 which have been generated.
ENERGY COSTS
Energy costs for a pressurized fluid bed combustion power
plant utilizing the one-step regeneration process have been computed
on the following basis:
e The cost of a 635 MW pressurized fluid bed combustion
power plant, excluding dolomite regeneration, has been
estimated as 228.77 $/kw (Includes interest during
construction).
L-33
-------
TABLE L-7
TOTAL INSTALLED REGENERATION PROCESS COST, $/kw3
Regenerator
Pressure
(atm)
10,
5
2
Conditions:
S02 Mole
Fraction
0.01
0.03
0.03
0.06
0.09
0.14
Coal
Gasification
27.5-56.5
13.5-29.0
20.0-42.5
13.0-28.5
13.0-29.0
10.5-23.5
Cost Range,
Oil
Gasification
22.0-31.0
11.5-17.0
18.0-29.0
12.5-23.0
12.5-24.5
10.0-20.5
$/kw
CH4
Oxidation
18.5-28.0
9.0-14.5
16.0-26.5
11.5-20.0
11.0-21.0
9.0-18.0
CH4
Reforming
15.5-27.5
9.0-15.0
16.0-26.5
12.0-20.5
12.0-23.0
10.0-19.5
635 MW fluid bed coal combustion plant
W [e-m X ] = 0.03
s s
L-34
-------
Curvg 650519-B
16
!"
i
10
Gas Production -Methane Partial
Oxidation or Reforming
Regenerator Pressure =10 atm
Energy Cost Without Regeneration Process Cost
- Base Line -
0.01 0.02 0.03 0.04 0.05 0.06
Mole Fraction So2 In Regenerator Product
Figure L-14. Energy Cost-Methane at 10 Atraosoheres
0.06
14
a
a.
Gas Production-Methane Partial
Oxidation or Reforming
Regenerator Pressure = 5 atm
Base Line
0.02 0.04 0.06 0.08 a 10 a 12
Mole Fraction SOj in Regenerator Product
Figure L-15 Energy Cost-Methane at 5 Atmospheres
a 14
Cur.i 6505^5-1
18
16
|
i
§14
te
•|
Gas Production-Methane Partial
Oxidation or Reforming
Regenerator Pressure = 2 aim
Base Line
0.03
0.01
0.02 0.04 0.06 0.03 0.10 0.12 0.14 0.16 0.18
Mole Fraction S02 In Regenerator Product
Figure L-lfr Energy Cost-Methane at 2 atmospheres
L-35
-------
Curve 6505U-B
Curve 6S05M-B
18
16
2
i
J
Q.
10
Gas Production - Oil Gasification
Regenerator Pressure = lOatm
Energy Cost without Regeneration Process Cost
- Base Line -
18
0 0.01 0.02 0.03 0.04 0.05 0.06 0.07 0.08
Mole Fraction SOg in Regenerator Product
Figure L-17 Energy Cost-Oil at 10 Atmospheres
i
i
16
14
12
10
Base Line
Gas Production-Oil Gasification
Regenerator Pressure = 5 aim
0 0.02 0.04 0.06 0.08 0.10 0.12 0.14 0.16
Mole Fraction S02 in Regenerator Product
Figure L-18 Energy Cost - Oil at 5 Atmospheres
16
14
s
•£
I
10
I I I I
Gas Production-Oil Gasification
Regenerator Pressure = 2 atm
Base Line
Process
.SulliK Load,,
0 0.02 0.04 0.06 0.08 0.10 0.12 0.14 0.16 0.18
Mole Fraction S02 in Regenerator Product
Figure 1-19 Energy Cost-Oil at 2 Almosoheres
L-36
-------
• Interest during construction is charged on the one-step
regeneration process investment at 7-1/2% for 3-1/2 years
construction period.
• Capital charges at 15%/yr
• O&M charges at 2%/yr
e Dolomite at $2/ton; make-up rate of 1 mole Ca/mole S fed
e 70% capacity factor
• No credit for recovered sulfur
• Plant heat rate of 8967 Btu/kWh (excluding inefficiency
of regeneration process)
• Fuel costs
- Coal at 45C/MM Btu
- Oil at 45C/MM Btu
- Methane at 80C/MM Btu
• Operating .water at 8c/H gal
Total energy costs for a 635 MW pressurized fluid bed combustion
power plant utilizing one-step dolomite regeneration is shown in the
following set of graphs, Figures L-14 through L-22. Each group refers
to a specific gas producer fuel and a specific regenerator pressure.
Energy cost ranges are shown for process sulfur loads of 0.01, 0.03,
and 0.06 as a function of the regenerator product SO- mole fraction.
General conclusions drawn from the set of graphs are that the
use of methane by either partial oxidation or reforming is the most
costly process -option due to the high cost associated with methane, and
coal gasification is the second most costly due to the high capital
cost associated with coal gasification. Oil gasification is the least
costly of the process options due to the low capital cost and low
fuel cost. These conclusions are largely dependent on the cost basis
used to compute energy costs,and other bases may change the relative
significance of capital costs and fuel costs. Care should also be
exercised in comparing the plant energy cost of the various fuel
options considered because the regeneration performance with the four
L-37
-------
Curv. 6505)7-8
18
Gas Production-Coal Gasification
Regenerator Pressure = 10 atm
16
Process
Sulfur^ Load,
0.06
Energy Cost without Regeneration Process Cost
-Base Line-
10
0.01 0.02 0.03 0.04 0.05 0.06 0.07
Mole Fraction S02 in Regenerator Product
Figure 1-20 Energy Cost-Coal at 10 Atmosoheres
18
E
I
SM4
12
10
Curv> 6505)6-8
Gas Production-Coal Gasification
Regenerator Pressure = 5 atm
Base line
Process
Sulfur Load
0.02 0.04 0.06 0.08 0.10 0.12
Mole Fraction SO. in Regenerator Product
Figure 1-21 Energy Cost-Coal at 5 Atmosoheres
0.14
Curvi 6S05J1-B
16
S
1
I"
"c
&
fl_
12
Gas Production-Coal Gasification
Regenerator Pressure = 2 aim
Base Line
0.02 0.04 0.06 0.08 0.10 0.12 0.14 0.16 a 18
Mole Fraction S02 In Regenerator Product
Figure L-22 Energy Cost - Coal at 2 Atmospheres
L-38
-------
Curve 65065
-------
fuel options Is not known. Thus, the S02 mole fraction possible at a
given regenerator pressure may differ among the four fuel options, and
it may be possible to operate one fuel option at the lover end of the
Indicated cost range at a given SO. mole fraction while another fuel
option may require operating at the high end of the cost range.
The energy cost has been computed assuming that 30 to 50% of the
fuel energy input into the gas producer element can be utilized to
generate power. This may be a significant consideration at low 502 m°le
fractions. For example, if none of the fuel energy input to the regen-
eration process can be utilized to generate power then the plant energy
cost will increase about 5% at an SOg mole fraction of 0.01. At high
SO. mole fractions this consideration becomes unimportant. For
example, at a mole fraction of S02 of 0.05 the increase in plant energy
cost is less than 1% if no useful energy is provided by the regeneration
process fuel input. If the one-step regeneration process can be operated at-
SO. mole fractions of about 0.10 or higher, then the recovery is not
required for economic feasibility. If S02 mole fractions no higher than
3 to 4% can be realized in practice, then every effort must be made to
recover useful energy from the regeneration process since the penalty
for not recovering the energy is significant.
The energy cost graphs indicate that, at a fixed SO. mole
fraction, decreasing the pressure increases the plant energy cost. The
SO. mole fraction may increase, though, as the pressure is decreased
to an extent large enough to make the plant energy cost lower at lower
pressures. Experimental evidence with each of the fuel options considered
is required in order to compute the advantage or disadvantage of going
to lower regeneration pressures.
The effect of the process sulfur load is also indicated in the
graphs. This parameter is extremely significant at low SO. mole fractions.
The energy cost of a pressurized fluid bed combustion power plant without
the cost of the dolomite regeneration process is included in the graphs
to provide a base line (11.3 mills/kWh).
L-40
-------
CONCLUSIONS OF GENERAL STUDY
Important Assumptions
• No change in boiler performance — no boiler modifications
needed
• Constant cost (base case) of make-up dolomite feeding and
handling
• Constant dolomite make-up rate of 1 mole Ca/mole S under
all conditions
• Low dolomite purchase and disposal cost of $2/ton
• No sulfur in reducing-gas producer fuel; coal and oil at
45C/MM Btu
• Constant Glaus plant sulfur removal efficiency of 95%
under all conditions
• Regeneration/sulfur recovery system thermal efficiency of
30 to 50%
• Temperature control requirements assumed small and not
costly.
Cost Comparison with Alternative
Alternative
• The alternative would be a conventional coal-fired power
plant utilizing limestone wet scrubbing for pollution
control — 635 MW with same process sulfur load.
Basis
• Highest cost in cost range selected for one-step regeneration
process
• Capital costs for- limestone scrubbing system taken from
"Some General Economic Considerations of Flue Gas Scrubbing
for Utilities," J. K. Burchard, G. T. Rochelle, W. R. Schofield,
J. 0. Smith, Control Systems Division, E.P.A., October, 1972.
L-41
-------
• Capital cost of conventional coal-fired plant and energy cost
basis taken from "Evaluation of the Fluidized Bed Combustion
Process"
TABLE L-8
CONVENTIONAL PLANT ENERGY COST
WITH LIMESTONE WET SCRUBBING
Process Sulfur
Load
0.01
0.03
0.06
TABLE L-9
Energy Cost,
(mills/kWh)
13.37
13.68
14.36
ENERGY COST COMPARISON3
Process
Sulfur
Load
0.03
0.06
0.01
Regenerator
Pressure
Atm
10
5
2
10
5
2
10
5
2
Reducing-Gas Producer Fuel
Coal
1.7
2.5
3.0
2.65
3.6
4.2
f Oil | Methane
0.95 1.85
1.25 2.5
2.3 2.8
1.75 2.95
2.35 3.65
3.5 4.3
< 1.0
S0_ break-even mole percent to meet energy cost of a conventional
coal-fired power plant with limestone wet scrubbing.
L-42
-------
Cost Conclusions
• At a fixed SO. mole fraction, decreasing the pressure
increases the capital and energy costs slightly.
• The energy and capital costs increase drastically as
the SO. mole fraction decreases below about 5%.
• As the regenerator pressure is decreased, the equilibrium
S0_ concentration increases much faster than does the SO.
concentration required to meet a fixed energy cost.
• The -energy and capital costs increase faster with increased
process sulfur load than does the cost of a conventional
coal-fired power plant utilizing limestone wet scrubbing.
• vThe cost feasibility of one-step regeneration at all
pressures, all gas producer fuels, and all process sulfur
loads considered is not limited by equilibrium, though
the ability to meet the cost of a conventional plant
utilizing limestone wet scrubbing becomes greater at
lower regenerator pressures and lower process sulfur loads.
• Low-pressure regeneration should be considered and has many
potential advantages over high-pressure regeneration:
- SO 2 mole fractions not much greater than 0.01 have been
obtained in batch tests at 10 atmospheres pressure while
mole fractions of 0.08 have been obtained at atmospheric
pressures. One percent SO. is too low at a process sulfur
load of 0.03 while 8% is a reasonable SO. level.
- Roughly, equipment sizes are proportional to (Ycr. P)
S02 _!
while the fuel requirements are proportional to (Yco ) .
so2
- At low S0? mole fractions (< 5%), the thermal efficiency
of the regeneration process must be maximized in order to
maintain cost feasibility, while at high SCL mole fractions
(> 10%), the recovery of useful energy from the regeneration
process is not required. Maximum thermal efficiency is
probably less than 507
L-43
-------
• The energy costs are optimistic for coal and oil gasification
at low SO2 mole fractions (< 3%) because the presence of
sulfur in these fuels is neglected. With high-sulfur fuels
the fuel requirements, Claus plant sulfur load, and energy
cost will increase drastically as the SO- mole fraction drops
below 3%. The alternatives of desulfurization of the reducing
gas or the use of low-sulfur fuels will also increase the
energy cost.
• Temperature control requirements (energy inputs or energy
withdrawals) are reduced at higher SO. mole fractions.
Design Conclusions
• The incentive to develop an integrated reducing-gas producer-
dolomite regeneration reactor is small unless performance
is improved by integration.
- Capital saving slight
- Ho savings in energy cost unless the SO. fraction is increased
- Possible introduction of complications
• Regenerator vessel modules rather than a single vessel
should be considered to aid turndown.
• Recycle of tail gas to the fluid bed boiler is attractive,
especially at high regenerator pressures. Further study is
required.
• Dolomite circulation at low regenerator pressures is about
four times the capital investment for dolomite circulation
at 10 atmospheres. Further design analysis is required.
L-44
-------
Problem Areas and Major Unknowns
Reducing-Gas Producer Element
• Composition of reducing gas
- Tars
- Trace metals, etc.
- Sulfur (or desulfurization of reducing gas)
- Particulate material
• Operability of gas producer (especially coal gasifier)
- Control
- Flexibility
-.Reliability.
Regenerator Element
• Effect of reducing-gas contaminants on regenerated stone
• Regenerator product gas composition and contaminants
S02, H2S, 02, H20, CO, H2, tars, metals, etc.
• Ash agglomeration
• Regenerated stone composition
• Attrition and elutriation
• Effect of pressure cycling on stone
• Temperature control.
Sulfur Recovery Element
• Effect of regenerator gas product composition (including
particulates) on sulfur removal efficiency of Claus plant,
catalyst life, fuel requirements, etc.
• Composition of Claus plant tail gas.
Tail Gas Handling Element
• Incineration option
- Sulfur and other pollutant emitted
- Fuel required and recoverable energy
L-45
-------
• Boiler recycle option
- Effect on boiler performance
- Modifications to boiler required
- Contaminants in recycle gas and effect on gas turbine.
Sorbent Circulation Element
• Operability and control
• Melt formation
• Stone cooling method
• Attrition.
Experimental Data Required
• Conditions to maximize S0« mole fraction
• Stone composition and activity; attrition
• Limestone or dolomite sorbent characteristics
• Maximum stone utilization capability
• Effect of pressure, temperature cycling
• Effect of reducing-gas composition (typical compositions
from commercial gas sources)
• Regenerator product gas composition
• Effect of ash in regenerator
• Temperature control data, operability, sensitivity
• Sulfur recovery from product gas-fuel required, poisoning, etc.
• Analysis of Claus plant tail gas composition
• Demonstration of stone circulation between vessels with large
pressure differential
• Demonstration of gas producer operability.
