PB 200 881
REMOVAL OF SO2 FROM POWER PLANT FLUE
GASES USING THE ALKALIZED ALUMINA PROCESS
PHASES I & II FINAL REPORT DATA EVALUATION
PRELIMINARY PROCESS DESIGN, OPTIMIZATION,
AND ECONOMICS
Kellogg Company
Piscataway, New Jersey
28 February 1970
DISTRIBUTED BY:
National Technical Information Service
U. S. DEPARTMENT OF COMMERCE
5285 Port Royal Road, Springfield Va. 22151
-------
THE M.W.
CpMPAN Y
U^pce^ii^iffd^ (v -^ '•"•"
PB 200 881
OaFROM
NATIONAL AIR PQUUH$N
-•-- •,-• v '>''•,«>', :'f... .-..:''•'• ,.'••'•-'..,. :• '• \, -i1'^ ' '"v.',". -.-• '• r" '•
, CONT|ACT NO! PH sf 68-8
-------
NOTICE
THIS DOCUMENT HAS BEEN REPRODUCED FROM THE
BEST COPY FURNISHED US BY THE SPONSORING
AGENCY. ALTHOUGH IT IS RECOGNIZED THAT CER-
TAIN PORTIONS ARE ILLEGIBLE, IT IS BEING RE-
LEASED IN THE INTEREST OF MAKING AVAILABLE
AS MUCH INFORMATION AS POSSIBLE.
-------
4. Title and Subtitle
Removal of SO? From Power Plant Flue Gases Using The Alkalized
Alumina Process - Phases I & II - Data Evaluation, Pre-
liminary Process Design, Optimization, and Economics
Research & Engineering Development
Piscataway, N. J.
12. Sponsoring Agency Nairn and Address
Department of Health, Education and Welfare
Public Health Service
National Air Pollution Control Administration
Washington, D. C. 20201
STANDARD TITLE PAGE
FOR TECHNICAL IMPORTS
• * nMOft No.
APTD-0688
V. Authorls)
ress
15. Supplementary Notes
57 Report L _.
February 28, 1970,
6."
lanlzation Code
Organization Rept. No
10. Project/Task/Work Unit No.
IT. Contwrt?5fSnrwoT *
PH 86-68-M
13. Type of Repert & Period Covered
Final
The'results of a contractual effort to correlate and evaluate data from several sources
on the development of the alkalized.alumina sorbent and process are reported in two
phases: P ase I - data evaluation and preliminary process design, and Phase II *,pro-
cess optimization studies. This report includes an analysis of available .attrition dat
and a summary of the attrition test methods used. The process design study includes an
evaluation of; sorption chemistry, current sorption data, and sorption modeling. Re-
generation reactions..and evaluation of the status of the daca o
-------
DISCLAIMER
This report was furnished to the Air Pollution Control
Office by
The M. W. Kellogg Company
Research & Engineering Development
Piscataway, N. J.
in fulfillment of Contract No. PH 86-68-86
-------
RESEARCH AND ENGINEERING DEVELOPMENT
REMOVAL OF SO2 FROM POWER PLANT
FLUE GASES USING THE ALKALIZED
ALUMINA PROCESS
PHASES I & II FINAL REPORT
DATA EVALUATION, PRELIMINARY PROCESS DESIGN;
OPTIMIZATION, AND ECONOMICS
Submitted to
DEPARTMENT OF HEALTH, EDUCATION, AND WELFARE
PUBLIC HEALTH SERVICE
NATIONAL AIR POLLUTION CONTROL ADMINISTRATION
Contract No. PH 86-68-86
APPROVED:
7 ~7 Ju
/^Wjy^^-x^^fe^*-^
Project Dire/tor
rch jL»€8x^OJj July, 1969)
Pro5SCtTT)irector
-------
R ft ED 20 LAB. 3-6S
THE M. W. KELLOGG COMPANY
A DIVISKM OF PULLMAN INCORPORATED
Research * Engineering Development
Page No. .
Report No.
70-1233
REMOVAL OF SO, FROM POWER PLANT
FLUE GASES USING THE ALKALIZED
ALUMINA PROCESS
PHASES I & II FINAL REPORT
DATA EVALUATION, PRELIMINARY PROCESS DESIGN,
OPTIMIZATION, AND ECONOMICS
FEBRUARY 28, 1970
Staff: M. C. Cambon, G. M. Drissel, R. M. Gialella, D.M. Masi, Jr.,
J. J. O'Donnell, C. J. Royce, A. G. Sliger, J. D. Turner
Period Covered: March 1968 to February 1970
1. 0. No.
6045
Distribution
NAPCA
J. B. Dwyer
B. wi Jesser
R. E. Vener
M. J. Wall
G. T. Skaperdas
W. C. Schreiner
A. G. Sliger
G. M. Drissel
J. J. O'Donnell
S. E. Handman/D. M. Masi, Jr.
J. C. Quinlan
P. T. Atteridg
R. I. D. (3)
AUTHORS:
Copy No
1-125
126
127
128
129
130
131
132
133
134
135
136
137
138-140
-------
REMOVAL OF S02 FROM POWER PLANT
FLUE GASES USING THE ALKALIZED
ALUMI1& PROCESS
PHASES I & II FINAL REPORT
DATA EVALUATION, PRELIMINARY PROCESS 'DESIGN j
OPTIMIZATION, AND ECONOMICS
Submitted to
DEPARTMENT OF HEALTH, EDUCATION, AND WELFARE
PUBLIC HEALTH SERVICE
NATIONAL AIR POLLUTION CONTROL ADMINISTRATION
Contract No. PH 86-68-86
by
THE M. W. KELLOGG COMPANY
RESEARCH & ENGINEERING DEVELOPMENT
PISCATAWAY, N. J.
February 28, 1970
-------
TABLE OF CONTENTS
Page No.,
PREFACE 1
INTRODUCTION 3
SUMMARY 5
CONCLUSIONS AND RECOMMENDATIONS 18
DISCUSSION 22
A. Status of Data 22
1. Attrition 22
a. Various Procedures for Generating 22
Attrition Data
(1) USBM Pneumatic Conveyor 22
Attriter
(2) Tyler Sieve Shakers 23
(3) USBM-Albany, Fluid Bed and Air 23
Lift Attrition Test Unit
(4) W. R. Grace Accelerated Air Jet 31
Attrition Test Unit
(.5) Standard of Indiana-Cracking 31
Catalyst Test Procedure
b. Distribution of Attrition Losses 32
c. Comparison of Experimental Data on 36
Common Basis
(1) Albany Data Converted to Circulating 37
Solids Basis
(2) Bruceton Data Converted to Inventory 38
Basis
d. Recommended Method of Attrition Testing 39
e. Methods to Improve Attrition Resistance 45
of Alkalized Alumina
(1) Prediction of Pilot Plant Performance 45
from Laboratory Testing
(2) Impregnation on UOP Catalyst Support 49
(3) Binder Strengthened Alkalized Alumina 49
(a) USBM Fillers and Mortar Types 49
(b) W. R. Grace Studies 51
(cj NaAlO- Pellets Prepared from 51
BSAC intermediates
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Pagej'Ho,
2. Sorption
55
a. Sorbent Selection
b. Sorption Chemistry
(1) Sorbent Preparation
(2) Sorption Reaction
59
(3.) Thermochemistry
59
(4) Sorption Kinetics
(5) Extraneous Chemical Effects 7o
(a) Chlorides Jj
(b)A Water 'I
(c) Oxygen,, L*
(d) Oxides of Nitrogen. '•*
c. Sorption Data '"
(1) Sources 76
(2) USBM -Albany 76
(3) USBM-Bruceton 89
(4) AVCO
(5) W. R. Grace 12*
(6) CEGB (Central Electricity Generating 126
Board )
d. Sorption Modeling 128
(1) -Kinetic Modeling 128
(2) Fixed Bed Modeling 1^3
(3) Fluid Bed Modeling |33
(4) Dispersed Phase Modeling 141
3. Regeneration I 46
a. Reactions 146
b. Status of Data 147
(1) . USBM Studies 147
(a) 55,000 CFH Pilot Plant Results 147
1 - Fixed Bed Regeneration 147
2 - Moving Bed Regeneration '147
11
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Page NOj
(b) Bench Scale Tests with Various 150
Reducing Gases
(c) Regenerations with>H2 and CO 153
1 - Bench-Scale Results 153
2 - Regeneration of a Composite 155
Sample from the Pilot Plant
3 - Half-Hour Regeneration Tests 156
(d) Albany Batch Regeneration Studies 158
1 - Data from Modified Thin-Bed 158
Test Unit
2 - Comparison of CEGB-Blyth 159
Regeneration Model with USBM-Albany
Data
(e) Evaluation of Various Sorbents 164
1 - Kaiser DN-112-F, Kaiser B, ^64
Kaiser DN-105, Peter Spence
2 - Grace No. 2, Kaiser PR-IS, 166
Kaiser DN-113-F
3 - USBM 423-X-4E, USBM X-3-F 166
4 - Grace No. 1 166
5 - Two-Step Regeneration Reactions 167
with Sorbents 1-4
6 - Bureau Prepared NaOH and Alumina 168
Hydrate Sorbents
7 - UOP Alumina Pellets Impregnated 169
with Active Compounds
8 - Kaiser PR-18-1 Alkalized Alumina 169
(15% Na and 0.6% Fe)
9 - Grace No. 1 and Grace No. 2 170
Comparison
(2) AVCO Studies 170
(3) W. R. Grace 172
B. Preliminary Process Design 174
1. Pilot Plant Studies 174
2. Commercial Scale Studies 177
a. Dispersed Phase 177
(1) Dispersed Phase Process Description: 177
Flow Sheet P 3197-D
(2) Design Notes - Dispersed Phase
iii
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Page N
183
(a) Preliminary Calculations _
1 - Effect of .Gas Velocity and
Recycle on Residence Time in
Absorber 188
2 - Effect of Sorbent Density on
Sulfur Removal 189
3 - Regenerator
4 - Solids Heater 19S
(b) Vessel Design
209
1 - Absorber 209
2. - Solids Heater **"
3,.- Regenerator
4 - Solids Transfer Lines
5 - Plant Layout f
6 - ^Vessel Analytical Sketches ZZJ
b. Fluid Bed 2zl
(1) Fluid Bed Process Description: Flow Sheet 221
P 3198-D
(2) Design Notes - Fluid Bed 22&
(a) Preliminary Process Calculations 227
1 - Minimum Fluidization Velocity 227
2 - Fluid Bed Sorber Gas Velocity 229
3 - Comparison of Bed Heights 231
Required for Fluid Bed Sorption
USBM-Albany vs. CEGB, Blyth
Models
(b) Vessel Design 235
1 - Sorber 235
2 - Regenerator 237
3 - Fluid Bed Heater 238
4 - Process and Mechanical Considera- 239
tions - Regenerator Train
5 - Plant Layout 241
6 - Solids Transport System 242
7 - Gas-Solids Separator 245
8 - Air System 245
c. Fixed Bed 247
(1) Fixed Bed Process Description: Flow 24?
Sheet P 3196-D
(2) Design Notes - Fixed Bed ?$7
IV
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(a) Process
1 - Selection of Base Case Design
for Estimate
2 - Safety Hazards
3 - Heating and Cooling Gases
4 - Insulation
(b) Mechanical
1 - Sorber Units
2 - Sorber Traiiis
3 - Gas Ducts
4 - Valves
5 - Flue Gas Fans
6 - Air Preheater Ductwork
d. Comparison of Processes
e. Reformer, Glaus Plant, and Precipitator
(1) Regeneration Gas Plant
(2) Claus Gas Pre-Treatment
(3) Claus Plant and Incinerator
(4) Electrostatic Precipitator
(5) Power Plant Air Preheater
f. Final Sorber Design
C. Investment
D. Economics
1. Base Case
a. Gross Costs
b. Net Costs
2. Final Process Design
APPENDICIES
A. Attrition Test Methods
B. Effect of Oxygen and Water Vapor on S02 Sorption
C. Nitrogen Oxide Chemistry
D. Fixed Bed Model
E. Dispersed Phase Model
F. Regeneration Data
Page No.
257
257
266
269
270
271
271
276
111
277
279
279
279
283
283
286
289
290
291
292
297
302
302
302
312
316
A321
B346
C354
D364
E367
F376
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Page NO
G. Process Calculations and Dispersed Phase Vessel G404
Sketches
H. Composition of Coal and Flue Gases M38
I. Fluid Bed Process Vessel Sketches I44S
J. Fixed Bed Sorber Designs J4SS
K. Sample Calculations - Heat Transfer of Alkalized Alumina K422
in Gas Stream
VI
-------
LIST OF TABLES
Table No. Title Page No,
1 Fluid-Bed Attrition Test Unit Data 27
2 Comparison of Sorbent Attrition Losses for 28
Air Lift and Fluid Bed Systems
3 Distribution of Attrition Losses 33
4 Variation in Attrition Resistance of RL-175 52
Composition Alkalized Alumina (Dawsonite Bound
with 23% Kaolin and 2% NASi03)
5 X-ray Diffraction Results on Loaded Grace 58
Alkalized Alumina
6 Heat of Formation Data 60
7 X-ray Diffraction Results for Grace Alkalized 69
Alumina Loaded in the Presence of Nitrogen Oxides
8 Effect of Water Vapor Concentration on S02 72
Absorption Rate
9 Comparison of Sorption Rates With and Without 86
NOX
10 Effect of NO- and NO on SO, Sorption Rate - 87
Grace Alkalized Alumina
11 Fluidized Bed (Sorption System) Tests of One 90
Day Duration
12 Evaluation of Fluid Bad Tests of One Day Duration, 91
for Steady State
13 Fluidized Bed (Sorption System) Test for 100 Hours 92
14 Evaluation of Fluid Bed Test for 100 Hours, 93
Steady State Operations
vii
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Table No. Title _____ S*S*I
15 Operating Conditions in Absorbing Sulfur DJOJide 10]
Under Countercurrent Gas-Solids Flow Compared
with Solids Entrainment With Solids Recyflla
16 Operating Conditions During Absorption
17 Operating Conditions During Sorption
18 Sorption Operating Conditions for Evaluation of
Grace No. 2 and Kaiser No. 1 Sorbents
19 Sorbent Loading Test Data 1
20 Summary Data of Initial Rate Studies 1
21 SO4 Determination of Sorbent Exposed to Flue 1
Gas Containing NOX After Attrition
22 Determination of Multiple Fixed Bed Parameters 1
Based on U. S. Bureau of Mines Fixed Bed Model
23 Sample Calculation of Fluid Bed Sorbent Inventory 1
Using the USBM Albany Model
24 Calculation of Sorbent Bed Height Using the CEGB
Blyth Model
25 A Comparison of the Bed Heights Required foir Fluid
Bed Sorption as Predicted by the USBM, Albany and
CEGB-Blyth Models
26 Various Regeneration Reactions 1
27 Heats of Formation at 25°C 1
28 Operating Conditions During Continuous Regeneration 1
29 Reduction of Spent Alkalized Alumina (2l4Naj With 1
Reducing Gases
30 Battery Limits Estimates and Total Capital Require- 1
ments (50 MW Plant)
31 Calculated Solids Residence Time in Bruceton Realtor 1
32 Moving Bed Regenerator Design - 1000 MW Plant 1
VIXi
-------
Table No. Title page No.
33 Effect of the Heat of Reaction on the Regenerator 196
Heat Balance
34 Comparison of Various Designs for Heating and 213
Cooling Solids in Regeneration Section
35 Identification of Lines on Layout Sketches 22°
Fig. 52 & 53 Dispersed Phase Sorption - Moving
Bed Regeneration Process Scheme
36 Minimum Fluidization Velocities at 600°F and 1 psig 228
37 Sorbent Inventory and Bed Height Versus Exit
Sorbent Loading
38 Sorber Operations Schedule 2->3
39 Heating Cycle Material Balance 254
40 Cooling Cycle Material Balance 256
41 Number of Parallel Fixed Beds Required for 90% 259
Sorption of S(>2
42 Effects of Sorption Time and Pressure Drop on 263
Fixed Bed Parameters
43 Effect of Sorption Time and Pressure Drop on 264
Sorber Dimensions
44 Major Process Factors - Alkalized Alumina 280
45 Relative Comparison of Fixed, Fluid and Dispersed 281
Phase Systems
46 Comparison of Calculated Temperature Equilibrium 296
Times, Dispersed Phase System
47 Approximate Distribution of Process Investment- 2^
Battery Limits Plant
48 Annual Costs - Base Case Plant (Excluding Attrition 303
Loss and Sulfur Credit)
49 Dispersed Phase - Process Economics 304
50 Fluid Bed - Process Economics 305
IX
-------
Table No. Title ._
*
306
51 Fixed Bed - Process Economics
52 Process Economics (Including Attrition Loss and
Sulfur Credit)
53 Total Investment Decrease Due to Deletion of the 317
Solids Cooler (Dispersed Phase and Fluid Bed
Schemes)
54 Change in Process Economics Due to Deletion 6f 318
Solids Cooler - Dispersed Phase (Excluding Attri-
tion Loss and Sulfur Credit)
55 Change in Process Economics Due to Deletion of 319
Solids Cooler - Fluid Bed (Excluding Attrition
Loss and Sulfur Credit)
A-56 Alkalized Alumina Attrition Loss Test Data A326
A-57 Attrition Resistance Data Sheet A331
A-58 Sample Worksheet Used in the Determination of &334
Attrition By the One Hour Method
A-59 Effect of Attrition on Particle Size Distribution A335
A-60 Effect of Catalyst Fines Content on Attrition A340
A-61 Precision of Attrition Test Data A341
A-62 Attrition of Fresh Ground Cracking A342
' Catalysts
A-63 Attrition of Fresh Microspheroidal Gracing A343
Catalysts
A-64 Attrition of Fresh Catalysts After Heatind at A343
1100°F for 4 Hours ' *
A--65 Summary of Sorption Data A347
A-66 Maximum Pore Radius Capable of Supporting KJ52
Capillary Condensation of Moisture
R"f'7 Calculated Values of Sorption Rate and Section E375
Rate Constant
-------
Table No. Title Page No.
F-68 Rat6|Of Consumption of Reducing Gas p
F-69 Regeneration of Grace No. 1 Spent Sorbent From F378
Pilot Plant Mixed Particle Sizes, 5.3%S
F-70 Half Hour Regeneration Tests - Original Grace
No. 1 Sorbent
F-7.1 Regeneration of Grace No. 1 Sorbent - Modified
Thin Bed Test Unit
F381
F-72 Test Data for Regeneration of Spent Alkalized
Alumina
F382
F-73 " Test Data for Two Step Regeneration of Spent
Alkalized Alumina
F383
F-74 Test Data for Regeneration of Sorbents Used
for SO2 Removal
F-75 Regeneration of Grace No.l and Grace No. 2 F384
Saturated With S02
F-76 Effect of H2 Partial Pressure on Regeneration F385
Time
F-77 Effect of Sorbent Loading on Regeneration Time F386
F-78 Effect of Carbon Monoxide on Regeneration Time F387
F-79 Regeneration Data for SO- Sorption F388
F-80 Effect of Temperature on Regeneration Time F389
F-81 Two Step Regeneration Summary-Grace No. 2 Sorbent F390
G-82 Regenerator Material Balance Calculations G405
G-83 Regeneration Reaction G407
G-84 Sample Calculation - Solids Heater Diameter G409
G-85 Sample Calculation - Solids Heater Bed Height ^410
G-86 Sample Calculation - Solids Heater Bed Height G413
G-87 Sample Calculation - Temperature History Curve G417
Fixed Beds Heating of Solids
XI
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Table No,
Title Page No:
G-88 Sample Calculation - Pressure Drop Thru Moving &421
Bed Solids Heater
G-89 Dispersed Phase Absorber Material Balance G423
G-90 Expression for Dispersed Phase Absorber Height -*425
; - *
G-91 Sample Calculation - Dispersed Phase Absorber rt426
Dimensions
G-92 Sample Calculation - Solids Heater
G-93 First Stage Temperature Profile for 3 Stagd
Fluid-Bed Solids Heater
H-94 Typical Coal Analyses H440
H-95 Sample Calculation for a 1% Sulfur ^Coal H441
H-96 Flue Gas Properties for Various Sulfur H444
Bearing Coals
1-97 Sample Calculation of the Minimum FLuidization H446
Velocity by the Method of Leva* et,al. ,.
J-98 Fixed Bed Sorber Dimensions J456
J-99 List of Vendors Contacted for Large Valves J457
Xll
-------
LISTS OP FIGURES
FIGURE NO. TITLE PAGE NO,
1 Particle Size Distribution of Activated 25
Grace Alkalized Alumina(1.63% Sulfur on
Catalyst)
2 Particle Size Distribution of Activated 26
Grace Alkalized Alumina(8.85% Sulfur on
Catalyst)
3 Particle Size Distribution of Grace #1 39
(Standard) Alkalized Alumina (Fixed Bed
Test Unit)
4 Particle Size Change During Absorption- 34
Regeneration Cycle
5 Particle Size of Solids Discharged From 35
Absorber
6 Particle Size Distribution of Grace #2 41
Alkalized Alumina (Before and After Attri-
tion Test)
7 Particle Size Distribution of Grace #2 42
and Peter Spence Alkalized Alumina (Before
and After Attrition Test)
8 Particle Size Distribution of Grace #2 44
and Peter Spence Activated Alkalized Alumina
(After 1 hr Attrition Test)
9 Variation in Particle Size Distribution With 4-7
Number of Operating Cycles Using a Spherical
Alumina Catalyst Support
10 Comparison of Laboratory Attrition Loss as 43
Measured by the USBM and NAPCA Methods. (Grace
#2 and Kaiser #1 Alkalized Alumina)
11 Relationship of Laboratory and Pilot Plant 50
Attrition Loss Measured by the USBM and NAPCA
Methods (Grace #2 and Kaiser #1 Alkalized Alumina)
12 Comparison of Attrition Loss of Various BSAC 54
Blends
13 Comparison of Predicted and Experimental Sulfur 62
Dioxide Removal from Effluent Gas for Surface
Reaction Controlling and Diffusion Controlling
(2" Packed Bed of Alkalized Alumina)
xiii
-------
FIGURE NO,
TITLE PAGE JJQ.
14 Comparison of Predicted and Experimental Sulfur 63
Dioxide Removal from Effluent Gas for Surface
Reaction Controlling and Diffusion Controlling
(3" Packed Bed of Alkalized Alumina)
15 DITTO 64
(4" Packed Bed of Alkalized Alumina)
16 The Effect of Oxides of Nitrogen on the 67
Sorption Rate of Sulfur Dioxide by Alkalized
Alumina
17 The Effect of Oxygen Concentration on the 74
Sorption Rate of Sulfur Dioxide by Alkalized
Alumina
18
Laboratory Sulfur Dioxide Sorption System (USBM) ''
19 Average Sorption Rate of Sulfur Dioxide by '
Grace Alkalized Alumina
20 Effect of Temperature on the Sorpticn Rate 80
of Sulfur Dioxide by Alkalized Alumina
21 Effect of Sulfur Dioxide Concentration on 81
Sorption Rate by Alkalized Alumina
22 Effect of Activation Temperature on the Sorption 83
Rate of Sulfur Dioxide by Grace Alkalized Alumina
23 Effect of Activation Time on the Sorption Rate 84
of Sulfur Dioxide by Grace Alkalized Alumina
24 Fluid Bed Sulfur Dioxide Sorption System (OSBM) 88
)
25 Pilot Plant Facilities for Flue Gas Generation 9S
and S02 Sorption of the Alkalized Alumina JPto-
cess
26 Pilot Plant Absorber (USBM-Bruceton) 96
27 Effect of Alkalized Alumina Particle Size *8
on Percent Sulfur Dioxide Removal
28 Effect of Gas Velocity on Percent Sulfur "
Dioxide Removal
xiv
-------
FIGURE NO. TITLE PAGE NO.
29 Effect of Residence Time (Baffled Absorber) 101
on Percent Sulfur Dioxide Removal
30 Sorber With Solids Recycle System (USBM- 104
Bruceton)
31 Dispersed Phase Sulfur Dioxide Removal pilot 1Q5
Plant (USBM-Bruceton)
32 Average Sulfur Dioxide Concentration of Gas 112
Versus Sorber Height
33 Sorber Bottoms Solids Recycle System (USBM- 12°
Bruceton Pilot Plant)
34 Sulfur Dioxide Sorption Rate Expression 121
Versus Sorbent Loading
35 Comparison of USBM Albany and AVCO Sulfur 123
Dioxide Sorption Rates on Alkalized Alumina
With Oxides of Nitrogen Present
36 Sorbent Loading Versus Time and Sulfur Dioxide ^2^
Concentration
37 Comparison of Sorption Rate Equations (Sorbent ^^
Loading Versus Concentration X Time)
142
38 Comparison of Calculated and Measured Sulfur
Dioxide Sorption by Alkalized Alumina in a
Fluid Bed System
'*.
39 ' Dispersed Phase System Sorption Equation. 144
Ratio of Sorption Rate to Sulfur Dioxide Com-
position of Flue Gas Versus Sulfur Loading (Test
Series C&D)
40 DITTO - (Test Series E&F) 145
41 Arrhenius Plots for t^ and CO Regeneration of 154
Grace II Sorbent (Rate Constant Versus Recipro-
cal Absolute Temperature)
42 Effect of Sorbent Loading on 100% Regeneration ^60
Time (Three Temp. Levels)
43 Arrhenius Plot for Reformer Gas Regeneration 161
of Grace fl Sorbent. (Rate Constant Versus
Reciprocal Absolute Temperature)
44
Time)
Comparison of CEGB-Blyth and USBM-Albany 163
Regeneration D^ta (Sorbent Loading Versus
xv
-------
FIGJRE NO,
TITLE PAGE! NO.
45 Comparison of Hydrogen Consumption Rate of TWO
Alkalized Aluminas With and Without Iron*
(Kaiser DN-112-F and Kaiser B)
46 Calculated Settling Velocity Versus Particle 184
Size for Alkalized Alumina in Dispersed Phase
Sorber
f
47 Comparison of Calculated and Measured Solids 19C
Density in Dispersed Phase Sorber
48 Relationship of Solids Heater Diameter and 200
Bed Height to Gas Inlet Pressure
49 Temperature - History Curves for Solids tittatejf 206
A Art
50 Solids Temperature Profile Versus Bed Height *UD
in Solids Heater
51 Fluid Bed Solids Heater Gas Temperature
Profile (First Stage)
52 Equipment Layout-Dispersed Phase 218
53 Dispersed Phase Plot Plan 219
54 Fluid Bed Sorption Equipment - 243
(Plan View)
244
55 Fluid Bed Sorption Equipment Layout
(Elevation View)
278
56 Fixed Bed Sorption Equipment Layout
(Plan & Elevation Views)
57 Regeneration Gas Plant — F/S Dwg. P1374-B 284
58 Flow sheet for Fixed Bed System-Claus Gad 288
Pretreatment - FS Dwg. P 1375-B
A-59 System for Pneumatic Conveying of Alkalized A322
Alumina (Sketch)
A-60 Accelerated Air Jet Attrition Apparatus A327
(Davison Chem. Co. Sketch)
A~61 Apparatus for Determination of Attrition &330
(Sketch)
A-S2 Attrition Resistance Test Results, A332
Weight % Over 20 Microns Versus Time
xv i
-------
FIGURE NO. TITLE PAGE NO.
A-O Attrition Resistance Test Results. Weight A332
% Attrition Versus Time
A-64 Apparatus for the Determination of Attrition A333
by the One Hour Method (Sketch)
A-65 USRM-Albany Attrition Equipment (Sketch) A338
A-66 Standard Oil Co. (Ind.) Attrition Apparatus A340
A-67 Fines Content During One Hour Attrition Test A341
(% Minus 325 Mesh Versus Time)
A-63 Fines Content During Successive One Hour A342
Attrition Test Periods .(% Minus 325 Mesh
Produced Versus Time)
A-69 Decrease in Attrition Rate of Ground Catalysts A344
During Extended Operating Period (Minus 325
Mesh Increase Versus Time) (Commercial Unit)
A-70 DITTO - (Pilot Plant) A344
B-71 Effect of Water & Oxygen in Flue Gas on B349
Sorption Rates (Data for Sample No. 6)
B-72 DITTO B350
B-73 DITTO B351
B-74 DITTO B353
E-75 Comparison of Measured and Predicted Sorber E369
Pressure Drop
F"?7?
E-76 Variation in Sulfur Oxides Concentration With r.j/*
Sorber Height
P "3 7 7
E-77 Sorption Rate Equation (Change* in Sulfur Oxide &.»/*
Concentration Per Unit Sorber Height Versus
Sulfur Oxide Concentration)
E-73 Sorption Rate Equation (Sorbent Loading Versus E373
Concentration)
F-79 Effect of Time and Temperature on Sulfur Removed F391
(Grace #1, 8.5%S)
XVll
-------
FIGURE NO,
TITLE
PAGE NO.
F-30
F-81
F-82
F-83
F-84
F-85
F-86
F-87
F-83
F-89
F-90
F-91
G-92
G-93
G-94
G-95
G-96
G-97
DITTO - (Grace II, 5.5% S)
DITTO - (Grace II, 2.3% S)
Percent Regenerated Versus Time (Grace II >
2.3% S)
DITTO - (Grace II, 5.5% S)
DITTO - (Grace II, 8.5% S)
Comparison of Hydrogen Consumption Rate
of Alkalized Aluminas Containing Iron
Effect of Temperature on Hydrogen Con-
sumption of Sorbent 423X-4E
Effect of Temperature on Regeneration
Gas Utilization for Grace #1 Sorbent
Effect of Temperature on Hydrogen Con-
sumption of LJOP-2 Sorbent Regeneration
Effect of Temperature on Hydrogen Con-
sumption of Kaiser PR-18-1 Regeneration
Effect of Reducing Gas Composition on Gas
Consumption for Regeneration of Kaiser
PR-18-1
Effect of Sorbent Loading on Regeneration
Time
Effect of Bed Height on Solids Heater
Temperature Profile
Vessel.Sketch - Dispersed Phase
Vessel Sketch - Fluid 3ed Solids Heater D-2
Vessel Sketch - Moving Bed Regenerator D~3
Vessel Sketch - Sorbent Make-up Hopper F-2
Vessel Sketch - Fines Recovery Cyclones
L-1A-1F J J
Vessel Sketch - Gas - Solids Separator L~2
F392
F393
F394
F395
F396
F397
F398
F399
F460
F401
F402
F403
G420
G432
3433
(3434
d435
(5436
G437
XVlll
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FIGURE rtO.
TITLE
PAGE NO.
1-99
1-100
1-101
1-102
J-103
J-104
J-105
J-106
J-107
J-108
J-109
J-110
J-lll
J-112
J-113
J-114
J-115
J-116
Vessel Sketch Dwg. AA-FB-3-A1 - Fluid-
Bed Sorber D-l
Vessel Sketches Dwg. AA-FB-5-A1, A2 -
Regenerator D-2
Vessel Sketch-Fluid Bed Solids Heater D-3
Vessel Sketch-Gas-Solids Separator L-l
Sketch - Countercurrent Gas Solids
Contacting Panel
Sketch - Dorfan Impingo Filter
Sketch - Cross Section View of Sorber Flow
Pattern
Sketch - Sorber Panel
Sketch - Sorber Front View
Sketch - Sorber Front View (Detailed)
Sketch - Sorber Plan View
Sketch - Top & Bottom Sorber Panel Cover
Detail
Sketch - Sorber Panel Frame Dimensions
Sketch - Sorber Panel Support Detail
Sketch - Sorber Screening Support Detail
Sketch - Sorber External Framing - Flat
Panel Design
Sketch - Sorber Cylindrical or Annular Design
Sketch - Sorbent Train - General Cross
Section
T447
1451
1453
1454
J458
J459
J460
J461
J462
J463
J464
J465
J466
J467
J468
J469
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J471
xix
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PREFACE
The work described in this report was performed under Contract
No. pH 86-68-86, Department of Health, Education, and Welfare, Public
Health Service, Consumer Protection and Environmental Health Service,
National Air Pollution Control Administration (NAPCA).
As originally scoped the contract specified three main phases,
viz., 1) preliminary studies; 2) process optimization studies; and
3) a conceptual design. This report covers the phases 1 and 2 work.
Phase 3 scope was modified substantially and will be a subject of
a separate report.
The plant investment and operating costs obtained as a result
of the phases 1 and 2 evaluation were unexpectedly high when compared
to values previously reported in the literature. Because of these
unusually high costs and the lack of success in sorbent development,
a reappraisal of the phase 3 work was made which resulted in Kellogg
recommending a scope change whereby a cost sensitivity analysis would
be substituted for the conceptual design. NAPCA subsequently agreed
to the proposed scope change since there appeared to be little value
in making a prototype plant design for a high cost process which re-
quires a sorbent that has not yet been successfully developed. On
the other hand, the proposed cost sensitivity study would delineate
the areas where major cost reductions might be possible and thus
provide useful information for guiding future research and develop-
ment of SC>2 removal schemes.
One of the principal results that should be obtained from the
phase 3 study will be the effect on operating costs of such items
as load factor, maintenance charges,;by-product sulfur selling price
and capital charge rate. Many of these items generally do not have
one specific "right" value since they vary from site to site, but
-1-
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a single value of each had to be chosen for the base case economic
evaluations of the three process schemes studied in the phases 1
and 2 work. Consequently, the values assigned to these parameters
will vary amoung different investigators so that comparisons of
different processes often are not an a common basis. By determining
the effect on plant costs of the major parameters over a range Of
values, as will be done in phase 3, a means for comparing different
processes on a common basis will be provided.
The results of the phase 3 work will be presented in thd form
of graphs and charts such that a range of values for essentially any
combination of major parameters may be evaluated. It is prtefcently
expected that the phase 3 final report will be published in the •
spring cf 1970.
- 2 -
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INTRODUCTION
The invention of the alkalized alumina sorbent and process
for removing sulfur oxides from gases is covered by United States
patent 2,992,884, July 18, 1961, assigned to the United States
of America by the inventors Bienstock and Field. Development
of the alkalized alumina sorbent and process resulted from a
study started in 1956 by the United States Bureau of Mines
(USBM) for the Department of Health, Education, and Welfare.
Objectives of the study were the screening of existing processes
and, if it appeared feasible, development of new materials or
processes which could be used for removing SO2 from flue gases
on a commercial scale. After several years of intensive research
during which many potential candidates were evaluated, the al-
kalized alumina sorbent was selected as having the most poten-
tial for commercial application. Consequently, development of
the alkalized alumina sorbent and process was undertaken to pro-
vide experimental data on which process designs and economics
could be based.
To facilitate the development of alkalized alumina, programs
were established by several different investigators, viz., AVCO
Corporation, W, R. Grace Company, and the USBM both at Bruceton,
Pennsylvania and Albany, Oregon. Various aspects of the develop-
ment were emphasized at the different locations, e.g., sorbent
development by Grace, kinetics by AVCO, etc., although some
overlap did occur. As the development of alkalized alumina
progressed, it became apparent from the mass of data being accu-
mulated that correlation and evaluation of these data by a single
contractor was needed. Consequently, The M. W. Kellogg Company
was awarded a contract to serve as prime contractor to evaluate
the alkalized alumina process. The contract comprised these
phases which can be briefly summarized as follows: phase 1 -
-3-
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data evaluation and preliminary process design; phase 2 - process
optimization studies; and phase 3 - a conceptual design for a tense
case based on the phase 1 and 2 results. This report covers
work done under phases 1 and 2, namely data evaluation, and
liminary process design including process economics. Phase 3
work, for which a scope change has been prepared, will be the!
subject of a separate report. A cost sensitivity study had
been substituted for the prototype plant design specified irt
original contract.
- 4-
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SUMMARY
The basic sorbent used in the alkalized alumina process is an
activated form of sodium aluminate (NaA10_) prepared by thermally
driving off water and carbon dioxide from Dawsonite [NaAl(CO3)(OH)2J
which in turn has been prepared as a precipitate by reacting liquid
solutions of sodium carbonate and aluminum sulfate. After activation
the sodium content is about 25%, the surface area is in the range
of 50-70 square meters per gram, and the pore volume is about 0.6
cubic centimeters per gram. Small amounts (<1%) of iron in the
sorbent are needed to provide adequate regeneration properties;
i.e., without iron, the regeneration is incomplete and the rate
of regeneration is very slow. Removal of sulfur dioxide by the
alkalized alumina is effected by sorption and oxidation to sulfur
trioxide and formation of the alkali metal sulfate; i.e., sodium
sulfate (Na_SO.). Regeneration is accomplished by reacting the
sodium sulfate with a reducing gas (e.g., hydrogen steam-reformed
methane) thereby converting the sodium sulfate back to sodium
aluminate and liberating the sulfur as gaseous hydrogen sulfide.
The standard sorbent exhibits one notable deficiency, viz., very
low attrition resistance, but sorbent development is the subject
of separate studies by others and is not part of the present con-
tract. All the work on the alkalized alumina process done to date
by M. W. Kellogg and included in this report is based on the standard
alkalized alumina sorbent as described above.
The mechanism of SCU removal by alkalized alumina has been
studied by various investigators (e.g., USBM, AVCO Corp., W. R. Grace)
and several kinetic models have been developed that appear to ade-
quately describe this mechanism based on laboratory and pilot plant
data. The various models all agree quite well for the lower range
of sorbent loading (i.e., percent sulfur on sorbent), but exhibit
some deviations in the upper range. There are some differences in
-5-
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sorption data reported by different investigators but in general
these differences are within the range of experimental accuracy.
Although the actual mechanism by which S02 removal is accom-
plished has not been completely resolved to everyone's satisfaction,
tne following is a composite of the generally accepted theories.
It appears that the S02 removal mechanism is pore diffusion con-
trolled when an actual flue gas is being treated, whereas both
diffusion and reaction rate are involved when a simulated flue
gas is used. The .nitrogen oxides (NOX) in the actual.flue gaa
apparently catalyze the Na2S04 reaction by increasing the SO^ to
SO3 oxidation rate,,such that a shell of Na2S04 is rapidly formed
on the outside-.surface of the sorbent particle. Further,'it had
been shewn.that above sorbent loadings of 1-2% sulfur, pure'dif-
fusion is controlling. Thus the rate of diffusion of the reactants
(i.e., S02, 02) into the interior of the particles essentially coh-
trols the rate of SO- removal from the flue gas since the time te-
quired for the reaction is negligible. Further, with.the higher
reaction rate the S02 backpressure in the particle interior is
essentially zero, thus enhancing the diffusion rate. An alternate
theory is that NO increases the reaction rate and weakens the shell
a
by causing the formation of a porous Na2SO^ shell. All of the! above
appear reasonable and it is quite possible that the actual Mechanism
is some combination of the various methods described above* It snould
be emphasized, however, that while the actual mechanism is important
for sorbent development work, for the present study it does hot affect
process design once adequate sorption models have been deiro loped, and
confirmed by experimental data.
As pointed out above, the rate of S02 removal from actual flue
gases is 2 to 3-fold that from simulated flue gases which dd not
contain NOX. Since actual flue gases from coal-fired power plants
will always contain NOX, the data and models based on gases con-
taining these two compounds have been used in the present study.
- 6 -
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The present goal of SCL removal efficiency has been arbitrarily
set at 90% based on the S02 concentration in the flue gas feed. Re-
ported experimental data show that this goal has been satisfactorily
achieved in the fixed and fluid bed systems but has not been adequately
demonstrated in the dispersed phase unit. Host of the dispersed phase
data show 50-70% SO, removal with about 90% removal being achieved in
but a few runs. However, an extended steady state run has not been
made, primarily due to the lack of an adequate supply of sorbent.
It is believed that 90% removal can be achieved with the proper
sorbent and perhaps some adjustment of operating conditions if
necessary.
The equilibrium level of sulfur on the sorbent after multicycles
of sorption-regeneration has not been clearly defined. The initial
sulfur content of the activated sorbent is generally less than 1%,
but builds up to a higher level after several cycles of sorption-
regeneration. The actual level does not significantly affect pro-
cess design provided (1) the design working level (i.e., the difference
in sulfur content between loaded and regenerated sorbent) can be
achieved, and (2) the sorption and/or regeneration kinetics are not
appreciably reduced. A multicycle test under anticipated operating
conditions is required to determine the equilibrium level of sulfur
on the sorbent. It is possible that the sorbent sulfur level con-
tinues to build up as the number of sorption-regeneration cycles
increases, in which case the working differential would be gradually
decreased. Long term tests would also resolve this question.
There are various reactions that could describe the regenera-
tion of alkalized alumina but to date, the exact mechanism has
apparently not been determined. Some experimental data have been
obtained which indicate the regeneration to be exothermic but
thermodynamic calculations show that it can also be endothermic.
At present there does not appear to be sufficient data available
to clearly define the reaction(s). Depending on which of the
t7-
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several possible reactions occurs, the heat of reaction can be
either endo- or exothermic and calculations of the heat bala*KS«
around the regenerator show that the temperature can vary by
15-50°F, the actual value being a function of the specific re-
generator design used. A better definition of the regeneration
mechanism would allow the regeneration system to be designed
with fewer safety factors than presently used.
The full effect of various process parameters on regenera*
tion has not yet been determined. One of the main problems in
present regeneration studies is the fact that a suitable Borbent
has not yet been found. Therefore, one of the major process
parameters, viz., the effect of regeneration conditions on
sorLent attrition resistance, cannot be properly evaluated* In
fact, the attrition losses of the presently available sorbenfefi
are so high that very little effort has been directed toward
measuring the effect of various parameters (e.g., water content
r
of the reducing gas) on attrition resistance since the results
would be of doubtful value. Once a suitable sorbent has bean
developed then optimization studies of regeneration conditions
can be made. Meanwhile, present regeneration studies have pro-
gressed to the point that several general trends are apparent,
as listed below.
The presence of about 1% iron in the sorbent both increase's'
the rate of regeneration gas consumption, and results in more
complete sulfur removal.. In fact, those sorbents without iron
exhibited very slow reaction rates and incomplete sulfur removal
to the extent that they would probably not be commercially aece"p-
table.
The gases tested 'for single step regeneration include hydi?o-
gen, carbon^monoxide, steam reformed natural gas, producer gals,
and methane. Two step regeneration tests have been conducted usin$
carbon dioxide as a secondary treatment after a primary?
-8-
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using H2, CO, and various combinations of their mixtures. Complete
reduction was obtained with H2 and mixtures of H2 and CO, but complete
sulfur removal was not obtained with CO or methane. It has been
found, however, that the reaction rate with CO is faster than with
H2% When C02 was used as a secondary treatment, sulfur removal was
increased, particularly for those tests where CO or mixtures of H2
and CO were used for the primary treatment. Reducing gas utilization
has not been clearly established, but some thin bed batch data at
a space velocity of 1000 hr~ show about 60% of the reducing gas con-
sumed, and fluid bed data show at least 70% utilization at a space
velocity of 37,000 hr"1. Thus, 80-90% utilization of reducing gas
might be obtained in a commercial scale regenerator and, in fact,
80% was the value selected for the present designs.
The effect of sorbent sodium concentration on regeneration rate
was evaluated over a range of 11-22% sodium. The lower sodium con-
centration produced a harder sorbent and exhibited higher maximum
reaction rates but the overall rates (basis: volume reducing gas
consumed per unit time) were about the same in all cases. The
effect of increased hardness on attrition rate was not determined.
The S02 capacity of the sorbents with the lower sodium level was
reduced but no data are listed for S02 sorption rates.
The regeneration of alkalized alumina has been shown to be
highly temperature dependent; e.g., a 1.7 fold increase in reaction
rate is realized for each 10°C temperature rise in the temperature
range'of 650-700°C (1202-1292°F). In addition, most of the ex-
perimental data show relatively slow reaction rates at 1200°F and
it appears that regeneration temperatures of 1300-1400°F are re-
quired if reasonable (i.e., 1-2. hours) holdup times in the regen-
erator are to be achieved.
All of the presently available experimental data from bench
scale and pilot plant measurements of attrition of alkalized alumina
indicate unacceptably high rates would be obtained in a commercial
unit. A standarized laboratory attrition test unit and method of
-9-
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reporting results have been developed to facilitate comparison Of
results from different investigators. The laboratory attrition
data are useful for screening purposes and in determining the
correlation between sorbent properties (e.g., hardness) and at-
trition rate, but insufficient data are available to permit pr$~
diction of actual plant losses to be made based on laboratory
results. Long term steady state data covering a range of operating
conditions and sorbents are needed to permit correlation of aatUal
losses with laboratory data.
Based on the results of the data evaluation study summarised
above, three different process designs for a commercial size plant
have been made for the alkalized alumina process, viz., fixed bed,
fluid bed, and dispersed phase. In all three cases the alkali*ed
alumina plant has been sized to treat the flue gas from a 1000
megawatt power plant burning coal with 3% sulfur, and a heat irate
•i
of 8600 Btu/kwh has been assumed. These designs were based on
available experimental data, supplemented by assumptions where
data were missing. The objective in using three different process
schemes was to determine which, if any, would be the optimum design
considering both economics and operability. Process descriptions
and design rationale are presented along with relative merits and
approximate economics for the three processes. It should be emphasized
that although a suitable sorbent is not presently available* the
design economics, and comparisons of the three processes are based
on the assumption that a commercially acceptable sorbent Will be
developed.
The fixed bed process comprises a large number of parallel
reactors containing stationary beds of sorbent through which flue
gas, heating gas, reducing gas, and cooling gas alternately flow.
The main advantages of this process are (1) the sorbent is Subjected
to substantially less mechanical stress since transfer otf the solids
between vessels is not required, and (2) the large change* itt flue
gas flow rates normally encountered do not adversely affect fe
-10-
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The main disadvantages are the need for very large ('«20' diameter)
valves to switch the gas flows between vessels for the different
operations (i.e., sorption, heating, etc.), and the potential explosion
hazard caused by carrying out the regeneration in the same vessel
as sorption, heating, and cooling, thus causing H_ and CO to come
into contact with gases containing small amounts of Oj at an elevated
temperature. The explosion limits of the mixed gases for the design
case were calculated based on available literature data but no firm
conclusions can be drawn from the results. The literature data
required extrapolation and the results were borderline with regard
to explosion limits. Therefore, it would be necessary to obtain
experimental data for the actual gas composition before building
a commercial plant, and if the results show that an explosion
hazard exists, an inert gas purge between cycles would be required.
The fluid bed system uses dense phase (30-40 Ib/cu ft) fluidized
solids reactors and regenerators characterized by moderate gas
velocities (2-5 ft/sec) and high solids residence times (1-2 hours
or longer if needed). The low gas velocities should help attrition
and the long solids residence times permit high sulfur loadings on
the sorbent thus minimizing solids circulation rates. Conversely,
the low gas velocities require a number of large diameter reactors
in parallel (e.g., 10 vessels @ 40 feet diameter) which results in
high investment and produces a complicated operation. Since the
stability of a fluid bed is dependent upon gas velocity, the reactors
and regenerators must be designed sucn that the lowest actual gas
velocity anticipated (e.g., a power plant turndown ratio of about
2), is higher than the minimum fluidization velocity. Under these
conditions the entire alkalized alumina plant will run continuously
and not require a shutdown of half the system when the power plant
output is cut in half which could occur as frequently as once per
day.
The dispersed phase design uses a high gas velocity (25-30
ft/sec) and very dilute solids loadings Ko.15 Ibs/cu ft). Con-
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sequently, only a small number (e.g., 2 £ 40 feet diameter) of reactor*
are required, but the solids residence time is low and the sulfur
loading on the sorbent is correspondingly reduced below that obtained
with the other two systems. It is expected that the high gas Velocities
will cause greater attrition losses than the other two processes but
since a suitable sorbent is not presently available, quantitative
measurements of these losses cannot be made at this time.