L-46
-------
REFERENCES
1. "Evaluation of the Fluidized Bed Combustion Process," Vols. I-III,
submitted to the Office of Air Programs E.P.A. by Westinghouse
Research Laboratories, Pittsburgh, Pa., November 1971.
2. "A Regenerative Limestone Process for Fluidized Bed Coal Combustion
and Desulfurization," R. C. Hoke, H. Shaw, M. S. Nutkis, L. A. Ruth,
Esso Research and Engineering Company, Monthly reports under
Contracts CPA 70-19 and 68-02-0617.
3. "Claus Technical and Economic Study," R. T. Shpall, Shell Develop-
ment Company, GAP Contract EH SD-71-45 Task Order No. 15.
4. J. Agosta, H. F. Illian, R. M. Lundberg, 0. G. Tranby, Commonwealth
Edison Company, Chicago, Illinois, "Status of Low BTU Gas as a
Strategy for Power Station Emission Control," A.I.Ch.E., 65th
Annual Meeting, November 1972.
5. Information provided in correspondence from Shell, February 28, 1972.
L-47
-------
M
-------
APPENDIX M
LOW-SULFUR COAL FLUID BED COMBUSTION
The economics of fluid bed combustion of low-sulfur coal has
been estimated. Two cases have been considered: the combustion of a
1 wt % sulfur coal which requires some desulfurization in order to meet
S0~ emission standards,and- the combustion of a coal requiring no
desulfurization. Capital investments and energy costs for the pressur-
ized fluid bed combustion power plant have been compared to the costs
of a conventional coal-fired power plant using limestone wet scrubbing
in the two cases.
CASE I: 1 WT % SULFUR COAL REQUIRING DESULFURIZATION
For the pressurized fluid bed combustion plant both once-
through and regenerative operations have been considered. In addition,
desulfurization efficiencies of 50% and 70%, giving SO. emissions
equivalent to 0.5 wt % and 0.3 wt % sulfur coal, respectively, are
included. The cost basis is listed in Table M-l.
TABLE M-l
COST BASIS
• Fluid bed combustion capital costs and energy cost basis are taken
from Reference 1.
• Conventional power plant capital costs are taken from Reference 1
and limestone wet-scrubber costs from Reference 2.
• Capital charges at 15%, fuel at 45 C/MM Btu, dolomite at $10/ton,
70% capacity factor, 635 MW plant.
• Once-through fluid bed combustion costs are based on 50% dolomite
utilization.
• Regenerative fluid bed combustion costs are based on the one-step
dolomite regeneration process at a pres'sure of 10 atmospheres.
M-l
-------
Limestone wet-scrubbing system Includes reheat, bypass, pond.
Plant heat rates (fitu/kWh) are as follows: once-through fluid bed
combustion plant — 9000; regenerative fluid bed combustion plant -
9300; conventional plant — 9400.
Table M-2 lists the results of the cost analysis.
TABLE M-2
COST RESULTS WITH 1 WT % SULFUR COAL
Desulfurization (%)
Plant Investment^ S/kw
Fluid Bed Combustion
Once- through 1 Regenerative
Conventional Plant with
Limestone Wet Scrubbing
50
70
50
70
265.0
265.5
Plant Energy
11.24
11.31
272.5
278.5
Cost, mills /kWh
11.58
11.86
330.5
331.4
13.29
13.36
These results indicate that the economic advantages of
fluidized bed combustion over the conventional plant using limestone
wet-scrubbing are large when a 1 wt % sulfur fuel requiring desulfviriza-
tion is being considered. A reduction of 15% in the energy cost is
projected.
CASE II: LOW-SULFUR COAL REQUIRING NO DESULFURIZATION
With no desulfurization required,the pressurized fluid bed
combustion plant can be operated without dolomite handling, storage,
feeding, and disposal and without dolomite regeneration. The conven-
tional power plant can be operated without limestone wet scrubbing but
must include an electrostatic precipitator for particulate control.
If the basic pressurized fluid bed combustion plant design is assumed
(1) then the figures in Table M-3 result.
M-2
-------
TABLE M-3
COST RESULTS WITH NO DESULFURIZATION
I Capital I Energy
_ J Cost. $/kw _ . I Cost, mill
ills/kWh
Pressurized Fluid 260.5 11.12
Bed Combustion
Conventional 316.3 12.51
Plant
About an 11% reduction in energy cost is estimated by applying pressurized
fluid bed combustion with the basic design condition used when desulfur-
ization was an important consideration. With no desulfurization
required many improvements in the pressurized fluid bed combustion design
conditions could be utilized. Some of these factors are listed in
Table H-4.
While it does not appear that the total plant cost could be
reduced significantly by the factors in Table M-4, the plant operation
with no sulfur removal would be considerably simpler than the operation
with sulfur removal.
M-3
-------
TABLE M-4
DESIGN FACTORS RESULTING FROM NO SULFUR REMOVAL REQUIREMENT
Factor
Effects
Bed Material ( type
of material, shape,
size, or size distri-
bution)
Higher Bed Temperature
(up to the ash
agglomeration tempera-
ture)
Easier Turndown,
Start-up, Control
Operate As an
Agglomerating Bed
2000°F-2100°F
higher or lower fluidization velocity:
may result in small cost savings even at
the optimum velocity (Figure M-l)
higher heat transfer coefficient: less
heat transfer surface could result in a
cost saving of only about $2/kw; would
insure against operating in the region of
low heat transfer where costs rise rapidly
(Figure M-2)
limited attrition and elutriation of bed
material: could result in a reduction in
particle removal equipment of about
$3.50/kw4
limited tube wear
limited solid waste material
higher gas-turbine inlet temperature giving
a high efficiency; increasing the inlet
temperature from 1600°F to 1800°F would
decrease plant heat rate from 9026 to
8820 Btu/kWh
less heat transfer surface in the bed
could produce a soot having a less damaging
nature for the gas turbine
operate at optimum cycle conditions without
sulfur removal restrictions
new design concept
higher gas-turbine inlet temperature
reduced heat transfer surface
reduced particulate removal
M-4
-------
Curve 6*f7l85-A
2.5
S 1-5
X
•w-
in
O
0
0.5
0
Bed Temperature 1750°F
Curve Tube Pitch/Diameter Ratio
Basic Design
1" 0. D. Tubes at
H=4"&V=2"
Basic Design
635 MW
1
Basic Design
318 MW
i
Curve 1
Curve 2
1
2
Cost reduction due to decrease in module"
diameter
Cost escalation due to increase in module
height
0
5 10 15
Fluidizing Velocity in Preevaporator, ft/sec
Figure M-l: Dependence of the Shell Cost on the Fluidizing Velocity
M-5
-------
Curve W7208-B
x 4
I/I
O
o
O)
0>
o
(O
*
to
Curve
1
2
3
Tube Pitch/Diameter Ratio
Basic Design
1-0. D. atH = 4"&V = 2"
2"0. D. atH=8"&V/4"
Bed Temperature 1750°F
Maximum Allowable Bed Depth 20ft
Bed Area
35ft'
10 20 30 43 50 60 70 80 90 100
Bed-Tube Heat Transfer Coefficient, Btu/ft2 hr - °F
110 120
Figure M-2: Effect of Bed-Tube Heat Transfer Coefficient on the Steam Generator
'Cost (not including erection)
M-6
-------
REFERENCES
1. "Evaluation of the Fluidized Bed Combustion Process," Vols. I-III,
submitted to the Office of Air Programs E.P.A. by Westinghouse
Research Laboratories, Pittsburgh, Pa., November 1971.
2. Burchard, J. K., Rochelle, G. T., Schofield, W. R., Smith, J. 0.,
"Some General Economic Considerations of Flue Gas Scrubbing for
Utilities," Control Systems Division, E.P.A., October 1972.
M-7
-------
N
-------
APPENDIX N
LIMESTONE WET SCRUBBING COSTS
The cost of power generation with a conventional coal-fired
power plant using limestone wet scrubbing has been estimated for
purposes of comparison with the pressurized fluid bed boiler power
plant. The following basis has been used:
• Wet-scrubber costs include
- By-pass piping
- Disposal pond
- No venturi scrubber
• New plant installation, 635 MW
• Coal-fired power plant cost without scrubber
= 306.3 S/kw1
• Other factors are the same as those used to compute the
capital and energy costs of the pressurized fluid bed power >
plant.
The limestone wet-scrubber capital investment is shown in
2
FigureN-l as a function of the weight fraction sulfur in the coal.
Table N-2 gives the energy cost breakdown for the conventional plant
for a limestone purchase plus disposal cost of $2/ton. The total energy
cost is shown in Figure N-l. The energy cost may be expected to vary
by as much as 0.5 mills/kWh from one plant to another, but the cost
in Figure N-l is a reasonable average.
N-l
-------
Curve 652794-A
15
814
OJ
c
c
ro
13
Conditions - 635MW Capacity
Sorbentat$2/Ton
Stoichiometric Make-up
95% Sulfur Removal
I
0 1234567
Wt% Sulfur in Coal
Figure N-l: Energy Cost of a Conventional Coal-Fired Power Plant
with Limestone Wet Scrubbing
N-2
-------
TABLE N-l
CONVENTIONAL PLANT CAPITAL INVESTMENT
Weight Percent
Sulfur in Coala
Capital Investment, $/kw
Scrubber System
Total
Power Plant
6.0
3.0
1.0
49.6
35.0
26.8
355.9
341.3
333.1
95% sulfur removal
TABLE N-2
CONVENTIONAL PLANT ENERGY COST
Fixed Charges
Fuel (45C/MM Btu)
Limestone ($2/ton)
O&M
TOTAL
Energy Cost, mills /kWh
Weight % Sulfur in Coal
.6.0 | 3.0
8.95 8.35
4.11 4.11
0.15 0.075
1.17 1.06
14.38 13.60
1.0
8.15
4.11
0.025
0.99
13.28
N-3
-------
REFERENCES
1. "Evaluation of the Fluidized Bed Combustion Process," Vols. I-III,
submitted to the Office of Air Programs E.P.A. by Westinghouse
Research Laboratories, Pittsburgh, Pa., November 1971.
2. Burchard, J. K., Rochelle, G. T., Schofield, W. R., Smith, J. 0.
"Some General Economic Considerations of Flue Gas Scrubbing for
Utilities," Control Systems Division, E.P.A., October 1972.
N-4
-------
0
-------
APPENDIX 0
PART LOAD OPERATION
The ability to change the electrical output of a power system
is necessary in order to meet the varying electrical demand. An analysis
of the pressurized fluid bed boiler power plant concept has been made to
determine its ability to follow load over the desired range. This
analysis considers alternative turndown techniques, process limitations,
and general performance projections.
To be competitive with conventional steam power plants, the
high-pressure fluidized bed boiler power system should have the
capability of at least a 75% turndown and a response rate of 5% per
minute.
The temperature-enthalpy diagram for the pressurized boiler
power system (Figure 0-1) shows the energy distribution among the
boiler functions, gas-turbine expander, and the heat recovery equipment
at a design point of 1750°F bed temperature and 17% excess air. The
power output from the system is a function of the following four process
parameters:
• Airflow rate
• Fuel/air ratio
• Bed temperature
• Bed depth.
Four modes of load control for the combined cycle fluidized
bed boiler have been considered and are summarized in Table 0-1. In
order to assess these alternative operating modes quantitatively, the
limitations imposed on the control variables by the process must be
considered. These limitations are summarized in Table 0-2.
0-1
-------
TABLE 0-1
ALTERNATIVE TURNDOWN PROCEDURES
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
with reduced airflow.
• Steam power decreases because of lower airflow and
pressure level,
Mode 2
• Constant airflow 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 because of reduced heat transfer
surface; gas-turbine power would be nearly constant.
0-2
-------
TABLE 0-1 (Continued)
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 with reduced temperature.
• Steam power decreases because of lower bed temperature.
• Gas-turbine power decreases because of lower gas temperature
and pressure level.
Mode 4
• Constant bed depth and temperature
• Variable airflow and fuel/air ratio
• Gas-turbine inlet pressure decreases to satisfy constant
turbine flow number with reduced airflow.
• Steam power constant; gas-turbine power decreases because
of reduced flow and pressure.
0-3
-------
Curve
4000
o
i_
CD
a.
i 2COO
1000
Feedwater
Heater
Stack
Losses
Balance of Boiler
System Losses
Piping
Losses
Energy to
Gas Turbine
Preevaporator
_u_
_L
J I
Evaporator
i '
-n _
0 1
3
5 6 7 8 9 10 11
Energy from MAP Fuel -1000 Btu/lb
12 13 14
Adiabatic Flame
Temperature
1750 (FBCD)
Tube Wall Temp.
Steam Temp.
Figure 0-1: Temperature-Energy Diagram for Pressurized Utility Boiler
-------
TABLE 0-2
PROCESS VARIABLE LIMITATIONS
Airflow Rate
Fluidized bed
Combustion
Gas-turbine
cannot operate at gas velocity
below minimum fluidization or
greater than terminal velocity
of majority of bed material —
ratio ^ 10 to 1
airflow such that excess air
greater than ^ 5%
variable inlet guide vanes
permit ^ 33% reduction in gas-
turbine airflow
Fuel/Air Ratio
Combustion
fuel/air ratio maintained such
that excess air >_ 5%, ^ 4 ft; upper limit not
clear; additional solids handling
and storage system required
0-5
-------
Within these limitations part load plant performance for two
combinations of these control modes has been estimated — modes 1 and
2 in sequence and modes 3 and 4 in sequence. In both instances, the
design point conditions are 1750°F and 10% excess air. These performance
estimates are shown in Figure 0-2 along with the part load performance
of a typical coal-fired plant. The combination of modes 1 and 2 in
sequence has a part load performance which is significantly better than
that of the conventional plant. The part load performance for operating
modes 3 and 4 in sequence is also superior to that of the conventional
plant over most of the range, but it is not as good as for operating
modes 1 and 2 in sequence.
Any mode with constant bed temperature requires a variable bed
depth to control steam power. As shown in Table 0-2, the minimum
fluidized bed depth is about 4 feet, so design point bed depths of 12 to
14 feet would give a good range. However, in practice heat transfer to
the steam cycle is not proportionate to bed depth. What really controls the
quantity of heat being transferred is the average heat transfer coefficient
of the total heat transfer surface, and the heat transfer coefficient
for the surface in the space immediately above the bed is nearly as high
as it is in the bed. Because of this, and since varying the bed depth
requires expensive materials handling and storage equipment, operating
modes which use constant bed temperature were eliminated from consideration
for the pressurized fluidized bed boiler system.
The two design point parameters which have the most effect on
plant turndown capabilities are bed temperature and percent excess air.
As shown in Figure 0-3 for operating mode 3, there is a nearly linear
relationship between plant power output and bed temperature, and the
slope of the operating line is independent of the percent excess air.
Operating lines for a range of design point temperatures were constructed
parallel to the 1750°F line. This indicates that the turndown capability
of the system under operating mode 3 is proportionate to the difference
between the design point bed temperature and the minimum bed temperature.
0-6
-------
Curve 65**051 -A
12,000
11,000
10,000
9900
9800
9700
9600
9500
9400
OQ
ja>
!!f 9300
9200
9100
9000
8900
Back Pressure = 3 in Hg
_Auxiliary Power = 3ft\
Overall Comp. PR = 10
\
Nonintercooled
Hammond #4 2400A
- 1000/1000
20
40
60
80
100
Figure 0*2: Part Load Characteristics
0-7
-------
Curve 655509-A
Operating Mode 3
1400 1500 1600 1700 1800
TBed" °F
Figure 0-3: Effect of Desiqn Point Temperature and
Percent Excess Air on Plant Turndown
Capability
0-8
-------
The turndown capability of the high-pressure I'luidized bod
boiler system under operating mode 4 is a function of the percent
excess air at design point. This is because the fraction of plant
power from the gas turbine increases as the percent excess air increases,
and the gas-turbine power is proportionate to airflow. At 10% excess
air, the gas-turbine power is 17.5% of the total plant power, and a one-third
reduction in airflow reduces the plant output by about 6%. At 90% '
excess air, the gas-turbine power is 29% of the plant output and a
similar reduction in airflow reduces the plant output by 9-1/2%. When
operating mode 4 is used sequentially with mode 3, the reduction attainable
in the second step is a percentage of the power .output at the end of the
first step, not of the design point power.
As indicated in Table 0-2, the upper limit on percent excess
air is about 300%. This is the amount of excess air required to give
an adiabatic flame temperature of 1750°F when burning bituminous coal
and air at a temperature of 650°F.
With a design point bed temperature of 1750°F and a minimum
bed temperature of 1300°F, the mode 3 turndown capability of the high-
pressure 'fluid bed system is about 50%. With mode 4 following mode 3,
the attainable turndown is as follows:
% Excess Air Turndown %
10 53
100 55
Adiabatic case 67
This shows that a single high-pressure fluid bed module will not meet
the 75% turndown objective. Therefore, two boiler modules will have
to be used. If the design point bed temperature has to be decreased
significantly and/or the minimum bed temperature has to be increased
significantly, more than two modules may be necessary to obtain 75%
turndown capability.
In the preliminary design of a high-pressure fluidized bed
boiler carried out in Phase I of this program, two gas-turbine
0-9
-------
modules are used, with two boiler modules per gas turbine. The gas-
turbine modules can be operated with only one boiler module with a bed
temperature of 1400°F and above. This system has an overall turndown
capability of about 82%. With the design point value of 10% excess air,
minimum bed temperatures* of about 1350°F are necessary to avoid dis-
continuity in the load curve. With a minimum bed temperature of 1400°F,
an excess air level of about 170% would be necessary to avoid discon-
tinuity in the load curve.
The proposed load control system which varies bed temperature
and airflow either simultaneously or sequentially is expected to
achieve the desired specifications and have the following characteristics;
• Bed depth is maintained constant eliminating the need for
rapid transport of solids into and out of the bed.
• Gas velocity through the fluidized bed is primarily a
function of the square root of the absolute bed temperature
even when the gas-turbine airflow is varied. This means
that the superficial bed velocity will vary only about 15%
in going from a design point bed temperature of 1750°F to
a minimum bed temperature of 1300°F.
• Excess air percentage increases during turndown so that
carbon carry-over will be minimized.
• Ability to change load rapidly by changing the fuel feed
rate and airflow rate.
• Particulate removal system volumetric flow variations are
the same as for the fluidized bed.
FT
With two boiler modules operating per gas-turbine module,
0-10
-------
REFERENCES
1. "Evaluation of the Fluidized Bed Combustion Process," Vols. I-III,
submitted to the Office of Air Programs, Environmental Protection
Agency, by Westinghouse Research Laboratories, Pittsburgh,
Pennsylvania, November 1971.
0-11
-------
-------
APPENDIX P
RECIRCULATING BED BOILER STUDIES
INTRODUCTION
The concept of a deep recirculating bed boiler was presented
in Volume I, Section 6.2. In order to study the feasibility of using
deep recirculating fluidized beds for coal combustion boiler power plants,
a two-dimensional cold model was constructed from transparent acrylic
sheets, to investigate solid circulation, operation characteristics, and
design criteria.
The recirculating fluid bed concept is illustrated in Figure
P-l. Gas is fed to the base of an open draft tube section. The superficial
velocity of the gas flowing up the riser may be 10 to 60 ft/sec. The
solids are picked up pneumatically in the draft tube. The effective
overall density of solids and gases is less in the draft tube section
than in the downcomer, which creates a solid circulation pattern upward
through the draft tube and downward in the downcomer. Solids and gases
from the draft tube pass into a fluid bed of the expanded cross-section.
Solids from the fluid bed above the draft tube flow into the downcomers
and enter the base of the draft tube. Gas is introduced at the base of
the downcomer at a rate necessary to permit the downward flow of solids.
The unit was operated with and without simulated imbedded heat transfer
surface in the downcomers.
Several recirculating fluid bed systems have been developed
123
and operated. The British Gas Council ' ' developed deep recirculating
fluidized beds for oil and coal gasification. In the gasifier this was
used to smooth out the local high temperatures at the point of oxygen
entry and, in the hydrogenator, to allow for the quick removal of particles
wetted by the oil in the vicinity of the oil inlet, thus avoiding the
P-l
-------
Dwg. 6.197A35
Downcomer
Draft
Tube
Air
Plenum
^ —
x
8.5"
— A-*)BH-C-»)B|— A —
i
~~JP
E
•^^"^
\
\
. —
I
/
X
X
/
/
y
/
x
X
X
x
x
y
X
X
X
X
X
^
x
(
ffn'
,
x
I T°
. — ^
Airt
®
3
•^
2
•^5\
®
^X
n
i
-*•
a
t
s
\
x
x
\
\
X
X
X
X
X
\
X
X
\
X
\
\
X
\
\
N
N
^
x
\
N
4;
^-^
1
CR)
1 -
Air Jet
Nozzle
Air
Solid Aluminum
Bar
Downcomer
A.B.C.D.E.F, and a
are adjustable
= Pressure Tap
Positions
•Air
Plenum
Figure P-l- Detailed Schematic of the Two-Dimensional Cold Model
P-2
-------
formation of agglomerates as well as closely controlling temperature Tor
process purposes. Experiments were performed in large-scale models up
to 5 feet in diameter and operated at pressures and temperatures up to
70 atmospheres and 7508C (1352°F).
The same concept was also used for vertical pneumatic transportation
A
of sticky and bridging powders. A variation of this concept was utilized
to promote solid mixing, circulation, heat and mass transfer in the bed,
and for pretreatment of caking coal. Taskaev and Kozhina utilized
a recirculating bed for the low-temperature carbonization of coals.
Product gas preheated to a temperature of 600° to 700°C (1112° to 1292°F)
was recirculated to entrain the coal particles of sizes up to 1/4 in
through a vertical draft tube 1 in in diameter and 20 in in length into
a concentric tube 6 in in diameter and 40 in in length. The coal or
char particles fell through this annular space (a downcomer) in a dense
descending bed and were then reentrained by the gas and recirculated
until carbonization was complete.
A recirculating fluid bed reactor is utilized in the multi-
stage fluidized bed coal gasification process for production of low-Btu
o
gas for power generation being developed by Westinghouse. A recirculating
bed operating near 1600°F is used in this concept to devolatilize the
coal and possibly simultaneously to desulfurize the fuel gas.
This experimental and theoretical study establishes both the
potential for utilizing a recirculating bed as a boiler and the initial
design considerations for developing the concept. The recirculating bed
boiler offers economical and environmental advantages as a second
generation pressurized fluid bed boiler.
P-3
-------
TWO-DIMENSIONAL MODEL DESIGN
A cold model was constructed from transparent acrylic shouts
to facilitate visual observation of the bed operation. The detailed'
dimensions are shown in Figure P-l. Only the bed cross-section, 8.5 in x
1.5 in is fixed; the other dimensions (A, B, C, D, E, F, and a) in
Figure P-l can be adjusted to study the effect of these design variables.
For the present investigation,, the following dimensions are used
A = 2.5 in; B = 1.188 in; C = 1.125 in; D = 0 in,
1 in, and 2 in; E = 18 in; F = 3 in; a = 45°.
Pressure taps are also provided for measuring differential
pressure drops between points 1, 2, 3, and 4. The distance between
pressure taps 1 and 4 (or 2 and 3) is 16 in. Pressure drops between
1 and 2, 1 and 4,and 3 and 4 were measured using U-tube manometers;''the
pressure drop between 2 and 3 was then obtained by difference. Water
was used as manometer fluid in 1 and 2 and 3 and 4 measurements; mercury
and alkazine with a specific gravity of 1.75 was used for 1 and 4
measurements'. The reactor pressure was measured with a separate U-tube
manometer (not shown in Figure P-l).
Air is supplied by four separate streams with their individual
controls and flow meters: one for each of the two air plenums supplying
air to fluidize the downcomer sections; and two for the air jet. The
air jet nozzle is an open 1/2-in.copper tube with a stainless steel
wire-mesh cap at the end to prevent weeping. The distributor plate
consists of 1/16-in.holes spaced 11/32 in apart in a staggered arrangement.
(The distributor plate is also covered with wire-mesh). The draft tube
and the downcomers are separated by the two solid aluminum bars as
shown in Figure P-l.
Some experiments were performed in a slightly modified two-
dimensional bed as shown in Figure P-2. Two separate fluidized beds (or
"boots") are created by inserting two aluminum bars of 3/4-in. thickness
underneath the draft tube. The objective of this modification is to
study the effect of the draft tube inlet design on the solid circulation
P-4
-------
Dwg. 6197A36
Downcomer
Draft Tube
Air Plenum
Downcomer
Solid Aluminum
Bar
Air Plenum
Jet Nozzle '
Figure P-2-Two-Dimensional Cold Model with Modified Draft Tube Inlet
P-5
-------
rate and the air distribution between the downcomers and the draft
tube. All the dimensions in the modified model are similar to the original
model (Figure P-l) except now D' = 5 in.,F' = 7 in. , and G = 3.0 in.
To simulate imbedded heat transfer surfaces in the downcomer
sections, serpentine solid copper rods of the dimensions shown in
Figure P-3 were immersed in the downcomers. The rods are supported by
pushing the ends of the rods into closely sized holes in the aluminum
bars. Two serpentine rods are immersed in each downcomer 4.5 in. below
the top of the draft tube with equal distances between the rods and
between the rods and the walls. The effect of the imbedded serpentine
rods on the solid circulation rate was studied by a similar method.
EXPERIMENTAL PROCEDURE
Ottawa sand was used as the bed material with the particle
size distribution as shown in Table P-l. The mean particle size was
calculated to be 606 y using the following equation:
p «t
where
X. = weight fraction of particles of size i and
d . = particle diameter of size i.
During operation, air is injected into the draft tube through
the air jet nozzle to provide an air velocity up to * 43 ft/sec in the
draft tube while the downcomer sections are minimally fluidized. This
creates a solid circulation pattern upward through the draft tube and
downward in the downcomers. The minimum fluid izing condition in the
downcomers is assured first by turning up the air supply to the down-
comers until air bubbles appear in the downcomers and then turning
down the air supply until the air bubbles just disappear.
P-6
-------
Curve 653461-A
1 "
•j copper rod
4
I11
r
9"
Figure P-3-Dimensions of the Serpentine Copper Rods
P-7
-------
TABLE P-l
PARTICLE SIZE DISTRIBUTION OF THE OTTAWA SAND
USED AS BED MATERIAL
U.S. Standard
Mesh
20
30
40
50
60
80
1
Opening in
U
841
595
420
297
250
177
Opening in
Inches
0.0331
0.0234
0.0165
0.0117
0.0098
0.0070
Wt %
0.28
74.82
14.10
8.95
1.10
0.75
The solid downward velocity in the downcomer was estimated
by following a tracer particle for 16 inches with a stopwatch. The
tracer particles are silica gel dyed with red pigment and of similar
size. For every operating condition, at least ten particles in each
downcomer were traced and the arithmetic average velocity was taken
to be the solid particle velocity in the downcomer. The solid recircu-
lation rate was then calculated by assuming plug flow in the downcomer.
More than 100 runs have been performed in the cold model.
The variables changed are static bed height, airflow rate, and the jet
nozzle position. With model configuration as shown in Figure P-l (Case I),
the experiments were divided into three series:
Test Series 1:
D = 0 in.;static bed height at 20 in..,25 in.,
30 in.,and 37 in.
Test Series 2: D = 1 in.;static bed height at 21 in.,24 in.,
27 in.,and 35 in.
Test Series 3: D = 2 in.;static bed height at 22 in.,25 in.,
29 in.,33 in.,and 37 in.
P-8
-------
With serpentine rods in the downcomers (Case II), the experiments
were also divided into three series. The variables changed are static
bed height, airflow rate, and the jet nozzle position:
Test Series 1: D = 0 in.;static bed height at 21 in,,25 in,,
29 in,,33 in.,and 37 in.
Test Series 2: D = 1 in.jstatic bed height at 21 in.,25 in.,
29 in,,34 in.,and 38 in.
Test Series 3: D = 2 in.jstatic bed height at 21 in.,24 in.,
28 in. ,33 in.,and 38 in.
With model configuration as shown in Figure P-2 (Case III), the air jet
nozzle position was fixed at the same height of the boots. The variables
changed were static bed height and air velocity.
EXPERIMENTAL RESULTS
The experimental data for Case I, Case II, and Case III are
summmarized in Tables P-2 to P-8. The largest observed solid recirculation
2
rate is 110,000 Ib/hr-ft of the downcomer area, which is comparable to
2 2
the figure 113,000 Ib/hr-ft reported by Horsier and Thompson.
All the.experimental data obtained so far can be approximated by equa-
tions based on the pressure drops between 1-2 (P,-P9 = AP-, 7) as shown
in Figures P-4 and P-5.