Either a moving, or fluid bed regenerator can be used with thfe
dispersed phase reactor. A moving bed design has been used in the
present case because calculations show very little difference in
cost between a moving and fluid bed, and for this particular design,
a moving bed should give better operability. Since the dispersed
phase design is dependent on gas velocity, the daily, one half
reduction in power plant flue gas flow normally encountered might
require a shutdown of one half of the alkalized alumina plant.
Consequently, the dispersed phase system has been designed as two
parallel trains such that one train can be independently shut down
if required. The frequent start-up and shutdown of a moving bed
is expected to be much easier than that of a fluid bed, and since
the costs are about equal, the moving bed design was selected.
A steam-methane reformer (for producing the reducing gas
needed for sorbent regeneration) and a Glaus plant (for recovering
elemental sulfur from the regenerator off-gas) have been included
in each of the three alkalized alumina processes discussed pre-
viously. Both of these items are considered to be standard units
and it is not expected that either will require development work,
The fixed and fluid bed schemes require a flue gas feed
containing as little fly ash as practical since the alkalized
alumina beds will act as filters and the sorbent could become
deactivated if large quantities of fly ash were added, to the
system. Therefore, the power plant electrostatic precipitator
has been relocated to a position upstream of the air preheater
-12-
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in the fixed and fluid bed designs. The resulting higher gas
temperature in the precipitator increases the gas volume and re-
quires a larger precipitator. The additional cost of the preci-
pitator is considered as part of the alkalized alumina process
investment. In the dispersed phase design, the precipitator
location is not changed but removal of the S02 from the flue
gas upstream of the precipitator lowers the collection effi-
ciency and again a larger, more expensive precipitator is re-
quired. As before, the additional cost of the precipitator is
included as part of the alkalized alumina investment.
The alkalized alumina plants have been designed such that
all of the effluent gas (transport air and heater flue gas) is
combined with the power plant flue gas for venting from the
system. This additional gas is hot so it is admixed with the
flue gas upstream of the air preheater. In addition, the hot
regenerated sorbent solids raise the flue gas temperature in the
reactor such that both the quantity and temperature of the hot
gas flowing through the air preheater is increased. Operation of
the alkalized alumina process therefore causes an increase in
the size and cost of the air preheater and this incremental cost
has been included as part of the S02 removal plant cost.
Estimates of plant investment for the three process designs
were prepared in two parts: part 1, comprising only the sorption-
regeneration section, and part 2, the entire plant including the
steam-methane reformer, the Claus plant, and the incremental costs
of the electrostatic precipitator and air preheater. Originally,
Phase 1 results were to be used to select the best process for a
prototype plant design as specified for Phase 3 by the contract
before it was changed. However, preliminary results from Phases 1
and 2 indicated that more useful information could be obtained
by making estimates for complete plants and then determining the
effect of various process parameters on economics. Accordingly,
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the latter approach will be followed in Phase 3. To provide reason"
ably sound bases for the estimates, approximate designs of major equip-
ment were made in sufficient detail to insure process .and mechanical-
feasibility, and approximate plant layouts were prepared from which
the quantities of large ductwork and piping were determined. With
the major equipment specified and approximate plant layouts .prepared,
standard Kellogg estimating techniques were used to determine them
plant investment. The plant location was assumed to be in Ohio, ands
a labor productivity typical for this area was used in estimating tr
construction man-hours. The accuracy of the present estimates ;may»
be considered to be in the range of -5% to +25% based on the current
flow sheets and using the techniques described above.
The total annual costs for ;the three processes have -Jaeen cal<- -
culated using capital charges of 13% with other cost items derived ?
from a variety of sources as shown in the main body of this report.
There does not appear to be a general agreement as to what the
capital charges should be (e.g., TVA suggests 14 1/2%), but for
the present evaluations 13% seems to be a reasonable value. >Fur-
ther, NAPCA suggested that for evaluation of area surveys for SO,
removal capital charges of 13% should be used. The effect of
capital charges on process economics will be studied in Phase 3
and, therefore, is notrincluded here.
The investments-for the complete plant (i.e., including the
regeneration gas producer, the Claus plant, and incremental cost
of the power plant .electrostatic precipitator and air preheater)
have been estimated'to be about $31,000,000, $36/000,000, -anc.
$40,000,000 for the dispersed phase, fluid bed, and fixed-bed
designs respectively.- Annual costs for the three cases are.
calculated to be $7,900,000, $9,300,000, and $10,400,000, .exclud-
ing attrition losses and credit for by-product sulfur. Costs have
also been calculated based on assumed attrition rates (0.1% of the
solids throughput for the dispersed case, 0.05% for the fluid bed$
-14-
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and a fixed bed sorbent life twice that in the fluid bed) using a
sorbent price of 25 cents per pound and a sulfur credit of $20 per
ton. The results show a net reduction in costs for the fluid and
fixed bed cases only. A brief summary of the overall economics is
shown below.
DISPERSED PHASE FLUID BED FIXED BED
INVESTMENT;1
Sorption-Regeneration $22,100,000 $27,630,000 $30,980,000
Complete plant (battery
limits) $30,720,000 $35,900,000 $40,210,000
COSTS1(COMPLETE PLANT)
$/YEAR;
Basic Process2 7,874,000 9,334,700 10,428,500
Assumed Attrition Loss
(l§ sorbent cost of 25$/lb) 926,600 249,400 158,400
Sulfur Credit
(@ $20/ton) (937,400) (937,400) (1,007,400)
Net 7,863,200 8,646,700 9,579,500
$/TQN COAL;
Basic Process 4.53 5.40 6.01
Net 4.53 4.98 5.52
1 Alkalized alumina plant sized to treat gases from 1000 MW power
plant burning coal with 3% sulfur.
2 Excluding Attrition, sulfur credit.
The above costs are based on a load factor (stream efficiency)
of 60% (5250 hours per year); i.e., it is assumed that the power
plant will operate at full capacity 60* of the time. This is.consis-
-------
tent with power plant practice based on a recent survey which
showed the largest plants (8.6% of industry) had a 63.6% load
factor, a selected group (89% of industry) had 58%, and the entire
industry averaged 54%. The effect of load factor on economics
will be evaluated in the Phase 3 study whereas only, the 60% case
is considered here.
The unit costs of 4.53 to 5.52 dollars per ton of .coal shown
above represent an increase in fuel cost of 70 to 85% based on the
6.50 dollars per ton (25C/MM Btu) coal used in this study. The
relatively minor effect (G<10% reduction) of by-product sulfur on
overall economics is due to two factors: (1) the low price of
sulfur ($20/ton) that can be expected since it is a by-product,
and (2) the small quantities produced (45,000-50,000 tons/year).
The cost of sorbent lost to attrition is based on the assumed
values listed above but it should be emphasized that to date
these low values have not been obtained experimentally; rather,
losses of about 7-8% (basis solids throughput) have been observed
in the dispersed phase pilot plant. Obviously, the economics
shown above are contingent upon the development of a sorbent with
much better attrition resistance than exhibited by those that have
been tested to date.
A comparison of the above tabulated costs for the three dif-
ferent processes shows the dispersed phase case to have both the
lowest investment and operating costs, followed in order by the
fluid bed and fixed bed case. it should be noted, hov;» ver, that
if the assumed attrition loss for the dispersed phase case (0.1%
of the solids throughput) is doubled, then the total net annual
cost becomes about the same as the fluid bed case. Thus it is
apparent that the relative economic position of the different
cases is very sensitive to attrition loss. Further, the limited
data presently available indicate that fixed bed attrition might
be small compared to dispersed phase attrition, and therefore, the
-16-
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assumed differential attrition losses used in the economics eval-
uation may be too low. When the uncertainties regarding attrition
losses are considered, and in light of the sensitivity of process
economics to attrition, a firm conclusion as to which of the three
processes offers the lowest costs cannot be made at this time.
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CONCLUSIONS AND RECOMMENDATIONS
As a result of the data evaluations and process designs reported
herein, which are based on information made available to Kellogg
under the subject contract, the following conclusions and recommenda-
tions are made regarding the alkalized alumina process:
• The process chemistry is technically feasible insofar
as the sorbent will satisfactorily remove SO2 from power plant..
flue gases and can be regenerated successfully, as demonstrated
in bench scale and pilot plant equipment.
• The outstanding technical weakness of the process is the
lack of a sorbent that has attrition resistance adequate for
commercial operation using fluosolids techniques. Considering
the effort that the various organizations (Peter Spence, W. R.
Grace, USBM, Kaiser) have spent on sorbent development without
producing a candidate with the desired properties, it is possible
that the nature of the alkalized alumina/SO? reaction precludes
the development of a sorbent suitable for use in a fluosolids
system.
• A less attrition resistant sobbent probably would be satis-
factory for a fixed bed system if an adequate sorbent life could
be obtained (vis-a-vis, chemical activity and/or pressure drop
considerations). However, test data on sorbent life and attri-
tion losses in a fixed bed system were not available:for- this
evaluation so conclusions regarding the suitability of the
present sorbents for this purpose cannot be -made.
» The unexpectedly high costs obtained for the alkalized alumina
process coupled with the lack of success in sorbent development
led to a reappraisal of the Phase 3 scope as specified in the
-18-
-------
present contract, viz., a prototype plant design. Subsequently,
Kellogg recommended to NAPCA that a cost sensitivity analyais
be substituted for the prototype plant design. It was the basic
similarity of all processes for removing S02 using a dry sorbent
which is chemically regenerated at an elevated temperature that
led to this recommendation. Since the S02 removal costs, exclud-
ing attrition, are largely a function of investment and since
any particular process design (hence investment) of the type under
consideration is not greatly affected by sorbent properties, it
appears that results and conclusions obtained from the alkalized
alumina cost sensitivity study will be applicable to processes
using other sorbents such as metal oxides. Accordingly, a series
of variables to be studied and their ranges was recommended to
and accepted by NAPCA as a new phase 3 scope. The phase 3 ob-
jective will be to discover which process parameters have the
most influence on plant costs thus providing a basis for guiding
future research and/or development of dry systems for SO_ removal.
Thus it is anticipated that many of the results obtained in phase
3 will be useful in the preliminary evaluation of other similar
processes.
A major finding of the phases 1 and 2 work was the unexpectedly
high capital investment required for the alkalized alumina pro-
cess. The enormous volume of flue gas to be treated combined
with its very dilute concentration of sulfur oxides produces
unusual process requirements and results in the need for ex-
tremely large and costly equipment. A special note should be
made of the large lengths of flue gas ductwork required by the
alkalized alumina plant. Normally, a power plant is designed
such that the gas ductwork is minimized, and the design of the
ductwork itself is based on a lower internal pressure than is
needed for the alkalized alumina process plants. The combina-
tion of high internal pressure (e.g., design value is 45 inches
of water for the fluid-bed case) and long runs (hundreds of feet)
-19-
-------
results in the gas ductwork becoming a major cost item in the
alkalized alumina process plants. Since the volume of gas
being handled is nearly a constant, other parameters being
equal, it is concluded that ductwork will comprise a major
cost item in practically all processes treating flue gases.
It is further noted that a sketch of the plant layout, however
rough, is essential for estimating ductwork and piping quan-
tities.
• The general economic analysis of the alkalized alumina,
process, or indeed any SO., removal process, involves several
factors for which values must be estimated but for which there _
is no single value that applies in all cases; e.g., load factor
(defined as the number of kilowatt hours generated per year as
a fraction of the installed generating capacity). Other factors,
such as maintenance, must be based on judgment since these are
first-of-a-kind plants and direct operating experience is not
available. One very important, conclusion that has been reached
regarding system economics is that credit for by-product sulfur
will not offset operating costs and therefore installation of
SO2 removal equipment will always cause an increase in power
costs. Based on the above considerations it is concluded that
without specific standards (e.g., SO- concentration in effluent
gas) to guide process design or definite economic goals to meet,
and since it presently appears that all of the SO,, removal pro-
cesses will cause an increase in power costs, then an economical
process is one which causes the smallest increase in power costs.
Thus it cannot be ascertained if a process is economical Without
comparing it .to other processes since the economics are on a
relative rather than absolute basis. Since many other SO2 re-
moval schemes are presently being developed, the relative merits
of alkalized alumina can be determined only when costs from these
other process are available. However, since the alkalized alumina
-20-
-------
process increases fuel costs by about 75%, it is concluded that
even if a suitable sorbent is developed, it is highly probable
that other processes can be developed which will be more economical,
-21-
-------
DISCUSSION
A. STATUS OF DATA
1. Attrition
The two major cost items in the alkalized alumina process
are capital charges and raw materials, these two comprising
about 75% of the total annual cost based on the process
economics of Katell (CHEM. ENGR. PROG., 62 (10) , 67 (1966)).
The sorbent make-up, based on an assumed attrition rate of
0.1% of the sorbent flow to the regenerator, accounts for about
25% of the total cost at an alkalized alumina charge of 25
cents per pound; the balance of raw material costs is due to
reducing gas generation. The major contributor to capital
charges is the absorber/regenerator combination which accounts
for about 17% of the total annual cost. Thus, major criteria
for process design should be an absorber/regenerator combina-
tion that will give minimum attrition rates and require mini-
mum investment. Since attrition plays a key role in process
design and economics, it is important to accurately determine
what level of attrition might be expected in a commercial plant.
Experimental methods of obtaining attrition data are discussed
below.
a. Various Procedures for Generating Attrition Data
Several different types of attrition test units have
been used to obtain attrition data. Analysis of the presently
available data along with a brief summary of the different
attrition test methods is included below. (see Appendix A
for details) :
(1) USBM Pneumatic Conveyor Attriter
This apparatus consists of a one-inch pipe 28
feet long through which the gas-solids flow, with the gas
velocity set at 29 ft/sec; the solids are separated from
-22-
-------
the air (in a large diameter chamber) and returned by gravity
flow to the attriter inlet. Attrition rate is reported as
the amount of fines (size not defined) collected in 24 hours
as a percentage of total weight of solids circulated in that
period. After 10 days of operation (19 cycles/hour or about
4500 cycles total), the attrition rate was 0.005% and the
particle size had decreased slightly, i.e., from 82% 8-10
mesh and 18% 10-12 mesh initially to 46% 8-10 mesh and 40%
10-12 mesh after 10 days. By converting these data to a
basis of inventory, rather than circulating solids rate, the
attrition loss becomes 2.28%.
(2) Tyler Sieve Shakers
Tyler "Ro-Tap" testing sieve shakers were used
to test Grace No. 1 (standard) alkalized alumina in various
physical forms (as received, after 600°C and 122°C air roast,
after 600°C H_ roast), and two different catalyst supports.
Attrition loss is defined as the percent reduced to minus
60 mesh. Results of these tests show that increasing the
hardness (crushing strength) of the material does not neces-
sarily increase attrition resistance.
(3) USBM - Albany - Fluid Bed and Air Lift
Attrition Test Unit
USBM at Albany Metallurgy Research Center has
developed both a fluidized-bed and an air-lift attrition
test unit. Reported gas velocity is 1.8 feet/sec in the
fluid-bed and about 25 feet/sec in the air lift. Attrition
loss is measured as the weight loss per 24 hours (reported
as percent of inventory per hour) with daily additions of
material to maintain the original sample weight (50 and 10
grams for the fluid and air lift units respectively). Size
distribution at the start and after 15 days are reported
along with the daily and total weight loss. Grace #1 absor-
bent with various levels of sulfur was tested.
-23-
-------
The fluid bed size distribution data for the
activated (1.63%S) and a loaded (8.85%S) sample have been
plotted on logarithmic probability paper, Figures 1 and 2.
As is evident from these plots, there is essentially no
increase in the quantity of fines which is not too surprising
since the method of testing eliminates the fines from the
system. After 15 days the particle size distribution has
been shifted slightly such that a more even distribution of
the various fractions is obtained; i.e., the quantity,of
minus 10 plus 35 mesh material has been increased. By in-
spection the other samples show the same characteristics and
therefore have not been plotted. This slight decrease in
particle size should help rather than hurt absorption rates
since an increase in surface area is obtained.
The attrition rates obtained in the fluid-bed
unit decrease rapidly the first few days and then level out
to essentially constant values after about 7-10 days; also
they generally show decreasing values with increasing sulfur
content (may be due to increasing density) as illustrated by
the data (Grace #1 absorbent) in Table 1.
The data in Table 1 indicate that regeneration
apparently weakens the absorbent and causes the attrition
rate to increase. These data are based on only one loading
and regeneration but verbal information from Albany indicates
essentially no change after 10 cycles. While these data pro-
bably cannot be used directly as design data, they are useful
for screening purposes and when compared to the air-lift data
could give an indication of the relative attrition: rates of the
two systems if. a common basis of comparison can be found. Listed
in Table 2 are data for the two systems for Grace #1 absorbent;
note that comparisons have been made on three different bases
the rates for the 1st 3 days, the rates for the 3rd day, and
the air-lift 3rd day rate with the last 5 days for the fluid-
bed rate.
-24-
-------
5-
3-
cn
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o
oc
9
2
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i ** I03 -
tO tf»
Ul
1 UJ
o
cr
<
a.
3-
io2-
o.c
TYLER STANDARD
SCREEN SCALE
^
— 10 MESH
— 1 4 MESH
— 20 MESH
/
/
— 28 MESH ^
'J
— 35 MESH ^ j£
==• 48 ME
1 /
SH A 6
! I
FIGURE 1
- •" GRA
ALBAN
START
\*
«'«'**
/ w^^A"l
V^
*r
PARTICLE SIZE DISTRIBUTION
CE ALKALIZED ALUMINA — ACTIVA
1.63% S ON SORBENT
Y, OREGON TEST DATA - FLUIDIZI
TABLE 4, LETTER DATED 4-11-61
^
LR 15 DAYS
3*
TED
ED BED
.
O RUN 1 START (46.7% -I- 10 MESH)
A RUN 2 START (56.5% + 10 MESH)
(
t
§ RUN 1 AFTER 15
k RUN 2 AFTER 15
DAYS (41% -»- 10
DAYS (44.4% -»-
MESH)
n MF*;H\
)l O.I 10 50 90 99 99.
PERCENT LESS THAN STATED SIZE
-------
Crt
O
X
o
UJ
> 2
to ,fl
a\
1 ui
K
e
2
TYLER STANDARD
SCREEN SCALE
4
- 10 MESH
IOZ
FIGURE 2
PARTICLE SIZE DISTRIBUTION
GRACE ALKALIZED ALUMINA - LOADED
8.85% S ON SORBENT
ALBANY, OREGON TEST DATA - FLUIDIZED BED
TABLE 4, LETTER DATED 4-11-68
RUN I START (67.1% + 10 MESH)
RUN 2 START (68.9% 4- 10 MESH)
% RUN I AFTER 15 DAYS (54.9% + 10 MESH)
A RUN Z AFTER 15 DAYS (55.7% 4- 10 MESH)
O.Ol
5O
TMAIM STATED SIZE
9O
99
IOO
-------
TABLE I
FLUID BED ATTRITION TEST UNIT DATA
AVERAGE ATTRITION LOSS*,% OF INVENTORY/HR
SAMPLE
Activated
Loaded
Loaded
Loaded
Regenerated
%
1
2
5
8
1
S
.63
.29
.87
.85
.38
1ST
0
0
0
0
0
5 DAYS
.151
.180
.106
.037
.279
2ND
0.
0.
0.
0.
0.
5 DAYS
078
092
068
036
163
3RD
0
0
0
0
0
5 DAYS
.064
.078
.063
.030
.144
* uSBM-Albany - April 1968 Monthly Summary
-------
TABLE 2
COMPARISON OF SORBENT ATTRITION LOSSES FOR AIRLIFT
AND FLUID BED SYSTEMS
ATTRITION LOSS*, % OF INVENTORY/HOUR
SAMPLE
%S
AIR LIFT
FLUID-BED
1ST 3 DAYS 3RD DAY 1ST 3 DAYS 3RD DAY LAST 5
Activated 1.63
Loaded 8.85
Regenerated 1.38
1.29
0.79
2.70
1.33
0.52
2.52
0.19
0.04
0.33
0.13 0.064
0.05 0.030
0.24 0.144
SAMPLE
%S
Activated 1.63
Loaded 8.85
Regenerated 1.38
RATIO: AIR-LIFT/FLUID-BED
1ST 3 DAYS
6.8
19.8
8.2
3RD DAY
10.2
10.4
10.5
3RD DAY/LAST
20.8
17.3
17 ,5
5
*USBM-Albany-April 1968 Monthly Report
-28-
-------
The comparison of the rates for the first 3 days
show more variation than do the other comparisons, possibly
because of varying amounts of fines initially present in the
different samples. The 3rd day rates are quite consistent but
the fluid bed test probably had not reached steady state con-
ditions, except perhaps for the loaded sample. Thus the most
valid comparison probably is the one between the air-lift 3rd
day rate and the last 5 days rate for the fluid-bed since these
data are the ones most likely to be for steady state conditions.
A direct comparison shows the air-lift rate to be about 10-20
times the fluid bed rate but this may not be a valid comparison.
Since the air-lift test conditions are such that apparently the
inventory is cycled about 6400 times per day (see section 5. C.
(1) below), the equivalent time in a commercial unit that this
represents compared to the equivalent time of the fluid bed is
needed.
The size distribution of Grace #1 sorbent with
various levels of sulfur are reported in the Albany attrition
test results discussed above. The samples were loaded using
a synthetic flue gas and regenerated in a fixed bed test unit
and any attrition produced during these tests would therefore
be due to chemical rather than physical action. The data have
been plotted (Figure 3) in the usual way and the results show
that there was no increase in the minus 60 mesh material for any
of the samples, and in fact, an increase in particle size was
obtained for the 8.85%S item. There appears to be a small
increase in the plus 48 minus 10 mesh fraction for the regen-
erated sample but the change is small. Based on these data
it is concluded that the attrition exhibited in the Albany test
units using a synthetic flue gas is due to physical rather than
chemical action, per se. However, it is quite possible, and
indeed, quite likely, that the chemical changes create stress
in the particles and weakens them such that subsequent physical
actions results in high attrition; rates.
-29-
-------
io* -
5-
3-
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Z
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,-^*°-
'*>*'
&•
O ACTIVATED (1.63% S)
X LOADED TO 2.29% S
A LOADED TO 5.89% S
$ LOADED TO 8.85%,S , ,
• REGENERATED TO 1.38% S
FROM 5.87% S LOADED
SORBENT
»
t
t
I
t
i
f
t
i
;
1
i
I
1
|
O.O1 O.I 1 IO 5O j , 9O 99 99.
PERCENT LESS THAN STATED SIZE
-------
(4) W. R. Grace Accelerated Air Jet Attrition
Test Unit
W. R. Grace and Company developed a new attrition
test for testing alkalized alumina sorbent, viz., the accelerated
air jet attrition test (AAJA). Basically it comprises blowing
air through the sample for a specified time, and then weighing
the material retained on a 14 mesh screen. Attrition loss is
reported as weight percent loss determined by difference between
the original and final sample.
\
Some values have been reported by Grace for both
Grace #1 (standard) and 12 (improved) but other than for com-
parative purposes (#2 worse than #1), these data presently have
little practical use. To date, the generally accepted defini-
tion of fines for this process is minus 60 mesh material and
since the Grace test reports all minus 14 mesh as fines, meaning-
ful comparisons with other data are essentially impossible
without knowing the percent of -60 mesh material in the 14
mesh fraction. Also, comparisons of results of different
materials tested on the Grace unit may be masked by the large
size screen used, in that two materials showing the same loss
could have quite different size distributions. For example,
a material with a 50% minus 14 mesh could also have 50% minus
60 mesh while another sample with 50% minus 14 mesh might
have 0% minus 60 mesh. These limitations severely limit the
usefulness of the Grace attrition data and modifications
either to the test procedure, or methods of reporting data
are indicated.
(5) Standard of Indiana-Cracking Catalyst Test
Procedure
A description of the attrition test method generally
used for determining relative attrition rates of cracking cata-
lysts is included in Appendix A; two variations are shown, the
-31-
-------
forty-five hour and the one-hour methods. The one-hour method
is basically the one developed by Standard Oil Company (Indiana)
and published in IEC, 41, 1200 (1949). The test does not re-
quire elaborate equipment, is easy to run, and the results are
reported in terms of an attrition index or attrition rate which
permits a rapid, direct comparison of the attrition resistance
of different catalysts.
Possibly this test could be used for testing the
alkalized alumina sorbent but since it has much larger particles
than cracking catalysts, modification of the equipment and/or
procedure may be required to obtain good fluidization with the
heavier particles (e.g., a larger size orifice may be needed).
However, any modification required probably could be determined
with 2-3 tests and the modified test procedure could then be
used by all the various investigators working on development of
the alkalized alumina. The attrition index of the standard
(Grace #1) sorbent could then be determined and used as the
reference point since some pilot plant attrition data are
available for this material. The attrition rate of the standard
appears to be unacceptably high so an improved sorbent would
have to exhibit a lower attrition index to warrant further
testing. A modified Standard of Indiana attrition test was
later developed by Kellogg and recommended for general use
(see section 1-d below).
b. Distribution of Attrition Losses
A one-cycle attrition test was run in the Brucetcn
pilot plant in an attempt to determine where the attrition
occurs. The pilot plant was cleaned of sorbent and about
1500 pounds of loaded, test sorbent charged to the regenerator;
this sorbent had previously been through about 9-10 absorption,
regeneration cycles and the size distribution was determined
prior to charging. The sorbent was regenerated for 10 hours
-32-
-------
with steam reformed methane, drained, the size distribution
determined, returned to the system, and pneumatically conveyed
to the top of the reactor. The size distribution of the con-
veyed material was then determined, next, sorption was carried
out in the usual manner at a feed rate of 180-190 Ib/hr so the
test took some 8 hours. After completion of the sorption run,
the system was drained (essentially a 100% weight balance was
obtained) and size distribution determined on each stream; i.e.,
the bottom discharge, the recycle and the overhead. These
data have been plotted in the usual manner (Figures 4 and 5)
and the attrition loss calculated for each step. Attrition
loss is defined as the percent increase of minus 60 mesh
fines compared to the original material. A summary of these
data is shown in Table 3.
TABLE 3
DISTRIBUTION OF ATTRITION LOSSES
After Regeneration
After Conveying
After Sorption
Total
* Attrition rate = (increase in -60 mesh) ,nn
(100-1.21)x 1UU
Note: Original sample contained 1.21%-60 mesh.
It is quite obvious from the Table 3 values and from
Figure 4 that sorption causes most of the attrition while con-
veying causes very little. It is also apparent that the total
loss of 6.82% is unacceptable high. The regeneration attrition
rate may not be a true value because of the method of filling
the regenerator, i.e., the sorbent was fed into the top in a
%-60 Mesh
an 2.65
3.06
7.94
-
Increase in
-60 Mesh
1.44
0.41
4.88
6.73
Attrition
Rate*
1.46
0.42
4.94
6.82
% of
Total
21.4
6.2
72.4
100.0
-33-
-------
I04
z
o
CE
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N |03
I UJ
o
(L
a.
FIGURE 4
CHANGE IN PARTICLE SIZE DURING ABSORPTION-REGENERATION CYCLE
BUREAU OF MINES, PITTSBURGH COAL RESEARCH CENTER DATA
TABLE 4, BM-19, QUARTER ENDING 12-31-67
TYLER STANDARD
SCREEN SCALE
^
- 8 MESH
<
• BEFORE REGENERATION
A AFTER REGENERATION
D AFTER CONVEYING "^t^^^
-,0 MESH + AFTER ^SORPTION ^^^
- 14 MESH
-20 MESH
-28 MESH
-60 MESK
i
sr
cvx
J& /
a/ /
y j*
Jr /
& /
/-/
y/
'/
sr f *
.s */ *
r /
/
/
/
/
^x
^r
-^S!l**~
---*^
4
DISTRIBUTION OF ATTRITION LOSS (-60 MESH)
CONVEYING - 6.2% 1
REGENERATION - 21.4% 1
ABSORPTION - 72.4% |
TOTAL ATTRITION PER CYCLE - 6.8%
O.Ol
O.I
10
5O
9O
99
99.99
-------
I04
i i 1 -r- r
FIGURE 5
PARTICLE SIZE OF SOLIDS DISCHARGED FROM ABSORBER
BUREAU OF MINES, PITTSBURGH COAL RESEARCH CENTER DATA
TABLE 2, BM-17, QUARTER ENDING 12-31-67
10*
TYLER STANDARD
SCREEN SCALE
-60 MESH
NOTE:
THE COMPOSITE SHOWN IS THE SUM OF THE
THREE DISCHARGE STREAMS. IT IS
IDENTICAL WITH THE "AFTER ABSORPTION"
CATALYST IN TABLE 4 OF SAME REFERENCE.
0.01
O.I
10 50 90
PERCENT LESS THAN STATED SIZE
99
99.99
-------
normal manner but since the regenerator was initially empty,
the particles' free fall distance was greater than with a
full vessel. Consequently, the force of impact was greater
which could cause higher attrition values. On the other hand,
a fixed bed rather than a moving bed was used for the regenera-
tion step and this would tend to cause lower attrition rates.
The relative effect of these two factors (greater free fall
distance, fixed bed) on the regeneration attrition rate is
unknown and the Table 3 values therefore may be misleading.
It should be noted that the bottoms product comprises
almost 60% of the fresh feed (Figure 5) which results in-a low
residence time for - the largest particles, thus causing low
su? fur pickup by'this fraction. Recycle of this bottoms-stream,
rather than the overhead, would permit higher sulfur loadings
to be obtained and could possibly reduce attrition losses; i.e.,
larger particles would be recycled and as these attrite they
would leave with the overhead stream to the regenerator. Thus
the particle size distribution should become narrower and a.
lower slip velocity could be used, thereby providing larger
contact time in the reactor. Pilot plant tests using recycle
of the bottoms stream could determine the effect on sulfur
loading and attrition loss.
c. Comparison of Experimental Data on Common Basis
An attempt was made to compare the Albany air-lift
attrition data with that reported for the Bruceton pneumatic
conveying test loop discussed previously. The ••pneumatic con-
veying results are reported as a percentage of circulating
solids whereas the Albany data are shown as a percentage loss
of inventory. Therefore, a direct comparison cannot be made
but with certain assumptions the data can be converted and
compared on a similar basis.
-36-
-------
(1) Albany Data Converted to Circulating Solids
Basis
Basis:
t
\
- solids density in the air-lift tube of 0.15
Ib/ft3, assumed based on pilot plant data for
similar operating conditions.
- average solids velocity through the air-lift
tube of 2 ft/sec, estimated based on particle
settling velocity and gas velocity.
- tube diameter is 1" and length of air-lift
tube is 12", based on sketch of equipment
- at 2 ft/sec and 1 ft tube length, time per
pass is 0.5 sec.
- sorbent inventory is 10 grams
Tube Volume = (ir/4) (1/12)2 (1) = 0.00545 ft3
Circulating rate = 0.00545 ft3(0.15 lbs/ft3) (24 hrs/day)
0.5 sec (1 hr/3600 sec)
= 141.26 Ibs/day
and (454 gms/lb)(141.26 Ibs/day) . 640Q cycles/day
10 gms inventory
Albany 3rd day air-lift data (April 1968 Monthly
Report) show 3.2 grams average weight loss/day for
Grace #1 activated absorbent.
•'* Attrition based on solids circulating rate =
3.2 gms/day (100 . 0^05%
(454 gms/lb)(141.26 Ibs/day)
This rate is in good agreement with the Bruceton pneumatic
conveying 2-4 day rate of 0.006% and the 10th day rate of
0.005%.
It should be kept in mind that the conversion
of the Albany data to a different basis is based on assumptions
-37-
-------
as to average particle velocity and solids density in the air-
lift tube. Bases for these assumptions were solids densities
reported for the Bruceton pilot plant, and calculated settling
velocities of 14 mesh particles. Thus, the assumptions should
be reasonably accurate. It should also be kept in mind that
different sorbents were tested in the two units: the,standard
(Grace #1) sorbent in the Albany unit, but the source of the-
conveying test sorbent is not clearly stated.
One other important point to be noted in comparing
the Albany air-lift and Bruceton pneumatic conveying test loop
data is that the air-lift tube is only one foot long whereas
the Bruceton unit is 28 feet in length. The close agreement
of the results when the difference in lengths is neglected,
as was done in the above calculation, indicates that either
1) the assumptions are wrong, 2) the end effects (i.e., solids
pickup and disengagement) cause most pf the attrition, or that
3) the two sorbents tested have different properties which off-
set the difference in apparatus. It is possible that for the
lengths of pipe tested, the attrition in the pipes themselves
is small compared to that resulting from entraining and dis-
engaging the solids. In this latter case the attrition would
be essentially independent of length but would vary directly
with the number of cycles to which the sorbent is subjected.
In any event, the results show the need for reporting data
on a common basis if meaningful comparisons are to be made.
(2) Bruceton Data Converted to Inventory B-.:~is
The data described in (1) above can also be
converted to an inventory basis and this has been done as
shown below. Assumptions are the same as those previously
used.
Bruceton Inventory - 19 passes/hr or hourly
circulating rate"was 19
x inventory.
-38-
-------
Reported attrition loss = 0.006% of 24 hr
circulating rate
' Conversion to inventory - 0.006(19)(24) = 2.74%
This is quite different from the 32% value obtained
with the air-lift. It should be noted, however,
that the air-lift inventory of 10 grams was cycled
267 times per hour (based on the calculations in
the previous section) compared to 19 passes per
hour for the Bruceton inventory. Thus, adjusting
for the number of times the inventory is circulated
through the equipment, the Bruceton loss becomes:
(267/19) (2.74) = 38.5%
This rate is in good agreement with the 32%
obtained in the Albany unit.
d. Recommended Method of Attrition Testing
It has been previously pointed out (section l-a-(5)
above) that the various laboratory attrition tests presently
used by the different investigators need to be standardized
to facilitate comparison of data from different sources.
Farther, modifications of the tests themselves would probably
give more meaningful results. Consequently, MWK undertook
the task of modifying the Standard of Indiana attrition test
so that it could be used for testing the alkalized alumina
sorbent. One sample each of Grace #2 and a Peter Spence
sorbent was obtained from Bruceton and used for the test
work. Concurrently, AVCO ran a similar Grace #2 sample
on a modified AAJA test apparatus.
The MWK modified Standard of Indiana test uses a
3/64 inch diameter orifice instead of the standard 0.0150
inch diameter and the orifice pressure drop is lowered to
-39-
-------
20 psi (vs. 55-60 psi standard). The length of the test
remains at one hour. The AVCO modified AAJA test increases
the opening of the flask to 3-1/2 inches and uses a 60 mesh
(U.S.) screen instead of a 14 mesh screen. Second degree
of severity conditions are used. Results of the AVCO test
were obtained from NAPCA for comparison with the MWK results.
The size distribution curves of the original and
attrited sorbent for both the MWK and AVCO tests are shown
in Figure 6. These same curves are also shown in Figure
7 along with the corresponding curves for the Peter Spence
sorbent. Prior to testing all MWK test samples had been
activated with H_ for 10 hoars at 630°C compared to 10 hours
at 650°C for the AVCO sample; therefore, the starting material
for both the MWK and AVCO tests is essentially the same.
Examination of the plots shows that in all cases the
attrited material has about the same size distribution .as the
original material down to about 14-18 mesh (U. S. Sieve Series),
but at this point the particle size of the attrited material
decreases rapidly. The actual quantities of fines (-60 mesh)
produced during the tests are:
AVCO test on Grace #2 - 25.7%
MWK test on Grace #2 - 26.7%
MWK test on Peter Spence - 21.1%
Thus, the results indicate that both the AVCO modified AAJA
and the MWK modified Standard of Indiana tests give e,~a:.valent
attrition losses on Grace #2. The MWK test on Peter Spence
material produced about 20% less fines than on the Grace #2,
but more data are required to determine if this difference is
significant. Further information is required on reproducibility
of both sorbents add test procedures as well as correlation of
test data with pilot plant data, but for screening studies, the
modified test appears to be superior to the methods previously
-40-
-------
I04
O
(E
O
1 £ 10'
T S
ui
o
(C
S
TYLER S
SCREEN
FIGURE 6
SIZE DISTRIBUTION OF GRACE NO 2 ALKALIZED ALUMINA
FANDARD
SCALE
•
\^
- 8 MESH
- 10 MESH
- 12 MESH A
- 14 MESH
- 1 6 MESH
- 20 MESH X^"
- 28 MESH '
- 35 MESH ^
/
- 60 MESd
i
(
X
O
CTIVATED GRACE
ORIGINAL
AVC
ACTIVATED SORBEC
ATTRITED SORBEr
ATTRITED SORBE
NO 2 ^, ~» >4*
/. _y
ff
0 TEST ->f / .
I K- M^
1
4T — BEFORE ATTRI
^T — MWK MODIFIED
Ml — AVCO MODIrlt
*
WK TEST
TION TEST
STD. OF JND. Tg
D AAJA TEST —
LSI. — 60 MSN. TEST
60 MIN. TEST
**£: *
I02
0.01
O.I
10 50 90
PERCENT LESS THAN STATED SIZE
99
99.99
-------
10*
TYLER STANDARD
SCREEN SCALE
- 8 MESH
to
1
FIGURE 7
SIZE DISTRIBUTION OF ALKALIZED ALUMINA
--X
ACTIVATED GRACE NO 2 — BEFORE ATTRITION TEST
ACTIVATED PETER SPENCE «— BEFORE ATTRITION TEST
ATTRITED GRACE NO 2 — MWK MODIFIED STD. OF IND. TEST — 60 MIN TEST
ATTRITED GRACE NO 2 — AVCO MODIFIED AAJA TEST — 60 .MIN. TEST
ATTRITED PETER SPENCE — MWK MODIFIED
STD. OF IND. TEST - 60 MIN. TEST
- 10 MESH
- 12 MESH
- 14 MESH
= 16 MESH
.20 MESH
ACTIVATED PETER SPENCE
ORIGINAL
AVCO ATTRITED
GRACE NO 2
. s 28 MESH
MWK ATTRITED _j;
PETER SPENCE
- 35 MESH
ACTIVATED GRACE NO 2
MWK ATTRITED
GRACE NO 2
o.ot
O.I
99.9
99.99
-------
used. Subsequent to these tests, the AVCO modified AAJA test
was adopted as the standard test.
The necessity of obtaining a particle size distribution
of the attrited material is illustrated by Figures 6 and 7. For
all three materials tested, the curves drop off sharply below
about 14-18 mesh, as previously noted, but the -16 mesh fraction
is still higher than the -60 mesh fraction; e.g., 38% vs. 26.7%
,'
in the AVCO test. Note that mesh sizes referred to are U. S.
Sieve Series. Thus reporting only the -16 mesh material, as
is presently done, would indicate significantly higher losses
than obtained based on -60 mesh material. If the attrition
characteristics of the sorbent were such that a flatter curve
were obtained, the difference in the -16 and -60 mesh fractions
would be greater. In fact, the shape of the particle size dis-
tribution curve might be used as a guideline in sorbent develop-
ment work; i.e., the desirable properties or processing conditions
would be those that tend to produce a size distribution curve
with a flat slope.
Visual observation of the attrited material revealed
that the -60 mesh fraction produced in the test was actually
much smaller than 60 mesh, so a particle size distribution was
obtained on the MWK attrited -60 mesh material and the complete
distribution curve is shown as Figure 8. The above two observa-
tions, viz., 1) fines much smaller than 60 mesh, and 2) above
about 16 mesh the attrited and original material are essentially
the same, lead to the conclusion that most of the fines produced
come from some particles that completely disintegrate while other
particles remain essentially unchanged. Thus, it would seem
that the sorbent might be a mixture of particles with significantly
different physical properties. Therefore, a 2-stage attrition
test is indicated to determine if the present sorbents are
indeed mixtures of particles having attrition resistance ranging
-43-
-------
FIGURE 8
SIZE DISTRIBUTION OF LABORATORY ATTRITED. SORBENTS
3 -
10s -
TYLEP
MWK MODIFIED STD. OF IND. TEST - 60 MIN. TEST
STANOAf
to
SCREEN SCALE
\f~
- 6
- 7
- 8
- 9
- 10
- 12
- 14
~-|6 -
- zo
- 1/16*
- 1/32'
- 1/64"
^•"•*"
r/V
f ?
I '
!
I*
•
I i
n
'
*£^
^^^^
PETER SPENCE — 2/ /£— GRACE NO. Z
A
J
//
ff(.
A
r
f
/ /
w
w
. 99.4% -
^
\**L^i<7C
100% - e
f UftiU
-^•*"_-, •
*• — ""
MESH
'
f —0— PETER SPENC6 — ACTIVATED
--&•- GRACE NO 2 - ACTIVATED
1 10 50 9
D a
a OQ a
PERCENT LESS THAN STATED SIZE
-------
from very poor to very good. The 2-stage test would comprise
running a standard attrition test, screening the attrited
material, and re-running the standard attrition test on the
coarser fraction, say + 16 mesh material. If the attrition
loss from the second test is lower than from the first test,
it would indicate a significant difference in physical pro-
perties among the different particles. In this case, the
sorbent manufacturing process could be revised whereby an
attrition step with fines recycle is included to produce a
more attrition resistant material as the final product.
Alternatively, the commercial plant could be designed such
that standard cyclones would recover tne fines (normally
100% recovery down to about 40 microns is obtained) which
could be re-balled and recycled back to the alkalized
alumina plant.
e. Methods to Improve Attrition Resistance of Alkalized
Alumina
Pilot plant evaluation of commercial sorbents subjected
to several cycles of intermittent operation—sorption followed
by regeneration—and also subjected to continuous pilot plant
operation indicated that attrition losses were greater than
acceptable for plant operation. Physical strength decreased
during activation and operating techniques to significantly
reduce physical degredation were not found. Consequently
development programs were initiated for the purpose of finding
a method of preparing an attrition resistant sorbent which still
retained its high activity.
(1) Prediction of Pilot Plant Performance from Laboratory
Testing
Since poor performance of the sorbents used in the
pilot plant had prevented correlation of laboratory attrition
test results with steady-state plant losses, it was decided to
-45-
-------
operate the pilot plant using a highly attrition resistant
inert catalyst support. The support, 1/16-inch diameter
porous alumina spheres, was purchased from Universal Oil
Products Corporation. Since S02 was not being sorted, air
rather than reformed natural gas passed through the regen-
erator and the regeneration level was controlled to give
three hours residence time. Attrition loss was determined
by periodically discharging and weighing the solids. At
each draining a representative sample was collected and
material was added to restore the original inventory.
Pilot plant operation was terminated after
95 sorption-regeneration recycles. Attrition loss averaged
0.12? of inventory per cycle. USBM states that part of the
loss was due to operation of blast gate valves below the
sorber and regenerator. Particle size distribution, Figure 9,
indicates that steady-state had been reached after about 44
cycles.
In previous pilot plant experiments attrition
loss was measured for Kaiser No. 1 and Grace No. 2 sorbents.
Even though steady-state values were not obtained, the initial
measurements were made at operating conditions comparable to
those used with the inert support. At these conditions, Grace
sorbent attrited at 1.8% per cycle and the Kaiser sorbeut at
2.3% per cycle. Attrition resistance of retained samples of
these sorbents and of the inert support were measured in the
laboratory using USBM and NAPCA accelerated air jet prc eaures.
It should be noted that the NAPCA test is the AVCO modified
AAJA test previously described while the USBM tests represents
a further modification to permit the use of a smaller sample.
A comparison of these two methods is shown in Figure 10. The
relationship of USBM and NAPCA laboratory-attrition loss with
pilot plant loss at approximately similar operating conditions
-46-
-------
0.10
.08
.06
tf -04
UJ
UJ
5
o
V-
cr
2
.02
.01
.008
.006
.005
1 I
O
o
A
A
O
O
o
At stort of test
After 12 cycles
After 24 cycles
After 44 cycles
After 55 cycles
After 70cycles
After 95 cycles
I
1
I
I
I
I
0.01 0.1
I 10 30 50 70
WEIGHT PERCENT SMALLER THAN
90
98
Figure 9-Porticle size distribution during pilot plant operation using a spherical alumina
catalyst support.
3-14-69
L-11149
-------
70
0 10 20 30 40
NAPCA ATTRITION LOSS.percent per hour
.FIGURE 10 -Attrition loss os measured by USBM and NAPCA
tests
3-13-69
LHII50
-48-
-------
is shown in Figure 11. It should be noted, however, that in
tests with active sorbents, USBM had used a 10/1 sorber recycle
ratio while Figure 11 indicates a recycle ratio of only 2/1.
Since about the same pressure drop was obtained in both cases,
the sorbent density in the reactor was the same and SO- re-
moval should be satisfactory. If a higher recycle ratio is
found to be necessary to obtain adequate SO2 removal, then a
greater attrition loss is likely. Despite this possible
deficiency, however, the data are quite useful in showing the
relationship between laboratory and pilot plant results. Also,
the data indicate that the attrition resistance of the UOP
catalyst support is in the range that will be required for a
commercially acceptable sorbent.
(2) Impregnation on UOP Catalyst Support
The high porosity of the UOP support suggested
that an active, attrition resistant sorbent might be prepared
by impregnation, and consequently sodium aluminate and copper
sulfate both are presently being evaluated as impregnants for
the UOP support. USBM has observed that regeneration at standard
conditions has been incomplete, resulting in loss of sorption
capacity during subsequent cycles but more study involving this
sorbent preparation is needed before any definite recommendations
can be made.
(3) Binder Strengthened Alkalized Alumina
(a) USBM Fillers and Mortar Types
Addition of binders was also studied as a
means of improving attrition resistance of alkalized alumina
filter cake. USBM tried fillers and mortar type materials
which do not require heat treatment to set. Preparation and
test procedure consisted of blending the binder with wet
filter cake and then forming approximately 3/32 inch ex-
-49-
-------
2.5
S2-°
CL
1.5
en
3
O
e u>
o.
fe -5
Koiser No I
I
• NAPCAtast
O USBMtest
Pilot plont operoting conditions -
Feed-200 Ib/hr
Recycle ratio*-2 to I
Particle size ~ 10 to 14 Tyler mesh
i
0 10 20 30 40 50
LABORATORY ATTRITION LOSS, percent
FIGURE 11-pilot plont sorbent loss predicted from lobo w«ory tests
3-13-69
Li f> i\ ffr t
-h J J1!
-50-
-------
trudates which were dried in air at 220 to 230°F. After
normal activation, the sorbents were exposed to flue gas
in the pilot plant for 24 hours and then regenerated with
hydrogen for 10 hours at 1200°F. Binders added to alkalized
alumina filter cake have improved its resistance to degrada-
tion, but long term stability and optimum binder concentra- .
tion are not yet known. Three materials were found which
need further study. One contains asbestos fibers, while the
other two contain a chemically setting potassium silicate
mortar. The best sample (USBM-CMBI) had an attrition loss
of 19.4% USBM basis ( 12% by NAPCA/AVCO test) after 2 sorption-
regeneration cycles.
(b) W. R. Grace Studies
Grace investigated sodium aluminate bound
with kaolin or meta-kaolin and concluded that the attrition
resistance was not satisfactory. Similar results were obtained
with cement bound Dawsonite samples for curing conditions of
100% steam up to 700°F. Also, variations in attrition char-
acteristics of different lots of Dawsonite bound by kaolin and
sodium silicate were obtained as can be seen in Table 4.