(Up,d)L + (Up,d)R = 0.107 (2gHp)2/3 for Cases I and II (2)
(Up,d)T + (Up,d)_ = 0.025 (2gHp) for Case III (3)
L K
62.43 (AP. ,)
"p • p. o~d> (4>
Since the inlet to the draft tube (point 2 in Figure P-l) is at
a lower pressure compared to that at outlets of the downcomers (point 1
in Figure P-l), the air from the air plenums favors the draft tube, and, thus
air by-passing into the draft tube results. The amount of air by-passing
P-9
-------
i
o
1.0
0.9
0.8
0.7
0.6
0.5
0.4
- 0.3
Curve 651038-B
0.2
0.1
Case I
Test Series 1
Static Bed Height
37"
30"
25"
20"
Test Series 2
Test Series 3
Case HI :
With Modified
Draft Tube Inlet
Design
Draft Tube Height
18"
456
2g Hp, ft2/sec2
7 8 9 10
20
Figure P-4-Experimental Data for Case land Case III
-------
Curve 653^64-8
1.0
0.9
0.8
0.7
0.6
0.5
£r 0.4
""rs
CL
2 0.3
Series 1
Series 2
Series 3
Static Bed Height
38"
34"
29"
25"
21"
x Data from experiments with no tubes
in the downcomers
0.2
2/3
0.1
' .1.1.1
456
2gHp, ft2/sec2
8 9 10
20
Figure P-5-Experimental Data for Case II
-------
depends on the pressure drop AP.. „ as shown in Figure P-6. The straight line
can be represented by the following equation:
F = 1.15 AP1-2 , (5)
where F, the air by-passing factor is defined as
_ m Total Air input into the Air^ Plenums _^
Total Air Required to Minimally Fluidize the Downcomers
When the jet from the draft tube exit penetrates the fluid
bed above the draft tube and the downcomers, the fluid bed behaves
essentially like a spouting bed. When the jet does not penetrate, the
fluid bed becomes a slugging bed. Due to this oscillating nature, the
solid downward velocities (obtained from timing with a stopwatch) in
the downcomers,at fixed operating conditions,vary for up to 10% of the
average value. The solid velocities in the two separate downcomers
can be adjusted to as close to each other as possible by varying air
input into their separate air plenums.
The manometer readings are generally stable when the
air jet penetrates the fluid bed. When the fluid bed becomes a
slugging bed, the oscillation frequency of the pressure readings can
be up to 20 cycles per minute. The range of pressure fluctuation is
up to 10% of observed value for AP,, and up to 20% for AP1 _.
After about 50 hours' operation, no visible erosion was
noticed on the draft tube sides of the aluminum bars, though a slight
blur appeared on the plexiglass walls.
P-12
-------
10
9
8
7
6
5
•2 4
o H
CXJ
a. 3
>s
OQ
oo
-------
TABLE P-2
EXPERIMENTAL DATA OF TEST SERIES 1 (CASE I)
Run No.
Static Bed
Height, in.
Superficial Fluid Velocity
in Draft Tube, ft/sec
Solid Circulation
Rate, Ib/sec
Specific Solid Loading
in Draft Tube,
Ib solids/lb fluid
1-1-1
1-1-2
1-1-3
1-2-1
1-2-2
1-2-3
1-3-1
1-3-2
1-3-3
1-4-13
1-4-2*
37
37
37
30
30
30
25
25
25
20
20
42.48
31.53
19.73
41.25
30.65
20.27
40.05
29.72
22.76
29.06
22.31
1.42
1.55
0.95
1.46
1.36
1.12
1.20
1.41
1.15
0.631
0.590
36.1
53.8
53.8
38.4
48.9
61.3
31.9
49.8
55.6
23.0
28.2
Downcomers are not at minimally fluidized condition.
-------
TABLE P-3
EXPERIMENTAL DATA OF TEST SERIES 2 (CASE I)
Static
Run No. Height
2-1-1 35
2-1-2 35
2-1-3 35
2-1-4 35
2-2-1 30
2-2-2 30
2-2-3 30
2-3-13 27
2-3-2 27
2-4- la 24
2-4- 2a 24
2-5-13 21
2-5-2a 21
Bed Superficial Fluid Velocity
, in. in Draft Tube, ft/sec
32.58
30.69
27.34
22.06
39.96
28.49
21.73
33.17
22.41
30.54
23.05
30.48
24.34
Solid Circulation
Rate, Ib/sec
0.912
1.21
1.27
1.04
1.50
1.49
1.34
1.12
0.521
0.759
0.895
0.830
1.00
Specific Solid Loading
in Draft Tube,
Ib solids/lb fluid
29.6
42.0
49.1
51.0
40.0
55.4
66.2
35.0
24.7
25.6
41.0
28.0
43.2
Downcomers are not at minimally fluidized condition.
-------
TABLE P-4
EXPERIMENTAL DATA OF TEST SERIES 3 (CASE I)
»o
er>
Run No.
Static Bed
Height, in.
Superficial Fluid Velocity
in Draft Tube, ft/sec
Solid Circulation
Rate, Ib/sec
Specific Solid Loading
in Draft Tube,
Ib sollds/lb fluid
3-1-1
3-1-2
3-1-3
3-2-1
3-2-2
3-2-3
3-2-4
3-3-1
3-3-2
3-3-3
3-3-4
3-4-1
3-4-2
3-4-3
3-4-4
3-5-1
3-5-2
3-5-3
3-5-4
37
37
37
33
33
33
33
29
29
29
29
25
25
25
25
22
22
22
22
38.29
32.05
24.95
40.54
37.11
32.74
28.74
39.63
33.95
30.54
27.43
36.82
33.21
28.75
25.42
36.59
32.28
27.39
24.74
1.53
1.39
1.27
1.42
1.47
1.42
1.29
1.33
1.37
1.31
1.22
1.14
1.15
.10
,04
,12
18
14
1.11
40.5
44.8
53.8
36.1
41.8
45.9
48.3
35.3
42.6
45.6
47.5
32.5
36.6
40.7
43.9
32.0
38.7
44.8
48.6
-------
TABLE P-5
EXPERIMENTAL DATA OF TEST SERIES WITH TUBES IN DOWNCOMERS (CASE II)
Run
No.
Static Bed
Height, in.
Superficial Fluid
in Draft Tube,
1-1-1 38 22.3
1-1-2 38 25.0
1-1-3 38 28.3
1-1-4 38 32.2
1-1-5 38 35.5
Velocity
ft/sec
Solid Circulation
Rate, Ib/sec
Specific Solid Loading
in Draft Tube,
Ib Solids/lb Fluid
0.709 35.3
0.711 31.4
0.726 28.0
0.744 25.0
0.718 21.4
1-2-1
1-2-2
1-2-3
1-2-4
l-2-5a
1-2-63
1-3-1
1-3-2
1-3-3.
1-3-4
l-3-5a
a
1-3-6
1-4-1
1-4-2
1-4-33
1-4-4
1-4-53
l-4-6a
1-5-1
1-5-2
1-5-3
l-5-4a
1-5-53
34
34
34
34
34
34
29
29
29
29
29
29
25
25
25
25
25
25
21
21
21
21
21
23.1
25.7
30.1
35.1
34.6
31.3
20.9
26.1
29.9
26.5
34.9
30.1
21.5
25.8
22.1
31.1
30.2
27.7
23.7
28.5
32.8
29.2
26.9
0.705
0.683
0.700
0.716
0.675
0.467
0.620
0.640
0.640
0.380
0.675
0.417
0.625
0.713
0.400
0.653
0.409
0.214
0.742
0.750
0.752
0.400
0.231
34.0
29.3
25.
21.
20.8
15.9
,3
,7
33.1
26.9
23.2
15.5
20.4
14.7
32.4
30.5
20.0
22.6
14.6
8.4
34.8
28.9
24.8
14.9
9.3
Downcomers are not at minimally fluidized condition.
-------
oo
TABLE P-6
EXPERIMENTAL DATA OF TEST SERIES WITH TUBES IN DOWNCOMERS (CASE II)
Run Static Bed Superficial Fluid Velocity
No. Height, in. in Draft Tube, ft/sec
2-1-1 38 28.7
2-1-2 38 31.7
2-1- 3a 38 29.7
2-1-4 38 37.2
2-1-53 38 33.8
2-2-1 33 27.6
2-2-2 33 30.7
2-2-3a 33 27.6
2-2-4 33 35.0
2-2-5a 33 31.6
2-2-6 33 31.3
2-3-1 28 27.9
2-3-2 28 31.0
2-3-3 28 28.6
2-3-4 28 35.0
2-3-5 28 33.2
2-4-1 24 27.1
2-4-2 24 31.0
2-4-3 24 29.4
2-4-4 24 34.7
2-4-5 24 33.1
2-5-1 21 27.0
2-5-2 21 31.2
2-5-3 21 34.1
Solid Circulation
Rate, Ib/sec
0.744
0.744
0.610
0.750
0.618
0.703
0.700
0.398
0.713
0.465
0.316
0.742
0.690
0.553
0.683
0,584
0.731
0.713
0.627
0.707
0.597
0.731
0.759
0.759
Specific Solid Loading
in Draft Tube,
Ib Solids /lb Fluid
28.3
25.4
22.2
21.5
19.5
28.1
24.9
15.8
21.8
15.7
10.7
29.2
24.3
21.0
20.9
18.7
29.7
25.1
23.2
22.5
20.0
29.8
26.6
24.0
aDowncomers are not at minimally fluidized condition.
-------
TABLE P-7
EXPERIMENTAL DATA OF TEST SERIES WITH TUBES IN DOWNCOMERS (CASE II)
Run
No.
Static Bed
Height, in.
Superficial Fluid
in Draft Tube,
3-1-1 37 26.9
3-1-2 37 31.1
3_1_3 37 35.5
3-1-4 37 34.1
Velocity
ft/sec
Solid Circulation
Rate, Ib/sec
Specific Solid Loading
in Draft Tube
Ib Solids /lb Fluid
0.977 40.4
0.977 34.7
0.977 29.9
0.834 26.5
3-2-1
3-2-2
3-2-3a
3-2-4a
3-2-5
3-2-6 a
3-3-1
3-3-2
3-3-3
3-3-4a
3-3-5
3-3-6a
3-4-1
3-4-2
3-4-3
3-4-4
3-4-5
3-4-6a
3-5-1
3-5-2
3-5-3
3-5-4
3-5-5
3-5-6a
33
33
33
33
33
33
29
29
29
29
29
29
25
25
25
25
25
25
21
21
21
21
21
21
26.5
30.9
32.7
31.0
34.5
33.4
26.2
30.8
33.0
30.3
33.5
32.4
25.3
27.2
29.8
31.8
34.0
33.1
25.5
26.6
29,
30,
34,
33.2
0.943
0.953
0.949
0.787
0.971
0.796
0.940
0.951
0.962
0.744
0.895
0.787
0.852
0.858
0.930
0.908
0.892
0.789
0.940
0.934
0.994
0.977
1.053
0.971
39.4
33.9
31.5
27.5
30.3
25.7
39.
33.
31.6
26.7
28.9
26.2
37.3
34.9
34.1
31.0
28.3
25.7
40.9
38.8
37.2
34.
33.
,5
.3
31.6
aDowncomers are not at minimally fluidized condition.
-------
to
O
TABLE P-8
EXPERIMENTAL DATA OF TEST SERIES WITH BOOT (CASE III)
Static Bed
Run No. Height, in.
B-l-1 35
B-l-2 35
B-l-3 35
B-l-4 35
B-2-1 32
B-2-2 32
B-2-3 32
B-2-4 32
B-2-5 32
B-3-1 29
B-3-2 29
B-3-3 29
B-3-4 29
B-3-5 29
B-3-6a 29
B-4-1 26
B-4-2 26
B-4-3 26
B-4-4a 26
B-4-5 26
B-4-63 26
Superficial Fluid Velocity
in Draft Tube, ft/sec
29.16
32.56
33.94
37.06
29.02
32.85
34.93
35.93
34.65
28.90
32.40
31.48
34.18
36.94
33.85
28.24
31.98
34.07
32.13
36.08
33.20
Solid Circulation
Rate, Ib/sec
0.681
0.638
0.586
0.662
0.679
0.662
0.614
0.601
0.467
0.685
0.631
0.566
0.598
0.610
0.413
0.694
0.644
0.640
0.484
0.681
0.311
Specific Solid Loading
in Draft Tube,
Ib solids/ Ib fluid
25.9
21.5
18.8
19.3
26.1
22.2
19.2
18.1
14.6
26.4
21.3
19.7
19.1
17.8
13.2
27.4
22.2
20.6
16.5
20.4
10.1
aDowncomers are not at minimally fluidized condition.
-------
DEVELOPMENT OF MATHEMATICAL MODEL
Pressure Drop in the Downcomer
During operation, the downcomers are minimally fluldized.
The pressure drop through a minimally fluidized bed can be calculated
Q
from the Ergun equation:7
(6)
AP - L
D " 8c
uuf(i
d o
p t
.2 „ 2,. ,
- O P*U* C1 " EJ)
d ..--ft a
22 lt/:t d + e .
e, p s d
> d f '
Yoon and Kunii found that the Ergun equation can
correlate the pressure drop through a moving bed if the fluid velocity,
U,, is substituted with the fluid-solid slip velocity, U - = U- + U ,.
f si t p,d
Thus the downcomer pressure drop can be estimated from the following
equation:
AP, = ±-
150
)j(Uf.d + Up.d)(1-
3 2,. 2 2
d
-------
AP, and AP. are usually negligible compared to other terms.
j can be estimated by the Fanning equation
2fp U 2L
The friction factor, f, is calculated from
PU D _ .,
.,
f = 0.0791 (---)--" for 3 x 103 < Re < 105 , (10)
and
f = 0.0008 + 0.0552 (-y2-)*"0'237 for 105 < Re < 108 . (11)
12
Jones et al correlated numerous vertical pneumatic
transport data and found out the solid friction term, AP, , can be
rs
expressed with an equation similar to the Fanning equation with a
solid friction factor, f :
s
2f p U 2L
with
C
fi A
f = 1.89 x 10"° —2— 6X . (13)
S . U. 3
The kinetic and the static energy terms, AP. and AP. , can
cCS nS
be expressed as
ps(1 - er>
ID _ •* 5
AP
ks 28C
P-22
-------
The solid kinetic energy term, AP, , is very important when
the pressure drop is measured over a section of pipe closely downstream
of the point of introduction of the solids, because the solid particles
are accelerated from zero velocity to their final velocities in this
13
section of the pipe. Kolpakov and Donat have made a detailed
experimental study of the pressure drop associated with the acceleration
of solid particles in vertical pneumatic conveying pipes. They conclude
that the simple expression for AP. of the form of Equation 14 is inadequate
and correlate their data through a complex empirical relationship
involving the terminal velocity of the particles, mean gas velocity and
the specific solid loading in the gas stream. Extrapolating their
correlation for the present application yields some unreasonable results
and thus is not used here.
Ignoring the negligible terms, the total pressure drop in the
draft tube can be represented by
2p,U 2 . p(l - e J U*
*P ' (f + '"-"' + ?> * >(1 - £
There are two unknowns in Equation 16, i.e., the voidage e and the solid
particle velocity U . The solid particle velocity in the draft tube
can be estimated through the correlation derived from the experiments
by Richardson and Zaki. The correlation is very similar
to that of Wen and Yu:
U.l-J&7DTDt-D£.r-Up,r (1?)
10 P
or
U e n-1
10 P
P-23
-------
18
To find n, the following equation can be used:
d
n = (4.45 + 18 j^) Refc~ ** for 1 < Refc < 200 , (19)
and
utVf
**t = —*— . (20')
For the present system (using Ottawa sand as bed material), n was cal-
culated to be 2.86.
18
The terminal velocity, U , was estimated from:
(P - P£) g d d U
s c
and
0.153(d ) * g (p p ) ' d U p
Ut PM 0.43° 0.29S for 2'° * -^T" * 5°° ' (22)
f
with a correction factor, n, for nonspherical particles,
*«
(23)
Voidage in the draft tube, er> can be calculated by writing a
material balance between the downcomers and the draft tube:
UP,r(1 - er)psAr = 2Up,d(1 ' ed)psAd = Ws '
Substituting Equation 18 into Equation 24, we have
a - •-' • • — — •
'r 10 P
P-24
Dt>
-------
The only unknown in Equation 25 is er; however, a trial and error procedure
is required to calculate e .
With f , U , and e calculated from Equations 13, 18, and
s p,r r
25, respectively, the pressure drop in the draft tube can now be
estimated from Equation 16.
19
Hinkle measured particle velocities experimentally by
means of a high-speed photographic technique and found the following
relationship correlated his data within + 5%:
i °'3P °'5> • (26)
He also proposed a way to calculate the solid friction term (AP ):
2 U GL
4pfS • £P —Sf- • <">
with
(28)
The largest specific solid loading in Hinkle's experiments was
6.02 compared to 23 to 66 in the present experiments. Equation 26 probably
has limited applicability because of the low specific solid loading covered in
his experiments; however, Hinkle's model was also utilized to estimate
the pressure drop in the draft tube for comparison with predictions
from other models.
Another model utilized to estimate the pressure drop in the
draft tube is the model by Curran and Gorin . The model is rather
complex; the readers should refer to the original reference for details.
The predictions by Curran and Gorin are also presented for comparison.
3pfCDD
fp ~ 2p d
F • Hs p
U,. - U
f,r p,r
U
P.r
P-25
-------
Jet Penetration
When the fluid-solid stream leaves the top of the draft tube,
it penetrates into the fluid bed above. The depth of this penetration
depends on the stream velocity, specific loading of the solid, and
particle and fluid properties. Due to the oscillating nature of the
bed operation, the penetration will vary for given operating conditions*
The average penetration depth, however, can be estimated by applying the
Bernoulli equation between the top of the draft tube and the end of the
penetration. The particle velocity was assumed to start to decrease at
the top of the draft tube, and the voidage in the jet starts to decrease
also. At the end of penetration, the solid particle velocity is essen-
tially zero and the voidage becomes equal to that of the fluidized bed
above the draft tube. The energy recovered from the deceleration of
the solid particles is dissipated by friction. The Bernoulli equation
gives
)()+ (1.E)_t .
The voidage, e, is the average voidage in the jet and is calculated from
the equation
er[J exp(-J) + exp(-J) - 1] - eb[exp(-J) + J - 1]
eave = J[exp(-J) - 1] " (30)
Equation 30 was obtained by assuming the voidage change in the jet can
be approximated by an exponential equation of the form
e = a + be"H (31)
with boundary conditions
at H = 0 e = e
(32)
at H = J e = e.
P-26
-------
Preliminary experimental data indicate.1 Lliat Jet peiu'Lratlnn
depth can be estimated from liquations 29 and 30 as preHfiited In Tables
P-9 to P-ll if the jet diameter, D , is taken to be the draft tube diameter.
Comparison with the Experimental Data
The experimental data were correlated with the mathematical
model described in the last section for the pressure drop in the down-
comers, in the draft tube, and for jet penetration depth. Hinkle's and
Curran and Gorin's model were also used to estimate the pressure drop
in the draft tube for comparison with the predictions by the present
model for Case I.
The pressure drop in the downcomer at minimum fluidization
was calculated from Equation 7 with e, = 0.467 and $ =0.86 and compared
with the experimental measurement in columns 11 and 12 of Tables P-12,
P-13, and P-14 for Case I. Comparisons between the experimental and the
theoretical values for Case III are presented in Table P-16. For Case II
(Table P-15) where additional serpentine copper rods were inserted into
the downcomers, the pressure drop calculations from Equation 7 were
corrected with the following correction factor:
(A, - A )
(H -H)+H % fc
* -- - • (33)
The pressure drop in the draft tube was calculated from Equation
16, from Hinkle's model, and from Curran and Gorin's model for Case I.
Their relative predictions were compared with the experimental measure-
ment in Table P-12 to P-14. Columns 1 through 6 were defined in the
previous section. Experimental observation shows that the pressure
difference between 3 and 4, AP3_^, can be taken as zero except when the jet
penetrates the fluid bed above the draft tube. (When the jet penetrates
the bed, AP^ has a value ranging from 0 to 0.5 H.O.) The pressure
drop between 1 and 2 can then be calculated from the difference between the
pressure drops in the downcomer and in the draft tube. The calculated
pressure drop in the draft tube, APn, from all models deviates consid-
K
erably from the experimental values, (AP ) - (AP J , especially
I/ 6Xp .L™ £ 6Xp
P-27
-------
TABLE P-9
COMPARISON BETWEEN EXPERIMENTAL AND CALCULATED
JET PENETRATION DEPTH (CASE I)
Run No.
1-L-l
1-1-2
1-1-3
1-2-1
1-2-2
1-2-3
1-3-1
1-3-2
1-3-3
1-4-1
1-4-2
2-1-1
2-1-2
2-1-3
2-1-4
2-2-1
2-2-2
2-2-3
2-3-1
2-3-2
2-4-1
2-4-2
2-5-1
2-5-2
Jet Penetration
Experimental (inches) J,
7-10
6-9
4-8
9-13
4-7
3-6
1- through
1-4
2-4
1- through
1- through
3-14
5-14
5-11
2-6
through
through
2-4
through
through
through
through
through
through
Deoth
Calculated (inches)
12. -9
9.5
2.18
12.9
8-.1
3.5
through
through
4.4
through
2.1
6.2
7.3
6.5
3.8
through
8.0
4.7
through
1.9
through
through
through
through
P-28
-------
TABLE P-9 (Continued)
Run No.
iJlet..Pene.tta£ion .Depth_
xperimental (inches) I Calculated (inches)
3-1-1
3-1-2
3-1-3
3-2-1
3-2-2
3-2-3
3-2-4
3-3-1
3-3-2
3-3-3
3-3-4
3-4-1
3-4-2
3-4-3
3-4-4
3-5-1
3-5-2
3-5-3
3-5-4
7-11
6-9
4. -5-7. 5
7-through
5-through
4- through
3.5-8
through
through
through
4- through
through
through
through
through
through
through
through
through
12.1
8.8
5.6
through
11.2
9.2
7.0
through
through
through
6.2
through
through
through
through
through
through
through
through
"Through" means the jet penetrates through the bed.
P-29
-------
TABLE P-10
COMPARISON BETWEEN EXPERIMENTAL AND CALCULATED
JET PENETRATION DEPTH (CASE II)
Jet Penetration Depth, Inches
Run
No.
1-1-1
1-1-2
1-1-3
1-1-4
1-1-5
1-2-1
1-2-2
1-2-3
1-2-4
1-2-5
l-2-6b
1-3-1
1-3-2
1-3-^
1-3-4°
1-3-5
l-3-6b
1-4-1
1-4-2
l-4-3b
1-4-4
l-4-5b
1-4-6°
Experimental
0.5-2.5
2.0-3.5
4.0-7.0
5.5-9.0
7.0-12.0
3.0-5.0
3.0-6.5
4.0-8.0
6.0-tlirouglia
9.0- through
8.0-tlirouf-h
3,0-7.0
4.0- tli rough
5.0-through
4.0-tlirough
6.0- tli rough
througli
through
through
through
through
through
through
Calculated
2.6
3.2
4.0
4.9
5.5
2.7
3.2
4.3
5.4
5.0
3.0
2.0
3.1
3.8
1.9
5.1
2.5
2.1
3.3
1.4
through
2.5
1.1
1-5-1
1-5-2
1-5-3
l-5-4b
l-5-5b
2-1-1
2-1-2
2-1-36
2-1-4
2-1-56
2-2-1
2-2-2
2-2-3b
2-2-4
2-2-5b
2-2-6b
2-3-1
2-3-2
2-3-3b
2-3-4
2-3-5b
through
through
through
through
through
4-9
6-14
7-through
7-through
8-through
5-9
7-through
7-through
9-tlirough
9-through
5-8
5-through
through
through
through
through
through
through
through
through
through
4.2
4.9
3.6
6.1
4.4
3.8
4.4
2.1
5.4
3.0
2.0
4.0
4.3
3.1
5.1
4.1
P-30
-------
TABLE P-10 (Continued)
Run
No.
2-4-1
2-4-2
2-4- 3b
2-4-4
2-4-5b
2-5-1
2-5-2
2-5-3
3-1-1
3-1-2
3-1-3
3-1-4
3-2-1
3-2-2
3-2-3
3-2-4b
3-2-5
3-2-6b
3-3-1
3-3-2
3-3-3
3-3-4b
3-3-5
3-3-6b
3-4-1
3-4-2
3-4-3
3-4-4
3-4-5
3-4-6D
3-5-1
3-5-2
3-5-3
3-5-4
3-5-5
3-5-6b
Experimental
through
through
through
through
through
through
through
through
4-8
5-10
7- through
8- through
5-9
5- through
8- through
7-through
5- through
7-through
4-through
6- through
5-through
through
5-through
5-through
through
through
through
through
through
through
through
through
through
through
through
> through
Calculated
through
through
through
through
through
through
through
through
4.9
6.1
7.4
6.0
4.6
5.9
6.5
5.0
7.1
5.6
4.5
5.9
6.6
4.6
6.3
5.3
3.8
through
through
through
through
through
through
through
through
through
through
through
a"Through" means the Jet penetrates through the bed.
bDowncomers are not at minimally fluldized condition.
P-31
-------
TABLE P-ll
COMPARISON BETWEEN EXPERIMENTAL AND CALCULATED
JET PENETRATION DEPTH (CASE III)
Run No.
B-l-1
B-l-2
B-l-3
B-l-4
B-2-1
B-2-2
B-2-3
B-2-4
B-3-1
B-3-2
B-3-4
B-3-5
B-4-1
B-4-2
B-4-3
B-4-5
Jet Penetration DC
Experimental
a
4.0- through
6.0- through
7.0- through
8.0- through
5.0- through
5.0- through
through
through
through
through
through
through
through
through
through
through
Ipth, Inches
Calculated
3.9
4.3
4.2
5.4
3.9
4.6
4.6
4.7
through
through
through
through
through
through
through
through
"Through" means the jet penetrates the fluid bed above the
draft tube.
P-32
-------
at lower air velocities in the draft tube. This Is because of the short
12 13 15 19
draft tube (18 inches) used in the cold model. Experimental data ' ' '
show that the solid concentration in the gas stream takes up to several
feet to attain the final equilibrium value. At lower gas velocities,
the solid acceleration is slower, and thus the voidage in the draft tube
increases to the final equilibrium value more slowly. Hence, the actual
average voidage in the draft tube is lower than the equilibrium value.
To obtain this average voidage, the exponential equation of the form of
Equation 31 is again assumed with boundary conditions
at H = 0 e = e,
(33)
at H = L e = e
Average voidage can then be found to be
ed[L exp(-L) + exp(-L) - 1] - er[exp(-L) + L - 1]
Eave = L[exp(-L) - 1]' (34)
The corrected static head term, (APhs)c> is then calculated through the
following equation
(APhs>c = ps(1 - eave)La + ps<1 ' er>
where
(3U - U )
La= (UQ-Ut)L ' (36)
For superficial gas velocity, UQ, larger than three times the terminal
velocity, no correction is necessary for the static head term based on
the experimental pressure measurements.
Theoretically, the distance required to accelerate a spherical
particle to a specific velocity should be evaluated from the following
integration equation:
P-33
-------
I
Ui
TABLE P-12
COMPARISON OF PRESSURE DROP PREDICTIONS IN THE DRAFT TUBE
(inches H,O)
1 (1)
Run No. 1 Up. r (ft/see)
1-1-1 (a)t 30.80
(b) 27.82
(c) 31.31
1-1-2 (a) 20.48
(b) 20.65
(c) 23.03
1-1-3 (a) 8.94
(b) 12.92
(c) 12.50
1-2-1 (a) 29.68
(b) 27.02
(c) 30.61
1-2-2 (a) 19.38
(b) 20.08
(c) 22.00
1-2-3 (a) 9.76
(b) 13.28
(c) 13.23
(2) TO) (4) '
0.975 11.25 1.02
9.18 1.02
0.975 11.57 1.00
0.959 8.04 1.64
8.18 1.64
0.962 9.37 1.51
0.942 2.15 2.31
4.50 2.31
0.957 3.10 1.71
0.973 11.24 1.09
9.32 1.09
0.973 11.71 1.08
0.962 6.66 1.52
7.14 1.52
0.965 7.82 1.39
0.937 2.80 2.52
5.19 2.52
0.952 3.89 1.91
(5) [ (6) (7) (8) (9)
5.26 17.53 1.02 17.53 3.40
0.88 11.08
2.29 14.86
4.29 13.97 5.85 18.19 4.40
1.31 11.14
2.34 13.22
1.68 6.14 6.43 10.26 2.60
0.86 7.67
1.80 6.61
5.27 17.60 1.09 17.60 3.20
0.96 11.38
2.30 15.09
3.69 11.87 6.34 16.69 2.90
1.17 9.83
2.11 11.32
2.02 7.34 6.95 11.77 3.20
1.04 8.75
2.04 7.84
(10) (11) (12)
z) JAPX 2) fll (APD5C3tp
-------
TABLE P-12 (Continued)
in
Run No.