Material from the fifty pound sample, CD-131, was subjected
to 84 cycles in the portable field test units. Attrition
of this material remained constant at 6% (AAJA 2nd degree
NAPCA or AVCO test) during this period. Calculations in
conjunction with Na2O and CC^ analyses show a possible dif-
ference in the Dawsonite base material from spray dried lots
1 and 2 and Grace states that X-ray confirms this possibility.
However, further data are needed to correlate this observation
with forming characteristics and attrition behavior.
(c) NaAlO? Pellets Prepared from BSAC Intermediates
A series of attrition tests were run in order
*
to evaluate various BSAC (Basic Sodium Aluminum Carbonate inter-
mediates) materials prepared at the Albany Metallurgy Research
-51-
-------
TABLE 4
VARIATION IN ATTRITION RESISTANCE
OF
RL-175 COMPOSITION ALKALIZED ALUMINA
(DAWSONITE BOUND WITH 23% KAOLIN AND 2% Na
DAWSQNITE BASE
SAMPLE POUNDS SINGLE CYCLE EVALUATION
LOT WO. MADE WT % S SORBED ATTRITION
1st Spray Dried Lot
2nd Spray Dried Lot
2nd Spray Dried Lot
CD-131
CD-132
CD-135
50
200
4000
7-8
6-8
6-8
6-10%
65-75%
75-80%
Ref:
W. R. Grace and Company, April 1969 Monthly Report, Phase I,
Table 1.
-52-
-------
Center. Modified 2nd degree AAJA test parameters were a 60
mesh top screen, 3.45 SCFM air velocity, 30 gram sample weight
and a 1 hour test time, these conditions are identical with
those used in the NAPCA/AVCO test. The samples represented
materials having crushing strengths from approximately 1 to
20 pounds. After each 1 hour test the samples were replenished
to 30 grams total weight. Results of these attrition tests
are shown in Figure 12. The following discussion is from a
USBM-Albany report and is believed to adequately describe
the results.
The apparent decrease in attrition loss
over the first three runs as displayed by the stronger
materials was expected. The pellets being made on a laboratory
scale at AMRC are right cylinders with length to diameter ratios
of approximately 1. Therefore, the corners wear off rapidly
during the initial attrition exposure and the true attrition
resistance of the materials cannot be evaluated until the
right cylinders are eroded to approximate spheres. This is
indicated by the leveling off of the attrition loss curves
shown in Figure 12.
However, the weaker sorbent materials (plots
#1, #2, and #3) displayed attrition loss behavior completely
different than previously discussed. The material represented
by Plot #1 was spherodized before the attrition tests were run
and the W. R. Grace material (plot #3) was in the form of
irregular spheres. Both material displayed increased attrition
losses following the initial one hour test. The attrition
behavior of the weaker sorbent materials is attributed to a
fracture mechanism instead of attrition.
The results of the above tests show that
modifying the basic structure of the alkalized alumina pellets
-53-
-------
c
o
0
t
CO
CO
o
o
h-
oc
h-
/ Blend (90% 80 A* 10% 60 A
AMCS-BSAC)
2 B'end (90% IOOA*IO% 50 A
AMCS-BSAC)
J Grace No. 2
4 90A AMCS-BSAC + app. 50%
Si02
5 Blend (60% IOOA«-40% 50A-
AMCS-BSAC)
5 Blend (70% 100 A + 30% 50A
AMCS-BSAC)
7 Blend (80% IOOA + 20% 50A
AMCS-BSAC)
6 90 A AMCS-BSAC+app. 25%
Si02
UOP AI203 (standard)
NUM B
RITIOMIMG F?LJrV/^ fl hi r S r u n )
-------
provides improved attrition resistance in some cases, at
least on the activated material. Additional tests are re-
quired to determine if these improvements are obtained after
the sorbent has been subjected to multi-sorption/regeneration
cycles. It can be seen from Figure 12 that the UOP alumina
sample exhibited very low attrition losses which is consis-
tent with the results of the Bruceton tests.
2. Sorption
a. Sorbent Selection
In 1956 the Department of Health, Education, and
Welfare initiated a study with the United States Bureau
of Mines (USBM) to screen existing processes and examine
the possibility of developing hew materials or processes
for the removal of sulfur dioxide (SC^) from flue gases.
After long and intensive study a new material (alkalized
alumina) was selected. In the intervening years, the
alkalized alumina chemisorbent was tested and developed
to the extent that a full scale process study was needed
to develop process design for a commercial size plant
based on this sorbent, with the objective of determining
technical and economic feasibility.
b. Sorption Chemistry
(1) Sorbent Preparation
Although many methods have been developed to
produce alkalized alumina, the basis of all these processes
is the production of Dawsonite, NaAlCCK (OH)2, as the inter-
mediate product. The production of Dawsonite by the original
Bienstock patent1 is typical:
Bienstock, D. and Field, J. H., Process for the Removal of Sulfur
from Gases, U. S. Patent 2,992,884, July 18, 1961.
-55-
-------
2A1(OH)3 + 3H2S04 + 3.6Na2C03 * 2NaAl(CO3) (OH)2
+ 2.6Na2S04 + 1.6H2C03 + 0.4H2SO4 -I- 2H20
The solid Dawsonite formed by the reaction is
filtered, washed to remove sulfur (S04), and dried. The
dried solids are then activated by heating in a 600°C reducing
atmosphere to remove any sulfur remaining in the Dawsonite
intermediate, and to drive off carbon dioxide and water
thus leaving the basic sorbent, sodium aluminate, NaAl02:
NaAlC03(OH)2 h£at> NaAl02 + C02 + H2O.
The solid product of the reaction, sodium aluminate, comprises
what is commonly known as alkalized alumina.
(2) Sorption Reaction
There are two chemical reactions which have
generally been accepted as representing the sorption:of SO2
by alkalized alumina. These reactions are as follows:
1. Sodium aluminate reacts with the sulfur
dioxide in simulated flue gas (does not
contain NOX) to form sodium sulfite and
aluminum oxide.
2NaA102 + S02 -> Na2 S03 + M203
2. Sodium aluminate reacts with sulfur dioxide
and oxygen in the flue gas to form sodium
sulfate and aluminum oxide.
2NaAl02 + S02 + 1/2 O2 •* Na2SO4 + A12O3
As well as being influenced by NOV, these
X
•eactions are highly temperature dependent. At 120°C in the
-56-
-------
absence of NOX a combination of the two reactions occurs
whereas at 330°C, the design temperature used in the sorption
process, or in the presence of NOX, only the second reaction
occurs. These conclusions have been verified by the analysis
of the reaction products leaving the various laboratory and
pilot plant sorbers.
The data in Table 5 (which were obtained using
a simulated flue gas) and the following discussion are taken
from USBM Report #RC-1331:
The data in Table 5 show that as temperature
is increased from 130° to 215°C the sorbed S02 tends to convert
from sodium sulfite to sodium sulfate and that at 300°C it is
all converted to sodium sulfate. However, whether the oxidation
of the sorbed SO- from sulfite to sulfate takes place before
sorption or after sorption has not been determined as yet. The
pellet loads obtained in 0.045 percent S02 flue gas show that
insufficient SO- has been sorbed in two hours to convert the
primary starting material NaAlO2 to either Na-SO, or Na-SO.
(Form III). As the SO2 gas concentration is increased
the amount of SO2 sorbed per unit time increases to nearly
the theoretical load, converting all the HaAlO- to either
sulfite or sulfate.
An additional point of interest as to the
exact sorption mechanism is that as the NaA102 is converted
to either Na2SO3 or Na2S04/ the alumina formed during the
reaction, as indicated by the equation, is not shown by x-ray
diffraction. When the loaded pellets are heated to 1,050°C
the alumina crystallizes sufficiently and
2NaAlO2 + S02 + 1/2 02+ Na2S04 -I- A1203
has been identified as iota-alumina. Whether or not the
amorphous alumina is chemically present in the Na2S04 has
-57-
-------
TABLE 5 _ x-ray Diffraction Results on Loaded Grace Alkalized-Alumina
CO
i
SO2, gas cone.
percent
0.049
.045
.045
.045
0.34
.34
.34
.34
2.01
2.04
2.07
1.98
10
10
10
10
Temp.
°C
130
215
300
400
130
215
300
400
130
215
300
400
130
215
300
400
Pellet
Load,
0.039
.054
.032
.029
0.130
.125
.097
.069
0.138
.140
.142
.136
0.119
.130
.142
.142
Primary
Na2Al204
Na2Al2O4
—
Na2S03
Na2SO3
Na2SC>4
Na2SO4
Na2SO3
Na2S03
N32SO4
Na2SO4
—
Na2SO3
Na2SO4
Na2SO4
Na2SO4
X-ray Diffraction
Secondary Minor Trace
Na2S04
Na2SO4
Na2S04
—
Na2SO4
Na2Al2O4
Na2SO4
Na2Al2O4
Na2S04
^_
Na2Al204
^a2Al20,
Ha2AloO,
NB2A12O,
I/ Gas-pellet contact .2 bours
Taken from USBM RI 7275, Table 7.
-------
not been determined at this time; however, additional work
is being conducted to identify the exact reactions that take
place and to determine the end products formed.
(3) Thermochemistry
Using the heat of formation data reported in
the J.A.N.A.F. Thermochemical Tables (see Table 6), the
heat of reaction, AHr, for the sorption of S02 at 330°C
(626°F) is -207, 350 Btu/lb mol S02. Using this value for
AHr and neglecting the heat picked up by the sorbent, the
adiabatic temperature rise of the gas in the sorbers is
calculated to be approximately 55°F. This temperature rise
is based on the gaseous effluent of a 1000 megawatt coal
fired power station using a 3% sulfur coal and the reaction
proceeding to the sodium sulfate form.
In the commercial designs this heat of reaction
will be used to compensate for heat losses in the system. There-
fore, the value of AHr should be verified by experimental data
to substantiate the assumption of an essentially isothermal system
which has been made in these commercial designs. At the time
of this writing, however, no data have been received to verify
the theoretical value calculated from J.A.N.A.F. data.
(4) Sorption Kinetics
In the early stages of experimentation the USBM
at Bruceton, Pennsylvania investigated the sorption kinetics
of alkalized alumina in two different types of gas-solids con-
tractors, namely a fixed bed, and a dilute phase falling bed.
A significant observation was the difference between the rate
constants as calculated for the fixed and falling bed reactors.
The constant associated with the falling-bed was greater than
that for the packed column.
-59-
-------
TABLE 6
HEATS OF FORMATION DATA1
MATERIAL
NaAlO,
600°K*
-271.568
-400.304
-330.874
-72.824
@ 700°K*
-271.436
-400.098
-330.386
-73.206
lf @ 626°F**
-271.564
^400.298
-330.859
- 72.835
1
*
**
Reported in Kcal/mol.
From JANAF Thermochemical Tables.
By linear interpolation.
-60-
-------
The difference is not entirely unexpected since
there is no reason to assume that the rate controlling mechanism
is the same in both cases. The effective solid density in the
falling bed is much less than in the fixed bed, so that possibly
gas diffusion is controlling in the former system. However, the
higher rate constant in the falling bed, together with the fact
that the fixed bed performed at a greater efficiency than pre-
dicted in the initial stages of operation, is significant. One
plausible explanation is that in the early stages of gas-solids
contact the reaction of SO, on the surface of the particle is
controlling. Later when the surface is covered with reaction
products, diffusion into the particle (which occurs at a much
slower rate than the surface reaction) becomes controlling. It
is this slower diffusion step which is important in the long run
behavior of the fixed bed reactor.
In the falling bed experiments the solids residence
time is very small (on the order of 15 seconds to a few minutes)
as compared to exposure time in the fixed bed system of 200 to
300 minutes, and fresh sorbent is constantly fed to the falling-
bed sorber while the sorbent in the fixed bed sorber is constant-
ly being loaded with sulfur. Thus it is conjectured that in
the falling bed contactor only the rapid surface reaction is oc-
curring, which in turn helps to explain the larger rate constant
calculated for this system.
In an attempt to test this hypothesis, some theo-
retical curves have been computed for the fixed bed reactor
assuming (1) only the rapid surface reaction is taking place,
(2) pore diffusion is controlling. These curves are shown
in Figures 13, 14, and 15. The curves indicate that the
hypothesis is correct. The initial sorption reaction is
accurately predicted by the assumption of a gas film
-61-
-------
4,000
Rapid surface
reaction
I
I
—Predicted
O Experimental
2 inch bed
3300C
1
50
100 150 200
TIME, minutes
250
300
FIGURE 13 -
Experimental removal of sulfur dioxide and predicted removal
uhen surface reaction is controlling and when diffusion is
controlling in a 2-inch packed bed of alkalized alumina.
10-8-65
L-9O78
-62-
-------
4,000
Rcpid surface
reaction
—Predicted
O Experimental
3 inch bed
330°C
tOO 150 200
TIME, minutes
250
300
FIGURE 14-ExperimentaI removal of sulfur dioxide and predicted removal
when surface reaction is controlling and when diffusion is
controlling in a 3-inch packed bed of alkalized alumina.
10-8-65
L-9079
-63-
-------
4,000
E
(X
sx
3,500
3,000
S 2,500
_j
u.
u,
UJ
? 2,000
UJ
o
x
o
o l,500|—
U.
_J
Rapid surface
reaction
— Predicted
O Experimental
I.OODh-
500h-
50
100 J50 200
TIME, minutes
250
300
FIGURE 15 - Experimental removal of sulfur dioxide and predicted removal
when surface reaction is controlling and when diffusion is
controlling in a 4-inch packed bed of alkalized alumina.
10-8-65
L-9080
-64-
-------
controlling mechanism while the data after the initial sorption
are supported by a pore diffusion controlling model. A further
proof of the initial surface reaction can be seen in that the
duration of the reaction is a function of the bed height/ i.e.,
the higher the bed and thus the larger the surface area, the
longer the applicability of the gas film control. These figures
also show a definite transition region between the two mechanisms
with the region defined by the bed height investigated.
Although the rapid surface reaction would be pre-
ferable in reactor design, the constraints of attrition, regen-
eration and the relatively small surface (excluding pore surface)
of the sorbent prohibit the use of this mechanism as the basis
for designing the actual process unit. The sorption reaction
is limited, therefore, to the relatively slower pore diffusion
mechanism. However, this does not demean the quality or
capacity of the sorbent, for it remains superior to the other
potentially commercial sorbents tested.
In the preliminary sorption studies most of the
investigators produced a synthetic flue gas by mixing the
appropriate major flue gas constituents of N,, 03, CC^, S02,
and H^O. However, the AVCO Corporation found that oxides of
nitrogen (NOX), which are present in small quantities in
actual flue gases and had been omitted by the earlier investi-
gators, improved the sorption rates of SO? on alkalized alumina.
Consequently, an evaluation of the sorption rates of SO- in
the presence of either nitric oxide, NO, or nitrogen dioxide,
NO2, was undertaken by the USBM at Albany, Oregon, to determine
the extent of this improvement.
Their results show that either nitric oxide (NO)
or nitrogen dioxide (N02) increased the sorption rate such
that a 0.10 gS/gP load was obtained in about 33 minutes as
compared to about 120 minutes for the standard runs made
-65-
-------
without nitrogen oxides (NOX) present. The actual sorption
rate obtained with NO., under standard test conditions (300°C,
*»
0.34% S02, 0.10 gS/gP) was 5.05 xlO~5 gS/gP/sec, whereas the
rate without NO was 1.44 xlO~5 gS/gP/sec. The effect of NOX
A
concentration on sorption rate was also determined by running
tests in which the level of NOX was varied from 250 to 2000
The results of these NOV tests are shown in Figure 16,
ppm.
x
which is taken from USBM RI 7275, pp 23-24, while the .standard
run results are shown in Figure 19, from page 17 of RI 7275.
The Figure 16 curve has been drawn through the 500 ppm points
and the data points for the other NO levels are shown. .The
A
results in Figure 16 show that while there is some variation
of sulfur load for each contact time, all of the points fall
along the 500 ppm line. Based on these results it is.-concluded
that, within the limits of the test conditions, the presence of
NOX increase* the SO2 sorption rate by 2 to 3.5-fold but the
concentration of NOX appears to have a negligible effect. Ad-
ditional tests were run at different SO- concentrations (0.045
to 10%) and similar results were obtained; i.e., sorption rate
increases of 2 to 3.5-fold over;those without NO .
The effect of temperature on sorption rate was
determined by running a series of tests at 215° and 130°C
with 500 ppm N0x in the flue gas. The results of these tests
are shown later in this report as Table 9 which is from a
USBM-Albany yearly report dated June 30, 1969. From Table 9
it is apparent that when the temperature is decreased not only
does the sorption rate decrease but the effect of NO on sorption
rate decreases also. For example, at 0.34% S02 and a pellet
load of 0.05 grams sulfur/gram pellet, the average sorption
rates for 300; 215; and 130°C are 7.94, 5.95, and 4.38 x 10"5
gS/gP/sec respectively, or a 1.8-fold decrease. The correspond-
ing ratio of sorption rates with NOV to those without NO,, are
« A
-66-
-------
0.18
0.16
0.14
FIGURE |6
EFFECT OF NOX ON S02 ^ORPTION AT 300°C
IN
0.34 PCT S02 FLUE GAS
0.34 PCT S02 AT 300*C WITH 500 ppm NOX
T . BARS REPRESENT ONE STANDARD
«» 250
> - 1,000
- 2,000
40 60 80
TIME, MINUTES
REFERENCE:
USBM RI 7275, Figur«» |4 ond 15.
100
120
-67-
-------
3.0, 144, and 0.78 or a 3.85-fold decrease. It is interesting
to note that the effect of NOX on sorption rate reverses at the
lower temperature; i.e., at 130°C, adding NOX causes the sorp-
tion rate to decrease.
USBM-Albany also did some work in an attempt to
determine the mechanism by which the presence of NOX increases
the sorption rate. The results are discussed in RI 7275, part
of which is reproduced below, and support the theory that NQX
increases sorption rate by increasing the oxidation of 602 to
S03 or Na2SC>3 to Na2SO4.
The increase in SO2 sorption rate in the presence
of NOX is believed to result from increased oxidation of S02
to SO3 or Na2SO3 to Na2S04. To substantiate this theory X-ray
diffraction analyses were obtained on samples loaded at 130°,
215°, and 300°C. The results in Table 7 show that the SO2
sorbed formed Na2SC>4 at all three temperatures. A comparison
of these data with the results in Table 5, wherein a 130°C the
sorbed S02 formed Na2S(>3, shows that the nitrogen oxides are
actually improving the sorption rate by increasing either the
oxidation of S02 to 803 or Na2S03 to Na2S(>4. The reason that
the sorption rate with NO is the same as the NO2 rate is shown
by the following equation:
2NO + 02 + 2N02
To further resolve the question of the effect of
oxidation on sorption rate, tests were conducted wnerein the
0.34 pet S02 flue gas was passed over a V205 catalyst to convert
the S02 to SO3 before sorption onto the alkalized alumina.
Because of the low oxygen content in the flue gas only 60 pet
of the SO2 was converted to S03 at 500°C. Results of the SO
test are given in Table 7 and show that with the SO3 in the flue
-68-
-------
TABLE 7 - X-ray diffraction results for Grace alkalized alumina loaded
in the presence of nitror.i-n oxides
Test con-
ditions
250 ppra NO.
500 ppm NO.
1,000 ppm NO.
2,000 ppm NO.
500 ppm NO.
500 ppm NO.
500 ppm NO.
500 ppm NO,,
500 ppm NOS
60 pet S03 . . .
St.Tidard run.
S02 con-
centration,
pet
0.34
.34
.3*
.34
.34
.34
.34
.34
.34
.34
.34
Temper-
ature,
0 C
300
300
300
300
300
215
130
300
500
500
500
Pel let
load,
F.S/«pi
0.157
.158
.158
.170
.152
.150
.113
.150
.080
.136
.092
X-rav diffraction analvsos
Primary
Na_SO.,
Nal,S04
Na-,50.,
Na?S04
NasS04
Na2 S04
NagS04
NaaS04
Na- S04 ,
Na8Al.,0.i
Na.S04
-
Sec-
ondary
-
-
-
-
-
-
-
.
-
-
-
Minor
-
-
-
-
-
-
-
Naj,Al204
NagAKSO.)..
Na-jAKSO* )3
-
Trace
-
-
-
-
-
-
Na2Al204
NaN03
-
-
-
1Gac-pellct contact time 2 hr.
Taken from USBM RI-7275, Table 9.
-69-
-------
gas, Na2S04 is formed and a pellet load of 0.136.gS/gP was
obtained in 2 hr. The sorption rate for a pellet load of
0.100 gS/gP was 3.9 x 10~5 gS/gP/sec which is less than the
rate of 5.05 x 10~5 gS/gP obtained for the standard runs using
NO . It should be noted, however, that the SO3 run was at
Jt
500°C instead of 300°C and as shown in Table 7, the .runs with
and without NOX at 500°C have lower pellet loadings (henqe
lower sorption rates) than at 300°C. For example, the >5QO
ppm N02 run shows a pellet loading of 0.150 gS/gP at 300°C
and only 0.080 gS/gP at 500°C, while the standard run,.without
N02 achieved a loading of 0.092 gS/gP at 500°C. Based on JOie
data in Table 7 it appears that S03 is sorbed faster than SO2
even when NO is present but no firm conclusions can be .drawn
X
because insufficient data are presented to fully .define the
temperature effect.
Further discussion of the NO effect will .be
contained in section 2-b-(5) of this report. Similarly,
the actual sorption data and eventual modeling of the
kinetic data will be presented in separate sections of this
report.
(5) Extraneous Chemical Effects
Aside from the basic sorption reactions mentioned
earlier, various other side reactions influence the -sorption
of SO2 on alkalized alumina. The chemicals causing these side
reactions were found to be:
(a) Chlorides
The reaction of chlorine compounds found
in flue gases has given some cause for concern throughout the
development of the sorbent. Experimental work by the Central
Electricity Generating board of Great Britian, C.E.G.B.,
showed that a synthetic flue gas containing hydrogen chloride
-70-
-------
and SO2 led to a rapid reaction of chlorine on to the sodium
aluminate. This chlorine was not displaced during the regen-
eration process so that the build-up was cumulative. Once
the chlorine concentration reached about 10% it severely reduced
the sorptive capacity of the sodium aluminate, while at about
12% chlorine, the sorbent became completely inactive toward
S02. Work at the USBM-Bruceton also confirmed this cumulative
build-up under laboratory conditions.
However, results from the C.E.G.B. con-
tinuous recycling rig at Blyth which operates on actual flue
gases, showed that the behavior of chlorine was completely
different from that in the laboratory. There was a small
uptake during the first few cycles, after which time no
further chlorine was retained on the reactant. The reasons
for this are still not certain, but it is possible that the
oxides of nitrogen may be important in oxidizing SO2 to 803.
It may be that sulfur trioxide can displace hydrogen chloride,
while sulfur dioxide is incapable of making this replacement.
Thus the catalytic oxidation of sulfur dioxide on the surface
of the reactant could be the reason why chlorine is not trouble-
some when power station flue gases are used. Regardless of the
actual mechanism, the important point with respect to process
design is that chlorine does not deactivate the sorbent.
(b) Water
The effect of water or steam concentration
on the sorption of SO2 by alkalized alumina is shown in Table
8. These data, which were obtained by the W. R. Grace Company,
indicate that a minimum 2.5% steam concentration, comparable
to the levels obtained in a power plant effluent, is necessary
during sorption to maximize SO2 pickup. A similar effect on
total capacity was also shown in work by the USBM. Tne precise
-71-
-------
TABLE 8
EFFECT OF WATER VAPOR CONCENTRATION ON SO^ ABSORPTION RATE1
H2O Vapor
Concentration
VHSV
Breakthrough
Time
Wt. % Gain
@ Break
Wt. % Gain
@ 150 min
0%
1%
2.5%
5.0%
4000
4000
4000
4000
41 min.
50 min.
86 min.
90 min.
3.95
5.02
8.27
8.69
7.47
10.16
12.86
13.34
W. R. Grace Co., Interim Report 10-31-67, pg. 22.
-72-
-------
role of water vapor in the sorption mechanism is unknown;
however, the data clearly show that its presence is required.
A detailed analysis of this effect by the AVCO Corporation
is contained in Appendix B, which, although this analysis
is not conclusive, provides further insight into the phenomenon.
(c) Oxygen
Despite the stoichiometric requirement
for oxygen in sulfate formation from sodium aluminate (see
section b-(2)), a direct demonstration of this requirement
was lacking. Consequently, W. R. Grace ran tests in which
the oxygen concentration was varied and the results, which
are plotted in Figure 17, show that oxygen concentrations
less than about 1% significantly decrease sorption rates.
Comparison of the 1% oxygen level with the standard synthetic
flue gas composition containing 4.5% oxygen shows little
change in sorption efficiency at the higher oxygen level.
The importance of oxygen in the nitrogen chemistry will be
discussed further in the next section of this report.
(d) Oxides of Nitrogen
During the early sorption studies
laboratory tests were run with synthetic flue gases, which
did not contain nitrogen oxides. In later pilot plant
testing it was discovered that reaction rates were many
times higher when actual flue gases were used. Therefore,
a study was undertaken to determine the cause of these
higher rates.
It was discovered that the nitrogen
chemistry played a vital role in the sorption of S<>2
by alkalized alumina in that oxides of nitrogen apparently
catalyzed the sorption reaction.
-73-
-------
FIGURE 17
EFFECT OF 0, CONCENTRATION ON ABSORPTION
RATE OF STANDARD ALKALIZED ALUMINA AT 660*F
10
20
60
80 loo
Time In Minutes
•74-
120
140
-------
In the combustion reaction a small fraction
of the nitrogen in the combustion air and/or in the coal is
oxidized to nitric oxide, NO, or nitrogen dioxide, N02, and
these nitrogen oxides are present in power plant flue gases
from a range of about 150 to 1500 parts per million. It was
found that NO2 reacted with S(>2 to form sulfur trioxide,
S03, which is much more reactive with alkalized alumina
than is SO2. Further, thermodynamic calculations show
that NO can be oxidized to NO2 in the proposed sorption
temperature range, thus catalyzing the sorption reaction.
The chain of reactions involved in the
sorption of S02 on alkalized alumina may therefore be
written as follows:
(1) NO + 1/2 O2 -*• N02
(2) NO2 + SO2 -»•-•> SOj + NO
(3) 2NaAl02 + SO3 •+ N
Reaction (1) oxidizes the nitric oxide to nitrogen dioxide.
Reaction (2) oxidizes the sulfur dioxide to trioxide and
returns the nitrogen dioxide to nitric oxide for further
oxidation. The sulfur trioxide formed by reaction (2)
then reacts with sodium aluminate by reaction (3) to
form sodium sulfate and aluminum oxide. The above
reactions represent the probable intermediate steps in
the basic sorption reaction mentioned previously (Section
B-2) . The effect of the oxides of nitrogen is a catalytic
one in that they are involved in the reaction but returned
to their former state. For a detailed discussion of the
thermodynamic equilibria involved in the nitrogen oxide
chemistry see Appendix C of this report.
-73-
-------
c. Sorption Data
(1) Sources
For the development of the various phases of the
alkalized alumina process, the Public Health Service (PHS)
of the Department of Health, Education, and Welfare awarded
contracts to several different organizations. These organi-
zations are the USBM at Albany, Oregon and Bruceton, Pennsyl-
vania; W. R. Grace and Company; and the AVCO Corporation.
The PHS has also worked hand in hand with the Central
Electricity Generating Board (C.E.G.B.) of Great Britain
who at one time considered piloting an alkalized alumina
sorption system, sized to treat the flue gases from a 50
megawatt power station, at Blyth, Northumberland. The
following discussions will describe the areas in which these
investigators have been working as well as present some of
the pertinent data and conclusions reached by them in the
area of sorption.
(2) USBM - Albany
The USBM at Albany, Oregon has been prolific
in both the amount and variety of its data and analyses in
the area of alkalized alumina sorption. Initially, their
responsibility involved the determination of both the sorption
rates of S02 on alkalized alumina and the variables which
affect these rates. This initial work was done on the apparatus
shown in Figure 18.
A synthetic flue gas consisting of N2, 02, H.O,
and SO2 is passed through a thin bed (i.e., one layer) of
alkalized alumina pellets. The pellet charge, approximately
0.5 grams, is placed in a stainless steel container and sup-
ported on an 80 mesh stainless steel screen which insures
-76-
-------
I
-J
-J
I
Sample temperature
both
2-inch tube
furnace
i
S02, 0.34 pet
standard
5-inch tube furnace
(gas preheater)
Heat gun
Steam generator
FIGURE 18 -S02 Sorption System.
-------
uniform contact between gas and pellets. After sorption the
loaded pellets are removed and analyzed for sulfur. Since
the small sample removes only a negligible amount of S02
from the gas stream, the inlet and outlet concentration of
S02 in the gas is essentially constant thus permitting the
measurement of the amount of S02 sorbed from a gas of fixed
concentration. Individual samples are contacted with the
flue gas for periods of time ranging from 1 to 120 minutes,
analyzed for sulfur, and the pellet load in grams of sulfur
per gram of pellet determined for each contact time. A plot
of these data is then used to determine the average sorption
rate for a given set of test conditions. Standard test
conditions are: a reaction temperature of 300°C, a pellet
load of 0.5 grams of alkalized alumina, and a superficial
gas velocity of 0.18 ft/sec which corresponds to a 31,100
hourly space velocity. Using this technique, results such
as those shown in Figure 19 may be obtained.
By changing the reaction temperature in the
apparatus the effect of temperature on reaction rate was
determined and the results are shown in Figure 20. They
show that as the sorption temperature decreases .the sorption
rate and sorbent loading increases. It has been hypothesized
that for the 25°C test runs the temperature was below that
required to provide the activation energy for the reaction
to proceed. The results of this inverse temperature relation-
ship is not surprising since the data in Table 1 shov that at
the lower temperatures the sorption is to the sulfite form;
therefore, the diffusion of oxygen to the reaction site is not
required thus increasing the sorption rate. It was subsequently
found that the addition of oxides of nitrogen to the flue gas
reverses this effect.
78-
-------
Q
LU
CD
rr
o
CO
0.14
.1 2
r»
; .1 0
.08
.06
i
UJ
SOg concentration =0.34 pet
Flue-gas temperature = 300° C
0.0936
.1024
.1047^
.1004
¥ .04
0.0150
0.0338
.0334
.0301
.0294.T
AXl
*S 0.04
0.0843
.0698
.0769
.0771
0.0437
.02
°443
/
.0222
.0179
.0212
,
I
1
1
1
0
20 30 40
50 60 70
TIME, minutes
FIGURE 19 - Reproducibility of Sulfur Loading
on Grace Alkalized-Alumina.
80 90 100 110 !<
68-99
-------
0.'6
GO
O
10 20 30
FIGURE 20
50 60 70
TIME, minutes
.-Effect of Temperature on Sulfur
Loading in O.34 percent S02.
100 110 \2$
—100
-------
0.16
I
a
c/)
o>
en
•»
Q
LU
00
Q:
O
cn
h-
I
o
LJ
.06
.04
.02
0.045 pet S02
I I I I
I
I
0
10 20 30 40
FIGURE
50 60 70
TIME, minutes
21 -Effect of S02 Gas Concentration
on Pellet Loading at 300° C.
80 90 100 110 120
66-103
-------
In like manner, the Albany test work showed the
effects of S02 concentration on sorption rate. The results
in Figure 21, as expected, indicated that as SO2 concentration
is increased the sorption rate also increases.
Because of the changes in surface area, porosity,
and the amount of carbonate removed in activation of the
Grace alkalized-alumina with hydrogen at different temperatures,
sorption tests were conducted on samples of hydrogen activated at
300°, 600°, 700°, 750°, and 900°C for 4 hours. Results of
these tests are plotted in Figure 22 and show that the initial
sulfur loading was only slightly affected by activation tempera-
ture, while with longer contact times the effect of activation
temperature became very pronounced. Activation temperatures
between 600° and 700°C produced sulfur loads of about 0.028
gS/gP with 20 minutes flue gas contact, while at 300°C the
pellet load was only 0.010 gS/gP and at 900°C, 0.017 gS/gP.
The low sulfur loads at 300°C resulted from incomplete removal
of the carbonate' and water from the sorbent, while at 900°C
the low sulfur loads were caused by sintering of the sorbent.
Thus an activation temperature between 600° and 700°C is near
optimum for high sulfur loading, which indicates that essential-
ly all of the Dawsonite in the precipitate must be converted
to sodium aluminate to obtain an active sorbent.
To determine if activation time had an effect on
pellet loading with different contact times, sorption tests
were made using samples of Grace alkalized-alumina tint hsd
been activated at 600°C for 2, 4, 6, and 8 hours. Test results
are plotted in Figure 23 and show that the pellet loads ob-
tained for given contact times are nearly equal with respect
to activation time. Hence, an activation time of 2 hours at
600° to 650°C is sufficient to prepare the samples for optimum
sorption.
-82-
-------
I
00
CJ
Q.
O*
\
cn
o>
Q"
UJ
00
o:
O
en
i-
X
S2
UJ
0.040
.035
.030
.025
.020
.01 5
.0101-
.005
_ Flue-gas temperature = 300° C
SO.
1 1 1 1 '
concentration * 0.34 pet
200 400 600 800 1,000
TEMPERATURE, °C
FIGURE 22 .-Effect of Activation Temperature
on S02 Sorption- Grace Alkalized-Alumina.
• •-•7
-------
00
*fc
1
Q_
o>
en
Q"
LU
CD
tr
o
X
UJ
£
.035
.030
.025
.020
.01 5
.OIQ'
.005
! 1 1 1 1 1 1
S02 concentration = 0.34 pet
Flue gas temperature * 300° C °
- 20 min . —
0
14 min a
a
a
8 min o
O -
o
• A
,__^., ^ ^- y ^ ^ Q 5 ^jn y
i 2 3 4 5 M 6 7 8 9
ACTIVATION TIME, hours
FIGURE 23 -Effect of Activation Time at 60O° C
on SO2 Sorplion-Grace Alkali zed-Alum in a. ,
-------
When it became evident that nitrogen oxides
(NOX) played a major role in S02 sorption, tests were made
to determine how the N0x-free results would be affected. As
shown in Figure 16 and Table 9, sorption rates increased by
approximately 2 to 3.5-fold when nitrogen oxides were added
to the flue gas. When the sorption temperature was lowered
from 300°C to either 215° or 130°C, the sorption rates were
reduced (e.g. by about 25% for each step) thus showing con-
siderable difference from the results reported in Figure 20,
wherein the 100 to 130°C sorption temperature in the absence
of nitrogen oxides showed a significant increase in rates
over the 300°C data.
Not only did the USBM-Albany study the various
parameters which affected the sorption of S0~ from flue gases
containing nitrogen oxides, but they also provided the only
extensive sorption data for tests of this type. These data
are summarized in Table 10.
The USBM-Albany is also the prime researcher in
the area of fluidized bed sorption. Their fluid bed sorption
system test apparatus is shown schematically in Figure 24 and
is operated in the following manner. The system is brought
to temperature while the chromatograph and Leco titration
units are being standardized. Grace No. 1 pellets are added
to the sorption system to obtain the desired fluidized-bed
depth. The sorption unit is allowed 30 minutes for the bed
to equilibrate. To begin the run the SO2 and NO flows are
started and the pellet feeder is set to feed 10 g/min of
pellets to the system. Gas analyses are obtained at 10
minute intervals and a representative sample (loaded pellets)
is removed from the underflow for sulfur analysis every 30
minutes. The underflow of the system is sent to a regenera-
tion unit and is recycled as feed to the sorber after sulfur
\
removal.
»-85-
-------
TABLE 9 - Comparison of Sorption Rates With and Without NOV'
Time required to load,
minutes
Sorption Rate,
gS/gP/sec.
Sorption ratio
Sorption SCU cone., Pellet load, Without
temp., °C percent gS/gP NOX
300
300
300
300
300
, 300
£ 300
300
300
300
215
215
130
130
0.045
0.34
1.0
2.0
10.0
0.045
0.34
1.0
2.0
10.0
0.34
0.34
0.34
0.34
0.10
0.10
0.10
0.10
0,10
0.05
0.05
0,05
0.05
0.05
0.10
0.';5
0. '.0
0.05
470l)
116
48
16
2.5
1851'
31
12
5.0
1.25
71
20
47
15
With
NOV
1811'
30
16
5
2.5
82
10.5
6.0
1.75
1.25
42
14
62
19
Without
NOV
0.35
1.44
3.47
10.4
66.6
0.45
2.69
6.94
16.7
66.6
2.35
4.17
3.55
5.56
With with NOX to
NOX without NOV
0.92
5.05
10.4
33.3
66.6
1.02
7.94
13.9
47.8
66.6
3.97
5.95
2.69
4.38
2.6
3.5
3.0
3.2
2.3
3.0
2.0
2.9
1.7
1.4
0.76
0.78
1) Projected times required to obtain load.
2) From USBM-Albany, Yearly Report, June 30, 1969.
-------
TABLE 10
EFFECT OF NO0 AND NO ON SO5 SORPTION RATE - GRACE ALKALIZED ALUMINA
1)
NITROGEN OXIDE NITROGEN
CONCENTRATE -OXYGEN
SO,
CONCENTRATE TEMP.
PERCENT-' - °C
CONTACT TIME, MINUTES
16 30 60
WEIGHT SORBED, gS/gP
120
1
GO
-J
1
*" ~^^^^' .. I,,.,, " .tt>*^M*-.i
300
500
1000
150
300
300
300
500
500
500
•auBj^-c^- — ~~ - ,^r-
N02
NO 2
NO2
NO 2
NO 2
NO 2
NO 2
NO
NO
NO
• - • • ^f
.34
.34
.34
.34
.34
.34
.34
.34
.34
.34
• - ^
300
300
300
300
500
215
130
300
215
130
0.0177
.0122
.0126
.0135
.0221
.0130
.0135
.0175
.0178
.0180
0.0369
.0205
.0227
.0206
.0224
.0225
.0226
.0357
.0296
.0336
0.0484
.0379
.0334
.0428
.0372
.0425
.0429
.0646
.0517
.0387
0.0476
.0558
.0672
.0650
.0553
.0650
.0610
.0955
.0751
.0643
0.1006
.0923
.1040
.0993
.0618
.0928
.0844
.1302
.1161
.1080
0.1459
.1311
.1395
.0800
.1292
.1117
.1524
.1257
.1262
1) From USBM-Albany, Quarterly Report, June 30, 1968, pp BM-44
-------
Feed hopptr
00
00
I
Dull tiller
ottembly
Thermocouple well
Monometer
High temperature
gloii reoetor lubt
Tlisrrnocoupl* w«lt
Boll vol»«
Oil lurnoc* a»»«mbly
To bed dump To p»ll«t ditchorg*
centaintr container
|l| Fiu'dir«d-bed attembly
FIGURE 24. .-Fluidi*ed-Bed SO8 Sorpfion System.
Air Otpirolor
Oil-got
llowmeltr
Th«rm«ceuplt
•nil
-------
Using this apparatus a series of sorption runs
were made using a typical power plant flue gas as feed. The
initial series of 17 one day tests shown in Tables 11 and 12
indicated that on the average 78% of the inlet S(>2 was removed.
Furthermore, if runs 7, 11, 13, and 14 are discounted, due to
the high sorbent sulfur loadings which are far in excess of
the loadings used in the commercial design, the average SOj
removal becomes 85%.
The encouraging results obtained in the one day
test series were then confirmed by an extended sorption run of
100 hours. The data obtained during this run are shown in
Tables 13 and 14. For this extended run, the average removal
of SO2 was 90.2%. The conditions used during this run are
comparable to the proposed commercial design conditions; there-
fore, the present commercial fluid bed sorption scheme which
is designed to remove 90% of the flue gas SO, appears to be
feasible. The other data presented provide a basis for deter-
mining the commercial bed's physical properties (e.g., density,
expansion, pressure drop, etc.). The present design was made
before these data were available but there does not appear to
be any large discrepancies between the assumed and the experi-
mental values so no adjustments to reactor design are planned
at the present time.
(3) USBM-Bruceton
USBM-Bruceton discovered the alkalized alumina
sorbent and did the original experimental work using first
a bench-scale fixed bed apparatus and later a downflow reactor.
The data from the bench-scale fixed bed unit were used as a
basis for designing the first pilot plant which used a down-
flow reactor. The following discussion is taken essentially
verbatim from a Bruceton report.
-89-
-------
TABLE 11 - rlutdited-ted (SoTpttai Sytua) Tests of Oat-Pur Duratloa
I
vo
o
I
fcn
No.
S
4
i
6
7
9
9
10
11 '
12
13
14
IS
16
17 '
1« '
19 '
Km
No.
3
4
S
6
)
a
»
10
IV
a
13
14
IS
16
17'
ia'
19 '
fen
hrs.'
s.s
4.0
5.0
S.O
5.2
S.O
S.O
5.0
5.0
S.O
S.I
S.O
S.O
S.O
S.O
S.I
4.0
velocity •
Ft/sec
„
-
-
-
-
2.4
2.4
3.4
3.6
3.6
2.9
2.6
3.6
3.6
3.S
2.6
Height.
inches
S '
S '
S •
5 •
5 •
7.5 «
7. •
7. •
7. •
S.
.3
.5
.0
.0
.3
Cyclone Product
a*raM
14
S
17
14
13
27
17
14
16
40
31
19
10
17
164
148
64
i/hr."
0.4
0.2
0.6
0.4
0.4
o.a
0.4
3.2
0,2
».»'
0.*
»M
fl.J
£ v
t ',
1.7
2.0
1 Msw Grace No. 1 activated at 1
J Superficial spec* velocity at
* 4P represents pressure drop ec
• Fnccnx of tad «*14ht
1 tnroucnput'
0.5
0.2
0.6
O.S
O.S
0.9
0.6
O.S
O.S
1.4
1.1
9.7
0.3
0.6
S.O
S.I
2.2
Expansion,
inches
.
-
-
-
-
-
-
-
1-1/4
-
1-3/4
1-1/2
'/a
3/4
1
3/4
-
Feed (as
ISO,
-
-
-
-
-
-
0.20
0.12
0.31
0.34
0.47
O.S7
0.20
O.SS
0.21
0.22
0.24
Jfcrny
layer* ture
TOSS fluidities plate end
•eight,
iras
574
573
553
650
713
677
96a
1125
i4sa
ao4
730
as4
645
60S
73a
7M
aoo
• «»*••
0.12
0.13
0.22
O.M
0.45
0.14
0.20
.21
.sa
.33
.44
.64
.20
.39
.24
.25
.25
pellet bed
ted
Density
I/CC
.56
.56
.54
.63
.69
.44
.63
.73
.94
.74
.67
.75
.73
.M
.60
.64
.72
Off gai
I SO,
0.02
0.02
0.04
0.12
0.19
0.02
< .01
< .01
0.17
0.09
0.22
0.21
< .01
0.12
0.02
0.03
0.02
» sulfur'
4.5
4. a
7.4
a.s
9.9
4.9
5.7
7.1
10.1
a. 4
10.9
10.4
S.I
a.2
5.6
6.6
-
Gas
i* analysis ,
1 sorbed
ai
as
M
6a
sa
as
>94
>9S
SS
73
SO
67
>9S
69
92
tt
92
' ted jejrej&ted becai
• Concentre* ton calcu
' co£t«ntratlen by g-ai
I?- '
340
340
340
340
340
340
325
330
375
330
335
315
336
3SS
326
324
309
I sorted''
as
7a
aa
10
67
91
107
101
72
77
67
72
102
73
19
•9
-
Pallet
f inches bulk density,
H.O ' tVcc
4
3-1/4
3-1/2
4
4
4
4-3/4
4-3/4
7-1/2
4-1/2
4-1/4 .99
4-s/a 1.02
3-3/4 .as
3-1/2 .92
4-1/2 .M
4-1/2 .ao
4-1/1 .7a
".
j»e of Snsuff tclaat fluUisttim ntelty.
lated inm ml^ fas velocity •! : 50,' Isyot
> chroBitocraph en^Ireie
1 on total Om^iput.
-------
TABLE 12 - Evaluation of fluid-bed test* of one-Ay duration, for »teady-«tat«
Solution
Parameters fixed
U»d in. wo. gS/SP
Feed r»te, w, gp/Bin
ted wight. M, gp
S>2 fe*d, s, gS/Bin
S>2 concentration. yo
Results
Load out, ji, gS/gp
SOj reBaining, y,/yo (gas analysis)
SOj roBtining, f1/yo (pellet analysis)
Calculated
Load out, o, gS/gp
S>2 mining. yx/yo
Regeneration
Parameters fixed
Load in.no. gS/CP
Feed ret*. «. gp/ua
Dad Might, N, gp
Results
Load out, a, gS/gp
Rate constant, K, Bin"1
Sorption
Parameters fixed
load in, no. gVgp
feed rate, w, gp/min
Bed weight, W, gp
SO? feed, %. gS/ain
SOj concentration. y0
Results
Load out, u, gS/gp
S02 raaaining, yx/y0 (gas analysis)
S02 raining, Y^Yo (pellet analysis)
Calculated
load out, i, gS/gp
SO? roaining, y«/yo
Regeneration
Parameters fixed
Load in,uof gS/gp
Feed rate, w. gp/ain
led wight, N. gp
Remits
Load out, ii, gVgp
Rate constant, (., Bin'1
Run Ho.
3
0.0096
9.4
$10.
0.41
0.0012
O.OS7
O.U
0.04
0.049
O.U
O.OS6
7.J
367.
0.013
0.067
12
0.02S6
11.1
635.
1.11
0.0033
0.100
0.27
0.30
0.098
0.32
0.099
9.9
432.
0.027
0.064
4
0.0097
9.7
504.
0.41
0.0013
O.OS1
0.14
0.02
0.041
0.09
O.OS1
1.9
0.01S
13
0.0292
9.6
S31.
1.6S
0.0044
0.129
O.SO
0.42
0.112
0.52
0.124
10.6
361.
0.033
O.Otl
S
0.014*
10.1
4S1.
0.72
0.0022
0.077
0.16
0.13
0.06*
0.26
0.077
9.3
361.
0.020
0.07S
An
14
0.0327
9.1
633.
1.65
0.0064
0.136
0.33
0.39
0.126
0.4S
0.13S
10. S
305.
0.036
0.096
6
0.0203
9.1
512.
1.1S
0.0031
0.111
0.32
0.2S
0.091
0.34
0.111
1.2
392.
0.023
0.010
No.
IS
0.014
13.4
S63.
0.49
0.0020
O.OS4
0.07
- 0.10
O.OS3
- 0.07
O.OSS
11. S
632.
0.014
O.OS4
7
0.0261
9.5
536.
1.47
0.0045
0.129
0.42
0.34
0.111
0.4S
0.129
9.0
416.
0.029
0.077
16
0.022
11.2
412.
1.17
0.0039
0.09S
0.31
0.30
0.092
0.33
0.096
10.0
502.
0.024
0.060
1
0.0146
10.6
S94.
0.47
0.0014
O.OSS
0.1S
0.09
O.OS6
0.07
O.OSS
1.4
479.
0.01S
0.046
17
0.009
11. S
635.
0.73
0.0024
0.060
0.10
0.21
0.069
0.06
O.OS2
9.7
580.
0.009
0.076
9
0.0170
10.4
130.
0.47
0.0020
O.OS9
0.07
0.08
0.071
• 0.19
O.OS9
9.4
410.
0.017
0.056
11
0.001
9.9
6SS.
0.74
0.0025
0.064
0.10
0.2S
0.077
0.01
0.061
9.0
5S7.
0.009
0.092
10
0.01SS
10.0
926.
0.74
0.0028
0.081
O.OS
0.15
0.095
- 0.03
0.081
10.0
347.
0.021
0.082
19
0.013
9.1
648.
0.63
0.0024
0.067
0.07
0.23
0.080
0.04
0.067
9.3
347.
0.01S
0.094
-91-
-------
TABLE 13 - Plutdlied-ied (Sorption 9rttm) Test for 100 HPUTS
Bid
I
vo
to
Shift
No.
1-1
1-2
2-1
2-2
3-1
3-2
4-1
4-2
S-l
5-2
6-1
6-2
7-1
7-2
Bed
S-l
8-2
9-1
9-2
10-1
10-2
11-1
11-2
12-1
12-2
13-1
13-2
tod
Ti»e.
hours
2.5
4.0
4.0
4.0
4.0
4.0
4.0
4.0
4.0
4.0
4.0
4.0
l.S
3.S
-
2.5
5.5
4.0
4.0
4.0
4.0
4.0
4.0
4.0
4.0
3.5
3.S
-
Velocity
ft/sec
3.3
3.3
3.3
3.3
2.8
3.1
3.3
3.2
3.2
3.2
3.2
3.2
-
3.6
-
3.2
3.2
3.2
3.2
3.2
3.2
2.7
2.7
2.8
2.8
2.1
2,1
-
'. Height.
inches
6.3
6.3
6.3
6.3
6.1
6.8
6.3
6.3
5.8
S.8
5.5
S.S
-
-
-
S.S
S.S
S.I
S.I
-
-
S.S
S.S
S.S
5.5
6.3
6.3
-
Height', Density,
(rams s/cc
890
915
91$
890
1060
1170
860
800
715
685
685
615
715
74S
666
615
715
615
615
685
655
715
715
715
715
775
775
712
.69
.71
.71
.69
.76
.84
.66
.62
.60
.57
.60
.60
-
-
-
.60
.63
.57
.57
-
-
.63
.63
.62
.63
60
.40
.60
Analysis! Temp.,
IS * C
3.9
6.6
6.2
6.8
5.9
6.2
7.4
7.0
S.3
S.2
5.2
5.1
6.4
8.9
8.1
4.5
5.4
6.5
7.0
S.S
4.9
4.8
S.O
5.1
6.S
6.6
6.9
6.6
322
329
329
324
330
333
333
326
324
322
320
319
326
324
-
310
320
322
324
333
329
309
312
317
319
319
319
-
4P-
inches
H,0
4-5/1
4-3/4
4-3/4
4-5/8
5-1/4
5-3/4
4-1/2
4-1/4
3-7/8
3-3/4
3-3/4
3-3/4
4
4-1/1
-
3-3/4
S-7/I
3-3/4
3-3/4
3-3/4
3-S/l
3-3/4
3-3/4
3-3/4
3-3/4
4
4
-
Pellet
Feed Rate
c/min
9.1
9.2
10.8
10.5
9.0
10.2
9.8
10.0
9.5
10.1
10.8
10.1
10.2
9.4
-
20.
20.
16.
15.
13.
14.
IS.
14.3
9.9
10.0
9.7
9.5
-
Cyclone Product
gS/gP
xlOO*
4.4
8.0
7.3
1.1
6.9
7.5
9.2
8.5
6.1
5.9
6.1
S.9
7.4
11. S
.
4.9
6.2
7.7
l.S
6.4
S.7
5.4
5.1
6.9
7.1
1.0
1.3
-
grass
60
98
94
94
36
36
55
56
45
46
62
63
11
44
.
51
12*
56
57
51
52
41
42
30
31
22
22
•
»/hr*
2.7
2.7
2.6
2.6
0.1
0.1
1.6
1.1
1.6
1.7
2.3
2.3
1.7
1.7
.
3.4
3.3
2.0
2,1
1.9
2.0
1.4
1.5
1.0
1,1
0.1
O.S
•
percent 1
throughput
4.1
4.4
3.6
3.7
1.7
l.S
2.3
2.3
2.0
1.9
2.4
2.4
2.0
2.2
-
1.9
1.9
1.4
1.6
l.S
l.S
1.1
1.2
1.3
1.3
1.1
1.1
-
SOit
0.22
0.22
0.23
0.23
0.23
0.21
0.20
0.20
0.17
0.17
0.17
0.17
-
0.36
.
0.31
0.32
0.32
0.32
0.21
0.21
0.23
0.23
0.23
0.23
. 0.23
0.23
'•
Feed Gas*
SOii
0.27
0.27
0.21
0.21
0.26
0.24
0.22
.
0.20
0.20
0.20
0.21
0.34
0.41
.
0.40
0.37
0.37
0.37
0.24
0.24
0.26
0.26
0.27
0.»7
0.27
n.27
-
Off Gas*
SO, I
0.02
0.02
0.02
0.02
0.01
0.01
0.02
.
0.01
0.02
0.02
0.02
0.01
0.14
.
0.03
0.04
0.04
O.OS
0.03
0.02
0.02
0.02
0.02
0.02
O.OS
A.AZ
•
SO i
Feed
cc/Bin
500
500
501
501
440
440
444
444
361
368
361
361
172
172
-
692
692
692
692
444
444
440
440
440
440
431
431
-
Gss
analysts
I sorbed
.
93
93
93
-
96
91
-
95
90
90
91
-
66
-
-
19
19
17
M
92
92
92
93
93
19
93
-
1 Superficial space velocity at
' led weights were estimated fro» bed pre* -.•"•' i'-rop
" Sased on average sulfur analyses of pel< •? •>*erflow
' ft? presents pressure drop across fluidizi- ^ plate anil jsllet bed
Srara sulfur per irmm pellet
Percent of bed weight
Percent of feed rate
Concentration calculated fnm ait IBS velocity end i
Concentration by (ax
SO, input
-------
TABLE 14 - Evtliation of fluid-bad teat for 100 houn. steady-state operation
9iift Nueber
Son* ion
Load in. u*,. gS/gp
Feed rate, H, gp/au)
Bed weight, «. gp '
S02 feed, i, gS/ain
ST>2 concentration, yo
Results
Load out, u, gS/gp
S02 renaining, yx/y0, (gas analysis)
»2 renaining, yx/yo> (pellet analysis)
Calculated
Load out , M , gS/gp
SOj renaining, yx/yo
Regeneration
Parameters fixed
Load in, u0, gS/gp
Feed rate, w, gp/«in
Bed Height. H. gp '
Results
Load out, v. gS/gp
Rate constant, K, ain'1
Sorptiof)
Paraaeters fixed
Load in, uo, gS/gp
Feed rate, v. gp/ain
Bed weight, W, gp '
SOj feed, s, gS/ain
St>2 concentration, ye
Results
Load out, S, gV(B>
S02 reaaining. y^/y,,. (gas analysis)
902 reaaining. yx/yo. (pellet analysis)
Calculated
Load out. it, gS/gp
S02 reaaining, yx/yo
Regeneration
Paraaeters
Load in. MO. *s/gp
Feed rate, w, fp/au
Bed weight, N, gp'
Results
Load out. ti, g6/gp
Rate constant. K, ada'1
1-2
0.009
9.2
743.
0.72
0.0027
0.0(0
0.07
0.0*
0.015
0.01
0.072
10.0
346.
0.010
0.170
2-1
0.010
10.1
751.
0.73
0.002*
0.073
0.07
.0.07
0.0*0
- 0.04
0.074
9.7
346.
0.011
0.170
2-2 3-2 4
0.012 0.012 0
10.5 10
701. 9)6
0.7) 0
0.002* 0
0.0(1 0
0.01 0
0.01 -0
O.OS1 0
0.01 -0
0.0*1 0
10.2 9
346. 346
0.015 >
0.13) 0
.2 9
719
.6) 0
.0024 0
.075 0
.06 0
.0) • 0
.0*1 0
.1) 0
.071 0
.7 9
346
.012 0
.132 0
-1 4-2
.01) 0.01)
.1 10.0
642.
.64 0.64
.0022 0.0022
.092 O.M5
.10
.21 -0.1)
.077 0.074
.02 0.05
.09) 0.0*6
.4 (.7
346.
.013 0.014
.170 0*.134
5-1
0.009
9.5
606.
0.53
0.0020
0.061
0.07
0.06
0.065
-0.01
0.062
9.9
346.
0.011
0.13*
$-2
0.009
10.1
576.
0.53
0.0020
O.OS9
0.09
0.04
0.062
0.00
O.OS9
10.2
346.
0.011
0.1)2
6-1 8-2 7-2
0.009 0.010 0.
10.* 10.1 *.
5*6. $76.
$31.
.011
,4
0.53 0.53 1.25
0.0020 0.0021 0.
0.061 0.059 0.
0.0* 0.
• 0.06 - 0.
0.059 0.
- 0.02 - 0.
0.059 0.
10.3 10.
346. 346.
0.010 0.
0.144 0.
01 0.
01 0.
060 0.
03 0.
059 0.
1 9.
346.
010 0.
13* 0.
.0041
115
34
23
100
34
112
•
014
207
Shift Ma*er
1-2
0.015
20.4
600.
0.99
0.0037
0.062
0.10
0.02
0.064
-0.01
0.0*2
20.3
346.
0.01S
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0.016
16.)
55).
0.99
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0.09
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15.1
346.
0.016
0.163
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544.
0.99
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346.
0.019
0.1SS
10-1
0.015
!).(
574.
0.64
10-2
0.013
14.6
556.
0.64
0.0024 0.0024
0.064
0.10
-0.06
0.061
0.00
0.061
14.6
346.
0.016
0.13*
0.057
0.09
-0.01
0.057
-0.01
0.057
14.4
346.
0.013
0.140
11-1
0.014
15.7
635.
0.63
0.0026
0.054
0.06
-0.01
0.05*
-0.11
0.054
14.6
346.
0.014
0.11*
11 -I
0.013
14.3
62$.
0.63
12-1
0.015
S.9
607.
0.63
0.0026 0.0027
0.05*
0.07
-0.01
0.061
-0.09
0.05*
14.2
34*.
0.014
0.132
0.069
0.0*
0.15
0.071
0.00
0.064
10. 1
M*.
0.01S
0.096
12-2
0.017
10.0
5*9.
0.6)
0.0027
0.07*
0.01
0.03
0.07*
0.02
0.077
10.3
346.
0.017
0.10*
1J-1
0.01$
9.7
6)5.
0.62
0.0027
0.0(0
0.10
-0.02
0.0*0
-0.02
0.010
».7
34*.
0.01*
0.110
1J-2
0.016
9.5
*36.
0.62
0.0017
0.0*3
0.09
-0.04
0.011
-0.01
0.0*3
9.7
Mo.
0.01*
o.as
1 Bed weights
-------
"The pilot plant consisted of a small pulverized
coal furnace to produce a flue gas of desired composition,
an absorber with its accessory equipment, and a separate unit
for regeneration. Gas and solids flows through the absorber
were continuous. A flow sheet of the flue gas generation
and absorption system is shown schematically in Figure 25.
The absorber shown in Figure 26 was fabricated
of 316 stainless steel. It was 26 feet long and.had an inside
diameter of 1.6 inches. Thermocouples and gas,.sampling ports,
180 degrees, apart, extended to the center of the absorber and
were placed at 4 feet intervals along the length of -the column.
Two pneumatically, operated slide valves could be closed simul-
taneously during absorber operation and the solids .caught
between them were isolated and measured to determine jthe
solids holdup. Baffles were made of 4-mesh stainless steel
screen welded onto a 1/8 inch rod. The baffle assembly could
be inserted into the absorber or removed as desired.
An induced draft fan located downstream of the
absorber provided the necessary draft for the passage of the
flue gas through the absorber. Linear velocity of the gas
through the absorber was 8 to 15 feet/sec in countercurrent
operation and 20 to 23 feet/sec in the entrained flow used
in some of the later experiments.
A pulverized coal furnace supplied 25C to 1,000
1
cubic feet of hot flue gas from the combustion of I ' t> 4
pounds of coal per hour. The products of combustion flowed
from the. furnace to two cyclone separators,- in. which about
60 to 70 percent of the fly ash was removed, and then entered
the bottom of the absorber. Solids leaving the absorber
were collected in a scales hopper. When a sufficient quantity
of spent absorbent was accumulated, the regenerator was charged
-94-
-------
FIGURE 25 - Flue-Gas Generation and Sulfur Dioxide Absorption in Pilot-Plant Studies
of the Alkalized Alumina Process.
"First vork cited in footnote 10.
-95-
-------
GO*
4-Thermocouple
5-Gas sampling tap
C -Slide valve
-Solids -feed top
£-Knock-out trap
r-Boffles
FIGURE 26 .Absorber.
-96-
-------
and the fixed-bed of solids was regenerated with reducing gas.
In countercurrent gas-solids operation the regenerated ab-
sorbent was fed to the top of the absorber. In tests with
the solids entrained by the flue gas, a recycle stream of
absorbent was returned to the absorber about 9 feet above
the bottom of the column.
In the initial pilot plant work crushed alkalized
alumina 8 to 24 and 16 to 28 mesh as well as the pelleted
absorbent, 1/16" x 1/16", was tested in free fall at a feed
rate of 3 pounds of absorbent per 1,000 actual cubic feet of
the flue gas. The flue gas flowed at a velocity of 8 ft/sec
countercurrent to the dropping solids. As shown in Figure 27,
the smaller particles, 16 to 28 mesh, completely removed the
sulfur dioxide. However, there was considerable loss of
absorbent through carryover of the smaller particles. The
absorption of the pellets was quite low in free fall even
though the activity of the pellets when tested as a fixed
bed in bench scale apparatus was equivalent to that of the
crushed absorbent. Aerodynamically the crushed particles,
containing platelets and slivers, are more readily buoyed
by the upward flow of flue gas resulting in a greater residence
time in the reactor and a higher degree of SC>2 absorption. The
crushed particles also offer a higher surface area than the
pelleted absorbent.
With increase in gas velocity, the buoyancy
of the solids is increased resulting in a gain in solids
holdup in the reactor and in the residence time. Figure 28
shows the effect of raising the gas velocity from 8 to 12 to
15 ft/sec upon the degree of SO2 removal. Accompanying the
greater pickup of S(>2 with higher gas velocities there is
a greater loss of absorbent, as solids carryover, as the
-97-
-------
FIGURE 27
REMOVAL OF SULFUR DIOXIDE WITH ALKALIZED ALUMINA
FREE FALL AT 330°C
100-,
90
80
so2 CONCENTRATION: 0.3%
GAS LINEAR VELOCITY: 8 FT/SEC
. SOLIDS-TO-GAS RATIO: 3 LBS/M
EXPT 16
(16 - 28 MESH)
EXPT 18
(8 — 24 MESH)
EXPT 59
(I/IS" X I/IS" PELLETS)
10 15 20
HEIGHT OF ABSORBER, FT
REFERENCE:
USBM Bruceton Quarterly Report, pq. BM-13, (June 30, 1963).
-98-
-------
70
t=6°
§
o
x
o
°40
30
b.
O
<
>20
UJ
DC
10
T
T
T
Absorbent mesh size: 8-24
concentration: 0.34-39%
A-Expts. 23 and 24: gas velocity -15 ft. /sec.
B-Expt. 26: gas velocity- 12 ft. /sec.
C-Expt. 25: gas velocity -8 ft. /sec.
10 15 20
HEIGHT OF ABSORBER, ft.
FIGURE 28 -Removal of sulfur dioxide as a function of liner
gas velocity with regenerated alkalized alumina
at a solids-to-gas ratio of 3 Ibs./M ft.3 of gas.
-99-
-------
gas velocity is raised. In addition to this effect of carry-
over the shape of the curves in Figure 28 indicates that much
of the reactor length is not effective in S02 removal. There
is greater removal of SO2 at the bottom and top of the absorber.
The fresh absorbent is injected at the top of the reactor where
its freshly regenerated surface is conductive- to rapid reaction
with S02. At the bottom of the reactor the entering flue gas
contains the highest percentage of SC>2 and thus the-greatest
gas-driving force. Little removal of S02 is achieved in the
midportion of the absorber, as shown by the flattened curves.
To achieve greater residence time, 100 stainless
steel screens (4 openings to the inch) were spaced 3 inches
apart in the absorber (Figure 26). Whereas in the free fall
reactor the solids were in very dilute phase with a reactor
residence of a few seconds, a 1-minute residence could now
be obtained in the baffled reactor. The baffles by restrict-
ing the escape of the solid particles considerably reduced
the loss in solids carryover. As shown in Figure 29, 64
percent of SO- was removed in the baffled reactor at a solids-
to-gas ratio of 3 lb/1000 ft3 and a gas velocity of 12 ft/sec.
In the absence of the baffles only 31 percent of the S02 was
removed.
To increase the residence time still far tier
the number of baffles was increased threefold. The screens
of 4 mesh were placed 1 inch apart. ^t a gas velocity of
12 ft/sec, and a solids-to-gas ratio of ^ ib/100'J rt\, the
residence time for the 1/16 inch pellets was 2 minutes and
the removal of SO2 was increased to 98 percent. The pressure
drop across the 26 ft of absorber length was 6 inches of
water.
In the previous tests, baffles of stainless
steel screens were used to increase the solids residence
-100-
-------
8
x
O
u.
o
o
5
UJ
or
70
60
50
40
3O
20
10
I T I T
Absorbent mesh size:8-24
S02 concentrotion 0.34-0.36%
_ Gas velocity-12 ft/sec ^
Solids-to-gos ratio: 3 Ib/M ft of gas
OWith baffles- 3 inches opart
_ A Without baffles
Expts 35, 36
and 37
Expts 31
and 34
I
I
25
30
29
5 10 15 20
HEIGHT OF ABSORBER, ft
Removal of Sulfur Dioxide With Alkalized Alumina in Presence and Absence
of Baffles.
-101-
-------
time. It was subsequently realized that these screens on an
industrially sized reactor would be expensive, would limit
the diameter of the absorber, and could cause excessive
pressure drop through eventual blockage of the screens. It
was observed that by increasing the flue gas velocity, solids
entrainment occurs. If these solids are returned to the
absorber, suitable solids-residence times can be achieved for
high SO2 removal, and greater flexibility in the usable range
of gas velocity and absorbent size will be obtained.
Consequently, a recycle system as shown on
Figure 30 was incorporated into the sorber design. The
sorbent used in the testing of this system was of a poorer
quality than had been used in the countercurrent tests, but
it was the only preparation on hand at the time. Despite
this drawback, the tests were sufficiently promising,to
warrant further study of a solids entrainment system. A
comparison of these two types of sorption schemes is shown
in Table 15.
As shown previously, the increase in residence
time showed a marked improvement on the sorption of SO9;
therefore, it was hypothesized that if sorber height could
be increased, to provide greater contact time, and the solids
to gas ratio could also be increased, acceptable exit SO-
! 2
levels would be obtained. To verify these assumptions the
construction of a 55,000 cubic feet per hour pilot plir-
which incorporated these ideas was undertaken. Construction
of this pilot plant which is depicted in the flow sheet in
Figure 31, was completed in the first quarter of 1967.
The sorption section of the 55,000 CFH pilot
plant was operated in the following manner. Coal is picked
up at the coal pit (1) and" pneumatically conveyed to a hopper (2).
-102-
-------
TABLE 15 - Operating Conditions in Absorbing Sulfur Dioxide Under
Countercurrent Gas-Solids Flow Compared With Solids
pntrainment With Solids Recycle
Gas-solids contact
Test No.
Gas velocity, ft/sec
Solids/gas, Ib/M ft3
SO2/ percent:
In
Out
Removal
Density, g/cm3:
Charge
Bottoms
Overhead
AP, inches of H2O
Carryover; pet of feed,
<20 mesh
Solids flow ratio:
Overhead/bottoms
Reflux/fresh feed
Sulfur content, wt-pct:
Charge
Spent bottoms
Spent overhead
Regeneration before run:
Temperature, °F
Time , hr .
Space Velovity, hr"1
Gas
Countercurrent flow Solids
1A
12
3.0
0.33
.23
30.3
0.670
.726
-
—
0.02
-
-
1.09
1.97
-
1,210
10
100
H,
IB
13
2.8
0.27
.15
44.4
0.670
.733
-
4.7
0.04
—
-
1.09
2.47
-
1,210
10
100
HO
1C 4A
14 23
2.4 2.2
0.32 0.36
.09 .19
71.9 47.2
0.670 0.698
.766 .770
.748
7.5 2.2
0.08
2.17
- 10.4
1.09 0.38
3.67 2.64
2.80
1,210 1,200
10 10
100 100
H- H2-CO-CO0
entrainment
6A
23
2.4
0.36
.18
50.0
0.705
.769
.736
1.8
—
1.94
10.4
0.47
2.65
2.38
1,220
10
100
H,-CO-CO0
77 13 10 77 13 10
1) USBM-Bruceton - RI-7021, pg. 23.
-103-
-------
Fresh solids
feed
Gas outlet
Solids knockout trap
Recycle solids feeder
-Gas inlet
Bottom drawoff
FIGURE 30-.-Absorber with solids recycle in alkalized alumna process.
12-10-64 L-B80I
-104-
-------
I CMl»l
2 PuMriiW «M4
S rVMnnt Mel kom«
9 C|CW» H»0»0'»r
4 ^Hhrtftft* Wye*
r
c
f f Furnoct
10 PUrt 9«» cooler
I I CyClOM uponloi
11 0«t comctut
IS AkHk*iil
14 ttMbtM lurg.
IS
14 AMofMut tnmfatt
IT ttWMfll COllICIOft
19 Sp««t ottovbtn* MMfo'or
2O SotO»"t
23
24
2S C«M>m« «ol comgrmo
2«
27
SYMBOLS
OPC-0<"Of">i« o
FC -
UC • It M< control
OtC-O»(«mcoMm
FIGURE 31 -S02-Removol pilot plant (55,000 cfh flue gas).
L-9294
-------
From here, the coal is fed, via a lock hopper (3), to the
pulverizer (4). A sweep gas entrains the coal, in-the
pulverizer and conveys it to a cyclone (5) where the coal
and gas are separated. The coal is then pneumatically con-
veyed by combustion air from the compressor (7), to the
furnace (9). The sweep gas exiting the cyclone is recycled
to the pulverizer via compressor (6). The combustion products
are cooled to the desired reaction temperature (10) and sent
through cyclones (11) for fly-ash removal. The gases enter
the sorber (18) where S02 is removed by contact with the
alkalized alumina sorbent, and exit via a dust filter (23)
to the stack.
Sorbent from the sorbent storage hoppers (13)
enters the system through the .sorbent feed hoppers (14) and
feeders (15). The bulk of the sorbent is entrained by the
flue gas and passes upward through the sorber (18) to the
two solids-disengaging sections (16) (later replaced with a
single-stage unit) at the top of the reactor. From-the
disengaging sections the solids are sent to collection
hoppers (17) at which point they are either recycled to the
sorber or sent by gravity flow to the regeneration system.
The unentrained sorbent falls to the bottom of the sorber
and is pneumatically conveyed to the regeneration system.
The regenerated sorbent is then pneumatically conveyed to
the sorbent1storage hoppers (13) completing the nciiis eyjie,
The initial data from the pilot plant described
above are given in Table 16 for the test series C, D, E, and
F. The data for test series C will not be discussed because
the location of the feed points of the feed and particularly
the recycle streams were much closer to the bottom of the
reactor than for the other tests. As a result the contact
time was probably much lower than the subsequent series
-------
I
!-•
O
--J
I
fiF. Inches RgO
Feed location, ft:
Fresh feed
Feed rates, Ib/hr:
Feed rates, lb/1000 SCF:
Recycle to fresh feed ratio, Ib/lb
•octoms to fresh feed ratio, Ib/lb
Solid density In absorber, lb/ft» .
Mean particle size. Inch:
Fresh feed
•ottoeis
Recycle
•ulk solid density, lb/ft» :
Fresh feed
Sulfur content, wt pet:
Fresh feed
tecycle
Bench scale activity, | S/100 ( ...
Sulfur oxides, BOB:
10 feet
20 feet
30 feet
40 feet
Outlet ,.
Sulfur oxides removal, ncti'
. 25.1
. 0.05
. 631
10
114
0
4.4
0
. 0
. 0.52
. 0.068
. 0.060
. 41.2
. 44.2
. 2.4
. 1*30
. 1720
. ••
TABLE
C201 C401
27.1 25.0
0.05 1.00
631 623
10 10
10
116 118
0 590
4.4 4.7
0 23.5
0 5.0
0.45 0.60
0.066 0.059
— 0.072
— 0.035
42.6
43.2
..
45.4
1.6
2.4
3.0
1740
1540 460
1090
52.9
16 -
C501
25.7
1.10
598
10
10
123
674
29.2
5.5
0.63
0.06*
0.062
0.048
0.020
44.1
47.2
48.5
46.6
1.7
3.3
3.3
540
640
64.0
Oaaratlne conditions darlna abaoi
P">*
23.4
597
20
20
115
470
20.3
4.1
0.59
..
..
--
.-
370
•3.9
D102 P103
23.5 23.6
603 608
20 20
20 20
115 115
470 470
20.3 20.3
4.1 4.1
0.58 0.62
..
.. ..
.. -.
.. ..
4.07
730 550
68.3 76.1
D104 D105
23.6 22.0
609 541
20 20
20 20
113 115
470 470
20.3 20.3
0.60 0.79
0.052
0.039
0.030
39. -•
40.
45.
37.
1.
2.
2.
•. ..
•40 870
63.7 62.4
f\t,oQ tests
22.4 22.6 22.*
546 551 552
20 20 20
20 2p 20
115 123 123
470 479 479
20.2 20.3 20.3
0.70 0.59 0.56
4.64
930 1730 1*70
59. • 25.5 19.2
22. 8
558
20
20
U3
479
20.4
0.49
••
1800
22.1
•D||0
22.8
559
20
20
123
479
20.4
0.49
-.
1200
48.1
Dm '
22.9
1.0
560
20
20
123
479
20.4
0.50
..
1350
41.7
22.*
1.1
559
20
20
123
479
20.4
0.58
6.61
1000
56.6
22.1
1.3
J57
20
20
123
479
20.4
0.6*
.-
-.
125ft
45.9
Continued on next paga
-------
Operating condition*
absorption te«t«
O
CO
1
(continued)
T«st No,
Foei location, ft:
Fresh Ned
(Ucyc ie :
Feed rjtct , Ib/hr:
Fr*»h feed
Rrcycl* ,
F«ed rate., lb/1000 SCF:
RecycU co iresh feed ratio, Ib/lb ....
Bettors to fresh f«d ratio, Ib/lb ....
Mean particle nit, Inch:
Fr«»h f«ed ,
1 o i c OBIS .
R*cyei*
Dl^4_
22.6
0.9
556
20
20
123
479
20.4
3.9
C.57
..
-•
D115
22.7
1.2
556
20
20
123
479
20-. 5
3.9
0.46
0.059
0.058
0.038
0.032
DH9
22.4
2.8
556
20
20
130
488
21.1
3.8
0.32
0.056
..
0.034
0.033
DI20
21.7
1.1
513
20
.20
177
488
20.9
0.84
0.063
0.057
0.035
0.023
6121
22.2
1.2
526
20
20
185
456
19.3
0.53
0.058
0.059
0.029
0.023
D122
21.7
1.2
516
20
20
179
460
19.8
0.60
0.064
0.058
0.029
0.025
D123
21.8
1.5
548
40
20
179
452
7.9
19.9
0.49
0.053
0.058
0,033
0.025
D124
22.2
1.5
548
40
20
173
486
7.5
21.1
0.03
0.052
0.060
0.029
0.024
D125
22.2
1.5
544
40
20
186
459
6.0
20.7
0.79
0.050
0.053
0.039
0.025
E101
26.7
2.1
604
40
20
166
161$
7.0
60.5
0,68
O.lt
0.070
0.066
0.025
E102
25.5
1.8
554
40
20
163
1572
6.9
59.1
0.25
.-
0.069
0.055
0.022
E103
25.9
1.5
576
40
20
187
1639
7.1
61.6
0.28
0.15.
0.068
0.054
0.021
E104
26.1
1.9
599
40
20
1JO
15J9
7.3
59.0
0.45
.'
0.068
0.052
0.018
E10>
26.2
1.8
,569
• 0
20
191
1169
7.1
43.5
0.54
>•
0.069
0.066
0.021
ElOf
26.1
2.2
580
40
20
186
13cl
7.0
51.2
0,54
»«
0.065
0.056
0.023
no?
26.0
:. i
57-
-0
20
186
1476
7.0
55.5
0.-.6
0.17
0.066
0.055
0.024
nca
26.1
2.6
566
40
• W
166
1476
6.9
54.9
7.9
0.54
--
-•
--
..
Bulk solid denilty, l»/ft»:
Fiffch feed ., . a...
SuU'-.r eanttr.r wi pet:
Sulfur oxide*, ppv:
42.1
43.7
0,4 0.5
... -- 0.9
2.T
44.3
47 5
46.7
0.6
3.1
1830
1950
42.2
45.1
47.6
45 0
0.4
1.0
2.3
525
1750
1250
44.;
43.
44
4*
0.
0.
2.
188(
150<
42.8
43.3
44 8
44 4
0.4
1.8
2.7
) 2110
1 1750
43.6
44,8
44 2
47.4
0.3
1.0
2.5
1870
1750
4.?. 2
44.3
47 4
ii.7
0.6
0.9
3.0
41A
1440
::co
43.3
46.2
44 7
44.:
0.6
2,S
1890
1510
48.7
47.9
30.4
0.9
1690
1590
47.9
48.5
49 9
0.9
1900
1469
48.8
46.9
52 0
O.I
1810
I44C
49.7
49.0
46 0
0.9
1640
!300
SO. I
44.0
48 9
0.9
1.6
•1820
1710
48.5
46.4
43 1
0.8
1.4
1740
1330
48.6
45.5
46 8
0.9
1.7
••
1910 2)2
UfO 155
oj'-;«t
lulfci Oslo*! removal, pet*'.
1220 1210
1150 12»0
1100 12lC
610 830
65.0 64.*.
1600
1C8C
590
74.6
710
1330
1270
550
76.2
1.'JO
10VO
440
1890
1510
i'SO
1C20
670
71.0
1690
1590
1150
110
too
95. t
1900
1469
liiO
1030
520
77.7
1810
I44C
ire-
me
$60
75.8
Coatlnf;»d
1640
!300
1C3C
«90
400
'12.5
en MX
•1820
1710
la-G
IG1C
530
76.4
MM ,
P40
1330
11*0
580
320
•6.1
1910
UfO
9.B
690
260
68.8
2)20
1550
1240
a 10
4)0
• 1.2
in
-------
I
(-•
o
T«»t No. __^___
Gas velocity, ft/sec
;P, inches HjO
Feed location, ft:
Feed rates, Ib/hr:
Fresh feed
Feed rates, lb/1000 SCF:
Freih feed
Recycle
Recycle to fresh feed ratio, Ib/lb ...
Bottoms to fresh feed ratio, Ib/lb ...
Solid density in absorber, Ib/ft3
Mean particle size, inch:
Recyc le-r.
Bulk solid density, Ib/ft9:
Recyc le
Sulfur content, vc pet:
Sulfur oxides, ppn:
10 f*et
20 feet
JO feet
40 feet
E109
25.0
567
40
20
187
1552
7.2
60.2
8.3
0.71
-.
1.3
1.9
1910
1430
990
420
240
89.6
EUO
25.0
567
40
20
187
1552
60.0
8.3
0.64
--
1.3
2.0
1910
1440
1190
960
1180
49.0
EUl
24. V
558
40
20
187
1552
59.7
8.3
0.73
--
1.4
1.1
9.8
1970
1340
1230
1030
410
62.1
(continued)
ZS.O
562
40
20
187
1552
59.9
8.3
0.65
..
1.2
1.4
2140
2420
-.
1760
1100
S2.5
EH*
25.4
577
40
20
187
1552
59.8
8.3
0.61
..
1.3
1.6
920
2030
1620
..
810
65.1
E114
25.3
591
40
20
187
1552
60.3
8.3
0.50
0.16
1.4
1.8
2170
1930
..
1660
750
67.}
. "Qi
26.8
0.3
602
40
..
187
0
0
0
0.60
0.03
0.5
l.l
2330
2430
2240
2040
--
**
26.8
0.1
583
40
..
187
0
0
0
0.53
--
0.6
1.1
2310
2270
2170
2080
1920
17.2
28.2
1.6
584
40
20
187
1851
65.9
9.9
0.84
0.13
0.5
1.6
2100
1710
1590
1300
160
93.3
29.0
1.6
596
40
20
187
1830
63.9
9.8
0.75
0.15
0.3
1.7
2060
2250
1400
620
280
88.1
27.9
1.9
565
40
20
187
1851
69.1
9.9
0.55
0.13
1.8
3.4
1920
1910
1390
1340
780
66.1
Plot>
28.1
1.6
548
40
20
187
1871
64.6
10.0
0.78
0.15
1.8
3.4
1330
830
1200
1180
700
69.7
26.6
1.5
513
40
20
167
1851
64.9
9.9
0.69
0.16
1.9
3.3
1620
1560
1116
1050
200
91.4
27.7
0.6
544
40
20
187
1851
64.5
9.9
0.54
0.09
1.8.
3.6
1850
2030
1580
1170
340
76.8
29.4
0.3
588
40
--
187
0
0
0
0.63
0.03
-•
"
2130
2400
2160
2170
2020
12.7
F112
26.6
1.5
494
40
20
187
1551
53.5
8.3
0.72
0.13
2.2
2.8
6.7
2080
1760
1210
1480
730
68.4
I/ Secondary separator not used during test aeries E and F
I/ Percent reaovcl hasad on average Inlet sulfur oxides concentration of 2314 ppa.
-------
of tests. Runs 1 through 19 of the D series were conducted
using a recycle ratio of approximately 4:1 whereas runs 20
through 25 had a recycle ratio of approximately 2.5:1. The
average SO2 removal for the entire D series was 58%, with
the runs at the 4:1 recycle ratio averaging 52% and the
runs at the 2.5:1 recycle ratio giving an average of 75%
removal. The higher removal for the lower recycle ratio was
not unexpected, however, because of the following reasons:
Although the recycle ratio was changed in the latter D series
runs the amount of recycle remained the same as in the earlier
runs while the amount of fresh feed was increased; therefore,
the total amount of sorbent cycling in the sorber increased
as did the percentage of fresh feed. The greater sulfur
removal in the latter D series runs reflects this and indicated,
as expected, that the SC>2 removal is a function of the solids
density in the sorber. This conclusion is further supported
by the E and F series runs with recycle, during which average
sulfur removals of 76 and 79%, respectively, were obtained.
In the E and F series, 4 runs, E-101, E-109,
F-103, and F-107 apparently removed 90% of the inlet SO-.
However, in two of these runs, E-101 and F-107, approximately
30% of the SCK was removed in the first 10 feet of the sorber.
Furthermore, the results of run E-109 could not be duplicated
in run E-110 even though apparently identical operating con-
ditions were used. It was also noticed that in 6 of the 24
E and F series runs, SO2 concentration increased bevi/rtr ,:hj
sample points in the reactor. The reasons for this are not
apparent but could be caused by stripping of SO2 from the sorbent
or possible inaccuracies in the gas sampling mechanism or SO2
analyses. Therefore, owing to the variation in SO, removal
•*v»
obtained in some cases for what are apparently essentially
identical operating conditions, and because only a few runs
-110-
-------
actually achieved 90% removal, it is felt that 90% removal has
not been adequately demonstrated.
Since in most of the test runs made in the 55,000
CFH pilot plant more than one variable was changed (e.g., gas
velocity, pressure drop, solids to gas ratio, etc.), the results
of several tests using similar operating conditions were com-
bined and averaged. Figure 32 shows the average SO2 removal
obtained under similar operating conditions as a function of
column height. For these runs, average sulfur removals of 80%
were obtained and a distinct relationship between SO- concen-
tration and a column height can be seen. On the whole, test
series £ and F runs with high recycle demonstrated average
sulfur removals of 70 to 80% and it is believed that when
an acceptable sorbent is developed, 90% removal can be ob-
tained in the dispersed phase type reactor.
In order to test the sorption model at higher
sorbent densities and sulfur loadings, a number of high recycle
ratio runs, series H, were made. The data for these runs are
tabulated in Table 17. Although recycle ratios of 15:1 were
attempted and the average recycle ratio was nearly 11:1 during
12 of the runs, no improvement in sulfur removal was obtained.
For the 12 series H runs at recycle ratios of 7:1 or greater
the average removal was 60.2%.
Even though the removal levels are lower in the
H series than in the E or F series, definite conclusions as to
the effect of high recycle operation cannot be made, for the
sorbent activity had decreased from 9.8 to 6.2 grams of sulfur
per 100 grams of sorbent and the sorbent sulfur levels at
which the runs were made were much higher than had previously
been used. The USBM-Albany in their testing program had shown
that sorption rate is a function of the sulfur loading on
-1JL1-
-------
BM-20
0.28
c
o>
o
o
z"
O
cc
h-
2
Ul
o
2;
8
UJ
or
u.
I
Fresh feed to gas ratio: 6.9 to 7.3 lb/l,000scf
Fresh feed to bottoms ratio: 1.5 to 3.6
Recycle to fresh feed ratio: 7.5 to 8.7
Test nos: EIOI to EI04, EI06 toEIOS, EUI to EI14.
0 10 20 30
HEIGHT OF ABSORBED feei
FIGURE 32-Average sulfur oxides concentration as function of absorber
height when total overheads are recycled at 20 foot level
•and fresh feed Introduced at 40 foot level.
-112-
-------
TABLE 17 - Qp«r«ttrn condition* ,7
SO.l
SO.S
3 13
l.SS-
4.S4
J.72
1430
880
18.6
1.19
HI 10
28 5
1.87
552
40
20
194
1602
54. 0
8.3
0.16
0 061
0.07S
0.027
O.OS3
48 6
44.4
SO.O
48.6
3 13
2.31
5.18
4.36
1640
870
47 i
l.\9
Mill
28.3
1.87
S49
40
20
199
2299
77. S
11.6
0.17
0 052
0.071
0.023
0.042
44.9
47.4
52.2
SJ.2
3 76
2.79
S.78
1. 39
16SO
650
60.7
3.96
M112
28.6
1.S3
563
40
20
181
2107
71.2
11.6
0.14
0.041
0.073
0.023
0.040
4}. 4
45.9
47.9
52. J
0 S9
1.80
3.16
3.78
1470
S70
61 1
).S1
H113
28.7
1.68
S62
40
20
181
2227
7S.1
12.3
0.16
0.041
0.074
0.020
0.049
4S.4
44.4
SO. 6
SI. 9
0 S9
l.SO
3.00
2.60
15*0
530
66 3
S.62
HU4
29.2
1.S6
577
«b
20
137
2181
73.1
IS. 9
0-1}
0.1S
0 041
0.070
0.020
0.047
4S 4
SO.l
49.6
SO. 9
0 Sv
1.72
3.02
2.98
1500
410
68 4
7.74
H11S
28.8
1.70
S67
40
20
212
2123
7 i
71. S
10.0
0 34
0.17
0.034
0.070
0.018
0.03<
53.0
47.2
48.2
SI. 4
2 21
1.8'.
3.61
1.7*
1630
14 '0
OCA
1130
ISO
Co 6
3.07
H1I6
30.4
0.3
610
40
...
212
0
7.1
0
0
0 29
0.03
0.034
0.079
0.072
2.JI
1.74
2.41
3.93
3. Of
».Jf
6.40
1510
1410
4 6
1.59
Htl7
26.6
1.88.
S64
40
20
208
219S
7.6
79.9
10.6
0.36
0.14
0.032
0.068
O.OIS
0.07*
ii.2
47.9
49.9
SO. 4
2 94
2.60
4.60
4.41
1670
1S90
I )}A
680
S9 S
S.23
HM8
25.2
1.V6
sss
40
20
209
2285
8.0
87. S
10.9
0.43
0.20
0.028
0.072
0.017
0.020
M.i
48.8
SI. 9
'.e.n
1 24
2.83
i.24
6.11
1990
1760
0SO
««
4.68
Hilt
26. S
1.93
S48
40
20
219
2378
7.9
66.1
10.9
0.46
0.1S
0.033
0.070
0.018
P. 021
S4.«
Sl.O
54.0
$5.*
J.»J
i.m
6.67
6.40
1480
1360
11m
780
* j 1
2. OS
NOT REPRODUCIBLE
-------
the sorbent (see Figure 16) and that as the sulfur loading
approaches the saturation level the sorption rate decreases
drastically. It should be noted also that in most of-the
runs (11 out of 15 for which the data are reported) the.
sulfur content of the bottoms stream is less than the fresh
feed. Further, since the bottom stream comprises from
about 20 to 60% of the fresh feed rate, it appears that a
substantial portion of the fresh feed is not removing any;
S02 and thus the effective fresh feed rate is much smaller
than the actual rate. Therefore, it is improbable that
high percentage removals of SO2 could be obtained at.these
conditions of low gas residence times, high sorbent.sulfur-
loadings, and by-passing of a significant amount of the
fresh feed, at which the H series tests were operated; The "
fact that even 60% removal levels were obtained under.these
conditions is, in itself, promising.
Further tests, series K and series M runs, were
made to evaluate properties of two new sorbents, Grace-2.,and
Kaiser-1, respectively. Data from these tests are listed
in Table 18. Many of the previously reported operating
parameters were not presented in these data, thus precluding
direct comparisons with earlier tests. The average removal
for the Grace-2 sorbent during the K series test was 51%,,
The average removal in the M series tests which used the
Kaiser sorbent was 47%; however, the last two runs cf this?
series, M-105 and M-106; gave sulfur removal levels cf -i
and 83%, respectively.
Because of the relatively small sulfur, pickup
by the bottoms-fraction, it has been suggested that a bottoms
recycle system be added to the present pilot plant. This
type of system would increase the solids residence time of
-------
TABLE 18- Sorptton operating conditions for evaluation of Grace No. 2 and Kaiser No. 1 »brbent»
1
t-l
m
Test No
Sulfur oxides concentration,
mole fraction:
Inlet, (y0)
Exit,
-------
TABLE I8- Sorption operating conditions for evaluation ofGrace Ko. 2 and Katser Wo. 1 sorbents. con.
Test No MIOOa MIOOb M101 M102 Ml03a M103b M105 M106
..... . Kaiser Kaiser Kaiser Kaiser .Kaiser Kaiser Kaiser Kaiser
Sorbent used ............ ^.... ...... „ . .. , _. , .. , ., , „ , „ . ., .
No. 1 No. 1 No. 1 No. 1 No. 1 No. 1 No. 1 No. 1
Sorption-regenaration cycle 1 1 2 3 4 4 5 6
Sulfur oxides concentration,
mole fraction:
Inlet, (y0) 0.00164 0..00134 0.00190 0.00184 0.00209 0,00157 0.00168 0,00144
Exit, (yf) 0.00099 0.00089 0.00127 0.00127 0.00133 0.00107 0.00019 0.00024
Flue gas flow rate, SCFH, (V) 25,200 '27,000 27,800 27,200 27,200 27,100 29,400 28,900
Sorbcr pressure loss, inches H20 (p). 0.82 0.88 0.68 0.74 0.66 0.71 3.11 1.93
Sulfur content of sorbent feed,
wt fract.S, (x0) , 0.0026 0.0015 0.0046 0.0049 0.0046 0.0057 0.0132 0.0135
Sorbent feed rate, Ib/hr, (Lf) ...... 200 198 204 201 226 206 230 .1326
Sorbcr temperature, CF 577 599 539 561 558 556 526 559
Solid density in sorber, Ib/cu ft ... 0.069 0.088 0.089 0.068 0.064 0.064 --- 0.178
Sorbent carried from reactor,
pet of feed « 0.60 * 0.96 --- * 1.10 •> 9.02 0.65
Sorbent <60 Tyler mesh carried from
reactor, pet of feed « 0.33—> 0.84 --- * 0.93 * 6.20 0.58
Sorption rate9 wt SOg reraoved/wt
sorbcnt/hr, (r)I/ ^ 1.14 0.78 1.46 1.19 1.78 1.07 0.80 1.02
Sorption rate constant. (K. ?/ 5.40 2.41 7.74 6.35 8.98 7.00 21.90 20.46
I/ Sorption rate defined a.' 2_/ Sorption rate constant defined as
0.0342
-------
the larger leas loaded sorbent pellets. Kinetic studies
have shown that these low sulfur content particles should
be more reactive than the smaller highly loaded particles
which have been recycled in the earlier tests. Based on
the above, the Bruceton pilot plant was altered so that
the effect of a bottoms recycle could be evaluated. The
new design is shown in Figure 33. At this time, however,
no data using alkalized alumina sorbent have been received
and the effect of bottoms recycle on S02 removal cannot
be evaluated.
(4) AVCO
The role of the AVCO Corporation in the study
of alkalized alumina sorption was directed toward kinetic
modeling. Their sorption model, which will be discussed
in a later section of this report, is based on sorption
data such as that shown in Table 19. Additional AVCO
data are available in their reports and are not included
here. A plot of these typical data (in the form of a
rate expression vs. sorbent loading) is presented in Figure
34 and a comparison of these data to those predicted by
calculations based on mass transfer to single spheres is
shown in Table 20. The variables used in Figure 34 are
as follows:
y - mol fraction of S02 in the vapor
W * sorbent loading, Ibs. of S02 gained/lb
sorbent.
Wc » sorbent loading @ saturation, Ibs. or S02
gained/lb sorbent.
With the exception of the initial loading region 0 - 0.02
Ib S02/lb ftorbent, the sorption rates were accurately described
by the model.
-117-
-------
TABLE 19
SORBENT LOADING TEST DATA
ISR-3
ISR-5
t.t*»i^H+- nf Qorbent = 217
ma
Time Weight gain (mg) Time Weight gain (mg). Time Weig
0 sec 0
8.3
15.0
22.5
30.0
45.0
1 min
1.5
2
3
4
5
7
10
15
20
25
30
35
40
45
50
60
80
85
0 sec
5
10
15
20
25
30
45
1 min
1.5
2
3
4
5
10
15
25
35
50
65
75
80
1.4
1.7
2.1
2.4
2.7
3.2
3.8
4.0
5.1
6.3
6.4
8.0
9.0
10.8
12.4
13.9
15.1
16.1
17.2
18.3
19.2
21.0
25.7
26.6
•
,j
ISR-2A
0
2.3
2.8
3.6
3.7
4.0
4.2
4.3
4.7
5.2
5.6
6.4
7.*
7.9
10.5
12.8
16.6
20.6
25.4
30.2
33.7
35.0
0 sec
5
10
15
20
25
30
40
50
60
0
.5
.8
1.0
1.6
1.9
2.0
2.0
2.5
3.1
1.25 min 4.4
1.50
1.75
2.0
2.5
3.0
5.25
8
10
15
20
25
35
45
55
65
75
85
T&T.^ V ff V* J» f^ £
0 sec
5
10
15
20
25
30
45
1 min
1.5
2
3
4
5
10
15
25
35
50
65
80
90
4.7
4.7
5.0
5.7
6.0
7.4
8.0
9.7
12.0
12.9
14.5
17.2
19.0
21.0
22.7
24.3
25.4
ejrt** Vion *~ — "3 *•* fi
S>\jL. L/CIi l_ J J O
ISR-41
0
2.3
3.7
4.3
4.6
4.9
5.1
5.7
5.9
6.4
7.0
7.8
8.9
9.6
12.6
15.6
20.6
25.7
32.0
38.2
44.7
48.7
0 sec
10
15
30
45
1 min
1.5,
2
3
5
10
15
20
25
35
45
55
65
75
79
ISR-51
0 sec
5
10
15
20
25
30
45
1 min
1.5
2.