1-3-1
1-3-2
1-3-3
1-4-1*
1-4-2*
(1)
Up. r (ft/sec
(2)
(a)
(b)
(c)
(a)
(b)
(c)
(a)
(b)
(c)
(a)
(b)
(c)
(a)
(b)
(c)
28.18
26.24
29.08
18.59
19.47
21.34
11.91
14.91
15.27
16.77
19.03
18.96
10.37
14.61
13.63
0.976
0.977
0.959
0.963
0.947
0.958
0.979
0.981
0.969
0.976
(3)
AP
fcs
(4)
(5)
(6)
(7)
8.71 0.94 4.15 13.79
7.54 0.94 0.85 9.33
9.14 0.92 1.99 12.05
6.60 1.64 3.53 11.77
7.24 1.64 1.20 10.09
7.86 1.47 2.19 11.52
3.51 2.12 2.31 7.94
5.49 2.12 1.20 8.81
4.59 1.67 1.95 8.21
2.70 0.82 1.59 5.12
3.48 0.82 0.58 4.89
3.13 0.76 1.15 5.04
1.54 1.23 1.15 3.92
3.07 1.23 0.62 4.92
2.10 0.96 1.18 4.24
(8)
(ii)
1.02 13.87
7.07 17.20
8.06 13.89
7.22 9.91
exp
2.60 5.47
3.40 2.61
3.60 5.32
6.99 11.28 0.90 2.26
0.90 3.63
(12)
exp
21.66 19.34
21.66 19.81
21.66 19.21
13.54
13.54
^ Row (a):Predictions by the present model.
(b):
(c):
Downcomers are not
the Hinkle's model.
" the Curran and Gorin'a model.
minimally fluldized .
-------
TABLE P-13
COMPARISON OF PRESSURE DROP PREDICTIONS IN THE DRAFT TUBE
(inches H,0)
(1)
Run No. Up. r (ft/sec)
2-1-1 (a)+ 20.59
(b) 21.34
(c) 23.18
2-1-2 (a) 19.21
(b) 20.10
(c) 22.33
.2-1-3 (a) 16.21
(b) 17.91
nj (c) 19.70
LO
°* 2-1-4 (a) 11.11
(b) 14.45
(c) 14.87
2-2-1 (a) 28.44
(b) 26.17
(c) 29.69
2-2-2 (a) 17.59
(b) 18.66
(e) 20.46
2-2-3 (a) 11.36
(b) 14.23
(c) 14.75
(2)
ET
0.975
0.978
0.966
0.969
0.957
0.964
0.948
0.961
0.971
0.972
0.954
0.959
0.936
0.949
(3)
AP
4.85
5.21
5.53
5.91
6.47
7.09
5.24
6.40
6.53
2.98
5.04
4.05
10.87
9.21
11.68
6.68
7.52
7.97
3.84
6.04
5.16
(4)
AP,
0.98
0.98
0.88
1.37
1.37
1.23
1.71
1.71
1.43
2.07
2.07
1.55
1.15
1.15
1.12
1.85
1.85
1.63
2.55
2.55
2.03
(5)
AP
2.55
0.79
1.56
3.19
1.05
1.92
2.95
1.10
2.03
2.00
1.06
1.85
5.13
1.04
2.36
3.60
1.29
2.28
2.50
1.33
2.27
(6) (7)
AP (AP.
8.38 4.
6.98
7.97
10.47 6.
8.89
10.24
9.90 8.
9.21
9.99
7.05 7.
8.17
7.45
17.15 1.
11.40
15.16
12.13 8.
10.66
11.88
(8) I (9) (10)
_)_ <&Vc1(&Pl 2> (M>1 S^alc
exp
68 12.08 2.20 6.61
20 15.30 2.60 4.07
94 17.14 2.60 2.32
67 12.64 2.10 6.34
26 17.27 3.50 2.76
13 18.42 3.60 1.57
8.89 7.84 14.18 3.40 5.47
9.92
9.46
(11) (12)
(&P_) , (AP_) ..
exj> D cai<
16.25 18.69
18.96 19.37
19.04 19.49
18.28 18.98
20.40 20.03
21.76 19.99
18.96 19.65
-------
TABLE 13 (Continued)
Run No.
2-3-1*
2-3-2*
2-4-1*
i 2-4-2*
OJ
•vl
2-5-1*
2-5-2*
(1)
Up. r (ft/sec)
(2)
(3)
AP
(a)
(b)
(c)
(a)
(b)
(c)
(a)
(b)
(0
(a)
(b)
(c)
(a)
(b)
(0
(a)
(b)
(c)
21.43
21.73
23.51
10.31
14.68
13.46
18.39
20.00
20.51
11.68
15.10
15.02
18.46
19.96
20.70
13.04
15.95
16.30
0.971
0.973
0.972
0.978
0.977
0.979
0.958
0.967
0.975
0.977
0.958
0.965
fcs
(4)
AP
fcs-
(5)
(6)
(7)
(8)
6.09 1.14 3.11 10.34 4.45
6.26 1.14 0.95 8.35
6.88 1.08 1.81 9.77
1.36 1.10 1.02 3.48 7.18
2.76 1.10 0.56 4.42
1.84 0.88 1.05 3.77
3.59 0.91 1.95 6.45 5.96
4.24 0.91 0.68 5.83
4.08 0.84 1.34 6.26
2.67 1.68 1.76 6.11 7.91
4.47 1.68 0.77 6.92
3.52 1.31 1.64 6.47
3.97 1.00 2.12 7.09 6.06
4.64 1.00 0.74 6.38
4.50 0.92 1.42 6.84
3.34 1.68 2.06 7.09 8.48
5.00 1.68 0.87 7.55
4.28 1.39 1.73 7.40
(9)
(10)
exp
13.66 3.40 2.66
9.56 1.90 2.68
11.50 2.10 3.46
12.34 2.40 5.34
12.15 2.10 2.13
13.89 2.80 3.11
(11;
(12)
16.32
12.24
14.96
17.68
14.28
17.00
| Row (a): Predictions by the present model.
11 (b): " by the Hinkle's model.19 ,,
" (c): " by the Curran and Gorln's model.
* Downcomers are not minimally fluldized .
-------
u>
00
TABLE P-14
COMPARISON OF PRESSURE DROP PREDICTIONS IN THE DRAFT TUBE
(Inches H.O)
Run No.
3-1-1
3-1-2
3-1-3
3-2-1
3-2-2
3-2-3
3-2-4
3-3-1
3-3-2
3-3-3
3-3-4
(1)
Up. r (ft/sec)
(a)t 26.87
(b) 25.08
(c) 28.38
(a) 20.75
(b) 20.99
(c) 23.18
(a) 14.05
(b) 16.35
(c) 17.26
(a) 28.93
(b) 26.55
(c) 29.94
(a) 25.66
(b) 24.31
(c) 27.34
(a) 21.43
(b) 21.45
(c) 23.82
(a) 17.50
(b) 18.83
(c) 20.34
(a) 27.95
(b) 25.96
(c) 29.01
(a) 22.52
(b) 22.24
(c) 24.61
(a) 19.22
(b) 20.01
(c) 21.77
(a) 16.19
(b) 17.97
(c) 19.12
(2)
er
0.969
0.970
0.963
0.966
0.950
0.959
0.973
0.973
0.969
0.970
0.964
0.967
0.960
0.964
0.973
0.975
0.967
0.970
0.963
0.967
0.959
0.965
(3)
UJ
10.44
9.10
11.35
7.32
7.49
8.44
4.56
6.17
5.72
10.64
8.96
11.11
9.58
8.59
10.50
7.71
7.71
8.84
5.69
6.59
6.86
9.65
8.32
9.82
7.84
7.64
8.59
6.38
6.92
7.25
4.98
6.14
5.94
(4)
1.24
1.24
1.20
1.46
1.46
1.35
1.98
1.98
1.63
1.09
1.09
1.08
1.25
1.25
1.20
1.44
1.44
1.31
1.59
1.59
1.43
1.06
1.06
1.00
1.32
1.32
1.19
1.48
1.48
1.31
1.63
1.63
1.39
(5)
4.78
1.12
2.35
3.71
1.18
2.14
2.69
1.09
2.09
4.79
0.96
2.23
4.64
1.12
2.27
3.96
1.19
2.21
3.21
1.12
2.03
4.49
0.95
4.28
3.96
1.13
2.09
3.43
1.13
2.04
2.88
1.07
1.92
(6) (7) (8) (9)
AP (iP ) (AP > CAP ,) ^
exp
16.46 2.02 17.24 3.10
11.46
14.90
12.49 5.39 16.41 3.10
10.13
11.93
9.23 8.91 16.16 3.30
9.24
9.44
16.52 1.09 16.52 2.90
11.01
14.42
15.47 2.53 16.75 3.10
10.96
13.97
13.11 4.95 16.62 3.30
10.34
12.36
10.49 7.76 16.66 3.30
9.30
10.32
15.20 1.30 15.44 3.4
10.33
15.10
13.12 4.17 15.97 3.2
10.10
11.87
11.29 6.38 16.20 3.2
9.53
10.60
9.49 8.82 16.68 3.5
8.83
9.25
(10) (11) (12)
{AP1 ?)ralr ^Vmm (AVr;,lr
2.85 19.04 20.09
3.36 19.04 19.77
3.33 19.04 19.49
3.31 19.60 19.83
3.20 20.30 19.95
3.21 19.60 19.83
2.88 19.43 19.54
4.2 19.9 19.6
3.7 19.4 19.7
3-4 19.4 19.6
2-7 19.4 19.4
-------
TABLE P-14 (Continued)
TJ
CO
VO
(1) (2) (3) (4)
Run No. Uo. r (ft/sec) e AP. AP.
3-4-1 (a) 24.99 0.975 7.37 1.01
(b) 24.12 6.86 1.01
(c) 26.38 0.976 7. 65 0.96
3-4-2 (a) 21.52 0.971 6.32 1.17
(b) 21.76 6.45 1.17
(c) 23.57 0.973 6.90 1.08
3-4-3 (a) 17.23 0.965 4.82 1.39
(b) 18.83 5.76 1.39
(c) 19.94 0.970 5.58 1.19
3-4-4 (a) 14.08 0.959 3.75 1.62
(b) 16.65 5.25 1.62
(c) 17.17 0.967 4.55 1.31
3-5-1 (a) 24.73 0.975 7.14 1.00
(b) 23.97 6.70 1.00
(c) 26.15 0.977 7.45 0.92
3-5-2 (a) 20.67 0.969 6.18 1.24
(b) 21.14 6.47 1.24
(c) 22.90 0.972 6.88 1.11
3-5-3 (a) 16.03 0.961 4.62 1.54
(b) 17.94 5.79 1.54
(c) 18.94 0.967 5.50 1.31
3-5-4 (a) 13.59 0.955 3.87 1.79
(b) 16.21 5.50 1.79
(c) 16.76 0.964 4.74 1.43
(5) (6) (7) (8) 1 (9) (10) 1 (11) 1 (12)
APf AP (AP ) (AP_) 1 (AP. ,) (AP. ,) . 1 (AP_) 1 (AP.) ,
exo
3.58 11.96 2.44 13.39 2.9 5.8 19.2 19.2.
0.90 8.78
1.79 10.40
3.27 10.76 4.46 14.05 3.1 5.1 19.4 19.2
0.98 8.60
1.77 9.75
2.72 8.94 7.61 15.16 3.1 3.9 18.9 19.1
0.97 8.12'
1.86 8.53
2.29 7.67 8.85 14.92 3.2 4.1 18.5 19.0
0.91 7.78
1.73 7.59
3.48 11.62 2.54 13.16 2.4 6.0 18.7 19.2
0.89 8.59
1-83 10.20
3.26 10.69 5.07 14.52 2.8 4.8 18.9 19.3
1.01 8.72
1-86 9.85
2.72 8.88 8.80 16.14 2.8 3.1 18.4 19.2
1.00 8.34
1.79 8.60
2.40 8.06 8.71 14.98 2.8 4.1 18.7 19.1
0.96 8.26
1.83 8.00
4 Row (a): Predictions by the present model.
" (b): " by the Hinkle's model.19
" (c): " by the Curran and Gorin's U model.
-------
TABLE P-15
COMPARISON BETWEEN THEORETICAL PREDICTIONS
AND EXPERIMENTAL VALUES IN THE DOWNCOMER (CASE II)
Run
No.
1-1-1
1-1-2
1-1-3
1-1-4
1-1-5
1-2-1
1-2-2
1-2-3
1-2-4
1-2-5
1-3-1
1-3-2
1-3-3
1-3-5
1-4-1
1-4-2
1-4-4
1-5-1
1-5-2
1-5-3
2-1-1
2-1-2
2-1-4
2-2-1
2-2-2
2-2-4
2-3-1
2-3-2
2-3-4
2-4-1
2-4-2
2-4-4
Pressure Drop in Downcomer, in. H.O
Experimental
15.4
15.2
15.4
15.6
15.4
15.4
15.4
15.6
15.4
15.0
15.2
15.4
15.4
15.0
15.2
15.4
15.4
15.2
15.0
14.7
15 .'6
15.4
15.4
15.4
15.2
15.0
15.6
15.6
15.4
15.6
15.4
15.4
Theoretical
14.8
14.8
14.8
14.9
14.8
14,8
14.8
14.8
14.8
14.8
14.6
14.7
14.7
14.8
14.7
14.8
14.7
14.9
14.9
14.9
14,9
14.9
14.9
14.8
14.8
14.8
14.9
14.8
14,8
14.8
14.8
14.8
P-40
-------
TABLE P-15 (Continued)
Run
No.
2-5-1
2-5-2
2-5-3
3-1-1
3-1-2
3-1-3
3-2-1
3-2-2
3-2-3
3-2-5
3-3-1
3-3-2
3-3-3
3-3-5
3-4-1
3-4-2
3-4-3
3-4-4
3-4-5
3-5-1
3-5-2
3-5-3
3-5-4
3-5-5
Pressure Drop in Downcomer,
Experimental
15.4
15.4
15.4
15.8
15.4
15.2
15.4
15.2
15.2
15.4
15.2
15.6
15.6
15.0
15.0
15.2
15.4
15.4
14.9
15.2
15.0
15.4
15.0
15.4
in. H20
Theoretical
14.9
14.9
14.9
15.3
15.3
15.3
15.2
15.3
15.3
15.3
15.2
15.3
15.3
15.2
15.1
15.1
15.2
15.2
15.2
15.2
15.2
15.3
15.3
15.4
P-41
-------
TABLE P-16
COMPARISON BETWEEN THEORETICAL PREDICTIONS
AND EXPERIMENTAL VALUES IN THE DOWNCOMER (CASE III)
Run No.
B-l-1
B-l-2
B-l-3
B-l-4
B-2-1
B-2-2
B-2-3
B-2-4
B-3-1
B-3-2
B-3-4
B-3-5
B-4-1
B-4-2
B-4-3
B-4-5
Downcomer Pressure Drop, in. HO
Predicted
18.2
18.1
13.0
18.1
18.2
18.1
18.0
18.0
18.2
18.1
18.0
18.0
18.
18.
18.
Experimental
18.9
18.4
18.4
18.4
18.9
18.9
18.7
18.2
19,
18,
18.
18,
19.
18.
18.
18.2
18.5
P-42
-------
AL -
(37)
4 dp pg 2 D
However, this requires a relation between the particle friction factor,
f , and the particle velocity, U . There is no proven and reliable
correlation existing so far. It is also difficult to isolate the solid
friction term from the pressure measurements of a short draft tube.
Because of these uncertainties, Equation 38 was not used in calculating the
pressure drop in the draft tube. Fortunately, this correction is
negligible when the draft tube is sufficiently long (i.e., > 10 ft).
The corrected static pressure drop, (AP ) , is presented in
us c
column 7 . Column 8 , (APD) , is the corrected total pressure drop
K C
in the draft tube by taking into account the term (AP, ) . Columns 9
. CIS O
and 10 compare the experimental and calculated values of AP, „• The
calculated pressure drop between 1 and 2, (AP, ~) i > is determined by
subtracting (APR) from (AP_) 1 . It is noticed that Curran and
Gorin's model predicts a higher solid kinetic loss, a lower solid
static loss, and a lower solid friction loss, resulting in a lower
overall pressure drop in the draft tube. The difference, however, is
usually less than 2.5 in H-0. Hinkle's model gives either a higher or a
lower kinetic loss, depending on the gas velocity range in question;
the friction loss, however, is much lower when compared to the predictions of
the other two models, resulting in a still lower pressure drop. The
difference in total pressure drop ranges up to 6.5 in H00 at high gas
velocities in the draft tube.
Excluding the runs with superficial air velocity in the draft
tube less than * 22 ft/sec, the calculated pressure drop, (APi_o) i »
is within + 37% of the experimental value in 70% of the data taken.
When the superficial air velocity in the draft tube is less than 22 ft/sec,
-------
TABLE P-17
COMPARISON BETWEEN THE CALCULATED AND
THE OBSERVED SOLID FLOW RATE
Run No.
1-1-1
1-1-2
1-2-1
1-2-2
1-3-1
1-3-2
1-4-1
2-1-1
2-1-2
2-1-3
2-2-1
2-2-2
2-3-1
2-4-1
2-5-1
2-5-2
3-1-1
3-1-2
3-1-3
3-2-1
3-2-2
3-2-3
3-2-4
Observed Solid
Flow Rate, Ib/sec
1.42
1.55
1.46
1.36
1.20
1.41
0.631
0.912
1.21
1.27
1.50
1.49
1.12
0.759
0.830
1.00
1.53
1.39
1.27
1.42
1.47
1.42
1.29
Calculated Solid
Flow Rate, Ib/sec
1.02
0.917
1.03
1.22
1.82
1.11
1.01
2.06
1.49
1.03
1.15
0.79
1.12
1.34
0.969
1.25
1.18
1.31
1.31
1.30
1.27
1.27
1.18
Deviation
%
- 28.2
- 40.8
- 29.5
- 10.3
+ 51.7
- 21.3
+ 60.1
+125.9
+ 23.1
- 18.9
- 23.3
- 47.0
+ 0
+ 76.5
+ 16.7
+ 25.0
- 22.9
- 5.8
+ 3.1
- 8.5
- 13.6
- 10.6
- 8.5
P-44
-------
TABLE P-17 (Continued)
Run No.
3-3-1
3-3-2
3-3-3
3-3-4
3-4-1
3-4-2
3-4-3
3-4-4
3-5-1
3-5-2
3-5-3
3-5-4
Observed Solid a
Flow Rate, Ib/sec
1.33
1.37
1.31
1.22
1.14
1.15
1.10
1.04
1.12
1.18
1.14
1.11
Calculated Solid
Flow Rate, Ib/sec
1.52
1.40
1.32
1.13
1.89
1.73
1.45
1.49
1.93
1.66
1.24
1.49
Deviation
:%
'+ 14.3
+ 2.2
+ 0.8
- 7.4
+ 65.7
+ 50.5
•f 31.8
+ 43.2
+ 72.2
+ 40.6
+ 8.8
+ 34.2
Includes both downcomers.
P-45
-------
the largest particles of the bed material which have terminal velocity
close to 22 ft/sec are likely to accumulate in the draft tube and form
a slugging bed. This phenomenon was actually observed in the experi-
ments. For those runs, the pressure drop prediction using a vertical
pneumatic transport correlation is no longer valid.
From the experimental relationship for the solid velocities
in the downcomers as shown in Equation 2, the solid circulation rate can
be calculated through the following equation:
2/3
W = 0.107 p A. (1 - e,)(2 gH ) ' ,for Case I,-
S S Q Q p *"
2/3
W = 0.107 p A (1 - e,)(2 gH ) ' for Case II, (39)
S S C Q p
and
W = 0.025 p A, (1 - e,)(2 gH ) for Case III,
s s a dp
where
A = F • A, .
c c d
By using the predicted pressure drop between 1 and 2, (AP, .) , , tne
calculated solid flow rates are compared to the observed ones in
Table P-17 for Case I. About 75% of the calculated solid flow rate comes
within + 35% of the observed solid flow rate.
The calculated jet penetration depth was compared to the
observed value in Tables P-9 through P-ll. Due to the oscillating nature
of the actual bed operation, the observed bed penetration depth has a
3 to 5 in range. The predictions from Equation 29 are good.
P-46
-------
DISCUSSION
By analogy with liquid, the flow rate of minimally fluidized
particles during outflow through a small orifice in a thin wall with
18
constant bed height above the orifice is given by Massimilla as
1/2
Ws = Co A PS (2 * V ' (40)
The solid discharge coefficient increases with the increase of d /d
o p
up to 40' and then levels off to a value of about 0.5. If the orifice
is a high vertical slot partially filled with fluidized particles, the
3/2
solid discharge rate is Wg a Hmf . If the orifice is profiled to a
quadratic hyperbola equation, the discharge rate is W a H f.
5 Hli
In the present set-up, the discharge of solid particles from
the downcomers into the draft tube is very similar to solid discharge
from a minimally fluidized bed. Thus, the solid recirculation rate in
the bed should be in a form as shown in Equation 40. Comparing Equation 40
with the experimental equations obtained, Equations 2 and 3 conform with
Equation 40 but with a different dependence on H _. This is understandable
mi
because of the different designs of the draft tube inlets and the much larger
orifice openings. However, when the solid, discharge coefficients (Co)
are evaluated from Equations 2 and 3 by using the open area between the
draft tube and the downcomers as the orifice openings (A), the coefficients
are * 0..05 for Cases I and II and ^0.01 for Case III, which are much
less than the 0.5 from Equation 40. Apparently, the solid recirculation
rate in a recirculating bed involves, some other complicated mechanisms
which are explained below.
In a recirculating bed, the solid recirculation rate is con-
trolled not only by the rate of solid discharge from the downcomers but
also by the rate of solid pick-up in the draft tube. At steady state,
P-47
-------
the solid discharge rate from the downcomers should equal the solid
pick-up rate in the draft tube. Thus, the solid recirculation rate can
be controlled either by the solid discharge rate from the downcomers or
by the solid pick-up rate in the draft tube, depending on the operating
conditions. The pressure drop in a minimally fluidized downcomer is
relatively independent of the solid circulation rate, because the actual
air velocity is much higher than the solid velocity (see Equation 7);
the pressure drop in the draft tube, however, is sensitive to the air
velocity in the draft tube and the total solid circulation rate as shown
in Figure P-7. There exists a choking air velocity for every solid flow
rate. An air velocity lower than the choking velocity will result in the
collapse of solid particles into a slug and will eventually choke the
vertical pneumatic conveying line. Thus, the solid loading at the
choking velocity represents the saturation carrying capacity of the
air stream.
The saturation carrying capacity of a gas stream can be cal-
culated from the empirical correlation by Zenz and Weil.20 The
saturation carrying capacities of the draft tube at different air
velocities are compared with the experimental solid recirculation rates
for Cases I and III in Figure P-8. Only the data for "minimally fluidized"
downcomers are included. It is interesting to notice that at lower air
velocities, the experimental solid circulation rate is much larger than
the saturation carrying capacity of the air stream. This can be easily
understood if we see the downcomers as solid feeders of capacities
depending on ^Pjo' When the air velocity in the draft tube is low,
A?1_2 is large. The downcomers feed the solid into the draft tube at
a rate much larger than that which can be handled by the air stream. Thus,
the solid transport in the draft tube takes a form of slugging transport
rather than pneumatic transport in the conventional sense. The slugging
imposes an additional pressure drop in the draft tube. This explains why
the pressure drop predictions for the draft tube from all existing
mathematical models are much less than the experimental observations at lower
P-48
-------
Curve 651039-A
u
ch
log u
f
Figure P-7- Schematic Representation of Flow Characteristics
in Vertical Pneumatic Transport^
From Fluidization and Fluid Particle Systems by F. Zenz & D. Othmer
© 1960 by Litton Educational Publishing, Inc.
P-49
-------
T
g
1000
800
600
400
200
100
80
xf 60
^ 40
Qj"
§i 20
6;ic-.o-a
o
00
10
8
6
4
Saturation Carrying Capacity
Draft Tube Height = 18 in.
10
20
40 60 80 10
20 40 60 80 10 20 40 60 80 10
Superficial Air Velocity in the Draft Tube, ft/sec
20
40 60 80 100
Figure P-8-Comparison between the Experimental Data and the Saturation Carrying Capacity
-------
draft tube velocities. In fact, the pressure loss due to this slugging
depends linearly on the theoretical pressure difference between the
downcomers and the draft tube. Figure P-9 plots the pressure difference
between the downcomers and the draft tube, (AP_) - (AP« „) , , and
D exp 2-3 calc
the slugging pressure loss, (APD>exp - (AP2_3)calc - (AP). The
straight line can be expressed by the following equation:
(AP)c = 3.23 + 1.025 t(AP) - (A1 for Cases l and
and
(41)
(AP)C = 2.10 -f 1.025 [(AP^ - (AP2_3)calc] for Case II.
If Equation 41 is used, but (APD)calc exchanged for (APD)ex , to predict
the solid circulation rate, most predictions are good to + 20% of the
experimental values for those runs with downcomers minimally fluidized.
The results are presented in Tables P-18 through P-20.
When the air velocity in the draft tube is increased, eventually
an air velocity is reached -.where the solid discharge rate from the
downcomers equals the saturation carrying capacity of the air stream.
This is the point of intersection between the experimental curve and the
curve for the saturation carrying capacity in Figure P-8. This point
also corresponds to the point of choking in Figure P-7 (point X) .
The. choking point does not coincide with the point of lowest pressure
drop, Y. Thus, a further increase in air velocity will result in a
slightly higher solid recirculation rate; the recirculation rate decreases
quickly, however, at even higher velocities because of the increasing
pressure drop in the draft tube. At, this region, the solid discharge
from- the downcomers becomes the controlling factor in solid circulation.
The optimum air velocity in the draft tube which corresponds to the
maximum solid circulation rate can be found by maximizing the pressure
drop between the downcomers and the draft tube, AP. _.
When the draft tube inlet design is as shown in Figure P-2, the
solid circulation rate is much smaller. This is because the air input
P-51
-------
Curve 653463-A
15.0
A Draft tube height 18 in.
o Draft tube height 18 in. with modified
draft tube inlet design
• Draft tube height 18 in. with serpentine
rods in downcomers
s 10.0
(AP) =
<
5.0
0
0
-Z. 10 + 1.025 [(APJ
-(APD) i
R calc.l
1.025[(APJ
(AP ) D exp
R; calc.J
5.0
10.0
Fig. 9-Pressure Loss Due to Slugging Transport
15.0
P-52
-------
TABLE P-18
COMPARISON BETWEEN THE PREDICTED
AND THE EXPERIMENTAL SOLID RECIRCULATION RATE (CASE I)
Solid Recirculation Rate, Ib/sec
Run No.
1-1-1
1-1-2
1-2-1
1-2-2
1-3-1
l«3-2
2-1-1
2-1-2
2-1-3
2^2-1
2-2-2
3-1-1
3-1-2
3-1-3
3-2-1
3-2-2
3-2-3
3-2-4
3-3-1
3-3-2
3-3-3
3-3-4
3-4-1
3-4-2
3-4-3
3-4-4
3-5-1
3-5-2
3-5-3
3-5-4
Predicted 1 Experimental
1.27 1.42
1.24 1.55
1.27 1.46
1.23 1.36
1.24 1.20
1.23 1.41
1.21 0.912
1.22 1.21
1.22 1.27
1.26 1.50
1.23 1.49
1.26 1.53
1.23 1.39
1.21 1.27
1.26 1.42
1.25 1.47
1,24 1.42
1.22 1.29
1.25 1.33
1.24 1.37
1.22 1.31
1.21 1.22
1.23 1.14
1.22 1.15
1.21 1.10
1.21 1.04
1.23 1.12
1.22 1.18
1.21 1.14
1.21 l.ll
Deviation
- 10.6
- 20.0
- 13.0
- 9.6
+ 3.3
- 12.8
+ 32.7
+ 0.8
- 3.9
- 16.0
- 17.4
- 17.6
- 11.5
- 4.7
- 11.3
- 15.0
- 12.7
- 5.4
- 6.0
- 9.5
- 6.9
- 0.8
+ 7.9
+ 6.1
+ 10.0
+ 16.0
+ 9.8
+ 3.4
+ 6.1
+ 9.0
P-53
-------
TABLE l>-19
COMPARISON BETWEEN THE PREDICTED
AND THE EXPERIMENTAL SOLID RECIRCULATION RATE (CASE II)
Solid Recirculation Rate, Ib/sec
Run
No.
1-1-1
1-1-2
1-1-3
1-1-4
1-1-5
1-2-1
1-2-2
1-2-3
1-2-4
1-2-5
1-3-1
1-3-2
1-3-3
1-3-5
1-4-1
1-4-2
1-4-4
1-5-1
1-5-2
1-5-3
2-1-1
2-1-2
2-1-4
2-2-1
2-2-2
2-2-4
2-3-1
2-3-2
2-3-4
2-4-1
2-4-2
2-4-4
Predicted
0.723
0.725
0.731
0.733
0.738
0.723
0.725
0.731
0.738
0.733
0.720
0.723
0.728
0.733
0.720
0.725
0.731
0.725
0.731
0.736
0.731
0.733
0.740
0.727
0.730
0.738
0.730
0.730
0.735
0.727
0.733
0.735
Experimental
0.709
0.711
0.726
0.744
0.718
0.705
0.683
0.700
0.716
0.675
0.620
0.640
0.640
0.675
0.625
0.713
0.653
0.742
0.750
0.752
0.744
0.744
0.750
0.703
0.700
0.713
0.742
0.690
0.683
0.731
0.713
0.707
Deviation, %
+1.97
+1.97
+0.69
-1.48
+2.79
+2.55
+6.15
+4.43
+3.07
+8.59
+16.13
+12.97
+13.75
+ 8.59
+15.20
+ 1.68
+11.94
-2.29
-2.53
-2.13
-1.75
-1.48
-1.33
+3.41
+4.29
+3.51
-1.62
+5.80
+7.61
-0.55
+2.81
+3.96
P-54
-------
TABLE P-19 (Continued)
Run
No.