3
4
5
10
15
20
30
40
60
80
-///-
ht gain
^^^*"^
0
1
1.2
1.5
1.9
2.2
2.8
3.1
3.7
4.4
5.3
7.0
8.7
9.7
11.0
13.5
15.8
17.9
19.2
22..0
23.0
0
2.!
4.)
6.1
7.1
8.1
9,!
9.8
9.)
10.?
11.3
12.1
13.1
13.1
16.1
19.8
21.1
25.8
29.8
37.1
44.1
1 AVCO, F^nal Rjport - 11-1-67, pp-141-14?
-------
TABLE 20
SUMMARY DATA OF INITIAL RATE STUDIES1*
Run No.
ISR-3
ISR-4
ISR-5
ISR-21*
ISR-41 >
ISR-51*
T(°C)
300
300
300
300
300
300
Gas
Velocity
(ft/sec)
1
1
1
1
1
1
.9
.9
.9
.9
.9
.9
so2
.0695
.0695
.0695
.0599
.0585
.0562
-Vol .
6
6
6
4
6
10
Percen
,42
.42
.42
.32
.42
.4
°2
3.12
3.12
3.12
3.19
3.12
2.99
Diar
.102
.102
.102
.102
.102
.102
:icle
neter
in.
in.
in.
in.
in.
in.
Initial
Rate From
Data, hr'1
4460
2260
5260
8632
2760
9260
Calculated*
Initial
Rate, hr~x
990
990
990
990
990
990
Computed as mass transfer rate to single spheres using the volume percent SO, in the gas
stream as the driving force for mass transfer.
1) AVCO Final Report, 11-1-67, pg. 94.
-------
STORAGE
HOPr'ilK
BAG
FILTER
RECYCLE
HOPPED
^— FLUE GAS
SOLIDS CCil
VEYING SYS1
S02 RcMCVA
PILOT PU
-------
10,000
IOOO
100
8
K>
X
o
O ISR-S 1
9 ISR-4 > 6.42% ICO
» ISR-5 )
O ISR- 21 (4.9% Hf)
* ISR-41 (6.42% HjO)
0 ISR-
* o-X, «
«'* °». .
AVCO MODEL
3OO*C
4p>aiO2ln.
VO.30
I
6
IOOW
FIGURE 34 INITIAL SORPTION RATE DATA
-121-
-------
These rates were then compared to USBM-Albany/
(N0x)-free sorption data. The initial USBM rates were
considerably lower than the initial rates measured at AVCO.
This may be due to differences in sorption properties,
particle diameter, resolution of the experiment, or mass
transfer limitations. Also, there was evidently some
depletion of SO2 in the gas during the initial part of
the sorption. The USBM-Albany rates at higher loadings
are also lower than those measured at AVCO but not to the
extent of the initial rates. The differences at the
higher loadings probably reflect differences in sorbent -
properties and particle size because mass transfer' and :
the resolution of the experiment should not present the
problems that they might have created at the lower loadings.
One factor which appeared to influence these
high initial rates was water vapor. Initially, as seen in
Figure 34, the rate increases with the moisture content
of the flue gas but appears to be independent of the moisture
level at loadings greater than 2 percent. The independence
of moisture content is in accord with the assumption of pore
diffusion control, an effect which is realized at loadings
greater than 2 percent.
When the role of NOX on sorption became c;pa rent,
the AVCO model's constants were adjusted and a new empirical
equation was defined. Although the new equation, represf->;V:ing
the AVCO NOX data, showed the same deviation at bov-h t e lower
and higher range of sorbent loadings when compared-with the
USBM-Albany's NOX data (Figure 35), the mid-range sorption
rates compared favorably. Since it is in this middle range
that commercial sorbers will operate, the ccrrcboration of
the Albany data by AVCO is deemed sufficient from a design
point of view.
-122-
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FIGURE 35
COMPARISON OF USBM ALBANY AND AVCO SORPTION RATES OF S02
ON
ALKALIZED ALUMINA WITH NOX PRESENT
O USBM ALBANY DATA W/NOX
(INITIAL LOADING DATA)
USBM ALBANY DATA W/NOX
(ALL DATA)
O AVCO DIFFUSION MODEL FITTED TO
NO DATA*
A AVCO EMPIRICAL MODEL FITTED TO N0»
DATA*
AS PRESENTED IN AVCO'S IIIU REPORT,
FIG. I.
FRACTIONAL S02 LOADING. LB S02/LB S02 AT SATURATION
-123-
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(5) W. R. Grace
W. R. Grace and Company is one of the major sorption
manufacturers; as such, their work has been directed toward
improved sorbent manufacture, mechanical properties, binder
and additive effects. In evaluating the effectiveness of these
improvements, however, they have necessarily been involved in
sorption studies. These bench-scale studies, in general, agree -
quite well with the sorption work of the USBMr-Albany and the
AVCO Corporation. For the most part, these studies were based
on NO -free flue gases and further discussion of these data
X
is not warranted in light of the improved sorption effects
found in the NO data.
X
Although the Grace results for NOX sorption
indicate sorption rate increases of the same magnitude as
were reported by Albany and AVCO, they found one effect which
had not been reported by either of the other two sorption
analysts. This effect was the high surface concentration of
sodium sulfate, NajSO^
It was hypothesized by Grace that the high
level of SO-, dictated by the NO chemistry would cause rapid
J A
conversion of the surface sodium aluminate to Na2SO4. This
theory is supported by the data shown in Table 21 in which*
the results of analyses of the sulfate content of the "fines"
and "cores" produced by attrition are tabulated. On the
average, sulfate concentration of the fines (as-sained to ^e
surface fragments) is higher by a faccor of 3.6 then that
which is found in the sorbent cores. For comparison an
NOx-free fun is presented. This run shows a fairly even
distribution between surface and interior loadings-
The increased reactivity in the presence of NO
was visualized, therefore, as the result of Na2SO4 crystals
crowding one another as sorption occurs. The crystalite
-------
TABLE 21
SO4 Determination of Sorbent Exposed to
Flue Gas Containing NOX After Attrition^-'
SAMPLE NO.
5526-28
5526-30
5526-32
5526-33
5526-34
"Cores"
Wt. % SO4*
8.43
7.24
5.96
13.14
7.12
"Fines"
Wt. % S04*
37.3
23.7
24.6
11.05**
17.83
* All samples contain from 2 to 3 wt. % SO* before sorption,
** No NOX in feed gas.
1) W. R. Grace, Monthly Report, Aug. 1968, Pg. 2.
-125-
-------
growth of Na_SC>4 at the sorbent surface causes fractures and
cracks to develop. The increased surface area created by
these cracks exposes more fresh sorbent thus increasing
sorption rate. The increased sorbent-attrition rates associated
with the NO sorption studies further support this hypothesis.
X,
To remedy this situation, Grace is studying the
possibility of impregnating the Dawsonite material on a
suitable support. Ideally, this support would increase the
mechanical properties of the sorbent by limiting the extent
to which surface cracking could proceed while not greatly
inhibiting the excellent reactivity or regenerability of the.-.
present material. In addition, Grace is also evaluating the...
use of binders to stabilize the sorbent structure as a means •
of improving attrition resistance.
(6) CEGB (Central Electricity Generating.Board)
The work of CEGB has been in the areas of sor-
bent evaluation'and pilot plant development, but not much
of the actual data has been reported. However, CEGB have
reported their conclusions obtained from these data. A
summary of these conclusions for alkalized alumina sorption
of SO- is given below.
In synthetic flue gas, S02 removal takes place
as sulfjlte at low temperatures and sulfate at high tempera-
tures. The temperature range studied is from belov IfO^C
to above 350°C. In actual flue gases, however, SO2 removal
is to the sulfate form regardless of the sorption temperatur.e.
This was not fully explained by CEGB but they did indicate
that nitrogen oxides appeared to exert a strong catalytic
effect on sorption, particularly in the surface density of
the sorbent. This last conclusion agrees quite well with
the W. R. Grace results discussed earlier and the effects
-------
0.16
PELLETS' 1/8 * 1/16
I3O°C
15 20
TlME.minutts
30 95
FIGURE 36 -S02 SORPTION CURVES OF CERL
-127-
-------
o
1 temperature on sorption in both synthetic and actual.
flue gases agree with those previously reported by the
USBM-Albany. Further conclusions drawn by CEGB were that
the high affinity of alkalized alumina for water at low-
temperatures was not present in the reaction temperature
range, the uptake of water having dropped to below 1% of
the sorbent's weight at a temperature of 300°C. Similarly,,
the accumulation of chlorine compounds found in flue gases,
which had given some cause for concern in bench-scale
studies with simulated flue gases, was found to be negligible
when actual flue gases were used. This is attributed to-the
catalytic effect of the nitrogen oxides, in that the sulfur.
trioxide is preferentially sorbed over the flue gas chlorides.
In an analysis of the available CEGB data typical
sorbent loading vs. time curves were prepared by AVCO and are
shown in Figure 36. When these results are compared to the
USBM-Albany's sorbent loading vs. time curves shown in Figure 20,
it can be seen that the CEGB loadings are about 62,5% of the
Albany loadings at similar conditions. However, the Albany runs
were made with a Grace sorbent whereas the CEGB runs were made
with a Peter Spence sorbent and in independent studies by the
AVCO Corporation, the Grace material was found to be approximately
30% more reactive in the loading range compared. If the CEGB
results are scaled up by this factor, it can be seen ,aat thay
agree quite well with the Albany results. In conclusion, the
sorption effects and loadings reported by CEG3 have bes', found
to be in general agreement with the preser.^iy available U. S-.
data.
d. Sorption Modeling
(1) Kinetic Modeling
The following discussion was abstracted from two
recent USBM-Albany reports (1). It is felt that this discussion
(1) USBM-Albany, Annual Report, June 30, 1969, pg. 4 and USBM-
Albany, Progress Report, September, 1968, pp. 2 and 3.
-------
adequately summarizes the present state of the art in the
area of kinetic modeling.
Initial studies to determine the reaction rates
for sorption and regeneration have been concerned with the
development of a mathematical model for sorption rate based
on batch thin bed studies. Results of these studies are
shown in Figure 37, wherein the four equations currently
being used are adjusted to fit the loading data for Grace
No. 1 sorbent at 300°C without NOV. The four equations
2C
which are compared are:
1. yt = 22.5 In (1 - «y. g--) , which was used by Amundsen
in his studies of adsorption (Amundson, Neal R.,
"Mathematics of Adsorption in Beds," J. Phys. Colloidal
Chemistry, v. 52, 1948, pp. 1153-7);
2. yt =3,000 u^ + 65 v, the tarnish equation presented in
Jost (Jost, W, "Diffusion in Solids, Liquids, Gases,"
Academic Press, Inc., New York, 1952, pp. 352-354, with
truncation at saturation);
3. yt = 691 [1 - (1 ;- g^yg- > 2/3]-2,815 n, which was developed
by AVCO (Removal of SO2 from Flue Gas," Final Report,
Contract No. PH 86-67-51, November 1, 1967, pp. 61-100);
4. yt = [154 In (1 - Q-''-££-) - 895 p] , a modification of the
Amundson equation which provides an identical fit to the
other equations at low loads;
where u = sulfur loading on the sorbent, IbS/lb sorbent
t = time, minutes
y = S02 composition of the gas, % vol
Although the scatter of the data is too great to allow a
definite choice among the models, some definite observations
can be made. The Amundson equations appear not to fit as
-129-
-------
0.16
o
I
0.045 pet S02
0.068 pet SO2
0.34 pet S02
0.97 pet S02
2.O7 pet S02
10.32 pet S02
EQUATIONS
Amundson yt « -22.5 In (I-
Josl yt « 3,000/4*
AVCO yl
Medifi«d Amundson yt • -[(54 In (I- jg) *
100 120 140
percent X minutes
180
200
of 3-.oa.dxng Equations at 3OO° C
-------
well as the others, although fitting to the data at higher
loads would reduce the error. At the highest loads the
AVCO equation obviously fits best, but at intermediate loads,
the Jost equation appears to be better than the AVCO equation.
In the loading range of 0-10 percent sulfur, which is the
range of principle interest, the Jost equation fits best while
in the 0-5% range all of the equations give essentially
identical results. However, the Amundson equation is the
easiest to handle while the AVCO equation is the most dif-
ficult. As such, the Amundson equation has been used
almost exclusively as the rate expression in sorption modeling.
(2) Fixed-Bed Modeling
A theoretical model representing the sorption
of SO- in fixed beds of alkalized alumina was derived by the
USBM-Bruceton. The derivation of this model is contained
in Appendix D of this report.
To determine the effect of changes in system
parameters, a computer program was developed utilizing the
Bruceton sorption model. A sample output from this program
is shown in Table 22. The output data lists the gas velocity,
bed pressure drop, and number of stages for an input range of
bed heights and column diameters. The number of stages are
the number of parallel fixed beds required for sorption at
the input conditions. The output was arranged in this fashion
for convenience but the results may be combined with the
column diameter to give the effective cross sectional area
required for sorption for configurations other than cylindrical,
Because the bench-scale studies at Albany and
AVCO used relatively small ratios of sorbent to S02, the inlet
and exit SO~ concentrations were essentially the same. There-
fore, the data they obtained could not be compared to the
-131-
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TABLE 22
DETERMINATION OF MULTIPLE FIXED-BED PARAMETERS
BASED ON U. S. BUREAU OF MINES FIXED-BED MODEL
Assumed Time for Effluent Gas to Reach 10% of
the Inlet Gas Concentration (i.e., cycle time):
24 hours
INPUT VARIABLES
Gas Flow Rate (Appendix I) 61150
Initial Gas Composition 0.0035
Sorbent Loading at Sat'n. 0.1920
Rate Constant 25.00
Temperature 625
Viscosity 0.03
Gas Density 0.0383
Solid Bulk Density 42.00
Void Fraction of Bed 0.5
Particle Diameter 0.06
CFS
Lb. SO2/lb. flue gas
Lb. SO2/lb. sorbent
1/hr.
op
CP (Centipoises)
Lb./ft.3
Lb./ft.-3
Dimensionless
Inches
OUTPUT
Bed Height
(Ft.)
1
2
1
3
1
3
Col. Dia,
(Ft.)
20
20
30
30
40
40
Gas Vel,
(Ft/Sec)
0.58
1.16
0.58
1.74
0.58
1,74
AP
(In. H3Q)
0.95
4.21
0.95
10.44
0.95
10.44
No. Stages
L09
149
49
-132-
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theoretically based Bruceton fixed bed model. Although
this model has not been confirmed by experimental re-
sults, as were the fluid bed and dispersed phase models,
it has been used to design the commercial scale fixed bed
unit, for it is the only available tool for that purpose.
(3) Fluid-Bed Modeling
In the area of fluid bed sorption, two sorption
models have been proposed. The first model, a theoretical
one, has been presented by the USBM-Albany. The second model,
an empirically derived one, was presented by CEGB.
The rate constant used in the USBM-Albany model
is derived from the experimental data shown in Table 10 as
follows:
* An integrated form g the Amundson equation is
used r = •& = ky (1 - -^-)
where k = rate constant
ius = sorbent loading at saturation
u0 = initial sorbent loading
rearrange, dy_ =
1- w •*
Ms
/-Si- =Ky^tdt
<
(1- V )
integrate,
-p ln (1- v ) + us In (1- UQ) * Kyt
"• »s
simplify,
-Kyt
w = JJS -
-------
• Since sorbent loading (n) versus time (t) data.
are available from Table 10, equation (1) may be evaluated
for w , MO, and K using least squares estimation of non-
linear parameters. Such a technique has been developed
(Marquardt, D. W. , "An Alogorithm for Least Squares- Es t ima
tion of Non-linear Parameters", J. Soc . Indust, and'Appl.
Math., U No. 2, (1963) 431-441) and the program is
available from the "Share Program Library".
• Using all 29, 300°C data points from Table 10,,
the following .values' are .obtained:
K = 0.00794 min"1
PS = 0.1637 Ibs sulfur Ib sorbent
vo = 0.0066 Ibs sulfur Ib sorbent
• As a check on the calculated value of K, it
was compared to that reported in Figure 37. The Amundsen,
equation shown in Figure 37 is
yt = -22.5 in (1 -
where — u
M
->-•> r , _ , -
s = -22.5 and y = 0.16
S
K
therefore, K = 0.0071 min"1 which is a good check on
the calculated value of 0.00794 shown above,
To compare these models, the required sorbent
bed heights were calculated by each of these models;;for
various exit sulfur loadings, sample calculations are shown
in Tables 23 and 24. Th ? results of these calculations are
shown in Table 25 and in licate that the two models produce
results that are in good agreement but with the Blyth model
being more sensitive to sorbent loading than the Albany model,
-134-
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TABLE 23
SAMPLE CALCULATION OF FLUID BED SORBENT INVENTORY
USING THE USBM-ALBANY MODEL
w
y0w
where: _
G - Gas flowrate, ft /min. = 3,720,100
k = Reaction rate constant = 0.00794
M = S content of SO2/ gm S/ft3 S02
w = Rate of sorbent addition, gm sorbent/min
W = Weight of sorbent in the bed, gms
y = SO- concentration in gas at inlet, volume %,
= ^0.2146
y = Arithmetic average SO2 concentration in gas,
volume %, = 0.11803 @ 90% SO- removal
y = Maximum of saturation S load on the sorbetnt
s = 0.1407 gm S/gm P
Basis: 6% sulfur at exit, based on total solids weight
SULFUR LOADING @ INLET, p
u=2.6wt%S= Average sorbent loading at inlet; assumed,
based on USBM-Bruceton data for X-3F & 423 X-4E sorbents
SULFUR LOADING @ EXIT,
-------
TABLE 23 (Con't)
If sorption occurs as
Ib SORBED
Ib SULFUR
1 MOL
1 MOL S
80.064 Ib S0? . MQL S
MOL 803 32.064 Ib
= 2.497 Ib SCU
Ib S
.".wt gain = 2.497 x
By material balance
SULFUR @ INLET + SULFUR GAINED
TOTAL WT @ INLET+ TOTAL WT GAINED
SULFUR 0 EXIT
TOTAL WT @ EXIT
OR
IQOu + x
100 + 2.497 X
100 + 2.497 X
x = 3.999 t Sulfur sorbed per 100# sorbent feed
BALANCE FOR GAS STREAM
y0 = 0.214h% S02 (MWK Progress Report #3)
= 288,583 mol flue gas (MWK Progress Report i-3)
Eif
hr
(S = 557.37 mol SO^ . 1 mol s . 32.064 Ib S
hr mol S00 tnoT~^"
,\S = 17871 Ib S ^ ' "
,SO2 = G(yQ - ye) = G(0
AS02 = 288,583 mol f.g. . 0.0019314 mol SO = 557.37 mol SO-/hr
TTg.
Continued
-136-
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TABLE 23 (Con't)
SORBENT FEED RATE, w
w = 17871 Ib S . 100 Ib Sorbent Feed
~"KF~ 3.99 Ib S sorbed
w = 446,887 Ib Sorbent Feed = 3,381,445 gm
Rr mm
AVERAGE S02 CONCENTRATION, y
_>
2 V,
y = 0.11803% S02
0.2146 (1+
^ \
WT OF S PER UNIT VOLUME SO2, M
M = A MOL S\/32 064 Ib Sl/'mol SO \j^920R \/^54 gms\
(mol SO2J ynol S yD5§f t3SO2 Jfl060<'R/' IFy
M = 18.82 gm S
Substitute the above vlaues into the USBM model to find
sorbent inventory W
3,381,500 = 0.00794 0.11803 , /). 11803
0.2146 W 0.1407 '0.2146 (p
-WO. 00794 A 0.11803 H
.78-3 ,720, lUO 7 0.2146 I
W - 201,583,000 gms sorbent
or W = 444,000 Ibs sorbent
-137-
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TABLE 24
CALCULATION OF SORBENT BED HEIGHT USING THE C .E.G.B. -BLYTH MODEL
n)
where b = function of temperature - 9.0 @ T = 330"C
n = constant = 0.13
m = constant = 0.22
w = S concentration, @ time = t, minus the inlet or
residual S concentration, lbs/100 Ib sulfur free
sorbent
c = log mean SO2 concentration in flue gas, mol" fraction
t = sorption time, min. ,,.
d = particle diameter = in = 0.0359 in. ( '
L = feed rate of solids by material balance = 446,887 Ib/hr ( '
Basis: S on loaded sorbent = 6 wt %
5 on regenerated sorbent = 2 . 6 wt %
Assume sorption occurs as SO-,
Total weight gain = #S sorbed x jjg|°3 = #Sx-|
= 2.497 x #S
Inlet: wt S = 2.6#; total wt = (2.6) (2.497) - 6.4922
and sulfur-free-sorbent (S-F-S) =100-6. 4922=93 .5078#
100 = 2.7805#S/100#S-F-S
Outlet: X _
93. 5078+2. 497X ~ °-06
X = 6.5992#S
.., w = 7.0574 - 2.7805 = 4 . 2769#S/100#S-F-S
0 90% removal C « 0
C in 0^002146
2 5TWCT2T4T
n - 0.0002146^ 3D2g«L
— ---- - -- , - __-A^-~-
C = 0.0008388 mol SQ^
mbl flue gas'
-139-
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TABLE 24 (Con't)
bstituting into the Blyth Model
loo (1 - 4:2!69 = -o 22
J.og U --- -- 0.12
t = 48.91 min
@ average residence time of 48.91 min., the bed inventory is:
Inventory = L-t = 446,887 Ib/hr -
Inventory = 364,287 Ib solids
Since @ 40' dia. 10 beds are required.
Inventory per bed = 36429 Ib solids
@ 45 lb/ft3, solids density & 40' dia. bed
36429 Ib solids
Settled bed ht.
Settled bed ht. = 0.644 ft
45 Ib -ir • (40 ft)
) Weight mean average calculated from data in Table 4, USBM-Bruceton,
Quarterly Report, June 30, 1968.
:) Table 23, this report.
-i39-
-------
TABLE 25
A COMPARISON OF THE BED HEIGHTS REQUIRED FOR
FLUID BED SORPTION AS PREDICTED BY THE USBM-
ALBANY, AMD C.E.G.B. - BLYTH MODELS
USBM-ALBANY MODEL CEGB BLYTH MODEL
BED HEIGHT1 BED HEIGHT1
i:XIT SORBENT LOADING, %S (ft> (ft)
6 0.785 0.644
7 0.905 0.872
8 1.080 1.38
For ten 40' Diameter Reactors
-14 Q-
-------
Although a direct comparison of these models is not quite
valid in a theoretical sense (the temperature in the CEGB
model was 626°F whereas the temperature used in the Albany
model was 600°F), the good agreement of these models (within
0.3 ft. on bed height) indicates that either should be adequate
for design purposes.
The vessel heights in the proposed commercial
fluid bed unit will be set by criteria other than sorption
(e.g., disengaging height above the bed). Therefore, the
good agreement between the theoretical and empirical models
indicates that sorption in the commercial unit may be cal-
culated by the easy to use empirical model rather than the
more tedious theoretical one, once this theoretical model
has been verified by comparison with pilot plant results.
This verification has been provided by the data
in Tables 13 and 14 from the USBM-Albany testing program.
Figure 38 compares the calculated and measured extraction
as indicated by both gas and pellet analysis to that pre-
dicted by the theoretical model. Although the data exhibit
a good deal of scatter, some of which is most likely due to
difficulties in measurement and sampling, the relative agree-
ment of the pilot plant data with the sorption model can be
seen. The average exit sorbent loading obtained in these
tests was 0.072 grams of sulfur per gram of particle. At
this exit loading the Albany model predicts that for 90%
removal, a bed height of about 11 inches is required (see
Table 25). The bed heights which were used in this testing
program ranged from 5.5 to 6.8 inches and with a gas velocity
of about 2.7 to 3.2 ft/sec, gave an average sulfur removal of
90.2%, thus verifying the proposed sorption models.
(4) Dispersed Phase Modeling
An empirical model to predict the sorption of
-141-
-------
o
i-
x
LJ
0
UJ
cr
UJ
120
10
100
2 90
J-
o
2 80
70
60
50
40
o By gas analysis
a By pellet analysis
• Theory
40
o o a ffip
50 SO 70 SO 90 100 l!0
CALCULATED EXTRACTION, percent
FIGURE 38 - Comparison of Calculated and Measured Extraction, Long Run
20
-------
SC>2 in a dispersed phase system was developed by the USBM-
Bruceton, based on their pilot plant studies. A derivation
of this model is contained in Appendix E.
Experimental validif ication of this model has
not been conclusive due to the scatter exhibited by much
of the pilot plant data. A specific example of this
variability can be seen when the average sulfur loadings
from the C and D and E and F runs, which were presented
in section 2-C-(3) of this report, are plotted vs. ratio
of the sorption rate to the SO, composition of the flue
gas.
If the sulfur loading expression is represented
in the form
— — 1
where: r = average sorption rate, hr
y » gas SO, concentration, mol fraction
— 1
K = rate constant, *hr"
XQ = sulfur concentration of the feed
sorbent, Ibs/lb sorbent
xf = . sulfur concentration of the exit
sorbent, Ibs/lb sorbent,
then the exponent, m, may be obtained from the slope of a
log-log plot of the data as was done in Figures 39 and 40
for test series C, D, E, and F, respectively. The results
of these plots indicate that for the C and D test series
the value of m is + 2/3, while the value of m obtained from
the E and F series test data is -2/3. Statistically, the
regression equation for the C and D test series runs is above
the level required for significance if test run D-115 is
-143-
-------
FIGURE 39
Loge (r/y) VS Loge
(X0 + xf) • 10'
(TEST SERIES C AND D)
7.1
7.O
6.9
6.8
o>
o
6.6
6.5
6.4
6.3
6.2
O
o
SLOPE = O.63I
NOTE:
TEST D-U5 WAS OMITTED
IN THE CALCULATION QF
THE REGRESSION LINE.
O
D-II5
2.3 2.4 2.5 2.6 2.7 2.8 2.9
|(X0 + X«)- I05
Loge
-144-
Z.O £..
E*o -»- xf)- IP5
2 I
-------
7.6
7.4
7.2
O
FIGURE 40
VS Lo<,>° + X<>"°3
(TEST SERIES E AND F)
SLOPE — -0.693
4>
O>
O
7.0
6.8
6.6
6.4
O
2.8 2.9 3.0 3,1 3.2 3.3 3.4 3.5
Loge
f(X0 ->-Xf). 10*1
-145-
-------
omitted from the data. For the test series K and F runs,
however, the correlation coefficient of the regression
equation is 0.43, which is well below the required level
for significance. Furthermore, if the two greups of test
data are combined, the resultant slope of the regression
equation approaches zero, which in essence indicates that
the effect of sorbent loading on sorption is negligible.
This last result is highly unlikely for it has been shown
by the various kinetic models that sorbent loading has -a
distinct effect on the sorption rate. Therefore, it
appears that all of the data presented for these .tests
are not indicative of'the actual conditions in the dis-
persed phase sorber; thus, evaluation of the mode} based
on these data cannot be considered conclusive.
Although no firm conclusions could-be reached
as to the accuracy of the Bruceton model due to the scatter
of the data, the model-was applied in the process studies
for the commercial scale unit. Typical calculations using
this model are given in Appendix G, Tables G-89 and G-92.
It is expected that, if necessary, future pilot plant data
can be obtained so that adequate evaluation and validifica-
tion of the model can be made. For the present studiesr
however, considering that a suitable sorbent has not yet
been developed, the present model is believod to be ^'equate
for the commercial scale reactor desiorn.
3. Regeneration
a. Reactions
There are several differi-\t possible reactions which
could occur during regeneration and thus the heats of reaction
could vary widely depending on which of the reactions are
-------
actually involved. Heats of reaction reported in the literature/
MWK's calculated heats of reaction, and heats of formation for
the various reactions and compounds are shown in Tables 26 and
27.
As can be seen from the tabulated data in Table 26,
there is a considerable discrepancy in the MWK vs. CEGB cal-
culated heat of reaction for equation 2. Also, depending on
which of the several equations is the actual valid one, the
reaction could be endothermic or exothermic and the heat of
reaction could have a significant effect on the regenerator
heat balance. For example, USBM has reported a 150°F temperature
increase in fixed bed regeneration with reformed natural gas
in the 55,000 CFH pilot plant.
b. Status of Data
(1) USBM-Bruceton Studies
(a) 55,000 CFH Pilot Plant Results
1 - Fixed Bed Regeneration
USBM reports for the case of fixed bed
regeneration with reformed natural gas in the large pilot plant
(Pittsburgh Progress Report - Sec I - 12/31/67), at a space
velocity range of 15-20 hr"1 and temperatures from 1175-1250°F,
an overall utilization of 75% of the reducing gas for a sorbent
sulfur content reduction from 2.5 wt % to 0.9 wt % in 10 hours.
2 - Moving Bed Regeneration
>For the case of continuous gravitating
bed regeneration with reformed natural gas in the large pilot
plant (Pittsburgh Progress Report - Sec I - 3/31/68), USBM
reports that even at lower than optimum temperature, the reac-
tion proceeded readily, with little additional sulfur removal
after 2-1/4 hours. For an average of 24 sets of data taken
-147-
-------
TABLE 26
VARIOUS REGENERATION REACTIONS
2)
3)
4)
5)
6}
EQUATION
REDUCTION WITH H.
H-S
2NaAlO,
t
2NaAlO,
4CO
REDUCTION WITH CO
+ 3CO + H20 -*•
+ 4CO + H^O •*
4C0
2H00
2NaAlO,
2NaAlO,
AHr (K cal/g mole)
77°F (25°C)
MWK
52.8 52. 51
-1.81 -23. 32
9.7
13.5
-31.3
-29.7
~ 1200°F
(649°C)
MWK
44.9
-11.7
0.83
12.6
-35.8
-31.3
AVCO Thermochemical Calculations.
CEGB, Dec. 1965 - Bomb Calorir.ietry and Thermochemical Calculations
-------
TABLE 27
COMPOUND
Na2S04(c)
Na2S03(c)
Na2°(c)
Al2°3(c)
NaAl02(c)
C°(g)
C°2(g)
H2S(g)
H-0, .
2 (g)
HEATS OF
(1)
-330.50
-261.2
- 99.45
-399.09
- 26.416
- 94.052
- 4.77
- 57.80
FORMATION AT 25 °C (K cal/g mole)
SOURCE
(2) (3) (4)
-330 -> -331.55 -327.4 -330.90
-260.4 -260.6
- 99.4
-400.4
-270.84 -272.88
- 4.82
(1) Perry, "Chemical Engineers Handbook", 4 ed., 1963
(2) "JANAF Thermochemical Data" Tables
(3) "International Critical Tables" - (Data at 18°C)
(4) National Bureau of Standards-Circular 500 - "Selected Values of
Chemical Thermodynamic Properties"
-149-
-------
during 2 days of steady state operation, feed and exit gas
composition and flow rate for a temperature range of 1100-1150°F
and a sorbent feed rate of 150 Ibs/hr are tabulated in Table 28.
Overall reducing gas utilization v: s 50% for a sorbent sulfur
content reduction from 2.7 wt % to 0.9 wt % in 2-1/4 hours. USBM
states that stepwise reduction occurs during regeneration; this
was reinforced by chemical analyses of sorbent samples taken from
bed heights corresponding to different residence times. Longer
residence times indicated higher sulfide concentrations with the
remainder of the sulfur present as sulfate--no sulfite was detected.
Table 28 indicates that the sulfur removed from the absorbent
is equivalent to 1.33 times the H-S recovered, USBM states that
this may be due to sample loss or elemental sulfur formation in
the regenerator, and therefore filters have been installed-in the
exit line. It is also reported that an overall material balance
does not account for 3.5% of the carbon present in the reformed
natural gas feed. Visual examination of the regenerator reve^ i.ed
carbon deposition only at locations where stainless steel probes
protruded into the vessel. The quantity of the deposited carbon
was not reported but it was quite small.
(b) Bench-Scale Tests With Various Reducing Gases
In decreasing effectiveness, (amount ;>f suJ.tur
removal) the reducing gases used for regenerating the spent ab-
sorbent are reformed natural gas, hydrogen, polucer ja- , and
methane. According to USBM, t.»e rainimun itn.^crdture Tc ,. ^eforraecJ
natural gas is 1200°F. Data are presented in Table 29 from which
the following conclusions can be made:
* Complete regeneration was achieved with H? at 1250*F, assuming
an initial sulfur concentration of about 1%.
• Regeneration with reformed natural gas resulted in a lower
sulfur concentration than with H2.
-150-
-------
TABLE 28
OPERATING CONDITIONS DURING CONTINUOUS REGENERATION
(TEST NO. Gl)
Temperature, °F HOO to 1150
Sorbent Feed Rate, Ib/hr 160
GAS COMPOSITION COMPOSITION RATE
AND FLOW RATE (VOL PCT) (LB MOLE/HR)
Reformed natural gas :
H-
CO
CO,
Regenerator exit gas :
H2
CO
C02
CH
ttA
2
74.2
15.8
10.0
40.1
7.6
17.2
0.6
8.0
26.5
0.592
.126
.079
0.313
.059
.134
.005
.063
.207
SULFUR
(WT PCT)
2.66
0.88
.74
RATE
(LB MOLE/HR)
0.133
.037
SORBENT SULFUR LOADING
RESIDENCE TIME
(HR)
0 (inlet)
2.25
4.50 (exit)
Three cycles of absorption-regeneration were completed.
These represent an average of 24 sets of data taken during two days
of steady-state operation.
REF: USBM-Pittsburgh-Sec. I
3/31/68 Quarterly Report - Table 1
-151f
-------
TABLE 29
.REDUCTION
Temperature,
1200
1250
1290
OF SPENT ALKALIZED ALUMINA (21 PCT Na) WITH REDUCING GASES
LENGTH OF TEST, 4 HR
HOURLY SPACE VELOCITY, 100 HR"1
°F Hydrogen Methane Reformed Natural Gas1 Producer Gas
32.3 1.1 2.3 1.6 2.3 0.6 2.3 1.2
6.9 2.5 6.9 4.0
2.3 0.9
6.9 1.1
2.3 °'9 2 3 1 6 2304 2'3 °'9
6.9 1.1 *"* 1'f> ^'J °'4 6.9 1.9
1 77 pet H2, 13 pet CO, 2,0 pet C02.
2 i; pet H2, 30 pet CO, 4 pet C02, 50 pet N2, 2 pet CH4
Figures denote weight percent sulfur at
REF: USBM, Rl 7021
July 1967
Table 5, p. 31
-------
• Regeneration with CH^ was incomplete. Based on a limiting
value of 0.9%, equivalent to 100% regeneration, 50% recovery
of the sulfur was obtained with CH^ after 4 hours at 1200°F
and 1290°F.
• With the sorbent containing 2.3% S and based on a final
sulfur value of 0.9%, complete regeneration was obtained
at 1290°F with producer gas flowing for 4 hours at a space
velocity of 100 hr"1. At 1200°F, 80% sulfur removal was
obtained. With the higher sulfur content of 6.9%, reduction
to 0.9% was accomplished at 1290°F in 4 hours, by increasing
the hourly space velocity from 100 to 200.
(c) Regenerations with H^ and CO
1 - Bench Scale Results
USBM Pittsburgh Sec. II Quarterly Report
of 12-31-67 states that the regeneration rate is independent of
flow rate but dependent on temperature. A tabulation of some
bench-scale results (Pittsburgh Quarterly 3-31-68) using H2 and
CO is listed in the Appendix in Table F-68, showing the rate of
consumption of reducing gas as cc/min, assumed to be at standard
conditions of 14.7 psia and 60°F as a function of temperature.
Assuming the reaction rate to be zero order, a plot of these
data in terms of the rate constant (reducing gas rate of con-
sumption, moles/hr) vs. reciprocal absolute temperature should
give a straight line on semi-log paper but the data do not
exactly fit a straight line (Figure 41). If some data points
are discarded and a straight line drawn, the calculated activa-
tion energy for CO is 44 kcal/g mole and for H2 is 53 or 63 kcal/g
mole, depending on the line drawn.
Based on the USBM tabulated data (Table
F-68), CO reacts at a higher rate than H2 but regeneration is
incomplete; that is, residual sulfur content is high in compari-
son to that obtained in H., regeneration. For fines, however,
the rate of reaction was not significantly different for CO or
-153-
-------
r
7 2
o
I
S
8
7
6
5-1
ex
O
- 3
9
8
7
6
5-
4 -
UJ
_i
o
2
z
o
o
111
t-
«J
V.
FIGURE 4 I
ARRHENIUS PLOTS
H2 AND CO REGENERATION QFjSRACEJjIOJjgggggl
-1- 2
(I256FJ
68QC
I
(I22GF)
66 OC
I.
I
CO - E
CO- E
- E
« 42 KGAL/MOLE
- 45 KCAL/MOLE
59 KCAL/MGLt.
44 KCAL/MOLE
53-63 KCAL/MOUf
(IS4SKJ
ffcQC
O
1.03
I
1.05
I
1.07
I
I.O9
I
I.II
1.13
1.15
TEMPERATURE, -n-' x io5
5
H2
i
0
A
0
CO
,
A
SOURCE OF DATA
USQM PITTSBURGH
QUARTERLY 3-38-6©
USBM PITTSBURGH
QUARTERLY 9-30-68
USBM PITTSBURGH
OUART£R?J* §-3Q-68
_«...™.,M
INITIAL % S
2.5
9.8
r - ~^»-~' u.^..-.^.,,....^ ^.-^
n^RGY OF ACTIVATION,
' KCAL/MOLE
H2 | CO
53 -S3 | «* 1
45
4
3.5 j 39
40
42
-------
H2. For a small particle radius it appears that only surface
reactions occur and diffusion controlled reactions are not
apparent.
2 - Regeneration of a Composite Sample from
the Pilot Plant
A pilot plant composite sample of Grace No.
1, equivalent to the feed to the regenerator, was regenerated
using H^ and CO reducing gases. This alkalized alumina had a
non-uniform size distribution (approximately .02 to .10 inch
diameter) and contained an average sulfur content of 5.3 percent.
A charge of 4.25 grams in the reactor occupied a 4cc volume and
had a bed height of 8 mm. A flow of nitrogen was used while
bringing the reactor up to temperature before the H2 or CO was
added. Appendix Table F-69 summarizes the results of regenera-
tion at different temperatures.
One hour reduction tests with H- and with
CO were compared at several temperatures and the overall percent
of sulfur removal was determined. At temperatures of 620° and
640°C, the consumption of CO or H2 was still going on after one
hour, whereas at 660° and 680°C there was little, if any, con-
sumption of reducing gas after 30 to 40 minutes. With hydrogen,
the amount of sulfur removed increased as the temperature was
increased but the CO experiments showed an inverse effect of
temperature on the sulfur removal, even though the consumption
rate of CO increased at higher temperatures. USBM stated that
this could be due to the possibility of (1) a more stable
sulfide formed at higher temperatures; (2) oxygen reaction
favored at these temperatures; and (3) carbon deposition takes
place at 660° and 680°C.
The maximum rates of consumption of the
reducing gases tabulated in Table F-69 were converted to moles/hr
and, assuming a zero order reaction rate, included as part of
-155-
-------
Figure 41 (6-30-68 data) . A least squares line was then
drawn through the data points. These data predict higher
regeneration rates but approximately similar energies of
activation in comparison to the previously discussed 3-31-68
data of Figure 41.
Reducing gas mixtures of H2 and CO
were run using the same pilot plant composite .sample. The
last three tests in Table F-69 show the results of these
experiments. H_ and CO mixtures removed about as much
sulfur as pure hydrogen. Regeneration with 2:1 and 1:2
mixtures of H- and CO gave practically the same percent
sulfur removal over a one hour period at 640°C but the
1:2 H2: CO combination used less reducing gas to -remove
the same amount of sulfur. When a Isl ratio was used
at 660°C, the sulfur removal was comparable to the 640CC
tests.
One further feature of= these H2 and
CO tests should be pointed out. The amount of -water vapcr
formed during the regeneration reaction is decreased.- USBM
states that when using a 1:2 H2:CO mixture at 640.°C, the
quantity of water formed is only about a quarter of the amount
of water formed during a H- regeneration at the same temperature.
A regeneration gas of high CO content would reduce the amount
of water vapor that contacts the sorbent at regeneraticr tem-
peratures .
3 - Half-Hour Regeneration _
Pure H2 and pure CO regeneration experi-
ments were conducted using the original Grace sorbent (Grace No.
1) which was spent in the laboratory (9.8 percent sulfur and 1.0
percent carbon) . Instead of the usual hour long test, the
reaction time was reduced to a half-hour so that. the regeneration
would be incomplete and differences in the relative extent of
sulfur removed could be demonstrated better as a function of the
-------
reducing gas and the reaction temperature. Table F-70 in the
Appendix lists the results of regenerations at various tempera-
tures. H2 consumption was completed after approximately a half-
hour for the 660° and 680°C tests while at 620° and 640°C, it was
still being consumed after one hour. H2 regeneration at 680°C
removed approximately\ three times as much sulfur as did CO re-
generation at 680°C. The percent sulfur removal at 30 minutes
varied from 82 percent at 680°C to 26 percent at 620°C for H2
regeneration, whereas the CO regeneration removed approximately
30 percent sulfur at three temperatures (680, 660, and 640°C)
and only 23 percent at 620°C. During the CO tests, the total
consumption of the reducing gas and sorbent weight loss increased
with reaction temperature. A few former runs using CO indicated
an inverse effect of temperature on sulfur removal (Table F-69).
The possibility of carbon deposition on the alkalized alumina
causing high CO consumption without removing appreciable amounts
of sulfur with increasing temperature was ruled out by the
results of these half-hour tests. Carbon analyses did not
indicate any increase above the one percent in the charge; in
fact, the lowest CO reaction temperature showed the most carbon
present. Oxygen removal is indicated.
Figure 41 also includes the data of
Table F-70 considering the maximum rate (moles/hr) of consump-
tion of reducing gas as the assumed zero order rate constant.
A least squares line was drawn through the points (9-30-68 data).
and the energy of activation for H_ and CO calculated as shown
on the plot.
In comparing the three sets of data plotted
on Figure 41, those obtained for the case of an initial sulfur
content of 9.8% indicated a higher reaction rate than those ob-
tained for 5.3 and 2.5% sulfur. In every case CO reacted at a
higher rate than H2 but sulfur removal was incomplete; that is
residual sulfur content is high in comparison to that obtained
-157-
-------
in H2 regeneration. It should be noted that the reaction rate
constants plotted on Figure 41 are for an assumed zero order
reaction. In any kinetic analysis for determining the order
of reaction, concentration-time profiles are needed. The only
information given in Tables F-68, F-69, and F-70 are rates of
consumption of reducing gas and initial and final sulfur content.
The rate of sulfur removal as a function of time is not included..
It was observed in Table F-69, however, that 1:2 and 2:1
tures of H2:CO removed identical amounts of sulfur for the same
maximum gas consumption rates. Since in a zero order reaction,
the rate of change of any reactant is independent of the con-
centration of that reactant, this was assumed to be the case for
H2 and CO regeneration for each set of data. Therefore, Figure^41
is a plot of the rate constants, assumed to be moles/hr reducing
gas consumption rate since the sorbent charge was the same for
each test (maximum where specified from the appropriate tables) ,
vs. reciprocal temperature on semi-log paper. These are
Arrhenius plots for an assumed zero order reaction. Although
there is a good deal of variation in the energy of activation
calculated from the different sets of data, in all cases the
reaction rate constant increases with temperature, indicating
that the higher temperatures should improve regeneration.
(d) Albany Batch Regeneration Studies
1 - Data from Modified Thin Bed Test Unit
Bate!: regeneration Vests vit,'. th^ ^.ctfified
thin bed test unit were carried out with 5 " •>•>'<. '1 =? t -z3 rM'orrr-a re-
ducer gas (50% N2, 10% C02 , 15% CO, 20% H2 , and 5% H_O) . The
results of these tests are tabulated in Appendix Table F-71.
Figures F-79, F-80, and F-31 are plots of regeneration time vs.
residual ?-lfur content v/ith various temperature parameters ca
semi-log paper, for Grace No. 1 sorbents having an initial
sulfur content of 8.5, 5.5, and 2.3% respectively. These plots
-158-
-------
can be resolved into two relatively straight lines which seem
to indicate that two parallel first order reactions are taking
place. There is one rate-limiting reaction of principal in-
terest. The second, slower reaction is not observed until the
sulfur load is below the sulfur load of the original sorbent
and this reaction represents the removal of residual sulfur.
The first order rate constants tabulated in Table F-71 were
obtained from the slopes of those portions of the curves of
Figures F-79, F-80, and F-81 representing the removal of sorbed
sulfur, the rate limiting reaction. A crossplot of these curves
indicating the time required for 100% regeneration of the sorbed
sulfur as a function of sorbent loading, for various parameters
of temperature, is included as Figure 42. Calculated percent
regeneration for all runs are plotted as Figures F-82, F-83, and
F-84. The horizontal lines drawn across the figures represent
the level at which all sorbed sulfur has been removed. The
results show that at 730°C (1346°F) all sorbed sulfur can be
removed in approximately 10 minutes.
The regeneration rate appears to be highly
temperature dependent; as the temperature is reduced the regenera-
tion rate drops very rapidly. The rate constants in Table F-71
drop from about 0.2 min"1 at 730°C (1346°F) to less than 0.005
min"1 at 650°C (1202°F) . An Arrhenius plot of these rate constants
(Figure 43) indicates that the activation energy is approximately
91 K cal/mole. In the temperature range of 650-700°C (1202-1292°F),
a 10°C rise in temperature increases the regeneration rate by a
factor of 1.66.
2 - Comparison of CEGB-Blyth Regeneration
Model with USBM-Albany Data
A study was made comparing the CEGB-
Blyth kinetic regeneration model to the USBM-Albany regeneration
data discussed above. The Blyth model relates the regeneration
-15?-
-------
FIGURE 42
100% REGENERATION TIME
VS
SULFUR LOADING
I
f 9
a
7 •
6
5
s
o
(T
O
Z
o
(C
Ul
Z
Ul
o
o
o
UJ
Z
I •
9
a
7
6
5'
9 3
REFERENCE:
USBM ALBANY PROGRESS
REPORT. FUCU^ES !„ 2
3, JANUARY -1969.
SUl FUR LOADIHC. T S
4
•160-
-------
TEMPERATURE, °C
730 700 680
650
o Staffing material 8 pet S
a Starting material 5 pet S
O Starting material 2 pet S
.002
0.96 0.98
1.02 1.04
1,000/T,
LIO
FIGURE 43 - Arrhenius Plot of Regeneration Rate Constants
•t-3t
-------
time to a log function of the sulfur concentration on the sorbent
as expressed in the following equation:
> = ~Kt
where t = regeneration time (min.)
W = sulfur concentration on sorbent at time t Ibs S/
100 Ibs sorbent. W excludes that fraction of the
sulfur which oc-cupies relatively inaccessible sites
and neither pai ticipates in regeneration, nor in-
fluences subsequent uptake.
W = sulfur concentration (excluding the inactive fraction
mentioned abovt;) on the reactant at the start of
regeneration (Ibs S/100 Ibs sorbent).
K = rate constant = f (temperature and sorbent. pro-
perties) = 0.011 at 700°C.
Two different inlet sorbent loadings were used, 8.85 and 5,67%Sf
at 700°C and the results obtained from the Blyth model were
plotted with the actual USBM-Albany regeneration data obtained
for these sorbent loadings (Figure 44). The Albany data were
taken from Table 1 of Albany Progress Report of January, 1969,
and the intersititial salfur loading used in the Blyth model
was taken as 0.22%S based on the Albany data.
From Figure 44 it car be i'een thc.t the
r ^generation model satisfies the basic characteristics o:: tne
regeneration reaction, namely an initial high rate which levels
off and approaches a zero reaction rate as the sulfur concentra-
tion approaches its interstitial valae. However, the model gives
? lower rate of regeneration than that which is predicted by tn^
data, possibly due to the different sorbents usxi. The empirically
derived value of K used in the Blyth model varies depending on the
grade of sorbent used, and the CEGd data were obtained using an
early grade Peter Spence sorbent whereas th'e Albany data are based
on Grace No. 1 sorbent.
-162-
-------
\
\
\
\
\
\
FIGURE 44
COMPARISON OF THE BLYTH REGENERATION MODEL
WITH
THE USBM ALBANY REGENERATION DATA AT 700°C
\
\
SORBENT LOADING (% S) VS TIME (MINUTES)
\
\
\
\
\
\
\
\
INTERSECTION CORRESPONDS
TO REMOVAL OF 100% OF
SORBED SULFUR.
USBM ALBANY THIN (FIXED) BED DATA
USBM ALBANY THIN (FIXED) BED DATA
. CALCULATED, CEG8 BLYTH MODEL
NOTE:
USBM DATA OBTAINED ON GRACE NO I SORBENT.
CEGB MODEL BASED ON DATA USING PETER
SPENCE SORBENT.
REFERENCE:
USBM ALBANY PROGRESS REPORT.
TABLE I, JANUARY, 1969.
SORBED SULFUR
, RESIDUAL SULFUR
30 40 50 60 70
TIME, MINUTES
90
100
NO
120
-------
(e) Evaluation of Various Sorbents
In the process of evaluating new S02 sorbents,
differences in gas consumption rate during the regeneration reac-
tion were found. In tests carried out at 1000 hr" space velo-
city (Table F-72), it was demonstrated that sorbents without iron
had slow rates of gas consumption, low weight losses, and there-
fore, little sulfur removal. On the other hand, iron-containing
sorbents showed fast consumption rates, high weight losses -and
more complete sulfur removal. High reaction rates may be re-
sponsible for a loss in hardness and resultant high attrition
rates. Rapid chemical changes could alter the large crystal-
lites by not allowing sufficient time for proper reorientation
of the crystalline structure. The regeneration of the sorbents
without iron is slower but the overall smoothness of the reaction
might cause less strain on the pellets and possibly could result
in smaller decreases in hardness.
1 - Kaiser DN-112-F, Kaiser B, Kaiser DN-1Q5,
Peter Spence
Figure 45 illustrates the differences in
the consumption rate of H2 for two Kaiser alkalized aluminas
with and without iron present, DN-112-F and B. The sorbent
containing iron, DN-112-F, reacted rapidly wiih H2 and was
apparently reduced completely within 45 minutes as indicated
by little or no additional H2 consumption. Kaiser B, contain^
ing no iron, had a slow rate of gas consumption and •?;,£ ?"r.:.ll
consuming H-, slowly even after 60 minutes. The oth^r £ ample?
without iron, Kaiser DN-105 and Peter Spence produced H- curves
similar to the Kaiser B sample. The data of Appendix .Table
F-72 show the sulfur removal obtained in the tests but do not
indicate the base level sulfur concentration; i.e., the point
at which all the sorbed sulfur is removed, Similarly, Figure 45
indicates point values of H2 consumption up to over 70%, but
the overall reducing gas utilization would probably be less
than 70%c or some integrated average of this curve, depending
-------
c
V
u
w
t>
a
z~
o
p
a.
2
D
(/I
2;
O
O
«a
Z
O
FIGURE 45
20
10
8
16
I ____ L
24 32
TIME, minutes
Differences in the H~ utilisation at 660°C and 1000 h:
ssace velocity of Ksiscr DN'-I12-F containing ircr. and
Kaiser (5.x withojt iron.
-165-1
-------
on the time required to remove the sorbed sulfur. For the
curve of Kaiser DN-112-F of Figure 45, in determining the
actual degree of H2 utilization, that portion of the curve
from 32 minutes to 60 minutes would be considered as wasted,
since the H2 consumption is very small, less than 10%.
2 - Grace No- 2, Kelsei
DN-113-F
Figure F-85 represents the H2 utiliza-
tion of iron-containing alkalized aluminas at 660°C and 1000
hr"1 space velocity. There was a slight difference in the
H2 consumption curve of Kaiser QN-113-F (similar to Kaiser
DN-112-F) when compared with Kaiser PR-15 and Grace No. 2
materials at 660°C. The DN-113-F seems to be more reactive-
than the PR-15 and Grace No. 2 sorbents; its H2 consumption
goes rapidly to a single maximum and then the rate drops off,
whereas the other two sorbents show two maxima before the rate
falls off. At 680°C, the rate of consumption of all three
sorbents showed curves similar to the Kaiser DN-I13-F curve
at 660°C, with a little higher maximum utilization of H^ .
3 - USBM 423-X-4E, USBMX-3-F
The Bureau preparations of X-3-F and
423-X-4E reacted with H2 in a manner similar to the other
alkalized aluminas containing iron. Figure F-86 shows a
family of curves of 423-X-4E regenerated with H2 at various
temperatures. At the hiahest t^-ipernture , 660°C,
consumption is rapid and the reaction i
-------
(H-1-5-74-A) spent in the pilot plant. The use of 1:1 H_:
CO mixtures increased the utilization of reducing gas when
compared to pure H- at 640-and 660°C. The types of curves
produced are apparently a function of both the gas composi-
tion and temperature. Grace No. 1 alkalized alumina reduced
with H2 and CO mixtures indicated as much if not more sulfur
removal than when pure H2 was used. The opposite effect was
found for the Kaiser PR-15 as shown in Table F-72. When it
was regenerated with H2 and CO mixtures at 660 and 680°C,
there was less sulfur removed than when the reducing gas was
pure H2 at these temperatures.
5 - Two-Step Regeneration Reactions with
Sorbents 1-4
The use of CO2 as a secondary treatment
after the primary reduction reaction seems promising. Table
F-73 demonstrates the effect of using a secondary treatment
with CO2, mixtures of CO- and water vapor, and mixtures of
N2 and water vapor, after primary reduction reactions with
H-, CO, and their mixtures. The first series of tests using
Kaiser DN-113-F and Kaiser DN-112-F indicates the effect of
temperature during the secondary reaction with CO2. To date,
the best temperature found for the CO2 treatment has been
300°C, when compared with the results at 600 and 100°C. A
C02 treatment at 300°C almost doubled the sulfur removal ob-
tained when using CO2 at 600°C. The C02 treatment at 100°C
was ineffective in removing any appreciable amount of sulfur
from the Kaiser DN-112-F sample, only 25 percent total sulfur
removal being achieved.
The addition of water vapor to the C02
gas improved the overall sulfur removal over the dry C02
treatment, but the presence of water vapor in N2 (3 percent
level) was not efficient.
-167-,
-------
The two-step regenerations carried out on
the peter Spence material showed poor sulfur rciroval properties.
However, this poor response could be accounted for since the
primary reduction reaction with CO was incomplete during ,the one
hour reaction time.
The remainder of Table F-73 indicates
the improvement in sulfur removal using C02 as secondary treat-
ment for various sorbents. The Bureau's X-3-F showed sulfur
removals as high as 86 percent when using a 1:1 H2:CO primary
reducing gas at 640°C followed by C02 at 325°C for a half hour.
6 - Bureau Prepared NaOH and Alumina Hydrate
Sorbents
These sorbents were prepared with and
without iron and their sodium content ranged from li to 22
percent. Preparations containing iron and those without iron
showed reaction rates similar to alkalized aluminas tested
previously. Table F-74 shows the results of regeneration
tests and crushing strength changes of sorbents prepared
with NaOH and alumina hydrate. Preparation No. 467 (1 percent
iron as red mud) showed 85 percent H_ utilization at the maxi-
mum consumption rate as compared to only 15 percent H» utiliza-
tion for No. 466 having essentially the same composition except
that there was no iron present. A comparison between prepara-
tions No. 470 (1 percent iron as Fe2Q.>) and No. 472 (no iron),
shows similar results.
The 2odium content Q/ these sorbents
also affected the reducing gas consumption rate. The lower
sodium content preparations have a lower capacity to react
with S02 but the material was harder and the regeneration
reaction gave a higher maximum rate than with the high sodium
soxbents. Sorbents Nos. 473, 474, 467 ar$ similar in composi-
tion except for the amount of sodium present. No. 473 (11
-ifcg-
-------
percent Na) had the highest maximum rate of consumption but
had the lowest, sulfur loading. No. 474 (17 percent Na) showed
a maximum reaction rate between No. 473 and No. 467 (22 percent
Na). The hardness dropped 50 percent for sorbent No. 467
(highest Na) but only 21 and 24 percent for No. 473 and No. 474,
after the H2 regeneration at 660°C. The sorbents without iron
showed little reaction and their hardness was retained.
Two step regeneration using CO and then
C02 on sorbent No. 473 indicated similar results to those ob-
tained previously, namely that a secondary CO2 treatment at
300°C increases the sulfur removal.
7 - UOP Alumina Pellets Impregnated with
Active Compounds
The sorbent formed by sodium aluminate
and ferric nitrate impregnated on alumina retained its high
hardness after H, regeneration, but the regeneration was in-
complete (sulfur removal only 60 percent - Table F-74). Figure
F-88 illustrates the similarity of the two rate curves for 640
and 6808C, and shows both regenerations to be completed within
30 minutes even though the sulfur removal was not more than 60
percent. The regeneration with H- was very fast and the reaction
did not show the temperature dependency demonstrated by the
alkalized aluminas. The UOP-V-1 sorbent (Na2SO4 and Fe2(SO4)3
impregnated on alumina) showed softness even before the regenera-
tion reaction, apparently due to some interaction between the
alumina carrier and the sulfate.
8 - Kaiser PR-18-1 Alkalized Alumina
(15% Na and 0.6% Fe)
This alkalized alumina reacts in a manner
similar to the other alkalized aluminas but the H2 regeneration
curves at 680° and 660°C tend to flatten out at the maxima,
rather than reaching a high peak and then falling off rapidly.
This flat-topping may indicate a rapid but controlled diffusion.
-169-
-------
The change in maximum consumption rate of the reducing gas' at
640*C reaction temperature is illustrated in Figure F-90. The
higher the CO content of the reducing gas, the greater the maxi-
mum consumption rate. Table F-74 indicates that the regenerated
pellets have low hardness values but thay are better thin Grace
alkalized aluminas (Table F-75) .
9 _ Grace No. 1 and Grace No. 2 Comparison
Table F-75 compares the regeneration of
totally spent Grace No. 1 and Grace No. 2. The two-step re-
generation with CO at 680°C followed by CO2 from 680°C to room
temperature was not effective in either case. Both Grace
alkalized aluminas when spent to SO2 saturation show similar
regeneration data at the same reaction temperature. At 700°C
the reaction rates are not much different than at 680 °C, other
conditions being the same. A red deposit of residual sulfide
was found on the pellets even at 700°C where the sulfur removal
was over 80 percent. Regeneration at 640 °C took two hours t
-------
of hydrogen pressure above 70 mm Hg.
l
Regeneration time increased rapidly with increased
initial sorbent loading. For a fixed loading, re-
generation time decreased rapidly with increasing
temperature (Table F-77 and Figure F-91).
In the presence of pure CO, the Reaction does not
go to completion; i.e., after constant weight is
reached in the H--CO mixture, exposure to pure
H2 will produce further weight loss (Table F-78).
As much as 25% of the sorbate is unremoved.
In the presence of the reaction products, H-S
and H2O, the regeneration rate was not signifi-
cantly lowered but ceased at 75% completion a,nd
pure H- was required to complete the process.
CEGB reports (June 1965) that rates measured at
654°C for a constant H2S pressure were found to
be independent of H-S partial pressure but con-
! ' ^
siderably lower than those determined in the
absence of H-S.
It appears that thermal regeneration decreases
as sorption time increases. At a given sorption
time, there seems to be no effect of flue gas SO2
concentration (Table F-79).
The effect of the humidity of the flue gas and
initial sorbent loading on the temperature depen-
dence of regeneration time was not measured. Few
data were taken and the temperature effect from
550-650*C was obtained on sorbents of low initial
loading on which the S02 was sorbed from dry flue
gas whereas the temperature effect from 650-700°C
was obtained on sorbents with a high loading, loaded
in moist flue gas (Table F-80).
-171-
-------
(3) W. R. Grace & Company Studies
A two-step regeneration procedure, consisting of
a high temperature reduction step with a reducing agent such
as CO and a low temperature hydrolysis step with water and CO2
to complete the regeneration, is undergoing evaluation presently.
The theoretical equations proposed are:
Reduction: Na2S04 + 4CO * Na2S + 4C02 @ 1200°F
Regeneration: 2Na2S + CO2 + 3H2O •* 4NaOH + H2S + COS
@ >212°F
The reduction rate of Na2S04 by CO is highly temperature .
sensitive. Grace.No. 2 beads loaded to approximately 5%
S were reduced at three * temperatures under two CO flow
rates (Table F-81) , and substantial exotherms, 104 to 160°F,.
were observed in all cases. Table F-81 also indicates that
regeneration of the reduced Na2SO. is accomplished more
rapidly with a C02 - H2O blend than with a N2 - H20 blend.
Two temperatures, 340°F and 500°F, and three CO2 to H2O
ratios were examined. Fastest rates were achieved with a
3/1 CO2 to H2O ratio at a temperature of 340°F. Exotherms
varied from 60 to 100°F, with the largest occurring under
the above conditions. Accordingly, a two-step regeneration
reaction shows promise in the reduction of the regeneration
time. Also, if CO and CO2 are used for regeneration, the
formation of steam during the re-Suction process wou.'.d fce
eliminated. Therefore, the amount of water va^c-c t-At con-
tacts the sorbent during regeneration would be reduced. In
order to eliminate steaming effects at high temperatures when
hydrogen is used as the reducing gas, reformed natural.-gas
is preferable to producer gas since the CO content is higher
and the H2 content lower.
-------
c. Summary of Status of Data
The exact mechanism by which the sorbent is regenerated
has not been defined clearly but sufficient data are available
to permit a preliminary design of a regenerator to be made. To
allow for the uncertainties in the regeneration step, safety
factors have been included in the regenerator design but actually
cause only modest increases in the vessel cost.
The general conclusions that can be drawn from the re-
generation data available are:
• Increasing the temperature increases the regeneration rate
and temperatures about 1300°F or higher appear to be indicated
to obtain a reasonable hold-up time in the regenerator.
• Pure H_ or mixtures of H2 and CO (i.e., reformed natural gas)
give satisfactory regeneration but pure CO does not. Natural
gas itself does not give complete reduction. Other reduction
methods (e.g. two-step using CO, CO_-H2O) are being evaluated
but no firm conclusions can be drawn at this time.
• If the sorbent does not contain a small amount of iron ( 1%)
then commercially unacceptable regeneration rates are obtained.
* The effect of repeated sorption-regeneration cycles on sulfur
build-up in the sorbent has not been determined. That is, the
life of the sorbent owing to loss of activity (sorption capacity)
rather than to attrition is presently unknown.
• There are some indications of sulfide formation but its effect
on the process has not been determined; possible effects
might be stripping of SO2 in the sorber or higher reducing
gas consumption.
* Some data have been obtained which indicate little, if any,
effect of H2S back pressure on the rate or degree of regenera-
tion. This is important since backmixing of the gas is likely
in a fluid bed regenerator.
-173-
-------
B. PRELIMINARY PROCESS DESIGN
1. Pilot Plant Studies
At the request of NAPCA, a quick process design and approxi-
mate cost estimate for several different demonstration-scale
alkalized alumina (PA} processes were prepared. Each of the AA
plants is sized to remove 90% of the sulfur from the flue gas
generated by a 50 megawatt (MW) power plant. The -two basic
designs differ primarily in reactor design; one case considers
the use of a dispersed phase reactor similar to the one in
the USBM-pruceton pilot plant, while the other utilizes a
fluid bed reactor.
Originally, the fluid-bed reactor was to have three beds
in series and the flue gas sulfur oxide content was that used
by the Central Electricity Research Laboratory (CERL) in their
fluid bed work. This sulfur concentration corresponded ,to a
flue gas obtained by burning a coal containing about 7% sulfur.
Further, the flue gas temperature was specified as 264°F, thus
requiring a flue gas preheater for evaluating sorption teir^p-cma-
tures higher than this. Regeneration conditions were also
based on CERL data.
The CERL regeneration rate correlation shows a relatively
slow rate at .650°C (1202°F) compared to that at 700°C (1292°F)
(e.g., about 25% sulfur removal after two hours vs. about 9^*>,
so to avoid excessive hold-up in the regenerator, the higher
temperature was used for the regeneration step in the fluid-
bed design.
The dispersed phase design was to be based on flue gases
obtained from burning a 3% S coal, with these gases available
at 600°F. Regeneration conditions were based on Bruceton pilot
plant practice, namely, a regeneration temperature of 1230*F
and a four hour holdup. Preliminary process designs ar.3 cost
estimates were prepared for these original cases. It was
•J.74-
-------
subsequently decided to extend the study and prepare a design
and estimate for each scheme (i.e., fluid-bed and dispersed
phase) that could be directly compared. Therefore, to provide
a case for direct comparison, a 700°C regeneration temperature
was also used in a dispersed phase case design. It was further
decided to use only one bed in the fluid bed reactor since
calculations showed that three beds were not required to achieve
the desired sulfur removal.
Flow sheets for the demonstration-scale unit are basically
the same as those developed for the base case design (1000 MW
power plant), which are discussed in other sections of this
report. (For 50 MW flow sheets, see MWK Progress Report #8).
The factors common to all estimates are (a) 90% removal
of the sulfur oxides from the entering flue gases, (b) natural
gas to generate reducing gases for sorbent regeneration; and
(c) a Glaus plant to convert hydrogen sulfide to sulfur. It
should be pointed out that approximate estimating techniques
have been used in developing the estimates. The experimental
data are incomplete and the process designs were not opti-
mized so more precise estimating procedures did not seem jus-
tified at the time. The battery limits estimates, and the
total capital requirements of the erected plants (including
offsites and sorbent inventory) for the basic flow sheets,
along with the alternatives, are given in Table 30.
Referring to the dispersed phase cases, it is obvious
from the tabulated plant costs that for these designs flexi-
bility in sorption temperature is costly, i.e., DP-2 costs
about §600,000 more than DP-1. It is also apparent that an
increase in regeneration temperature from 1200°F to 1292°F
causes a significant increase in plant cost: about $500,000,
comparing cases DP-2 and DP-3. This latter increase is due
primarily to more costly vessels required arising from the
decrease in allowable metal stress at the higher temperature.
-175-
-------
TABLE 30
BATTERY LIMITS ESTIMATES AND TOTAL CAPITAL REQUIREMENTS* (50 MW PLANT)
Dispersed Phase (3% S in coal)
1. Base Case:
600°F sorption
1200°F r.egeneration-4 hrs
2. 264-655°F sorption
1200°F regeneration-4 hrs
3. 264-655°F sorption
1292°F regeneration-2 hrs
COST, THOUSAND DOLLARS
BATTERY LIMITS TOTAL
4,000
4,535
5,030
4,905
5,515
6,080
Fluid Bed
(includes electrostatic precipitator)
1. Base Case:
3 beds, 7% S in coal
264-655°F sorption
1929°F regeneration-2 hrs
2. 1 bed, 3% S in coal
264-655°F sorption
1292°F regeneration-" h>:s
7,560
5,735
8,695
6,895
* Includes offsites and sorbent inventory
-------
Although the vessel design has not been optimized, a brief
look at various alternatives (e.g., brick lined) showed no
obvious substantial savings to be realized from different
types of design. Therefore, the conclusion that the higher
regeneration temperature causes a substantial cost increase
may still be valid for optimized designs.
The investments for the two fluid bed cases show that a
substantial decrease ($1,800,000) is obtained by designing
for one fluid bed (rather than 3) and 3% sulfur in the coal.
The investment required for a fluid bed plant is about 13%
higher (-$800,000) than required for a dispersed phase plant
when the two are compared on a common basis, i.e., DP-3 and
PB-2.
Considering the approximate nature of the estimates, the
difference in the plant investment required by the two dif-
ferent comparable designs (DP-3 and FB-2) is too small to
indicate a clear superiority of one system over the other.
The purpose of preparing these 50 MW plant designs was
to determine the approximate cost of building and operating
a pilot plant which at one time the Central Electricity
Generating Board (CEGB) was considering. CEGB was looking
for partners to share the expense of such a plant and NAPCA
wanted an independent estimate of the cost in case they decided
to join CEGB in the venture. The project was later abandoned,
presumably owing to their lack of success in developing a com-
mercially acceptable sorbent.
2. Commercial Scale Studies
a. Dispersed Phase
(1) Dispersed Phase Process Description;
Sheet '
The following process description is for the
Alkalized Alumina process scheme shown on flow sheet P 3197-D,
-177-
-------
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The scheme is based on removing 90% of the SO- in the flue
gases from a 1000 MW power plant burning coal containing 3%
sulfur. It should be noted that the quantities shown are
for only one train of equipment whereas two parallel trains
are required when the power plant is running at full capacity.
Hot flue gases (about 600°F) from the power plant enter
the bottom of the dispersed phase reactor, D-l, and flow upward
through the reactor at about 25 feet per second with a total
gas residence time of about three seconds being achieved. In
the reactor the flue gases contact the sorbent particles under
conditions designed to effect removal of 90% of the SO, present
in the entering flue gas. The treated flue gas effluent from
the top of the reactor passes through a one-stage cyclone to
recover sorbent particles greater than about 60 mesh, and then
it sent to the power plant stack after first flowing through
the air preheater and electrostatic precipitator. An alter-
native design would relocate the electrostatic precipitator
upstream of the S02 removal plant. The air preheater and
electrostatic precipitator are standard pieces of equipment
in power plant operations and are not part of the SO2 removal
process.
Sorbent is introduced into the D-l reactor at three dif-
ferent points: a heavy particle recycle stream near the top
of the reactor, a fines recycle stream which enters near the
bottom of the reactor, and a stream of regenerated sorbent
also near the bottom. The heavy particle stream comprises
particles whose settling velocity is greater than the super-
ficial gas velocity in the reactor, and thus these particles
fall downward through the reactor and are withdrawn from the
bottom. The regenerated solids consist of a range of particle
sizes, including particles with settling velocities both
greater and less than the reactor gas velocity. The heavier
particles settle to the bottom and are withdrawn through the
bottom exit line while the lighter particles are carried
-179-
-------
upward and are recovered in the reactor cyclones. The fines
recycle stream is comprised of a portion of the fines recovered
in the reactor cyclones.
The bottom exit solids are transported via a dense phase
riser line to the top of the reactor where the fluidizing air
and solids are separated in the gas-solids separator, L-2.
The air leaving L-2 is combined with the flue gas leaving
D-l for eventual discharge to the atmosphere via the power
plant stack. The solids in the bottom of L-2 are maintained
in a fluidized state and are withdrawn through a standpipe
and split into two streams: a recycle stream to D-l, and a
side stream which is sent to the regeneration section.
The lighter sorbent particles carried overhead with
the flue gas move through the reactor at different velocities
depending on the slip velocity of the individual particles
(i.e., slip velocity is the difference in gas velocity and
particle settling velocity), and therefore have different
residence times in the reactor. Thus the heavier of the
overhead particles have the lowest slip and hence highest
residence time. The lightest particles have the shortest
residence time. However, these short residence time particles
have a high surface area to weight ratio. Therefore, it has
been assumed that most of these overhead particles will ha/e
either sufficient residence time or sufficient exposure of
surface area such that they will be loaded to design capacity
during one pass through the reactor and will not hsve to ce
recycled. In case one pass is not sufficient, however, pro-
vision for recycle of some of the fines back to the reactor
has been included on the flowsheet.
Entrained fir.es in the flue gas effluent from the top of
D-l are recovered in cyclones L-1A-1F. The large volume of
gases to be handled require the use of several cyclones in
parallel. The flue gas effluent from these cyclones is sent
-180-
-------
to the power plant stack, while the recovered fines (+60
mesh particles) may be recycled to the reactor or may be
combined with the side stream of heavier particles from
L-2 and sent to the regeneration section as discussed
above.
The regeneration section comprises a solids heater,
D-2 and a regenerator, D-3. These vessels are operated
at a nominal pressure of 30 psia (selection of the operating
pressure is discussed later), whereas the D-l reactor pres-
sure is essentially atmospheric. Thus, it is necessary to
feed solids from a system at low pressure into one at high
pressure and vice versa. The latter condition can be met
easily by taking a suitable pressure drop across the solids
feeder, but the former requires a build-up of pressure in the
solids transfer line between the D-l reactor and solids heater
D-2. This has been done by use of a dense phase solids stand-
pipe extending from the bottom of the elevated L-2 gas-solids
separator to grade. The height of the standpipe is such that
the pressure at grade is equivalent to that in the solids
heater. Air is injected into the line at grade such that
the solids are pneumatically transported to the solids heater
and fed into the top bed through a feed nozzle designed to
separate the air and solids so that bed stability is not
affected. The transporting air mixes with the gases leaving
the top bed, the combined gas stream then passes through an
internal cyclone for fines removal, and finally joins the
main flue gas exit stream. If a single bed heater were used,
the solids from D-l could be injected directly into the bed
and the transporting air would then supply part of the com-
bustion air thus reducing the air compressor horsepower. A
two stage heater proved to be more economical, however, since
the reduced fuel requirements more than offset the horsepower
savings.
-181-
-------
Sorbent solids are heated from 658°F to 1400°F in solids
heater D-2, which contains two fluidized beds: (1) a top bed
which preheats the cold solids by exchange with hot exit gases
from the bottom bed, thereby effecting heat economy, and (2)
a bottom fluidized bed in which the solids are heated to the
desired temperature of 1400°F by direct combustion of natural
gas in the bed itself. Hot solids leave the bottom bed through
an overflow withdrawal line and are fed into the moving bed
regenerator, D-3. A slide valve in the discharge line from
D-2 and a splash baffle underneath the end of the line in
D-3 provide a seal to prevent the spent regeneration gas from
flowing into the D-2 solids heater.
Sorbent is regenerated in D-3 by countercurrent contact
with upflowing regeneration gas. The regeneration gas is
supplied to the regenerator from the steam-methane reformer,
L-3. The reformer gas generation process flowsheet (P1374-B)
and description are given in Section B.2.Q* part (1) of this
report.
The sorbent which is loaded with sulfur in the form of
SO4 or SO^ is converted back to its original state (NaAlO2)
by the reducing atmosphere in the regenerator. The 1400°F
regenerator off-gas which contains the reduced sulfur in the
form of H2S exits from the top of the regenerator and after
passing through the waste heat boiler, C-2, it is routed to
the Claus Plant for the recovery of elemental sulfur. A
description of the Claus Plant is given in section 6, part v3)
of this report.
The regenerator is designed to provide (1) a bed which
moves down through the vessel in a plug-type flow pattern,
and (2) uniform gas distribution throughout the bad. These
objectives are achieved by (1) installing suitable baffles
neir the vessel bottom, and (2) providing gas distribution
-------
rings spaced so that equal bed heights are obtained above
each ring by allowing for build-up of sorbent due to a
single inlet line and the angle of repose of the sorbent.
Hot, regenerated sorbent is withdrawn through a standpipe
and is pneumatically transported to the bottom of reactor D-l,
thus completing the solids cycle. Make-up sorbent is stored
in hopper F-2, and added to the system as required by feeding
it in through the regenerated solids transfer line.
The air required for the solids heater D-2 and that needed
for fluidizing and transporting purposes, is supplied by the
centrifugal air compressor J-l. The J-l discharge stream is
split into four separate streams: the three streams used for
transporting the solids and the stream to the solids heater.
2. Design Notes - Dispersed Phase
a. Preliminary Calculations
(1) Effect of Gas Velocity and Recycle on
Residence Time in Absorber
An attrition ran reported in the Bruceton Quarterly
for the period ending 12-31-67 contained a sorbent size distri-
bution ranging from 246u to 2362y (60 to 8 mesh). This dis-
tribution was based on a composite of streams (bottoms, recycle,
and overhead) from the absorber. Although this distribution is
for a very friable sorbent it was used in the preliminary design
because no other pertinent data were available and the distribution
per se, in the range considered, does not greatly affect process
design since both overhead and bottom recycle lines are provided.
For each of the average particle sizes tabulated, theoretical
settling velocities were calculated: based on physical properties
of the system, a function of the Reynold's Number was calculated
from which the Reynold's Number and hence the particle settling
velocity could be calculated. Figure 46 is a plot of this cal-
-183-
-------
T
I
9
8
7
6
5
4
-- 1
Q I _
8
7
u
UJ
t
O
O
z
nn
(
O
FIGURE 46
SETTLING VELOCITY VS PARTICLE SIZE
PARTICLE DENSITY =* 70 LBS/CU FT
4 1-
PARTICLE SIZE. MICRONS
3 4 5 6 7 B 9 I
X 10s — 1
3 4 &
X IQ4
r <
-------
culated settling velocity vs particle size. As can be soon
from this plot, the settling velocity of Lhe average (i.e.,
50% size) particle size, 1050 micronsf is approximately 16
ft/sec but the settling velocity of the different size fractions
ranges from 33 ft/sec for the largest particles (2362 micron)
down to 3 ft/sec for the smallest, (2^6 microns). This wide
range of settling velocities results in large differences in
residence time in the absorber for the different size particles.
The preliminary process evaluation for a 1000 megawatt power
plant will use a 25 ft/sec gas velocity in the absorber. However,
a lower gas velocity may increase the solids residence time in
the absorber and improve efficiency since the settling velocity
calculations indicate that at 25 ft/sec gas velocity the heavier
particles will fall rapidly through the absorber while the fines
(-60 mesh, 240 microns) will be rapidly blown out, thereby not
achieving maximum usage for sulfur removal. The lower velocity
would, of course, require larger diameter reactors which would
be more expensive.
Based on the feed cited above, the theoretical residence
time for each size fraction has been calculated for superfi-
cial gas velocities in the absorber of 20 and 25 ft/sec. The
weighted average solids residence time in the reactor was cal-
culated from the weight fraction and residence time of each
sorbent size. The results of these calculations are presented
in Table 31 and show that for the conditions used, lowering
the gas velocity from 25 to 20 ft/sec will increase solids
theoretical weighted average residence time by about 16%.
The longer residence time should provide increased sulfur
pickup, other conditions remaining constant.
It should be noted that while the residence time of all
particles greater than 10 mesh is decreased by the lower gas
velocity, all the particles less than 10 mesh have longer
residence times. By recycling the bottoms to the absorber
-185-
-------
TABLE 31
CALCULATED SOLIDS RESIDENCE TIME IN BRUCETON REACTOR
CO
a-.
TYLER AVG. SIZE,
MESH MICRONS
WEIGHT
FRACTION
+ 8
-8 +10
10 +14
14 +20
20 +28
28 + 60
-60
2362
2006
1410
1000
711
418
<246
O.OC74
0.2127
0. 240?
0.1373
0.105"
0. 2171
0.079/
THEOR. SETTLING
VELOCITY FT/SEC
33
31
22
16
11
6
3
RESIDENCE TIME,* SECONDS
VV
25T
-8
-6
+ 3
+ 9
+14
+ 19
+ 22
FT/SEC.
20f
-13
-11
-2
+ 4
+ 9
+ 14
+ 17
OVERHEAD
25f 20f
— —
18.35 —
6.11 13.75
3.93 6.11
2.89 3.93
2.50 3.24
BOTTOMS
25 20'
5.00 3.08
6.67 3.63
— 20.
— —
— —
— —
— —
Weighted Average Residence Time
(RESIDENCE
25f
0.04
1.42
4.41
0.84
0.42
0.63
0.20
7.96
L . ) A
TIME)
20f
0.02
0.77
4.80
1.89
0.65
0.85
0.26
9.24
•,otc: Residence time is sascd on tnc overhead stream particles being fed into the absorber at
the 2.C ft. level aM iroppinc; another 10 ft. so that they travel a total distance of
55 ft. The oottor.is p ir tidies are assumed to be fed into the reactor at-the 40 ft. level
and thus travel oriy 40 ft. total.
Sas velocity in ft/sec,
-------
(instead of once through operation as is now used), the
residence time of the larger particles could be increased.
This has been discussed with the Bruceton people and a
bottom recycle line was eventually installed but operation
of the pilot plant was stopped before data could be obtained
using this new recycle line on particles having a substantial
size distribution. Some mechanical performance data were ob-
tained using a catalyst support material (UOP alumina spheres)
of uniform size but the results were inconclusive since only
a few runs were made.
The H-series data (Table 17) show that runs
H-106+H-111 and H-115^H-119, or a total of 11 out of 15 runs,
indicate lower sulfur concentrations in the reactor bottoms
than in the fresh sorbent feed. These data seem to substantiate
the theory that under the present method of operation, the
heavier particles remove very little sulfur from the flue gas
and thus a large portion of the fresh sorbent feed is essential-
ly inactive. The feed comprises a range of particle sizes with
different sulfur content but only the average is reported.
However, the heavier particles fall out the bottom and if their
sulfur content initially is lower than the feed average, and
if they pick up very little sulfur in the reactor, the sulfur
content of the exit stream could be lower than that in the
fresh feed. This would explain the H-series results cited
above where the bottoms exit sulfur concentration is lower
than the fresh feed. To better define the amount of sulfur
pickup as a function of particle size, it would seem worth-
while to separate (screen) the different streams (i.e., feed,
recycle, bottoms, sorber) into several size fractions and
analyze for sulfur in each reaction. This procedure would
furnish data to either confirm or disprove the theory that
the bottoms fraction contributes very little to sulfur removal.
-JB7
-------
(2) Effect of Sorbent Density on Sulfur Removal
The USBM Section I Quarterly Report of 6-30-68
states that a venturi throat was installed near the sorber
bottom and as a result the solids residence time was increased.
It also states that in tests using a sorbent density of 0.30
lbs/ft3 in the absorber nearly all of the S02 was removed from
the flue gas. However, the only sorption data included are
for the H-series which are tabulated in Table 1 of the USBM
report and the maximum sulfur oxides removal shown in 68.4%.
Apparently, the venturi was not used in the H-series but no
data are included to support either the value of 0.3 lbs/ft3
solids density or removal of nearly all of the sulfur oxides.
In previously reported data on hand, only a
relatively few runs have actually shown 90% sulfur oxides
removal; i.e., the only runs showing at least 33% S02 removal
are E-101, E-107, E-109, F-103, F-104, F-107 (See Table 16).
Because all process designs are based on '>39, SO- removal , it
would seem desirable that a steady state pilot plant run of
appreciable duration (say 1-3 days) should be made to demon-
strate that 90% removal is practical under commercial condi-
tions. This run could, of course, be combined with a run
designed to measure sorbent attrition losses when improved?.
sorbents become available.
Since density in the sorber is an irn^rt.ant
variable in degree of sulfur removal from the fi^c cas, the
H-series runs were used to determine the correlation between
pressure drop and density. The absorber pressure drop measure-
ment is direct whereas the density is determined by simul tane-
ously stopping the flue qas, fresh solids feed, and recycle
solids flows. The sorbent in the reactor column is then collected
and weighed and the density calculated based on column volume
and sorbent weight. It is felt the pressure drop measurement
-------
is the more accurate and consequently it has been used to
calculate the sorber density. These calculated values have
been plotted against measured values in Figure 47. A 45°
line through the origin shows most of the measured points
are higher than the calculated points. A possible reason
for this could be that there are solids in the feed lines
downstream of the shut-off valves and these solids, which
do not contribute to pressure drop, are collected along with
the solids in the sorber itself. A least squares fit for
all of the data points from tests using recycle solids has
been made and a line obtained with slope of 1.07 and an
intercept of 0.006, compared to a theoretical slope of 1.0
and an intercept of 0. Thus the correlation is considered
good enough to justify the use of calculated values for
either density or pressure drop in design studies.
(3) Regenerator
The preliminary process calculations for the
moving bed regenerator are summarized below. It should be
noted that the discussion below is based,on a regeneration
temperature of 1200°F which was later changed to 1400°F.
However, the following discussion which is based on the 1200°F
regeneration is valid except for a few minor points for a
1400°F operation and therefore has been included even though
it does not represent the final design. The rationale for
switching to 1400°F is given in the section on vessel design-
solids heater.
Reformed natural gas, consisting of 76.9% 1^,
13.1% C(>2/ and 10% CO, was selected for the regeneration of
alkalized alumina, with an arbitrarily specified utilization
of 80% (i.e., 25% excess) for design purposes. Since very
little experimental data on product distribution (both gas
and solid phases) were available at the time of this design,
-189-
-------
0.20-
0.19
FIGURE 47
SORBER DENSITY
CALCULATED VS MEASURED
0.13
SLOPE
1.07
0.17
0.16
0.15
0.14
3
O
0.12
O.I I
LEAST SQUARES FIT-)
O O/
X
o
0.10
REFERENCE:
USBM-8RUCETON DATA. SERIES M.
JUNE 1968 REPORT.
CALCULATED SOLID DENSJTY.IN SOMBER. LB/CU FT
JJ
0.18
0.10 0.11 0.12 0.13 0.14 0.15
0.16 , 0.17
-190-
-------
it was decided that product distribution based on thermodynamic
calculations was a reasonable choice for a design basis. Further,
since these calculations already had been done by AVCO it was
decided to use their results, modified as required by slightly
different conditions. As a check on AVCO's values, a product
distribution was calculated using an internal Kellogg computer
program and for the conditions selected, the results were
generally in good agreement. The AVCO 16th Monthly Summary,
April 29, 1968, Table VIII, p. 50 lists the product distribution
at chemical equilibrium for alkalized alumina with the above
reformed natural gas composition for a temperature range of
527+927°C (980-^1700°F) at 100°C intervals, and a sorbent mole-
cular formula of 1 Na20 + 2.435 Al-03, which is equivalent to
20 weight percent (wt %) Na20. The solid phase consists of
NaAlO- and excess A120,, with a residual content of Na-SO.
for temperatures below 727°C. The gas phase is primarily
H-O, CO-, and H_S, but above 700°C, S- SO-r and COS concen-
trations increase. For the present MWK case, the inlet solids
feed to the regenerator contains 3 wt % S and is to be regen-
erated at 1200°F and 15 psig, to a residual level of 0.5 wt % S
whereas AVCO's regenerated product at 627°C (1161°F) contains
0.45 wt % S. The difference in sulfur content is probably too
small to have a significant effect on the results and therefore
the AVCO equilibrium product distribution has been used as a
basis for our calculations, adjusted to 0.5 wt % residual sulfur.
Using the absorption reaction comprising 2 moles of NaAl02
yielding 1 mole of Na2SO4 plus 1 mole of A12O3, and the specified
condition of 3 wt % S as regenerator feed, the moles of Na2S04
and therefore the moles of A1203 formed in the absorber were
calculated. This material balance calculation is illustrated
in Appendix Table G-82.
AVCO's reformed gas composition has been adjusted
somewhat to suit our conditions. The AVCO off-gas composition
was then slightly adjusted in order to achieve the balanced
-------
reaction. Table G-83 contains the balanced regeneration
reaction along with the amount of regeneration gas needed
corresponding to 80% utilization.
The regenerator was initially designed at
1200°F and 15 psig with 80% reducing ga? utilization for a
range of conditions, namely a superficial inlet gas velocity
of 0.2+0.5 ft/sec, and a solids hold-up time of 1+4 „hours.
A maximum vessel diameter of about 40 ft was assumed and the
bed density was taken to be 45 Ibs/ft based on reported
values.
Table 32 lists the specified regeneration
conditions along with the various dimensions which have
been calculated for these preliminary conditions. As
can be seen from the table, the space velocity depends
solely on the solids hold-up time, and the bed height
is directly related to the hold-up time for each vessel
diameter and superficial gas velocity. This it is ap-
parent that several different combinations of gas velocity
and vessel size can be used for the same space velocity.
For our preliminary process design case, it was decided
to choose a 2 hour hold-up time and a superficial inlet
gas velocity of 0.3 ft/sec since this approximates an
actual USBM 55,000 CFH pilot plant moving bed regenera-
tion run*. To allow for process uncertainties, however,
a 25% safety factor on height has been included. A.dditioAai
allowances are needed for gas connectiors, disengaging
height, etc., so 20 feet was added to the calculated bed
height resulting in overall dimensions for each of the
two regenerators, under the conditions of Table 32, of
33 feet diameter x 43 feet high.
USBM Quarterly Report, Section I, March 31, 1968.
-------
TABLE 3 2
MOVING BED REGENERATOR DESIGN—1000 MW PLANT
vo
U)
I
Reducing Gas Reformed Natural Gas (76.9% H_,
13.1% C02, 10% CO)
Reducing Gas
Utilization 80%
Total Inlet ACFM...30,950
Packed Solids
Density 45 lbs/ft3
Total Solids Flow Rate.705,792 Ibs/hr.
REGENERATOR DIAMETER AND BED HEIGHT AS A FUNCTION OF GAS VELOCITY AND HOLD-UP TIME
Temperature 1200°F
Pressure 15 psig
Inlet Superficial
Gas Velocity 0.2-^0.5 ft/sec
Solids Hold-up
Time l-*4 hrs
Cross
Sectional Solids
Area/ Hold-up Volume/
VG Vessel No. of Time Vessel
(ft/sec) (ft2) Vessels Diameter (hrs)
0.2 1290 2 40'-6" 1
2
3
4
0.3 360 2 33' 1
2
3
4
0.4 1290 1 40'-6" 1
2
3
4
0.5 1032 1 36'-3" 1
2
3
4
(ftj)
7
15
23
31
7
15
23
31
15
31
47
62
15
31
47
62
,842
,684
.526
,368
,842
,684
,526
,368
,684
,368
,052
,736
,684
,368
,052
,736
Space
Velocity
ft3 gas/hr
ft3 solids
118
59
39
29
118
59
39
29
118
59
39
29
118
59
39
29
.4
.2
.5
.6
.4
.2
.5
.6
.4
.2
.5
.6
.4
.2
.5
.6
Bed Ht/
Vessel
(ft)
6.
12.
18.
24.
9.
18.
27,
36.
12.
24.
36
48
15.
30.
45.
60.
Bed
25%
Ht f
S.F.
(ft)
1
2
1
4
1
2
3
4
2
4
.6
.8
2
4
6
8
8
15
23
31
12
23
34
46
15
31
46
61
19
38
57
76
.5
.5
.5
Preliminary Case:
0.3 ft/sec superficial inlet gas velocity )
2 hrs hold-up )
2 Regenerators each
33'<(! x 43' overall height
-------
A heat balance around each of the regenerators
showed the temperature of the outlet streams to be 1129°F,
based on the following:
- All inlet streams at 1200°F.
- 80% utilization of regeneration gas.
- Heat of reaction is endothermic and equals about
509 cal/g mole Na2SO4 at 1200°F, based on the
product distribution in Table G-83.
- 3 wt % S regenerator feed is regenerated to
0.5 wt % S
- All of the SO2 removed from the flue gas is
transferred to the regenerator off-gas.
- Heat loss based on a 25°F ambient temperature,
a wind velocity of 20 miles/hr, and the regenerator
size, preliminary case, specified in Table 32.
The calculated heat of reaction for the assumed regeneration
reactions is endothermic, but because of the small number of
moles reacting compared to the total flow rate, its effect
on the overall heat balance is small. Therefore, the major
factor contributing to the temperature decrease during re-
generation for this particular design is the estimated .heat
loss.
In these preliminary process design calculations,
regeneration has been assumed to occur at 12CCr; , oa.:ed o:j.
Bruceton pilot plant data. The haat of .-reaction calcu,T ;.cad
depends on both the reaction temperature and the regeneration
reaction (i.e., product distribution). In our regeneration
calculations the reaction used was that obtained from the
16th Monthly AVCO report, April 29, 195B, and adjusted for
our conditions, it should be noted that adjusting the product
distribution caused the h«at of reaction to change from exo-
thermic to endothermic. This calculated heat of reaction at
-------
1200°F was used in the regenerator heat balance from which
an outlet solids temperature of 1129°F was obtained. In
order to maintain the regenerator at a minimum of 1200°F,
the solids heater duty must have additional heat for a
AT of 71°F, i.e., the solids enter the regenerator at 1271°F
and cool to 1200°F by the time they are discharged. There-
fore, the solids heater duty must be capable of heating the
solids from 600-"1271°F. Since there is some doubt as to
which of several possible regeneration reactions is the
correct one, it is desirable to know what effect each of
the different reactions would have on the heater design.
This has been done as described below.
Table 33 lists heats of reaction at 1200°F
for various regeneration reactions. Each of these heats
of reaction has been used in the regenerator heat balance,
other conditions remaining the same, to calculate the outlet
temperature in each case. Table 33 summarizes the results
of the preceding calculations and shows that, based on a
1200°F regeneration temperature, the outlet solids temperature
will vary from 946 to 1278°F, depending on which of the
several regeneration reactions is used. Therefore, it can
be seen that it is important to know which of the various
regeneration reactions is applicable so that the solids
heater is designed for the duty required to maintain the
desired regeneration temperature.
(4) Solids Heater
Since sorption occurs at 600°F and regenera-
tion at 1200°F, heating and cooling of the solid sorbent
particles through a 600°F range is required for each sorption-
regeneration cycle. Thus heat exchange between solids and
gases comprises a major process step; however, no experimental
data have been reported for this operation by the various
alkalized alumina investigators, and it is necessary to design
-1.9 5-
-------
TABLE 33
EFFECT OF THE HEAT OF REACTION ON THE REGENERATOR HEAT BALANCE
2,
3.
*
MWK AHH Calculated Re-
@ 1200*F generator Outlet
EQUATION* (kcal/g mole) Temp** (°F)
Na2S04 + 4H2 --Na20 + H2S + 3H20 44.9 946
Na2S03 + A12O3 + 3H2 -» 2NaAl02 + H2S + 2H2O -11.7 1179
Na2SO4 + A12O3 + 4H2-" 2NaAl02 + H2S + 3H20 0.83 1127
Na2S04 + 4CO + H20 -» 4C02 + H2S + Na2O 12.6 1079
Na2SO3 + Al2°3 * 3c° + H2°">' 3C02 4 H2S "*" 2NaAloo -35.8 1278
Na2S04 + A1203 + 4CO + H2O^ 4C02 + H2S -»• 2NaAl02 -31.3 1260
AVCO adjusted reaction used for preliminary process
design calculations — 10;>0 MW plant. 0.51 1129
15th Monthly AVCO Report; 4/29/68, p. 50
MWK Progress Report No. 4-5, Table 3, p. 29
Fi-cim irrel iminsyv r)T"AC;e>s
-------
the solids exchangers based on other information. To ascertain
what information is available, a literature search was carried
out for heat transfer correlations for gas-»solid systems, par-
ticularly for the case of heat transfer to a moving bed of
solids. The sources surveyed included the following:
• International Journal of Heat and Mass Transfer,
Vols. 1-11
• ASME Index, 1940-1956
« Engineering Index, 1962, 1967
October 1968
• Theoretical Chemical Engineering
Abstracts, 1964-1967
• Chemical Abstracts, 4th & 5th
Decennial 1937-1956, 6th Collective
1957-1961
• Assorted Review Articles
A review of the above sources showed that there
is a paucity of design information in the published literature
relating to equipment for transferring heat to or from granular
(i.e., nonfluidized) materials. To further complicate matters,
two of the correlations reported by different investigators
give opposite results; i.e., one correlation shows bed height
increasing as bed diameter increases whereas another correla-
tion shows a decrease in bed height as bed diameter increases.