2-5-1
2-5-2
2-5-3
3-1-1
3-1-2
3-1-3
3-2-1
3-2-2
3-2-3
3-2-5
3-3-1
3-3-2
3-3-3
3-3-5
3-4-1
3-4-2
3-4-3
3-4-4
3-4-5
3-5-1
3-5-2
3-5-3
3-5-4
3-5-5
1 Predicted
0.727
0.735
0.738
0.738
0.743
0.753
0.735
0.743
0.745
0.751
0.735
0.743
0.745
0.743
0.730
0.733
0.740
0.743
0.745
0.733
0.735
0.743
0.743
0.753
Experimental
0.731
0.759
0.759
0.977
0.977
0.977
0.943
0.953
0.949
0.971
0.940
0.951
0.962
0.895
0.852
0.858
0.930
0.908
0.893
0.940
0.934
0.994
0.977
1.053
Deviation, %
-0.55
-•3.16
-2.77
-24.46
-23.95
-22.93
-22.06
-22.04
-21.50
-2-2.66
-21.81
-21.87
-22.56
-16.98
-14.32
-14.57
-20.43
-18.17
-16.57
-22.'02
-21.31
-25.25
-23.95
-28.49
P-55
-------
TABLE P-20
COMPARISON BETWEEN THE PREDICTED
AND THE EXPERIMENTAL SOLID RECIRCULATION RATE (CASE III)
Run No.
Solid Recirculation Rate, Ib/sec
Predicted
Experimental
Deviation, %
B-l-1
B-l-2
B-l-3
B-l-4
B-2-1
B-2-2
B-2-3
B-2-4
B-3-2
B-3-4
B-3-5
B-4-1
B-4-2
B-4-3
B-4-5
0.640
0.640
0.640
0.645
640
642
642
642
0.640
0.640
0.640
0.645
0.640
0.640
0.642
0.647
0.681
0.638
0.586
0.662
0.679
0.662
0.614
0.601
0.685
0,631
0.598
0.610
0.694
0.644
0.640
0.681
- 6.0
+ 0.3
+ 9.2
- 2.6
- 5.8
- 3.0
+ 4.6
+ 6.8
" 6.6
+ 1.4
+ 7.0
+ 5.7
- 7.8
- 0.6
+ 3.1
- 5.0
P-56
-------
into the air plenums appears as air bubbles which occupy almost the
whole cross-section right above the distributor plate. These bubbles
either by-pass into the draft tube and contribute to slugging or ascend
into the downcomers and interfere with the solid downward flow. This
phenomenon also creates severe pressure fluctuation in the downcomers
and the draft tube.
Even when the length of the draft tube is increased, slugging
is still expected to occur in the lower part of the draft tube at low
air velocities to make up the additional pressure drop. Predictions of
solid recirculation rate, at low air velocities, from pressure drop
calculations of the vertical pneumatic transport without taking into account
the slugging will give high values when compared with experimental data.
Thus, in designing and operating a recirculation bed, it is preferable
to operate at around the maximum solid recirculation rate, not only to
take advantage of the high recirculation rate but also to prevent
slugging in the draft tube.
CONCLUSIONS
• A mathematical model has been developed which can be used to predict
recirculation or to select operating conditions to achieve a given
recirculation rate for a given reactor configuration.
• The draft tube inlet design, draft tube height, downcomer draft
tube area ratio are critical design variables for a recirculating
fluid bed reactor.
P-57
-------
ACKNOWLEDGEMENT
Calculations of the pressure drops in the draft tube using
Curran and Gorin's model were performed by Dr. J.P.L. Chen. His help
is hereby acknowledged.
P-58
-------
REFERENCES
1. Dent, F. J., "Methane From Coal," 9th Coal Science Lecture, BCURA
(1960).
2. Horsier, A. G., and Thompson, B. H., "Fluidization in the Development
of Gas Making Processes," Tripartite Chem. Engr. Conf., Montreal
(Sept. 1968).
3. Horsier, A. G., Lacey, J. A., and Thompson, B. H., "High Pressure
Fluidized Beds," CEP, 65 (10), 59-64 (1969).
A. Decamps, F., Dumont, G., and Goossens, W., "Vertical Pneumatic Conveyor
with a Fluidized Bed as Mixing Zone," Power Technol., 5_, 299 (1971/72).
5. Buchanan, R. H. and Wilson, B., "The Fluid-Lift Solids Recirculator,"
Mech. & Chem. Eng., Trans. (Australia), 117, May (1965).
6. Curran, G. P., and Gorin, E., "The CO. Acceptor Gasification Process -
A Status Report - Application to Bituminous Coals," Paper presented
at the Third International Conference on Fluidized Bed Combustion,
Hueston Woods, Ohio (1972).
Currati, G. P., Pasek, B., Pell, M., and Gorin, E., "Fretreatment of
Bituminous Coals for Pressure Gasification," Paper presented at
the Fluidized Bed Combustion Symposium, American Chemical Society
Meeting, Chicago, Illinois (1973).
7. Taskaev, N. D., and Kozhina, M. I., Trudy Akad. Nauk Kirgiz S.S.R.,
]_, 109 (1956).
8. Archer, D. H., Vidt, E. J., Keairns, D. L., Morris, J. P., and
Chen, J. L. P., "Coal Gasification for Clean Power Production,"
Proceedings of the Third International Conference on Fluidized
Bed Combustion, Hueston Woods, Ohio, October 1972.
P-59
-------
9. Ergun, S., CEP 48, 89 (1952).
10. Yoon, S. M. and Kunii, D., "Gas Flow and Pressure Drop Through Moving
Beds," Ind. Eng. Chem. Process Des. Develop. 9^ (4), 559 (1970).
11. Curran, G. P. and Gorin, E., "Studies on Mechanics of Flow-Solids
Systems," report prepared for Office of Coal Research by Consolidation
Coal Company (1968).
12. Jones, J. H., Braun, W. G., Daubert, T. E., and Allendorf, H. D.,
"Estimation of Pressure Drop for Vertical Pneumatic Transport of
Solids," AIChE Journal, 13 (3), 608 (1967).
13. Kolpakov, V. M. and Donat, E. V., "An Investigation of the Pressure
Drop in the Acceleration Zone of a Vertical Pipeline for Conveying
Solid Particles," Intern. Chem. Eng., 10 (3), 394 (1970).
14. Papai, L., "Velocity and Pressure in Vertical Pneumatic Conveying,"
Acta. Tech. Aca. Sci, Hungaricae, 69, 83 (1970).
15. Stemerding, S., "The Pneumatic Transport of Cracking Catalyst in
Vertical Risers," Chem. Eng. Sci., 17. 599 (1962).
16. Richardson, J. F. and Zaki, W. N., "Sedimentation and Fluidization:
Part I," Trans. Instn. Chem. Engrs., 3_2, 35 (1954).
17. Wen, C. Y. and Yu, Y. H., "Mechanics of Fluidization," CEP Symposium
Series 62, 62 (1966).
18. Davidson, J. F., and Harrison, D., Fluidization, Academic Press (1971).
19. Hinkle, B. L., Ph.D. Thesis, Georgia Institute of Technology, Atlanta,
Georgia (June 1953).
20. Zenz, F. A. and Weil, N. A., "A Theoretical-Empirical Approach to
the Mechanism of Particle Entrainment from Fluidized Beds," AIChE
Journal, j4 (4), 472 (1958).
21. Leung, L. S., Wiles, R. J., and Nicklin, D. J., "Correlation for
Predicting Choking Flowrates in Vertical Pneumatic Conveying,"
Ind. Eng. Chem. Process Des. Develop., 10 (2), 183 (1971).
22. Zenz, F. A. and Othmer, D. F., Fluidization and Fluid Particle Systems,
New York: Reinhold Publishing Co., (1960).
P-60
-------
NOMENCLATURE
2
A = area of orifice opening, ft
A = specific surface area of the solid particles and can be
expressed as A = 6/ d , ft
2
A, = cross-sectional area of the downcomer, ft
Q
2
A = cross-sectional area of the draft tube, ft
A = cross-sectional area occupied by the serpentine tubes
C = solid discharge coefficient defined in Equation 40
C = drag coefficient on solid particles
d = orifice diameter, ft
o
d = mean particle'size, ft
d . = particle diameter of size i, ft
D = draft tube diameter, ft
D. = diameter of the jet, ft
f = Fanning friction factor
f = solid friction factor from Equation 28
P
f =' solid friction factor from Equation 13
s
F = air by-passing factor
F = correction factor for downcomer pressure drop
2
g = gravitational acceleration, ft/sec
2
G = solid mass flow rate, Ib/ft -sec
H = length in downcomer occupied by the serpentine tubes
H = distance between the two pressure taps in the downcomer
H , = bed height at minimum fluidizing condition, ft
P-61
-------
J = jet penetration depth, ft
L = height of the draft tube, ft
AL = distance required to accelerate a spherical particle
to a specific velocity, ft
n = empirical constant defined in Equation 19
2
AP, _« = pressure drop between 1 and 2 (see Figure P-l) , Ib/f t
2
(APj_2) = calculated pressure drop between 1 and 2, Ib/ft
2
(AP ) = experimental pressure drop between 1 and 2, Ib/ft
«L— fc e icp
2
AP._, = pressure drop between 1 and 4 (see Figure P-l) , Ib/ft
2
AP = pressure drop in the downcomer, Ib/ft
2
(AP_.) , = calculated pressure drop in the downcomer s, Ib/ft
2
(AP ) = experimental pressure drop in the downcomers, Ib/ft
2
AP = pressure drop due to gas-pipe wall friction, Ib/ft
AP,. = pressure drop due to solid-solid friction and solid-pipe
wall friction, Ib/ft2
2
AP = pressure drop due to the static head of the gas, Ib/ft
AP = pressure drop due to static head of the solid particles,
us
(AP. ) = corrected pressure drop due to static head of the solid
ns c -
particles, Ib/ft
AP = pressure drop due to kinetic energy change of the gas,
8 2
Ib/ft
AP = pressure drop due to kinetic energy change of the solid
ks _
particles, Ib/ft
2
Ap = pressure drop in the draft tube, Ib/ft
K
2
(Ap ) = corrected pressure drop in the draft tube, Ib/ft
Re = p,U D/JJ, the Reynolds number
P-62
-------
Re = Reynolds number based on the terminal velocity of the
solid particles
U = superficial fluid velocity, ft/sec
o
n = actual fluid velocity in the bed; Uf = UQ/e, ft/sec
IT = actual fluid velocity in the downcomers, ft/sec
f ,d
U, = actual fluid velocity in the draft tube, ft/sec
f »r
U = solid particle downward velocity in the downcomers, ft/sec
p,d
U = solid particle velocity in the draft tube, ft/sec
P»r
(U ,)T, = solid particle downward velocities in two downcomers,
p,d L
NpA ft/sec
U = slip velocity between fluid and solid particle in the
S-L
draft tube, ft/sec
U = terminal velocity of the solid particle, ft/sec
W = solid flow rate, Ib/sec
s
2 3
X = empirical exponent; X = 1.0 when A < 6300 ft /ft
X. = weight fraction of particles of size i
e = voidage of the fluid bed above the draft tube
e, = voidage in the downcomers
d
e = voidage in the draft tube
y = viscosity of the fluid, Ib/sec-ft
p = density of the fluid, Ib/ft
3
p = solid particle density, Ib/ft
S
n = correction factor for nonspherical particles (Equation 23)
6 = specific solid loading, Ib solids/lb fluid
= sphericity of the solid particle
P-63
-------
BIBLIOGRAPHIC DATA
SHEET
1. Report No.
EPA-650/2-73-048a thru -d
3. Recipient's Accession No.
4. Title and Subtitle
Pressurized Fluidized-Bed Combustion Process
Development and Evaluation, Volumes I, n , m, and IV
(Vol n, Appendices; Vol m, Boiler Development Plant Design;
5. Report Date
December 1973
6.
\Tn1 TV
•m /TVieii1fiii«ii7«»fir>T^
7.
D. L
Keairns. D. H. Archer et al.
8. Performing Organization Rept.
No.
9. Performing Organization Name and Address
Westinghouse Research Laboratories
Pittsburgh, Pennsylvania 15235
10. Project/Taik/Work Unit Nu.
ROAP 21ADB-09
11. Contract/Grant No.
68-02-0217
12. Sponsoring Organization Name and Address
EPA, Office of Research and Development
NERC-RTP, Control Systems Laboratory
Research Triangle Park, North Carolina 27711
13. Type of Report & Period
Covered
Final
14.
15. Supplementary Notes
16. Abstracts Tne report presents: results of a process evaluation of the pressurized
fluidized-bed combustion (FBC) system for power generation; preliminary plans and j
a cost estimate for a 30-MW pressurized FBC boiler development plant; identification;
of a project team and program to demonstrate FB oil gasificatioh/desulfurization for
power generation on a 50-MW plant; and evaluation of pressurized oil gasification for
combined-cycle power generation. It identifies no problems which preclude the
development of pressurized FBC combined-cycle power plants and FB oil gasification
power plants which can generate electrical energy within environmental goals at lower
energy costs than competitive systems. Work reported here, a continuation of earlier
FBC process evaluation efforts, is aimed at the development and demonstration of
these FB fuel processing systems.
17. Kej Words and Document Analysis 17o. IV-icnpu>rs
Air Pollution
Combustion
Gasification
Fluidized Bed Processing
Desulfurization
Oils
Fossil Fuels
Wastes
Electric Power Generation
17b. [dentificrs/Open-Ended Terms
Air Pollution Control
Stationary Sources
Fluidized-Bed Combustion
17e. COSATI Field/Group
21B
IB. Availability Statement
Unlimited
19. Security C\a;s (This
Report)
UNCLASSIFIED
20. Security Class (This
Page.
UNCLASSIFIhlD
21. No. of Pages
258
22. Pncr
FORM NTIS-33 (REV. 3-72)
P-64
USCOMM-DC M8S2-P72
------- |