However, all of the correlations used to calculate the effec-
tive exchanger length (defined as the axial length in which
all of the heat is transferred) for moving bed exchangers
give very short bed heights (i.e., 1-3 feet). The large
volume of gases and solids being handled requires a very
large cross-sectional area ( 6000 ft2, equivalent to a vessel
diameter of 88 feet) and the feed distribution systems will
likely be very costly, particularly for the solids. Thus if
the bed height calculated from the published heat transfer
-------
correlations is increased by a 100% safety factor, the
resulting increase in vessel cost will probably be small
compared to the total cost. Therefore, the time and cost
of an experimental program to obtain heat transfer data
does not appear to be justified.
Process calculations for the regeneration
solids heater are described on the following page.
The solids heater duty was calculated to be
75.6 MM Btu/hr/train based on the following conditions:
• Regenerator feed heated from 600-*1271°F. The regen-
erator is to operate at a minimum temperature of 1200°F
and the additional 71°F AT is necessary to maintain
this temperature based on the regenerator heat balance.
• One exchanger train required for each regenerator, or
a total of 2 trains.
• Heat loss based on an assumed AT of 100°F, wind velocity
of 20 miles/hr, and an estimated solid heater size, v?aa
essentially negligible in comparison to the required
duty. For calculation purposes, the heater vessels are
assumed to be cylinders.
• The duty specified above, 75.6 MM Btu/hr for each of
2 trains, includes a 10% excess over theoretical for
process uncertainties.
To avoid sintering of ^.h^ ^crbent particles,
the maximum allowable inlet gas temperature tc the solids
heater was assumed to be 1500°F. With the inlet temperature
fixed, the required amount of gas necessary to satisfy the
duty will vary with the exit gas temperature. The vessel
diameter in turn is dependent on the actual gas flow rate
(hence pressure, since temperature is fixed) and the super-
ficial gas velocity. For this particular case, a 10°P and
-------
50°F approach (outlet gas-inlet solids) was chosen, along
with a 0.5 ft/sec average superficial gas velocity. Therefore,
assuming a 1 psig pressure drop through the exchanger, the
variation of cross sectional area, and hence diameter, with
pressure can be calculated. A sample calculation of the solids
heater diameter for a 2 psig inlet pressure and a 10°F and
50°F approach is tabulated in Table G-84. Vessel diameters
have also been calculated for various assumed pressures and
the results are shown in Figure 48 along with the required
bed height for each case. (The method of calculating bed
height is discussed later on.) Calculations indicate that
the smaller temperature approach, namely 10°F, will be more
economical yielding a lower required gas flow rate and hence
a smaller vessel diameter. The 50°F temperature approach
case has been included in the event that a 10°F approach
cannot be achieved in the actual exchanger.
With duty, gas flow rate, and vessel diameter
specified, the remaining unknown is the required solids bed
height necessary to accomplish the desired heat transfer.
As previously mentioned, the published literature contains
relatively little design information relating to the case
of countercurrent heat transfer between a flowing gas-and
a moving bed of solids but some information is available.
Using the information obtained from the literature, the
required bed height for several different cases has been
calculated. The pertinent references are discussed below.
• "Heat Transfer to Granular Materials-Settled Beds
Moving Downward through Vertical Tubes," Brinn, M.S.,
Friedman, S. J., Gluckert, F. A., & Pigford, R. L.,
IEC, 4_0, 1050 (1948).
• The method for transferring heat to the granular solid
consisted of passing the granular material, as a settled
bed, downward through vertical jacketed tubes. Data
-199-
-------
FIGURE 48
SOLIDS HEATER
VESSEL DIAMETER AND BED HEIGHT
VS
PRESSURE
2 TRAINS REQUIRED
75.6 mm BTU/HR/TRAlN
10" APPROACH:
SOLIDS - 600 -»•
GAS -610*-
50* APPROACH:
SOLIDS - 600 -»
CAS - 650 «»-
0.5 FT/SEC AVERAGE 6
VELOCITY
I psig PRESSURE DROP
ASSUMED
FURNAS
I0« APPROACH
FURNAS
50* APPROACH
BRINN
10* APPROACH
30 40 50 6O 70 8Q90OO
PRESSURE. p«ia
O.I
200
Ind.
42, I48II
-------
collected indicated that the beds moved substantially
as rods. It was shown that heat transfer to fluids
in rodlike flow in a tube is mathematically identical
to the unsteady state heating of a long cylinder with
negligible surface resistance. Theoretical equations
were developed and design charts prepared for the
general case of countercurrent heat transfer to a
material in rodlike flow in a tube. Equations were
also developed for the case of parallel flow.
The design charts presented are plots of the Graetz
Number (we /kL) vs a temperature ratio, with a para-
meter B, defined as the ratio of the resistance to
heat flow of the jacket fluid to the granular material.
These charts are given for three values of z, defined
as:
we
wr^- where
wcp
w = solids weights flow rate per tube
W = jacket fluid weight flow rate per tube
c - solids specific heat
Cp= jacket fluid specific heat
Examples are shown for a particular set of heat
transfer conditions, illustrating the use of the
design charts in calculating the required flow
rate per tube and, from this and the total sclids
flow rate, the total number of tubes required.
Although heat is transferred in the present
tentative design of the solids heater by direct
countercurrent contact of gas and solids, rather
than indirectly through jacketed tubes, a modified
calculation procedure can be developed for our
-201-
-------
particular case. That is, the correlation
can be used to calculate the -total tube length
required for our set of conditions and depending
on the number of tubes required, which is ob-
tained directly as a function of vessel diameter,
tube spacing, and pitch, the required bed height
can then be calculated. This procedure is
tabulated in Table G-85 for the low pressure,
90.5 ft diameter, 50°F approach case. Other
calculations were done for lower diameter and
higher pressures for both a 50°F and 10°F
approach and these results are plotted on
Figure 48. For a given pressure, the bed
height and vessel diameter can be determined
for the two temperature approaches. The
Brinn correlation predicts an increase in bed
height with a decrease in vessel diameter
and the 10°F approach predicts a slightly
higher bed height than the 50°F approach for
a comparable pressure. The actual calculated
bed heights range from 0.33 feet to 3 feet
for the pressure-diameter-temperature approach
spread studied as indicated on Figure 48.
"Solid-Fluid Heat Exchange in Moving Beds",
Munro, W. D. and Amundson, N. R., IEC, 42,
1481 (1950).
The problem of heat exchanger was considered
for the case in which the solid is in the form
of uniform spheres. The resistance to heat
transfer by conduction in the solid vas taken
into consideration. It was assumed that con-
duction in the fluid in the direction of flow
is negligible and that the exchanger itself
is adiabatic, so that no temperature gradient
-------
in a radial direction in the fluid exists. An
exact solution of the theoretical equations
involved in the problem of heat exchange between
solid and fluid was developed for both parallel
and counterflow exchangers. The authors felt
that in some cases the assumption by other in-
vestigators of negligible resistance to heat
transfer by conduction in the solid could lead
to an undersizing of exchanger length. However,
the equations presented require a computer
program for use, and it appears that the as-
sumption of negligible resistance to heat trans-
fer within the particles is valid provided the
particles which make up the bed are small enough
and have a sufficiently high thermal conductivity.
Therefore, it was decided to use an approximate
method, referred to in the article, for the
calculation of bed height.
The Furnas* approximate solution considered the
case in which conduction in the solid is neglected
and the problem is considered as an ordinary
countercurrent heat exchanger in which the area
for heat exchange is the total external area of
the particles in the bed at one time. This method
is outlined, along with a sample calculation, in
Table G-86 for the low pressure, 90.55 ft diameter,
50°F approach case. Included in Figure 48 is the
variation in bed height, determined by the Furnas
approximate equation, as a function of pressure and
vessel diameter for both a 50°F approach and 10°F
approach. For a comparable pressure, the 10°F
approach predicts a slightly higher bed height than
* Furnas, C. C., IEC, 22, 26 (1930)
Furnas, C. C., TRANS AIChE, 24, 1942 (1930)
Furnas, C. C., US. BUR. MINES, BULL. 361 (1932)
-203-
-------
the 50°F approach, as does the Brinn correlation.
However, the Furnas equation predicts an increase
in bed height with an increase in diameter, which
is the reverse of the height-diameter relationship
obtained from the Brinn correlation. The Furnas
equation (Table G-86) indicates that the temperature
ratio
is a function of the exponential ratio of the
product of the heat transfer coefficient at the
solid-fluid interface (hf) and the bed height (x)
to the solids mass velocity (Gs):
(T-t0)/(TQ-t0) = f(exp [(hf/Gs) x]
The correlation used to calculate heat transfer
coefficients used in the Furnas equation is the
one developed by Leva and is shown in Table G-86.
It is possible that other correlations for cal-
culating heat transfer coefficients might produce
a different relationship between bed height and
diameter as predicted by the Furnas equation, but
the Leva correlation has been used here since it is
generally accepted as a valid correlation. As can
be seen from Figure 48, the bed heights calculated
using the Furnas method range from 0.26 feet to
1.6 feet for the pressure-diameter-temperature
approach spread studied, vs a 0.33 feet to 3 feet
range predicted by the Brinn correlation for the
same process conditions.
M. W, Kellogg Correlation
This correlation is used for calculating the
t-amp-srature history curves during heating cf a
-------
fixed bed of solid particles. The calculation
consists of first determining the convection
coefficient and then, by use of the design charts,
determining the temperature at several points in
the bed after an amount of heating gas correspond-
ing to a selected time interval has passed through
the bed. If the average temperature of the bed
so calculated is not that required, another inter-
val of time is selected and the calculation repeated.
This calculation is intended for use in the case of
a fixed bed and not a slowly moving bed. However,
it was decided to use this procedure in the follow-
ing variation for the case of a slowly moving bed:
Assume it is a fixed bed and that for a specific
hold-up time, a definite amount of material would
be placed in the vessel and allowed to come to
equilibrium. Various hold-up times were then
chosen and from the solids flow rate, packed
density, and vessel diameter, the bed height was
calculated. Temperature history curves were then
plotted to determine at what bed height the de-
sired average solids temperature would be 1271°F.
Table G-87 contains a sample calculation for the
88 feet diameter, 10°F approach case for a hold-up
time of 0.075 hours. The complete set of tempera-
ture history curves for various hold-up times and
corresponding bed heights are plotted on Figure 49
for the low pressure case and in Figure G-92 for
the high pressure case, both for a 10°F approach.
(See Figure 48 for pressure-diameter relationship.)
As can be seen from the average solids temperatures
obtained for the corresponding hold-up times, the
average required temperature of 1271°F was exceeded
in less than a 1 foot bed height for both the 88
-205-
-------
FIGURE 49
TEMPERATURE-HISTORY CURVES FOR SOUPS HEATER
MWK METHOD
0.49
0.40
2 TRAINS REQUIRED
88 FT DIAMETER VESSEL
16.7 psia INLET
0.5 FT/SEC GAS VELOCITY
I psig PRESSURE DROP ASSUMED
io« APPROACH:
SOLIDS - 600 -*• I 27 IF
GAS - 610 —- I500F
75.6 mm BTU/HR/TRAIN
A SOLID
TEMPERATURE
.GAS
TEMPERATURE
0.300 WRS HOLD-UP
O.225 HRS MOLD-UP
O.ZO -1—0.194
0.150 HRS HOLD-UP
0.097*
BED HT
0.075 HRS HOLD-UP
TAV - I306F
TEMPERATURE,
,000 8,(00 l.S'fj© S
-------
feet (low pressure) and 33 feet (high pres-
sure) diameter case.
In all previous discussed calculations, a
pressure drop of 1 psig was assumed. Happel, J., IEC, 41,
1161 (1949) developed a correlation for pressure drop due
to vapor flow through moving beds. The correlating equa-
tion was obtained by applying the Navier-Stokes equations
to the viscous motion of fluid through a group of uniform
spheres. Flow through a granular material was accounted
for by a new function of the fractional void volume. Table
G-88 is a sample calculation for pressure drop for the 90.5
feet diameter, 50°F approach case. The results indicate
that for the conditions studied, an assumed pressure drop
of 1 psi is sufficient, since this provides an additional
0.5 psi for the gas distribution system.
The inlet gas temperature to the solids
heater was taken to be 1500°F in the previous calculation
and consequently part of the sorbent was heated to this
temperature. In the event that the sorbent temperature
might have an allowable upper limit of about 1300°F (e.g.,
to avoid sintering), an additional case has been evaluated
in which the inlet gas temperature is fixed at the outlet
solids temperature, namely 1271°F. The MWK correlation was
used to calculate the average solids temperature for specific
hold-up times and bed heights for a low pressure, 10°F ap-
proach case. Figure 50 is a plot of the average solids tem-
perature vs bed height, and indicates that the required bed
height necessary to accomplish the desired heat transfer
(solids heated from 600+12710F) is small, being only about
two feet. As the bed height is increased, the average solids
temperature rapidly approaches and becomes asymptotic to
1271°F for the conditions chosen. After only two feet, how-
ever, the solids temperature is within about 2°F of the inlet
-207-
-------
FIGURE 50
AVERAGE SOLIDS TEMPERATURE VS BED HEIGHT
2.0-1
1.9-
1.8-
1,7-
1.6-
1.5-
1.4-
1.3-
1.2-
I. I -
1.0 -
0.9-
0.8
O.T -
Q.5
0.4
0.3
O.I
n
X
o
UJ
X
a
u
03
2 TRAINS REQUIRED
99 FT DIAMETER VESSEL
16.7 ptia INLET
0.5 FT/SEC GAS VELOCITY
I P»i« PRESSURE DROP ASSUMED
IO*F APPROACH
SOLIDS -600 -*"I27IF
GAS -610 —1271 F
75.6 mm BTU/HR/TRAIN
AVERAGE SOLIDS TEMPERATURE, »F
I
DESIRED
— SOUDS
TEMPEKATU
JL
i,200 1,210 1,220 1,230 1,240 1,250 1,260 1,270
-------
gas temperature so by increasing the inlet gas temperature
by two degrees, the desired solids temperature could be
obtained.
b* Vessel Design
(1) Absorber
A material balance, Table G-89, around the
600°F dispersed phase absorber establishes the inlet and out-
let flow rates. A heat balance around the absorber showed
the temperature of the outlet stream to be 603°F, based on
the following:
* All inlet streams at 600°F
• Heat of reaction is -115.7 kcal/g mole
SO2 at 600°F, assuming sorbent is NaAlO-
• 90% removal of sulfur oxide in entering gas.
• Heat loss based on an assumed AT of 300°F (i.e.,
vessel wall temperature of -325°F), wind velocity
of 20 miles per hour, and an estimated absorber size.
Thus the temperature rise in the absorber is
negligible due to combination of large flow rates and low
concentration of SCK in the feed gas, hence relative small
total heat release.
For the calculation of the absorber dimensions,
25 ft/sec was chosen as the inlet gas velocity and it was
decided that initially the maximum specified diameter per
vessel would be 40 feet. USBM Quarterly Report, Section I,
March 31, 1968, included a series of equations relating the
sorption rate constant with absorber height and other known
quantities. Table G-90 is a consolidation of these equations
into one equation explicit in column height, for the case of
90% SO2 removal. A sample calculation for obtaining the
absorber dimensions is included as Table G-91.
-209-
-------
(2) Solids Heater
A major process step in the alkalized alumina
process is heat exchange between solids and gases in each
sorption-regeneration cycle since sorption occurs at 600°F
and regeneration at 1200°F or higher. Previous calculations
have indicated that a moving bed heater requires a very large
cross sectional area, approximately 6000 square feet ("90 feet
diameter), due to the low gas velocity of 0.5 ft/sec necessary
to avoid fluidization. On the other hand, various correlations
used to estimate the required bed height have indicated that
this is very small, in the range of 3 feet or less. This com-
bination of large diameter and extremely small bed height re-
sults in an impractical vessel design with very complicated
and expensive gas and solid distribution systems. As a result
of the above factors it was decided to evaluate -other heat
exchange systems.
A fluidized bed concept would permit the use of
higher gas velocities, thereby decreasing the vessel diameter,
and a direct combustion fluidized solids heater would minimize
volumetric gas flow rate, thus further lowering the vessel
diameter.
Therefore, it was decided to evaluate the use
of a direct combustion fluidized bed design as the solids
heater. Since at the time these calculations were made the
process design included a solids cooler downstream of the
regenerator, the optimization described below was for both
the heating and cooling operation rather than only heating.
Elimination of the solids cooler does not change the relative
comparisons of the different cases, however, and the design
selected for the heater also would be picked without the
cooler. As a result of the fluidized bed study, a direct
combustion fluidized bed heater was selected for the final
design.
-210-
-------
The regeneration temperature was originally set
at 1200°F based on pilot plant data available at the time
but later data indicated that it would be desirable to regenerate
at a higher temperature, viz, in the range of 1300°F. Since it
was shown that a direct combustion fluidized heater gave the
lowest costs, and since the minimum temperature at which a
direct combustion will occur without excessive carbon deposi-
tion is about 1400°F, it was decided to increase the regenera-
tion temperature to this level. As indicated above, a regenera-
tion temperature of about 1300°F would be satisfactory so it was
not necessary to heat above 1400°F to allow for heat losses as
was done with the 1200°F design. That is, if heat losses and/or
endothermic heat of reaction caused the temperature to fall
below 1400°F, reaction kinetics would not suffer as long as the
regenerator did not drop below about 1300°F.
To determine the optimum (or nearly optimum) design
for these heat exchange vessels, several schemes were evaluated,
namely, various combinations of 1, 2, and 3-stage fluid bed
heaters and coolers, and a direct combustion fluidized solids
heater combined with either a 1, 2, or 3-stage solids cooler.
Further, each combination was evaluated at pressure levels of
20, 30 and 50 psia. Final selection was based primarily on
economics, calculated as follows.
For each scheme an energy balance and fuel require-
ment were calculated (based on a regeneration temperature of
1400°F in all cases) and preliminary vessel designs prepared
from which the approximate capital investment was estimated.
The annual cost of fixed charges was taken to be 15% of the
erected plant cost. For utilities, natural gas was charged
at 27C/MM Btu* and electric power was assumed to be 5 mills/kwh.
* Average fuel cost for the U. S. as reported in "Steam-
Electric Plant Factors — 1968 Edition," published by
the National Coal Association.
,-211-
-------
A stream efficiency of 60% or 5250 hours/yr was chosen as
being a reasonable basis for a 1000 MW power plant upon which
our evaluations are based. Total annual costs of each scheme
were then calculated and, along with process operability con-
siderations, used to select the optimum heating and cooling
arrangement for the sorbent solids. Approximate capital in-
vestment arid total annual cost for each scheme studies are
tabulated in Table 34. Annual cost of fixed charges and
total annual cost for the moving bed schemes were not included
in Table 34 because of the impractical vessel requirements, as
previously discussed. Of the remaining fluid bed cases studied,
Table 34 indicates that the use of a 2-stage direct combustion
fluid bed solids heater has a distinct economic advantage over
the fluid bed heaters using externally heated gases as the
source of heat. For the solids coolers, the total annual
costs shown in Table 34 indicate that economics favor increas-
ing the number of stages in the fluid bed solids cooler. The
largest decrease in total annual cost occurs in going from
1-stage to 2-stage while a smaller decrease is evident in
going from 2-stages to 3-stages. It can also be seen from
Table 34 that an increase in pressure tends to increase the
annual cost due to the larger compressor requirements.
It should be noted that both the investment and
total annual costs shown in Table 34 are approximate in nature
since detailed vessel designs were not made. For example,
vessel costs were estimated directly from the approximate
vessel weight with no allowance for additional incremental
costs due to multiple beds. Likewise, secondary effects
were neglected (e.g., running the regenerator system at about
30 psia eliminates the need to compress the gas stream feed
to the Glaus plant) since their influence on the total annual
costs is probably not significant. Therefore, because of the
approximate method of determing costs, small differences in the
values shown in Table 34 should be neglected. Thus, a combination
-212-
-------
TABLE, 34
COMPARISON OF VARIOUS DESIGNS FOR HEATING ACT COOLING SOUPS IN REGENERATION SECTION
Conditions;
1) Vessels - carbon steel with
1/2- thick wall.
2) 5 ft/sec average superficial
gas velocity.
3) 1 psig pressure drop for
fluid bed (FB) cases and 4
psig pressure drop for 2-
stage direct combustion
fluid, bed (DCFB) cases.
4) Vessel heights calculated a*
follows:
• S ft plenum chamber
• 5 ft between beds
• Bed height - For FB cases,
1 ft/stage, for 2-stage
DCFB 10 ft for first
stage and 1 ft for second
stage.
• Total Disengaging Height
(TDK) taken as function
of vessel diameter -
Ref: Frantz, J. F. , CHEM.
ENG. , 69., 161 (1962)
», Vessel Diameter, (ft)
10
20
35
TDH (ft)
3.5 X #
2 x 4
1 x 4
5) Cost figures chosen:
Carbon Steel - 50«/lb
Natural Gas - 27«/MM Btut
Heating Value of 943.7
Btu/SCF
Electricity - S raiU»/KW-hr
«) 60% steam efficiency or 5250
hrs/yr i
m (2) <3)
PR£SSURE
METHOD * (PSIA)
SB - I Stage FB.I
SC - 1 Stage FB
SH - 2 Stage FB
SC - 2 Stage FB
SH - 3 Stage FB
SC - 3 Stage FB
SH - 2 Stage DCFB
SC - 1 Stage FB
SH - 2 Stage DCFB
SC - 2 Stage FB
SH - 2 Stage DCFB
SC - 3 Stage FB
20
30
50
20
30
50
20
30
50
20
30
50
20
30
50
20
30
| 50
SIZE
SOL. HTH. SOI. COOL.
t X HT.
(FT>
!£_«
55 x 61
44 x 50
34 X 40
.
-
-
38 x 56
31 x 58
24 X 59
12 X 60
10 x 54
7 x 46
12 x 60
10 X 56
7 x 56
12 x 60
10 x 56
8 x 47
24 X 47
19 x 45
IS X 47
.
-
-
20 x 58
16 x 60
12 x 56
25 x 48
20 x 46
16 x 48
21 x 52
17 x S3
13 x 52
20 x 58
16 x 60
12 x 56
FIXED ANNUAL
ERECTED CHARGES ANNUAL ELEC.
TOTAL PLANT AKNUAL FUEL FUEL POWER
WT. COST COST*' (MPH COST COUP. COST
(M LBS.) (MS) (M S/YR) N.G.) M S/YR) BMP (M S/YR)
552
378
244
__
-
-
36C
289
208
197
144
105
178
139
98
181
142
96
828
567
366
_
124
85.1
54.9
-
-
-
549
434
312
296
216
158
267
209
147
272
213
144
82.4
65.1
46.8
44.4
32.4
23.7
40.1
31.4
22.1
40.8
32.0
21.6
1080
M
840
11
M
520
"
11
263
m
"
240
«
«
232
M
•
519
H
403
"
250
"
**
126
H
m
115
n
H
111
t«
•
13,300
23,230
38,700
-
~
"
5,100
10,100
17,000
6,500
11,300
18,800
3,600
6,200
10,300
3,000
5,200
8,700
260
455
758
-
~
100
198
333
127
221
368
70.5
121
202
58.7
102
170
TOTAL ANNUAL
COST -
(1 + 2*3) ,
(M S/YR; '
I4e;
1165
1893
> 403 '>
m
43:
513
630
297
37*
518 i
I
226
267
335
211 i
245 1
303 '
UGBND
SH - Solids Heater
SC - Solids Cooler
FB - Fluid Bed
MB - Moving Bed
DCFB - Direct Combustion Fluid Bed
* Regeneration temperature is 1400*F unless noted otherwise.
*• 15% of erected plant cost/year.
t 3 x basic equipment cost
+t For regeneration temperature of 1200T; 1400°F case would be higher
-213-
-------
was selected for the present dispersed phase design (viz.,
a 2-stage direct combustion heater and a 2-stage solids cooler,
both operating at 30 psia), that does not have the lowest annual
costs, but the difference between it and the combination with the
lowest total annual cost is relatively small. Further, the combi-
nation selected does exhibit the lowest investment compared to the
other combinations that have lower annual costs, and it also has the
advantage of easier operation: two beds instead of three in.
some cases, or no need for a compressor on the regenerator off-
gases in other cases. Therefore, in terms of both economic
and practical considerations, the best regenerator heat exchange
arrangement for the present design was the one mentioned above,
i.e., a 2-stage direct combustion fluid bed solids heater is com-
bination with a 2-stage fluid bed solids cooler, both operating
at 30 psia. A preliminary vessel design calculation for this
combination is illustrated in Table G-92.
As previously noted, the above discussion and the
sample calculation shown in Table G-92 are for the process design
which included a solids cooler. In the final base case design
the cooler was deleted which caused the heater duty to increase
(preheated air was no longer available) and resulted in increased
gas flow rates and slightly different vessel dimensions. The
preliminary design discussed above produced a heater which was
10 feet in diameter by 56 feet high where?.? the final design requizx
11 feet by 45 feet vessel. A sketch of the final design of the heatei
is shown in Appendix G, Figure G-94.
In the previous discussed heat transfer calculations
involving a fluid bed concept, the bed height was assumed to be
one foot per stage. Literature articles, in general, have indi-
cated that heat transfer in a gas fluidized bed occurs over a
very shallow portion of the bed and that the axial length in which
all of the heat is transferred is very short. Temperature homo-
geneity is one of the characteristics of a fluidized bed and it
is belfeved that heat transfer takes place almost instantaneously
-------
As a numerical check to see if the allowed one foot per stage
bed height was sufficient, a calculation was made to estimate
the bed height required for a specified heat transfer rate.
Table G-93 shows a sample calculation for a 3-stage fluid bed
solids heater with the temperature profile for the first stage
plotted in Figure 51. It is evident from Figure 51 that the
desired heating of the solids to 1400°F by 1500°F inlet gas
can be accomplished in a bed height of less than 0.3 inch per
stage, well under the assumed height of one foot per stage.
(3) Regenerator
Regeneration of the loaded sorbent is assumed to
take place in a moving bed regenerator, similar to the one
presently used in the USBM pilot plant at Bruceton. Process
calculations and several different vessel configurations for
the regenerator were discussed previously (Table 32), The
preliminary calculations were made for a regeneration tem-
perature of 1200°F whereas the final vessel design is based
on 1400°F and consequently has slightly different dimensions
than those shown in Table 32, viz, final dimensions are 35 feet
diameter by 36 feet straight section (59 feet overall)vs. 33
feet x 43 feet preliminary size.
Since relatively little regeneration data are
presently available, only tentative designs can be made at
this time and assumptions are required for space velocity,
reaction temperature, hold-up time, and superficial gas
velocity. For the present design it has been assumed that
for a nominal 1400°F regeneration temperature and 0.3 foot
per »econd gas velocity, a two hour hold-up time is required.
However, the regenerator has been designed so that solids
hold-up time can be increased to three hours if necessary
to obtain the specified level of regeneration: i.e., 0.5%
-215-
-------
FIGURE 51
FLUIDIZED BED TEMPERATURE PROFILE
(FIRST STAGE)
1.500
1,475
ui
cc
oc
UJ
a.
UJ
1450
1,425 --
O
a.
1,400
SOLIDS
1,375
600F
990F
990F
I240F
GAS
I I240FJ M400F
I400F1 JI500F
3 STAGE FLUID BED SOLIDS
30 psio - 31 FT DIAMETER VESSELS
2 TRAINS REQUIRED
SOLIDS: 600 —- I4OOF
GAS: 990 —I500F
90.1 mm BTU/HR/TRAIN
DISTANCE ABOVE BED BOTTOM, FT
0.01
(0.12")
0.02
(0.24")
0.03
(0.36")
0.04
(0.48")
REFERENCE:
Frantz. J.F., Cfienn. £ng. frog.,
57, No. 7, 35 (1961).
-216-
-------
sulfur at exit. The cost of the extra vessel height to obtain
the additional one hour hold-up is small and therefore the
overall regenerator cost is increased only slightly. The
flexibility provided by the additional hold-up time is felt
to justify the extra cost.
(4) Solids Transfer Lines
The transfer of sorbent particles between the
reactor and regenerator (and vice versa) is done pneumatically,
using dilute phase gas-solid mixtures. Line sizes and trans-
port gas requirements have been calculated based on throughput
solids loadings of one pound per cubic foot (derived from
erosion considerations), whereas pressure drop calculations
have used a slip factor of two, which has been estimated from
the particle size distribution. The recycle transfer line from
the bottom of the top of the reactor has been designed as a
dense phase system, however, since it is essentially a vertical
line and very large flow rates are involved. Use of a dense
phase system with its lower velocity should result in lower sorbent
attrition rates than a high velocity dilute phase system.
(5) Plant Layout
A rough layout of the process equipment was pre-
pared in order to ensure a feasible basis for the estimate of
plant investment and process economics. The layout is re-
quired because of the unusual nature of the process. That is,
because of the extremely large volume of gases being handled,
the very large gas ductwork required is a major cost item
and therefore it is necessary to determine the approximate
quantities needed. By preparing a rough layout of the process,
the ductwork requirements can be readily obtained, and in
addition, a better estimate of piping, structures, foundations,
etc., can be made. Figure 52 shows the plan and elevation
for one train of equipment and Figure 53 is a plant layout for
-217-
-------
FIGURE 5Z
EQUIPMENT LAYOUT
OlSPERSEP PHWE PROCESS
1000 MW PLAMT
/// "'f'/•";/!'////
/x/' FLUE v ////
GAS INLET
0' GRADE
" 77/777
ELEVATION
., STACK,
-------
FIGURE: 53
PLOT PLAN
7OO't -
JDOO MW POVNER ST7
)) - 1
a.?'
DISPERSED PHASE PROCESS
1DOO M.W PLANT
(a TRAINS)
>
".TICH ,„ „
| *_«-o> | _
HI F.D. HEADER 11
y.-t
i
\i-i--
':/
"I
«-*le<
rruM
"<
PRCHEATER
ct>Qtft>o
(INIER
EA
r N
+ i
V ,-'
;: !
|
T.,l
; /
I
*rf.*
ff) /
"N^
GAS INLET*
S INLETS \
©
["SCALE: i'=4o'
-------
TABLE 35
IDENTIFICATION OF LINES ON LAYOUT SKETCHES - FIGURES 52 AND 53
DISPERSED PHASE SORPTION - MOVING BED REGENERATION PROCESS SCHEME
LINE NO.
(FIGURES 52 & 53) DESCRIPTION
1 Ductwork - Flue Gas Inlet & Clean Gas
Outlet
2 Natural Gas Inlet to D-2 Solids Heater
3 Regeneration Gas Inlet
4 Regeneration Off-Gas TL.
5 Air Line from Compressor J-l to
Solids Heater D-2
6 Flue Gas from D-2 Solids Heater to Clean
Flue Gas Duct
7 Gas TL Gas-Solids Separator (L-2) to Clean
Flue Gas Duct
8 Sor-ber Solids (Dense Phase) TL to L-2
9 L-2 Solids Outlet Standpipe
10 Solids Transfer Line from Grade to D-2
Solids Heater
11 L-1A-L-1F Cyclone Cluster Solids Outlet to
Standpipe
12 Sorber Solids Recycle Line from L-2
13 L>-3 Regenerator Solids TL to sorber
Note: TL = transfer line
-------
2 trains. The pipe and ductwork numbered on Figures 52 and
53 are described in Table 35.
The sketch in Figure 52 shows the bottom of the
reactor about 50 feet above grade. This height was determined
by a pressure balance since dense phase transfer of recycle
solids is used and the pressure needed to elevate the solids
is built-up in the reactor solids withdrawal standpipe. Thus
the length of the standpipe, hence vessel elevation above grade,
is set by the recycle solids requirements.
(6) Vessel Analytical Sketches
The process equipment sketches are included as
Appendix Figures G-93 thru G-98.
b. Fluid Bed
(1) Fluid Bed Process Description; Flow Sheet
P3198-D
The following process description is for the
alkalized alumina process scheme shown on flow sheet P3198-D.
The scheme is based on removing 90% of the SC>2 in the flue
gases from a 1000 MW power plant burning a 3% sulfur coal.
The equipment shown on this flow sheet is de-
signed to handle 50% of the power plant effluent. The remain-
ing 50% of the power plant effluent is fed to a second train
of equipment identical to that shown on this flow sheet. In
actual plant construction, however, the gas flow will not be
split until after the flue gas booster fans. This will enable
the use of a single inlet header, thereby decreasing duct
construction costs. The same practice will be used for the
sorber effluent gases with a singleduct being provided at
this point for the same reason as stated above. The flow sheet,
therefore, schematically represents only the flows in one train
-221-
-------
0
22*34
J
nrw
-A
-------
ttorex: I CtuHMTtrifs &*fot*/iv *
-2
-------
^xiSsijii^r
r/rifg &*towft tMtr /ncwe t?>vf mtitf* TWO
* *f*f fftlts/Jt&£>r t/ivt.iF3S tve"rf£> t?rw£/r*¥i
THE M. W. KELLOGG COMPANY
DAT* •¥
DMAWN: p.- X,
CO«MTHUCTIO*I
60-fS
I I
I -.!«-
-222-
-------
of equipment.
The SO2-rich power plant effluent enters the flow
sheet at 600°F and 10 in. H2
-------
pressure (30 psia) to match that in the regeneration train.
The sulfur loaded sorbent leaves the gas-solids
separator, L-l, through a dense phase fluidized standpipe
and flows down 90 feet to a reverse loop where it is joined
by a transport air line for pneumatic conveyance (@ 1 Ib/cu
ft loadings, throughput basis) to the regeneration train. The
elevations of the regeneration train's solids heater and the
gas-solids separator were set by a pressure balance around the
pneumatic transport loops. A solids loading of two pounds per
cubic foot was used in this calculation. This agrees with the
slip factor of two which was estimated from the particle size
distribution and theoretical settling velocities. The trans-
port gas used in both loops is air at 35 psia.
The sulfur loaded solids enter the upper bed of
the solids heater, D-3, at approximately 634°F. Two fluidized
beds were used in this vessel to allow the 1400°F gaseous ef-
fluent from the lower bed to preheat the solids in the upper
bed. This arrangement minimizes the natural gas and air re-
quirements for the system by decreasing the duty required in
the lower bed. The preheated solids leave the upper bed at 726°F
via an overflow standpipe and are fed to the lower bed. The
major part of the heating duty is supplied in this lower bed
where the solids are heated by the direct combustion of natural
gas and air. The solids leave the lower bed at the gas-solids
equilibrium temperature, 1400°F, and exit via an overflow stand-
pipe which feeds the sorbent regenerator, D-2.
The effluent gas from the upper bed of the heater
mixes with the pneumatic transport air in the vapor space above
the upper bed. These gases are then passed through an internal
cyclone to separate the entrained solids from the gas stream.
The mixed gases leave the vessel at 702°F and are let down by
a "alve from approximately 25 psia to 10 in. H20. The combined
-224-
-------
heater effluent then flows to the main flue gas exit header
for heat recovery and ultimate disposal through the power
plant stack.
The sorbent leaves the solids heater at 1400°F
and 30 psia and enters the fluid bed sorbent regenerator, 0-2.
In this vessel the sulfated alkalized alumina sorbent is re-
generated with a reducing gas to desulfurize the sorbent from
7 weight percent sulfur to 2.6 weight percent. The sulfur-rich
reducing gas effluent formed in the regenerator is passed through
a cyclone to remove entrained solids and then sent to a Glaus
unit for sulfur recovery as elemental sulfur. A separate dis-
cussion of the Glaus unit is given in section B.2.e, Part (3)
of this report.
The regeneration gas is supplied to the regenerator
from the steam-methane reformer, L-3. The reformer gas genera-
tion process flow sheet (P1374-B) and description are given in
section B.2.e.,Part (1) of this report. The regeneration gas
at 1400°F and 25% excess over stoichiometric proportions enters
the bottom of the vessel and passes through the fluidized sor-
bent bed, removing the sulfur from the sorbent. The products
of the reduction reactions, based on a reformed natural gas
feed are: sodium aluminate (NaAl(>2), hydrogen sulfide (H-S) ,
carbonyl sulfide (COS), sulfur dioxide (S02), elemental sulfur
(S,), and water. These gaseous products of the reactions leave
overhead, while the sulfur depleted solids are withdrawn through
a fluidized standpipe.
The solids leave the sorbent regenerator, D-2, at
an average temperature of about 1400°F via five overflow stand-
pipes. These standpipes are fluidized and are provided with
control valves to regulate the flow in each line. The lines
serve as the dense phase dipleg for a pneumatic air transport
-22J-
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system which carries the sorbent from the standpipes to the
fluid bed sorbers, one line being used to feed each sorber.
The regenerated sorbent then enters the sorbers and the
cycle is completed.
Along with the major process loop of sorption,
heating and regeneration, two other systems are shown on the
flow sheet. These are the air and sorbent make-up systems.
In the air system, air is compressed from atmospheric condi-
tions to 40 psia by the air compressor, J-2. The air flow
is then split into four streams: solids transport to the sorbers,
solids transport to the gas-solids separator, solids transport
to the solids heater and combustion air to the solids heater.
The design temperature on the compressor effluent is 355°F.
Thus the transport air modifies the temperature of the solids
somewhat. For example, the regenerated solids are cooled from
1400°F to 1361°F by the transport air.
A make-up sorbent storage hopper, F-2, is provided
to replace solids lost to the flue gas effluent through attri-
tion. The vessel was designed to provide 1000 cubic feet of
active storage capacity and thus be able to receive truck load
quantities of purchased alkalized alumina sorbent. The solids
discharge from the hopper is connected to the pneumatic trans-
port system of a sorber feed line. Flows of make-up sorbent
shown on the flow sheet are estimated since attrition data for
this type of system are not available.
(2) Design Notes - Fluid Bed
Since at the time the fluid bed design was made
the fluid bed test program (USBM-Albany) was just getting
started and little experimental data were available, the pro-
cess design is, of necessity, based mostly on literature cor-
relations and various assumptions as to gas velocity, sulfur
loadings, etc. The resulting design should provide a valid
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preliminary basis for estimating the plant investment and
calculating process economics, however, since any modifi-
cations resulting from differences in actual vs. assumed
process variables will probably be small.
(a) Preliminary Process Calculations
1 - Minimum Fluidization Velocity
Initial investigation of the fluid
bed system was undertaken using both empirical design methods
and theoretical considerations. The range of the minimum
fluidization velocities (uMp), obtained from various correla-
tions is shown in Table 36. With the exception of the Frantz
correlation, all the calculated values for UMF fall in the
range 0.55 to 0.75 ft/sec. The average minimum fluidization
value is 0.638 ft/sec plus or minus a 0.064 ft/sec standard
deviation, if the Frantz correlation results are excluded.
This exclusion is probably justified since the range of particle
diameters studied by Frantz was well below the range of particle
diameters encountered in the alkalized alumina system, and the
resultant empirically derived "theoretical" equation developed
by Frantz reflects this particle diameter range in the equation
constant.
Values for the process variables used to
define U „ in the various correlations were obtained from pub-
MF
lished literature wherever possible or values were assumed when
no information was available. Since equilibrium particle size
distribution data from steady operation of fluidized beds were
not available .at the time this design was made, a size distribu-
tion had to be assumed. It was decided to use the particle size
distribution reported in Table 4 of USBM-Bruceton Quarterly Re-
port, June 30, 1968, for Grace #2 after the 3rd regeneration
rather than that used in the dispersed phase design because the
former has a slightly smaller amount of fines. The lesser
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TABLE 36
MINIMUM FLUIDIZATION VELOCITIES AT 600°F AND 1 PSIG
CORRELATION
Ergun, S.
Leva, M.
Leva, M., et. al.
Miller, C. 0. & Logivinuk, A. K.
Frantz, J. F.
Wen & Yu #1
Wen & Yu #2
REFERENCE
4
1
1
1
3
2
2
UMF
(ft/sec|
0.744
0.559
0.671
0.622
1.117
0.600
0.634
Leva, M., Fluidization, McGraw-Hill Book Company, Inc.,
New York, 1959. Tpp~r"63, 64 and 69).
Kuni, D. and Levenspiel, O., Fluidization Engineering,
John Wiley & Sons, Inc., New York, 1969. (p. 73)".
Frantz, J. F., Design for Fluidization, Chemical
Engineering, September 17, 1962;Minimum Fluidization
Velocities and Pressure Drop in Fluidized Beds, Chemical
Engineering Progress Synposium Series, No. 62, Vol. 62,
(p. 29).
Ergun, S., Fluid Flow through Packed Columns, Chemical
Engineer ing Progress, 48:89 (LWl) , (p. 81).
-228-
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quantity of fines would be expected for fluid bed operation
because of the lower gas velocity.
The weight mean particle diameter (D )
was calculated from the cited particle size distribution and
a value of 0.0359 inches or approximately 900 microns was ob-
tained. For some of the correlations the particle shape factor
and voidage at minimum fluidization were required and the as-
sumed values of 1 and 0.4, respectively, were used. The bulk
and particle densities of the sorbent were taken as 45 and 70
lbs/ft3 respectively, based on recent USBM data. These values
give a fixed bed voidage of 0.357. Since a slight expansion
occurs in the bed at incipient fluidization, the earlier as-
sumption of a 0.4 voidage at UMp is supported. The remainder
of the variables used in the UMF calculations are gas proper-
ties which have been fixed by the system. A typical calcula-
tion for UU1, is given in Appendix Table 1-97.
Mr
Although the computed values agree fairly
well, the actual minimum fluidization velocity should be verified
in pilot plant or bench-scale work. The particulate matter used
to determine the correlations mentioned earlier was for the
most part non-vesicular and since the alkalized alumina particles
are very porous, it is possible that the correlations could give
erroneous results.
2 - Fluid Bed Sorber Gas Velocity
To provide a basis for determining a
reasonable operating range of superficial gas velocity in the
fluid bed sorber, the terminal velocities of the maximum,
minimum, and weight mean diameter particles were calculated.
The maximum, mean, and minimum particle sizes of the sorbent
cited previously (i.e., Grace #2 after 3 cycles) were deter-
mined and the terminal velocities, Ufc, of these particles were
calculated:
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1. Maximum particle diameter, 0.081 in
(2057 microns):
U. = 28.8 ft/sec
2. Wt. mean particle diameter, 0.0359 in
(912 microns):
U. = 13.6 ft/sec
3. Minimum particle diameter, 0.008 in
(203 microns):
Ut - 2.2 ft/sec
Although it would be preferable to run the fluid beds in a
velocity range below the terminal velocity of the smallest
size particles, this would be impractical due to the large
volumes of gas being processed which would result in many
large diameter reactors. For example, at a 2 ft/sec operat-
ing velocity, eleven 60" diameter beds are required. There-
fore, it is desirable to use as high a gas velocity as prac-
tical and since no specific data are available, the design
velocity has been arbitrarily set at 5 ft/sec. Under these
conditions, the total cross-sectional area required (hence
the number of beds) is greatly reduced while the gas velocity
remains below the terminal velocity of the bulk of the particles,
thereby keeping entrainment at a reasonable level. A range of
superficial gas velocities should be evaluated in the pilot plant
to determine the practical upper limit which could be used in a
commercial design. Selection of the final design velocity would,
of course, involve an optimization whereby reduced equipment
cost resulting from higher gas velocities must be balanced
against increases in bed height, pressure drop, and attrition
losses.
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3 - Comparison of Bed Heights Required
for Fluid Bed Sorption - USBM-Albany
vs. CEGB-Blyth Models
Once a superficial gas velocity is esta-
blished, the bed height required to achieve the specified 90%
SC>2 removal can be determined. This is a function of the
kinetics of the system of which sorbent loading is an impor-
tant variable. A kinetic model based on the Amundsen rate
equation has been presented by the USBM-Albany1 to represent
the sorption reaction in fluidized beds of alkalized alumina.
Using this model the sorbent inventory and therefore bed
height were calculated for various assumed sulfur loadings.
The level of sulfur on the regenerated
sorbent after it has been through enough sorption regeneration
cycles to reach a equilibrium value was assumed to be 2.6% by
weight. This value was based on regeneration data generated
by USBM-Bruceton for various sorbents including Grace #1 and
#2, Kaiser PR-15, and two USBM sorbents designated as 423X-4E
and X-3F. The regeneration data reported naturally covered
a wide range of temperature, initial sulfur loading, reducing
gas composition and flow rate, and regeneration time. Since
a commercial sorbent has not been developed, it was felt that
for design purposes a satisfactory residual sulfur level could
be determined from those test data that were obtained at ap-
proximately the design conditions. On this basis a concentra-
tion of 2.6% was selected as a reasonable design value since
it was about mid-range of the pertinent test data. Other values
in the range 2-3.5% could have been used instead but the process
design and economics would not be changed appreciably from the
present case.
The other sorbent and gas properties
used were the same as those in the fixed bed study with the
1 USBM-Albany, Progress Report, November 1968
-------
exception of the reaction rate constant which was obtained from
the USBM-Albany report cited previously. Table 37 presents a
summary of the results obtained for various exit sorbent load-
ings.
The results given in Table 37 indicate
that the theoretical settled bed heights required for sorption
are small; i.e., the maximum predicted bed height is only
1.20 feet. However, this is a theoretical model and does
not allow for the non-uniformity and physical non-ideality
in the bed. The bed heights used in the actual sorber design
were a static bed height of 2 feet and an expanded bed height
of about 4 feet. This range of bed heights should provide
sufficient contact time for sorption under actual operating
conditions using a superficial gas Velocity of 5 ft/sec.
The final sorber bed height is deter-
mined by the examination of many criteria. The sorption
kinetics indicate a bed height of from 0.79-1.1 feet for
a reactor 40 feet in diameter. These low bed heights are
the result of the use of large diameter vessels which in
turn are caused by the fact that large volumes of gas are
being processed, and because the gas velocity must be kept
at a reasonable level, say 2-5 feet per second. To provide
good operability at 50% turndown (corresponding to off peak
power operation), the lower value of gas velocity was set
at about 4 to 5-fold above the minimum fluidization velocity.
At the other end of the spectrum, gas velocities much above
5 feet/second cannot be tolerated due to the excessive entrain-
ment they would incur. Attrition losses probably would increase
also at higher velocities. Thus, the design velocity of 5 feet/
second was chosen as a compromise between vessel and process
requirements and, although an optimization of velocity could
be made after sufficient data are available, it Ls unlikely
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TABLE 37
SORBENT INVENTORY AND BED HEIGHT VERSUS EXIT SORBENT LOADING
USBM-ALBANY KINETIC MODEL-NOV. 1968 PROGRESS REPORT
30RBENT SULFUR LOADING @ EXIT, -- 0.06 0.07 0.08
lb P
50RBENT INVENTORY, lb 444,015 511,750 610,710
JED HEIGHT REQ'D, ft, @ GAS VELOCITY
OF 5'/SEC:
40' - DIAMETER* 0.785 0.905 1.08
50' - DIAMETER* 0.837 0.965 1.15
60' - DIAMETER* 0.872 1.01 1.20
DIAMETER - No. of beds req'd.
40' - 10
50' - 6
60' - 4
-------
that the final velocity would differ much from the value
used. However, even at the operating velocity of 5 feet/
second, large diameter vessels are required; e.g., 10 vessels
@ 40 feet diameter, 6 vessels @ 50 feet diameter, etc. The
selection of 40 feet as the vessel diameter was made to mini-
mize the number of vessels used and yet not incur the operation-
al problems of extremely large vessel diameters (40 feet diameter
FCCU's are not uncommon in the petroleum industry).
With the vessel diameter, number of
vessels, and bed heights required for sorption set, the pres-
sure drop criterion must be tested. If a 30 inch E^Q pressure
drop is allowed across the vessel (as in the fixed bed case)
and the cyclone, distributor plate, and inlet losses are taken
into account, the actual bed pressure drop becomes about 18.5
inches of H-O. At this pressure drop the maximum allowable
static bed height is 2.14 feet, well above the bed heights
indicated by the kinetic models.
In order to improve gas distribution
across the bed and thus operate under more stable conditions,
and to provide a sufficient safety factor for the kinetic
models, the static bed height was set at 2 feet. At this
bed height and superficial gas velocity of 5 feet/second
through the bed, a gas residence time of 0.4 seconds is
obtained. This residence time compares favorably with
that which is presently being studied in the USBM-Albany
4 inch fluid bed reactor.
In the area of fluid bed design, many
empirical correlations and theoretical models have been pre-
sented in the literature. Due to the great variation which
may be experienced in a particular fluidized system, however,
the authors of these correlations recommend the physical veri-
fication of their work by pilot plant or bench-scale study.
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The particular areas of study should be the verification of
sorption kinetics at a 5 feet/second gas velocity over a range
of sulfur loading on the sorbent, and the determination of the
physical characteristics of fluidized alkalized alumina under
process conditions. Some of the physical characteristics which
should be determined before the final design of this type of
system is undertaken are: bed height, pressure drop, bed expan-
sion, transport disengaging height (T.D.H.), and the size and
amount of solids carry-over above T.D.H. For the current studies
however, it is believed that a design adequate for the present
needs can be made based on available data and literature correla-
tions. Necessary adjustments can be made as required when more
experimental data become available.
(b) Vessel Design
In the fluid bed sorption process, all of the
major process vessels, namely the sorbers, regenerators, and
heaters, are designed as fluid bed equipment. The sorbers and
regenerators are fluid beds by definition in the study, so as
to compare this type of operation to the fixed and dispersed phase
bed types studied separately. The solids heaters became fluid
beds through the optimization of equipment design and process
conditions and their selection as such bears no relationship to
the definition of the type of sorption used.
1 - Sorber
The fluid bed sorber was designed using
various empirical and/or theoretical correlations; the correla-
tion used for each specific section of the vessel has been noted
and referenced. For the 1000 megawatt power plant design basis,
ten 40 foot diameter fluid bed sorbers were used, giving a
superficial gas velocity at full load of approximately 5 feet/
second. As noted, previously, the static bed height will be 2
feet. Since no bed expansion data are presently available, the
-235-
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expanded bed height was assumed to be 4 feet, i.e., a 100%
expansion. The solids dust loading was set by means of the
particle size distribution given in a USBM-Bruceton Quarterly
Reportl*and the catalyst entrainment correlation of Zenz and
Weil2. With a solids dust loading set, the cyclone pressure
drop was calculated for a 60 feet/ second inlet velocity
based on an internal MWK correlation. The inlet gas dis-
tributor was designed as a baffle type to minimize the
height required for the inlet. The sizing and pressure
drop calculations for this distributor were done using MWK
correlations. The catalyst support and distributor plate
was sized to give the optimum hole spacing at the maximum
pressure drop using the correlation recommended by Zenz^.
Once the inlet, distributor and cyclone pressure drops were
calculated, a check was made to insure that the actual total
pressure drop agreed with the assumed total pressure drop of
30 inches H~O.
The cyclone size and dimensions were
approximated from the typical proportions given in Perry's^.
All cyclone dimensions given are comparable to this reference
with the exception of the top cylindrical section. Because
of the extreme size of this section a check will be made through
cyclone vendors to determine whether it is applicable to this
system or not. Alternatively, it may be desirable to use
several smaller cyclones in parallel as in FCCU practice.+ In
any case all cyclone dimensions will be verified by a vendor
before an estimate of this system is made. The cyclone dip-
leg was sized using an MWK correlation.
The sorbent withdrawal line from the
reactor is a dense phase, 30 Ib/ft3, fluidized line which
feeds into a dilute phase, 1-2 lb/ft3, solids transport line.
The sizes of these lines were determined by an internal MWK
correlation as was the solids feed line. The solids inlet
Cumbers refer to references at the end of this section
+ Subsequently it was decided to use smaller cyclones in'
parallel in the final design.
-236-
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nozzle was sized to give a velocity of approximately one half
of the transport velocity and was elevated 20 feet above the
top of the bed to provide sufficient room for the solids to
disengage from the transport gas. A detailed vessel sketch
and notes are given in Figure 1-99 as drawings AA-FB-3-A1
through A-4.
2 - Regenerator
The regeneration reactor for this system
is a fluid bed operating at 1400°F and an exit gas pressure of
15 psig. A regeneration temperature of 1400°F was selected
for the reasons given previously in this report under section
B,2,a. (2) (b)2- Dispersed Phase Vessel Design-Solids Heater;
viz., at 1400°F better reaction kinetics and a more economical
solids heater design are obtained. For the design purposes the
CEGB-Blyth kinetic regeneration model was used along with a
50% safety factor on residence time. This 50% safety factor,
plus the added safety factor obtained by using a sorbent with
higher regeneration rate characteristics, should provide a re-
generator design capable of handling minor temperature varia-
tions and varying intersititial sulfur loading without adverse-
ly affecting sorbent regeneration. Based on respective inlet
and outlet sorbent loadings of 7 and 2.6 weight percent sulfur,
a solids residence time of 30 minutes was obtained.
Two regeneration sections comprising a
regenerator and heater, were used in the design with each
regenerator section handling the solids flow from five sorbers.
This arrangement minimizes both vessel sizes and length of
transport lines since the regenerator sections can be located
symmetrically on the plot plan.
The regenerator vessel was sized using
the following bases: 1) a superficial gas velocity of 2 feet/
second, and 2) a sorbent inventory of 84,000 Ibs (30 minute
holdup). The gas velocity in the vessel was set at 2 feet/
second to allow for 1 foot/second superficial velocity operation
-237-
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during non-peak load (50% flow). At this latter velocity, the
flow is approximately double the theoretical minimum fluidiza-
tion velocity. Under these design conditions the vessel diameter
was found to be 12 feet and the static bed height for normal
operation was calculated to be 16-1/2 feet. In order to pro-
vide for uncertainties in the regeneration kinetics (e.g.,
effect of back-mixing of solids in the fluid bed), a 50%
safety factor based on sorbent inventory is provided, so that
a total bed height of 25 feet is obtained.
The gas residence time obtained by using
a 16-1/2 foot bed at 2 feet/second is approximately 8 seconds.
This value corresponds to a space velocity of 950 hr~^. Space
velocities in the range of 1000 hr"-*- have been run successfully
in the Bruceton thin bed tests but these values should be con-
firmed in the Albany fluid bed regeneration tests.
For the actual vessel design, e.g.,
cyclone, gas distributor etc., refer to section (b) 5.
3 - Fluid Bed Heater
The solids heater used in the fluid
bed sorption system is the same type which was used in the
dispersed phase system. It is composed of two fluidized
beds arranged in series in a single vessel. In the lower
bed natural gas is burned with air to heat the sulfur laden solids
to the 1400°F regeneration temperature. The bed height in
the lower bed was set at 10 feet to satisfy combustion require-
ments. Based on the fact that the exit solids and gas tem-
perature are in equilibrium, the combustion products leave
the first stage at 1400°F and pass through the upper bed
where they preheat the solids from 634°F to the equilibrium
gas-solids temperature of the upper bed which is 726°F. The
gas temperature drops to 702°F at the vessel exit as a result
c,^ mixing the transport air with the heater gas above the top
bed.
-238-
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The solids heater was sized on the
basis of a 5 feet/second superficial gas velocity for the
effluent of the lower stage. Using this criterion, a
vessel diameter of 9'-0" was obtained. The bed height for
the lower bed was specified at 10'-0" and a 5'-0" spacing
was left between the upper and lower beds to de-entrain
the bulk of the solids carried out of the lower bed. Since
the upper bed is a gas-solids heat transfer stage with no
combustion occurring, an expanded bed height of one foot
was provided (based on the dispersed phase bed calculations,
this should provide more than sufficient height for heat
transfer). The minimum fluidization velocities were then
calculated for each of the heater sections and compared to
the actual superficial velocities being run in these sec-
tions. In all cases the actual velocities exceeded the
minimum fluidization velocity by at least a factor of 4,
thus providing a safety factor of 100% during non-peak
gas flow.
As in the case of the regenerator the
various correlations and parameters used in the vessel design
and equipment specification are given in section (b) 5.
4 - Process and Mechanical Considerations
Regenerator Train
The major process consideration for
setting operating pressure in the regeneration train was
the overhead composition and temperature of the fluid bed
regenerator. The pressure drop through the Glaus plant
including any pretreatment equipment that might be required,
sets the pressure of the regenerator off-gas. Two methods
for increasing the regenerator effluent pressure were con-
sidered: 1) provide a compressor for the regenerator ef-
fluent, or 2), increase the inlet pressure in the regen-
erator. The first alternative would require costly materials
of construction for the off-gas compressor in that the regen-
-239-
-------
erator effluent contains a high quantity of sulfur compounds
at 1400°F. Larger sized vessels for the solids heater and
regenerator also would be required at the lower pressure.
If not required by process considerations, therefore, a low
pressure (0 psig) regeneration system with compression of the
off-gas has little if any advantage over a high pressure (15
psig) system. On the other hand, the second proposal has
some significant advantages in that, by increasing the
operating pressure of the regenerator, the gas is compressed
up-stream of the regenerator thus decreasing the vessel dia-
meter and the line sizes to and from the regenerator. Another
beneficial effect derived from increasing regenerator operat-
ing pressure is the reduction in size of the solids heater
since this vessel is run at nominally the same pressure as
the regenerator due to the solids overflow feed system.
Therefore, the higher pressure alternative was used in the
design and the regenerator effluent pressure was set at
30 psia.
The solids feed system for the regen-
erator train is a dense phase standpipe arrangement with the
solids overflow from the beds used as the standpipe feed.
The vessel internals for the regenera-
tion train distributors, cyclones, grids, etc., were sized
using empirical and theoretical correlations.
Vessel diameters, bed heights, tem-
peratures and pressures were set by process considerations
(see specific vessel section for details of the above).
Once these parameters were specified the transport disengag-
ing heights, TDH, were calculated using the correlations of
Zenz and Weil2. The solids loading at TDH was then calculat-
ed using the particle size distribution given in a USBM
quarterly report , the superficial gas velocity for each
-240-
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vessel, and the catalyst entrainment correlation proposed
by Zenz and Weil2, with the solids dust loadings set, the
cyclone pressure drops were calculated and cyclone selection
was made using vendor supplied data.
Two types of gas inlet distributors
were specified for the regeneration train: a baffle type
for the regenerator, and slotted pipe types for the solids
heater and solids cooler. The sizing and pressure drops
for these distributors were done using MWK correlations.
The catalyst support and distributor plate for each vessel
were sized to give the optimum hole spacing at the maximum
pressure drop using the correlation recommended by Zenz3.
The solids inlet line for the solids heater was sized to pro-
vide a velocity equal to half of the transport velocity and
thus decrease the penetration of the solids into the bed.
Slots are provided in this line to allow at least a partial
gas-solids separation before impingement on the bed, which
further decreases solids penetration. Solids withdrawal from
the beds in the regeneration train is handled by dense phase
standpipes with cone top feeders to capture the overflow solids
Three concentric pipe rings are used to feed natural gas to
the bottom bed of the solids heater. These rings were designed
to provide uniform heat transfer by maximizing the gas distri-
bution in the bed.
The vessel sketches and design data for
these vessels, the regenerator and solids heater, are shown
in Figures 1-100 and 1-101, respectively.
5 - Plant Layout
A preliminary plant layout study was
undertaken to determine an optimum vessel configuration so
as to minimize both the amount of the large diameter ducting
required, and the plot size and field fabrication problems
associated with this ducting. Consequently, seven different
-241-
-------
layouts were made and evaluated. The study results indicated
that there was little difference between the configuration in
terms of total cost; the maximum cost was about 5% higher than
the minimum cost, based on the approximate estimating techniques
used for this study. The case which was finally selected had
an estimated cost for materials and construction that was midway
between the highest and lowest but gave the minimum plot area
and simplest layout. A plot plan showing this layout is pre-
sented as Figure 54. Since all the vessels are at nominally
the same pressure, they are stacked to allow gravity feeding
between vessels. A tentative layout of the proposed plant,
Figure 55, shows the elevations required to Obtain this type
of solids feed system.
6 - Solids Transport System
The transfer of sorbent particles
between the reactors and regeneration train is done pneumati-
cally, using dilute phase gas-solid mixtures. However, because
of the difference in operating pressure between the reactor
and regenerator train (i.e., 15 psia vs. 30 psia respectively),
the reactor exit solids cannot be fed directly into the solids
heater without some means of increasing the pressure in the
solids feed line. This is done by use of an intermediate feed
hopper sufficiently elevated so that the necessary pressure
buildup is obtained in the dense phase leg, thus providing
the solids driving force for the pneumatic elevation of the
sorbent to the solids heater.
The line sizes and transport gas re-
quirements have been calculated based on a solids loading of
one pound per cubic foot for the dilute phase lines, and 30
pounds per cubic foot for the dense phase standpipes. However
for the pressure drop calculations a slip factor of two was
used, this value having been estimated from the particle size
-242-
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AJR FBOU
FOPCtO DRAFT
FLUE C,AS
NOTE:
QA.S OUTLET UNtS FROM
SCLIOS HEATER AWO 0,4.5- SOLIDS
5CP6,R4,TOB MOT SMOWKi .
FLUE
REFORMER AREA.
\
RE5EME3A.TQR
uv 0;
FIGURE. 54
FLUID BED 50RPTIOM EQUIPMENT IAVOUT
(F'wA-Vt V1E.W)
-243-
-------
150 ff -|
SQUD5
HEATER
^r^
^^
SOUO^i
FROM OOMPRE.ft*O
NC3TE:
5TROCTU«ft.L STEEL V1OT *MOWM.
FIGURE 55
FLUID BED SORPTION EQUIPMENT LAYOUT
(ELEVATION)
rr-
50 fT-
* 77V
-244-
-------
distribution and theoretical settling velocities. Compressed
air at 35 psia was used as the transport gas in the pressure
drop calculation and the intermediate feed hopper elevation
was calculated to be 115 feet.
7 - Gas-Solids Separator
The intermediate feed hopper, mentioned
in the previous section, was designated as the gas-solids
separator. This vessel has the dual function of providing
a gas-solids disengaging region for the dilute phase risers
and providing a six minute surge capacity, based on solid
flow, to insure a fairly constant head on the dense phase
standpipe. The vessel diameter was set at 9'-0" by assuming
an L/D ratio of approximately 1. A cyclone was provided to
minimize solids losses. The vessel is fed by risers from
each of the five sorbers in the train. As in the case of
the fluid bed designs, these inlet lines are designed for a
reduced velocity and are fitted with slotted ends. The vessel
is provided with a steep cone bottom to insure drainage into
the standpipe at all times. This vessel is shown as Figure
1-102.
8 - Air System
Compressed air at 35 psia is provided
as combustion air to the solids heater and as a transport
medium for the solids. A single air compressor per train
is furnished for this service. The compressor discharge is
split via appropriate controls into the four separate streams
required.
-245-
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REFERENCES
1. USBM-Bruceton Quarterly Report, June 30, 1968, Table 4.
2. Zenz, F. A. & Weil, N. A., A Theoretical-Empirical Approach
to the Mechanism of Particle Entrainment from Fluiaizetf
Beds, AIChJE Journal, December 1958, Vol. 4, No. 4, pp.
373^479.
3. Zenz, F. A., Bubble Formation and Grid Design, Tripartite
Chemical Engineering Conference, Symposium on Fluidiza-
tion II, Session 32, September 25, 1968 (Montreal)
4. Perry, R. H., et. al., Chemical Engineer's Handbook,
McGraw-Hill Book Company, N. Y. 1963, pp. 20-69.
-246-
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c. Fixed Bed
(1) Fixed Bed Process Description;
Flow Sheet P3196::D~ ~
The S02-rich flue gas from a 1000 megawatt power
plant burning a 3% sulfur coal enters the flow sheet after
leaving the power plant's precipitator at 600°F, as stream 1.
Stream 1 is compressed by the flue gas booster fan, J-l, from
10 inches of water to 45 inches of water to overcome the line
losses and bed pressure drop in the SO2 removal scheme, cal-
culated to be 35 inches of water; thus the flue gas is returned
to the power plant circuit at the same pressure at which it
was withdrawn. The compressed gas then enters a header, common
to the four fixed bed sorbers, D-l thru 4. During steady state
operation, two of the four sorbers will serve to remove S02
from the flue gas, while the remaining two sorbers will be in
the regeneration, heating or cooling cycle. A sorber by-pass
line to the power plant stack via the air preheater is also
provided on the header. This line is designed to accept 50%
of the stream 1 flow should maintenance require one of the
sorbers to be shut down temporarily.
The stream 1 flow from the header is split and
fed to each of the two reactors in which sorption is occurring
(in the case shown on the flow sheet sorbers Dl and D2) as
streams 1-A & 1-B. Each sorber is composed of 84 annular sor-
bent beds of alkalized alumina enclosed in a rectangular chamber.
The beds are 3'-7" O.D. by 2'-4" I.D. by 22'-0" high, giving
a bed thickness of ^7"and a bed pressure drop of 30 in. H20,
and are described in the design notes.
The flue gas from stream 1 flows into the rectan-
gular sorber chamber then through the annular sorbent beds where
S02 is removed from the gas stream. The clean flue gas then
flows up the center cylinder, formed by the annulus, to a
-247-
-------
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-248-
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-------
common sorber exit gas chamber and out of the sorber to the
effluent flue gas header, stream 2. The clean flue gases flow
through this header and combine with the sorber heating or
cooling gas from a third sorber. A side stream, stream 4, is
withdrawn from these mixed gases for use as a heating or cool-
ing medium for the sorbent beds. The remainder of these mixed
gases, stream 3, leaves the process area and flows to the power
plant air preheater for heat recovery before being discharged
through the stack.
Because the process utilizes fixed bed sorption
and regeneration, the operation is on a batch rather than a
continuous basis. The cycle time for this process was set by
allowing two hours as the time required for the regeneration
step (and one hour each for heating and cooling the beds to
and from the regeneration temperature of 1200°F. Thus a total
heating-regenerating cooling time of four hours is obtained.
Based on the above, plus a four hour sorption time, a complete
cycling of the sorbers is made in eight hours. Four sorbers
and one hour heating and cooling time were chosen to allow
major process equipment (such as the recycle flue gas blower,
J-2, the regeneration gas generation and purification sections,
and the Claus-gas purification section) to operate on a con-
tinuous basis.
As mentioned earlier, the sorbent beds must be
preheated from the 600°F sorption temperature to the 1200°F level
required for the regeneration reaction. Recycled flue gas
from the sorber exit header is used as the heating medium
for this purpose, and is taken off as a side stream from the
effluent flue gas heater, stream 2. This side stream of
recycle flue gas, stream 4, is compressed from 10 in. H20
to 25 in. EjO, by the recycle flue gas blower, J-2, to over-
come line losses and bed pressure drop, which are smaller
-249-
-------
than for the flue gas circuit because of the smaller gas flow.
The recycle gas at 602°F is then passed through the convection
section of a natural gas fired furnace where the furnace's
combustion products are mixed with the recycle gas. This
furnace, the recycle flue gas heater, B-l, was designed to
give an effluent flue gas temperature of 1225°F, when using
10% excess air. Two blowers, the air blower, J-3, and the
pilot air blower, J-4, provide the air required for combustion
in this furnace. Both blowers are designed to provide a
positive pressure of 25 in. H20, which is sufficient to move
the heating gas through the fixed sorbent beds.
The pilot air blower, J-4, serves the function
of providing combustion air, stream 5-A, during the cooling
cycle when no process flow is required through the furnace.
The combustion products from the pilot burner serve to main-
tain the operation temperature in the furnace, thus preventing
thermal cycling in the furnace during the no-flow condition.
The effluent from the pilot burner is fed directly to the stack
through its own exhaust duct which ties into the clean flue
gas duct, stream 3, downstream of the sorbers.
The recycle heating gas effluent from the furnace,
stream 7, enters the common sorber inlet duct, stream 1-D, and
flows through the sorbent beds, preheating them to 1200°F for
regeneration. Heat transfer in the fixed bed is very rapid with
the temperature profile in the bed appearing as a sharp wave-
front. Because of this, the exit gas temperature is 600°F, the
initial bed temperature, for almost the entire heating cycle.
It is only during the end of the cycle that the wave-front begins
to break through causing an increase in the exit gas temperature.
The average bed exit temperature is 625°F, whereas the final bed
exit temperature is approximately 1200°F. The spent heating gas
then leaves the sorber via the common sorber effluent duct,
stream 2-D, and flows into stream 2 for disposal to the stack.
-2 50--
-------
The cooling cycle is essentially the same as the
heating cycle; i.e., clean flue gas from stream 2 is recycled
to the sorbers to cool the beds from 1200°F to the sorption
temperature. There are, however, two differences in this
cycle; first, the recycle flue gas heater, B-l, is by-passed;
second, the bed exit temperature for the cooling gas is 1200°F
for the greater part of the cycle. Due to the latter difference
the average effluent gas temperature is increased which means
that the recycled gas cannot cool the beds down to the sorption
temperature of 600°F, but rather to the mix temperature of the
gas in the effluent header, 675°F. To compensate for this, the
gas flow rate is increased slightly thus causing an earlier break-
through. Once break-through occurs, the effluent gas temperature
drops off sharply, giving a mix temperature in the effluent gas
header of approximately 600°F. Therefore, during the remainder
of the cycle gas inlet cooling gas to the sorber is at the desired
temperature. This cooler gas removes further heat by forming a
second wave-front in the bed. This wave front precools the front
section of the bed at 600°F for sorption and propagates itself
as further 600°F gas is passed through the bed. Even if the
entire bed has not been cooled to 600°F at the end of the cooling
cycle, the introduction of 600°F flue gas for sorption will con-
tinue to cool the sorbent by extending the cooling wave-front
ahead of the reaction zone.
The last operation cycle is that of regeneration.
The sulfated alkalized alumina is regenerated with a reducing
gas to desulfurize the alumina and form a sulfur-rich gas which
can be fed to a Claus unit for sulfur recovery as elemental
sulfur. On the present flow sheet, a 1200°F reformed natural
gas, containing 10% CO, was used as the reducing agent. The
regeneration gas generation and purification stages are shown
on a separate flow sheet, P1374-B, section B.2.e. part (1) of
this report. Similarly, the regeneration gas effluent from
-251-
-------
the sorters has to be pre-treated prior to entering the Glaus
unit; the Claus-gas purification stage is also shown on a
separate flow sheet, P1375-B, and discussed in section B.2.e.
part (2) of this report.
The assumed regeneration gas, containing a 25%
excess over stoichiometric proportions, flows as stream 8 to
the common sorber inlet duct, stream 1-C, at 1200°F. Leaving
the inlet duct, the gas passes into the sorber and through
the annular sorbent beds to remove the sulfur from the sorbent.
The products of this reduction reaction, based on a natural
gas reformate, are: sodium aluminate (NaAlO2), hydrogen sul-
fide (H2S), carbonyl sulfide (COS), sulfur dioxide (SO2>, ele-
mental sulfur (S2), and water. This sulfur-rich gas flows out
of the common sorber exit duct, stream 2-C, to the regeneration
gas effluent header, stream 9. This header feeds the Glaus gas
purification stage where the gas is pre-treated (H2O removal, etc.)
and sent to the Glaus unit for sulfur recovery as elemental
sulfur.
The high process temperatures encountered in the
large ducts require water-cooled closure rings (i.e., valve
seats) to minimize thermal stresses in the duct walls. To
avoid scale formation in the rings, treated cooling water is
required. This water will be made available from off-site
facilities.
It should be noted that the sorber operation shown
on the flow sheet corresponds to the first set of conditions
given in Table 38 which outlines the proposed operation schedule
for each of the four sorbers. The material balance given in
Table 39 shows the process flows for the major process streams,
also for the first set of conditions in Table 38. To show the
changes in stream flow which occur during a cooling cycle, Table
40 was provided as an alternate to Table 39. Table 40 shows only
those flows which would be affected by the cooling cycle; all other
£u.ows are. the same during both the heating and cooling cycles.
-252-
-------
TABLE 38
SORBER OPERATIONS SCHEDULE
TIME
(MRS)
0
1
2
3
4
5
6
7
8
D-l
S
S
H
R
R
C
S
S
S
D-2
S
S
S
S
H
R
R
C
S
SORBER
D-3
R
C
S
S
S
S
H
R
R
D-4
H
R
R
C
S
S
S
S
H
NOTES: S denotes a train which is sorbing S02
R denotes a train which is being regenerated
H denotes a train which is being heated
C denotes a train which is being cooled
0~7 represents one complete cycle; 8 is identical to 0
and is the start of another cycle
-253-
-------
TABLE 39
HEATING CYCLE MATERIAL BALANCE
STREAM NO,
COMPOSITION M.W.
CO 28.011
1 (INLET FLUE GAS
FROM POWER PLANT) ;
LB/HR
CO, 44.010 1.755,024
COS 60.075
H, 2.016
H,0 18.015
R,S 34.080
423,792
KPH
39,877.84'
23,524.42
N, 28.013 6,024,234 215,048.30
0, 31.999
S| 64.123
SO, 64.063
CHi 16.043
TOTAL
STREAM Nd.
COMPOSITION M.W.
CO 28.011
CO, 44.010
COS 60.075
H, 2.016
H,0 18.015
H,S 34.080
N, 28.013
O, 31.999
S, 64.128
SOj 64.063
TOTAL
304,411 1 9,513.14
39,674 619.30
'
1-A (FLUE GAS INLET
TO SORBER £-1)
LB/HR
877,512
211,896
3,012,117
15J ,206
19,837
I
8,547,135| 288,583.0c| 4,273,568
MPH
19,938.92
11,762.21
107,524.15
4,756.57
309.65
144,291.50
2 (SOKBHR EXIT HEAUEkj i-A (FLUE GAS EXIT
LB/HR
1,923,638
' 483,814
6,*51,18B
322,185
MPU
43,709.11
26,856.18
237,428.79
10,068.60
9.380.8-25 318,062.68
FROM SORBER 0-1)
I.B/HR
877,512
211,896
3,012,117
MPH
19,938.92
11,762.21
1-B (FLUE GAS INLET
TO SORBER D-2J
LB/HR | MPH
877,512
19,938,92
211,896 I 11,762.21
3,012,117
i-c (REGEN. GAS INLET] I-D (HEATING GAS INLET
TO SORBER D-3) | TO SORBER D-4)
i _
LB/HR
9628.50
17758.40
5338.07
1414.00
MPH
343.74
404.19 1
2647.85
78.49
LB/HR
68,614
MPH
3831.27
60,022 3331.76
107,524.15 J626.953
152,206 4,756.57 : ; 27,682
19,837 309.65. j
' I 7.06 0.44 !
.
4,273,568
144,291.50
1
22380.49
865.11
34176.03 3474.71 883,271 30408.63
i-B (FLUE GAS EXIT ' i-C (RiJEN. GAS EXIT i 2-Q CHEATING GAS EXIT 1
FROM SORBER &-2) FROM SORBER D-3) FROM SORBER D-4) '
LB/HR
877,512
211,896
107,524.15 3,012,117
147,251 4, 601. 75 j 147.251
,... . I .,
4.248.77S 143.827.03
4.248.776
MPH LB/KR I MPH
i
2492.42
19,938.92 ,28527.28
645.81
! 1128.07
11,762.21
107,524.15
4,601.75
'29056.75
18875.55
479.68
254S.22
I4J.827.Q3 83757.8^
I
88.98
648.20
10.75
459.56
1612.92
553.86
7.48
39.73
3521.^2
LB/HR
168,614
60,022
626,953
27,682
3S3rJ71
MPH
3831.27
3331.76
22,380.49
865.11
-254-
-------
TABLE 39 (COH'T)
STREAM NO.
COMPOSITION M.W.
CH. 16,041
CJL 30.070
CjHg 44.097
C^Hlf. 58.124
CSH1? 72.151
CSH14 86.178
CO 28.011
CO, 44.010
COS 60.075
H2 2.016
Hto 18.015
H'S 34.080
N, 28.013
Oj 31.999
Si 64.128
SO, 64.063
TOTAL
3 (CLEAN FLUE GAS
TO STACK)
LB/HH
1,779,734
447,631
6,153,531
296,045
MPH
40.439..31
4 (RECYCLE FLUE GAS)
LB/HR
143,904
24,847.66 36,183
il9,663.85 497,656
9,314.21 24,140
i.
94.265.03 701,883
MPH
3269.80
2008.52
5 (AIR TO RECYCLE
FLUE GAS HEATER)
LB/HR
3,734
17764.94 116.286
754.39
35,096
23797.65 155.116
MPH
207.27
4151.09
1096.78
5-A (PILOT AIR TO RECYCLE
FLUE GAS HEATER)
LB/HR
415
12,921
3,899
5455.14 J17.235
MPH
23.04
461.24
121.85
606.13
6 (NATURAL GAS TO RE-
CYCLE FLUE GAS HEATER
LB/HR
8221. 3S
4C?.04
101.42
99.39
77.20
45.67
89.92
9037.03
MPH
512.46
13.37
2.30
1.71
1.07
0.53
3.21
534^ 65
STREAM NO.
COMPOSITION M.W.
CO 28.011
CO- 44.010
COS 60.075
H, 2.016
HjO 18.015
HjS 3". 080
N, 28.013
Oi 31.999
S, 64.128
SO. 64.063
CB4 16.043
TOTAL
7 (RECYCLE FLUE GAS 8 (REGEN. GAS FEED ' 9 (REGEN. GAS EFFLl^
HEATER EFFLUENT) TO SORBER) , FROM SORBER)
LB/HR
168,614
60,022
626,953
27,682
••3.271
MPH LB./KR
3831.27
3331.76
9628.50
17708.40
5338.07
1414.00
22380.49 !
865.11
10401. (3
7.0C
94176.03
MPH
343.74
404.19
2647.85
78.49
0,44
3471.7;
1 LB/HR
2492.42
28527.28
£45.81
-L1.J8.07
29056.75
18875.55
479.68
2545.22
83757.84
MPH
88.98
648.20
10.75
559.56
1612.92
553.86
7.43
39.73
„
3*21.92
T
-255-
-------
TABLE 40
COOLING CYCLE MATERIAL BALANCE
STREAM NO.
COMPOSITION
co2
H20
n2
°2
TOTAL
M.W.
44.010
18.015
28.013
31.999
2(SORBER EXIT HEADER
LB/HR
1,948,485
470,541
6,688,372
327,011
9,434,409
MPH
44,273.68
26,119.43
238,756.17
10,219.43
319,368.71
3 (CLEAN FLUE GAS TO STAC
LB/HR
1,757,495
426,176
6,037,164
294,857
8,515,692
MPH
39933.99
23656.74
215509.86
9214.56
^a«,31!>.15~
STREAM NO.
COMPOSITION
CH.
C286
C3H8
C H
C5H12
dL14
H20
N
°2
TOTAL
M.W.
16.043
30.070
44.097
58.124
72.151
86.178
44.010
18.015
28.013
31.999
4 (RECYCLE FLUE GAS) 1
LB/HR
193,461
46,749
664,138
32,509
936,857
MPH
._ .
6 (NATURAL GAS TO PILOT
BURNER OP B-l)
LB/HR
|_ ._ — __ .
822.14
40.20
10.14
9.94
5 7.72
4395.84
2595.01
23,707.87
4.57
8.99
1,015.94
31,714.66 J9d3.7d
MPH
51.25
1.34
0.23
0.17
0.11
0.05
0.32
53.4?
NOTES:
1.
The flow and composition given for Stream 4 are the same
as that for Streams 1-D, 2-D and 7, which are to be changed
from the values given in Table 39.
2. There is no Stream 5 flowing during the cooling cycle.
-256-
-------
(2) Design Notes - Fixed Bed
(a) Process
1 - Selection of Base Case Design for Estimate
Although the actual sorption reaction has
not been completely defined, kinetic data were presented by the
USBM-Albany. These data coupled with the Amundsen rate equation
provided a reaction rate constant which was used in the USBM-
Bruceton sorption model to calculate sorber size and operating
parameters. At the time it was felt that this procedure would
provide a basis for a fixed bed design comparable to the fluid
and dispersed phase cases. However, additional data for a fixed
bed sorption would be required to confirm the present design
before an actual plant was built.
A computer program (discussed previously
in the Sorption Section) using the above model has been developed
so that the effect of system variables upon design can be ascer-
tained. Under initial consideration was the use of multiple,
low-height, fixed-beds, with a large number of beds contained
in a single vessel. Parallel trains of multiple vessels would
be required since a cyclic method of operation is necessary, i.e.,
one train is regenerated while the other train is being loaded.
Two main advantages derived from this type of operation are:
1. the reactor and regenerator are contained in the same
vessel, and
2. attrition losses are minimized.
The problem areas in a system of this type are pressure drop,
gas distribution, piping design, and catalyst life. It would
also be necessary to remove the fly ash from the flue gas
upstream of the absorber to prevent plugging of the sorbent
beds.
-257-
-------
Pressure drop considerations keep the
bed heights low (1-5 feet) and since a. large amount of gas
must be processed, either the number of beds or the bed
diameters become large. At low pressure drop and large diam-
eters gas distribution becomes increasingly more difficult.
On the other hand, should the number of beds be increased, to
avoid distribution problems the piping arrangement would
become more complex and pressure drop would be increased.
Sorbent life in fixed bed systems is
an area in which little information has been reported. In
order to make the fixed bed system economically feasible the
sorbent should have no marked loss in activity or physical
degradation over several hundred in-place regenerations.
As a first step, cylindrical vessels
were assumed and the number of sorption stages calculated
for various process conditions as summarized in Table 41.
The information presented in Table 41 was obtained for each
of two stream times and bed pressure drops, respectively
12 and 24 hours and 2 and 30 inches of water. The higher
bed pressure drop was set by allowing this pressure drop to
equal the normal AP of 2 inches of H_0 plus a one psi N28
in. H20) increase, which is about the upper limit on pressure
that can be obtained from a booster fan. Due to the problems
of gas distribution and extremely low L/D ratios ( pancake type
vessels) it was decided to evaluate "panel type" arrangements
for the sorber configuration rather than cylindrical shapes.
Consequently, a thin rectangular shaped panel design was
developed as the first basic design in the fixed bed evaluation.
Two different methods of contacting gases
and solids which may have merit in the alkalized alumina process
are offered by the Dorfan and Squires contactors. However, both
the Dorfan and Squires contactors1 differ somewhat from our
1 AVCO Phase III Reports, Nos. 1-5
-258-
-------
TABLE 41
NUMBER OF PARALLEL FIXED-BEDS REQUIRED
FOR
90% SORPTION OF SO.,
TIME (hrs) 24
AP (IN. H20) 2
BED HEIGHT (ft) 0.42
SORBENT VOLUME (ft3) 76,710
f STAGES* 145
24 12
30 2
1.53 0.30
76,710 39,075
40 103
12
30
1.
39
29
07
,075
* Number of parallel fixed-beds required at a bed diameter of
40 feet.
-259-
-------
sorption panel design in that they use a moving bed design,
requiring a mechanical and/or pneumatic sorbent collection,
feeding and transporting system as well as a separate regen-
erator. A brief description of the two systems is given below.
The Squires' sorber contains louvered
panels to (1) direct gas flow through the sorbent, and (2)
keep the sorbent from spilling out of the sorber (Figure J-103 ),
Since these louvers are inclined to the horizontal the length
of the path of the gas through the sorbent is relatively long.
Thus in order to operate with a low pressure drop, a low gas
velocity through the sorber is required; therefore to process
the large volume of flue gas from a 1000 MW power plant, which
is the basis for the present design, the contactor cross-sec-
tional area must be large.
The Dorfan design (Figure J-104) uses
louvers to direct the sorbent flow downward and yet provide
a nearly true cross-flow design. The use of the louvers de-
creases the available cross-section for gas flow, however,
necessitating larger or more numerous units than a true cross-
flow sorber using a static bed.
Due to the attrition characteristics
of the present sorbent, the concept of a static bed rather
than a moving bed was the alternate chosen for study along
with the fluid or dispersed phase systems. Since the design
calculations for a fixed or static bed and a moving bed in
plug flow are quite similar, it is felt that the fixed bed
design will establish a sound reference point from which
the Squires and Dorfan designs can then be evaluated if
desired. In addition, it would appear that the present work
on the fixed bed design should achieve the benefits of in-
expensive containment of the sorbent in the optimum configura-
tion that seen to be the objective of the Squires and Dorfan
-260-
-------
designs. Accordingly, a detailed evaluation of the Squires
and/or Dorfan systems was not made.
The basic static bed design uses a thin
bed of sorbent to minimize bed pressure drop, but a higher
pressure drop alternative, requiring a flue gas booster fan,
was also evaluated. The first sorber design comprises a
number of sorbent filled, rectangular shaped panels in
parallel through which S02-rich flue gas flows (Figure J-105).
Each panel consists of a center core of sorbent held in place
by the vertical walls of the panel which consist of a 100 mesh
screen backed by a 2 mesh screen and a "subway" grate; in short,
a sandwich arrangement in which the screening and grate form the
sandwich wall. The top of each panel is blinded to act as a
seal thus minimizing channeling of the gas should settling of
the sorbent solids occur (Figure J-106). The panels are 40
feet in length with an active bed height of 20 feet plus an
additional 2 feet for the seal, giving an overall height of
22 feet. The panels are inclined 23° from the horizontal
to allow for gravity filling and draining of the sorbent
(Figure J-107} . A cylindrical design was later adopted as
the basic unit because of mechanical requirements (see section
(b)-l below). This discussion on rectangular beds is included,
.however, since it applies equally well to the cylindrical
design.
In order to prevent frequent start-up
and shutdown of the regeneration and air compressors, a con-
tinuous operation schedule, for one complete sorption and re-
generation cycle, utilizing four parallel trains of equipment
was derived (Table 38). This operations schedule is arranged
such that at any given time, two trains are in the sorption
stage, one train is being regenerated, and one train is being
heated or cooled. The operation schedule in Table 38 is broken
down into eight steps of equal time increments, the increments
-261-
-------
being a function of the onstream time required for sorption (e.g.,
if the elapsed time required to load the sorbent in the fixed bed
units is 24 hours, the time increments are 6 hours each; similar-
ly for a 12 hour sorption time the time increments are 3 hours
each).
Two additional design cases (i.e., 6 and
4 hour sorption times) were added to those previously reported in
Table 41. A summary of the design for each of these four cases
is shown in Table 42. The number of on-stream sorption stages
for each case shown in Table 42 was determined by obtaining the
total cross-sectional area required for flow and dividing this
area by the area per panel (401 x 20'). Under these conditions
the number of cycles per year, number of panels per train, width
per panel, length per train and plot area required were calculated
and the results shown in Table 43. From the results shown in
Table 43 it is seen that increasing the pressure drop by about
1 psi (from 2 to 30 inches of water), decreases the number of
panels required by a factor of about 3.5. Since the higher pres-
sure drop cases require the use of a flue gas booster fan, however,
the resulting increase in operating costs must be balanced against
investment savings to determine if there are net savings to be
realized by increasing the pressure drop.
It is also apparent from Table 43 that
as the sorption time decreases, the number and size of the panels
(hence the sorbent inventory) decrease, but the number of sorption
regeneration cycles per year increases. Comparing the number of
cycles per year for the various sorption times shows a linear
relationship; i.e., halving the sorption time doubles the number
of cycles per year. A similar comparison of sorbent inventory
and sorption time shows that the relationship is not quite linear;
e.g., reducing the sorption time by a factor of 6 only reduces
the inventory by a factor of 5.5. Thus as the sorption time de-
creases the sorbent would require more frequent replacement,
-262-
-------
TABLE 42
EFFECTS OF SORPTION TIME AND PRESSURE DROP
ON FIXED BED PARAMETERS
TIME
(hrs.
AP
BED THICKNESS
(In.)
GAS VELOCITY
(ft/sec)
"NUMBER OF
ON-STREAM
SORPTION PANELS
• • «
24
24
12
12
6
6
4
4
30
2
30
2
30
2
30
2
18-1/2
5
13
3rl/2
9
2-1/2
7
2
1.23
0.339
1.71
0.476
2.27
0.64
2.70
0.75
64
230
46
164
36
122
30
104
*
'.*Total number of panels is twice the number of on-stream sorption
'•panels.
-263-
-------
TABLE 43
EFFECT OF SORPTION TIME AND PRESSURE DROP ON SORBER DIMENSIONS
i
K>
SORPTION1
TIME
(hrs)
24
24
12
12
6
6
4
4
AP
(In. H20)
30
2
30
2
30
2
30
2
SORBENT
INVENTORY
Thousand Pounds
6904
6904
3517
3517
1823
1823
1259
1259
CYCLES/
YEAR2
(-)
182.5
182.5
365
365
730
730
1095
1095
#PANELS/
TRAIN3
(-)
32
115
23
82
18
61
15
52
THICKNESS/
PANEL
(In.)
36*5
23
31
21*5
27
20%
25
20
LENGTH/
TRAIN
(ft-)
97
221
60
147
41
105
32
87
PLOT
AREA4
(ft2)
20,370
46,410
12,600
30,870
8,610
22,050
6,720
18,270
1 Time required for one complete sorption-regeneration cycle is twice the sorption time
2 Based on a 100% stream factor (365 days/year
3 Four trains required
•* Width of plot is 210 feet in all cases
-------
although the difference in sorbent life is small. For example
if a 24 hour sorption time design had a one year sorbent life,
use of a 4 hour sorption time would require a life of 12/5.5
or 2.2 months to give the same sorbent costs. Since the labor
and time required to drain and fill the sorber system would be
appreciable, a sorbent life of only 200-300 cycles would prob-
ably require a long sorption time. Thus it is necessary to
know the number of cycles the sorbent can withstand before
complete economic evaluation of the fixed bed systems can be
made. Since these data were not available, assumptions were
required to complete the current design.
Calculations have shown the relationship
between sorption time, bed pressure drop, bed thickness, and
number of sorption panels required. Examination of these com-
binations shows that the lowest cycle time-high pressure drop
case will produce the lowest investment in equipment and sor-
bent inventory. Although the lowest investment will not neces-
sarily produce the lowest total annual costs, very brief, approxi-
mate estimates indicate that there probably is not & very large
difference in operating costs (exclusive of capital charges) for
the other high pressure drop cases. Brief approximate estimates
also indicate that the high pressure drop cases will probably
give lower total costs than the low pressure drop cases. There-
fore, the high pressure drop case using a four hour sorption
time has been selected for the base case design and estimate.
should be noted that the four hour case (eight hours for a
»
"complete sorption-regeneration cycle) is probably the minimum
time that can be used since it provides only two hours for
regeneration and one hour each for the heating and cooling
operations, including the time required for switching gas flows.
Using the cylindrical design that was
^selected for the base case (see section (b)-l below), bed
-thickness and number of beds were determined by use of the
-265-
-------
sorption model proposed by the USBM-Bruceton. The bed surface
area obtained from the model was increased by 20% due to blockage
of the bed surface by the support steel, a 10% safety factor was
then included to obtain the actual sorbing surface area, and the
number of units were adjusted to allow for a symmetrical arrange-
ment; as a result, the total surface area is about 50% greater
than that calculated from the model. However, this allowance
is not considered too large considering the lack of data for
the fixed bed system.
2 - Safety Hazards
With regeneration and sorption occuring
in the same vessel, a potential explosion hazard is created in
that reducing gas, composed mostly of combustible compounds
(H_ and CO), will periodically come into contact with gases
containing oxygen. That is, after the heating cycle ends,
the ducts and sorption vessel will contain heating gas which
has an oxygen content of about 3%. When the reducing cycle
is started, therefore, combustible gases first entering the
system will contact oxygen-containing gases at a temperature
of about 1200°F, which could possibly result in an explosion
or fire. Consequently, a study was undertaken to determine
if the fixed bed system could be operated as presently conceived
or would some modification be required, such as inert gas purging
between cycles.
A literature search was conducted to
determine if there is enough published information available
to evaluate the present system without having experimental
data. Essentially all of the references which were found
that contained data on inflammibility of gases referred to
experimental work done by the Bureau of Mines and published
in Bulletins 279 (first edition, 1928, revised and enlarged,
1930 and 1938), and Bulletin 503, dated 1952. The latter was
-266-
-------
used for the evaluation of the fixed bed system.
Referring to Bulletin 503, Table 44, page
130 shows, for various gases or vapors, oxygen percentages below
which no mixture is inflammable at "ordinary temperatures and
pressures." For H2 and CO, the combustible compounds in our
reducing gas, these oxygen limits are 5.0% and 5.6% respectively,
when the air is diluted with N2. The oxygen limits are increased
to 5.9% when the air diluent is CO2. Thus, at ambient tempera-
ture our mixture would not be inflammable since the oxygen con-
tent is well below the minimum required for combustion, i.e.,
3% actual vs 5% in 02-N2 mixtures. However, since our gases
will be at 1200°F, the elevated temperature may cause the mixture
to be hazardous and the problem becomes one of determining what
effect temperature has on flammability limits.
As a first step, it was assumed that when
the reducing gas is fed into the system and contacts the residual
heating gas, combustion will occur until all the oxygen present
is consumed. Knowing the volume of the equipment, the 02 concen-
tration, the temperature, and the pressure, the total heat liberated
and the adiabatic temperature rise can be calculated. For the
fixed-bed conditions considered, this temperature rise was found
to be about 720°F, and the corresponding pressure increase would
be 6.8 psi, assuming a pressure tight vessel. Since the vessels
arid ducts are not presently designed for pressure operation
(maximum design pressure is 45 inches of water), a much more
expensive design would result if the system had to withstand the
above pressure increase. Therefore, it is important to the fixed
bed evaluation to know if the reducing gas will burn when it
enters the sorption vessels. It became necessary, therefore,
to estimate the effect of temperature on flammability limits.
The effect of temperature on the H2-N2
flanunability limits was determined as follows. Figure 6, page 19,
-267-
-------
Bulletin 503 shows the influence of temperature on limits of
flammibility of H2 in air for temperatures up to 400°C (752°F).
The data in Figure 6 show a linear relationship between percent t
H2 and temperature and it was assumed that these data could be
extrapolated to 1200°F (550°C) for both the upper and lower
limits of H2 concentrations. Using the extrapolated limits*,
for H2 and air from Figure 6, the adiabatic temperature rd.se ,
for these mixtures was then calculated and compared :to the ><*
actual temperature rise. For the upper H2 concentration limit
the calculated temperature rise of the fixed bed process gases
was essentially the same as that for the H2~air mixture2 concentration on flammability limits is unknown. .
For the upper H2 concentration case, the adiabatic temperature
rise calculation indicates that a flammability problem may exist
but no firm conclusion can be drawn at this time. That is, the
present calculations are based on an assumed reducing gas com-
position and an extrapolation of data obtained from-H2 burning
in air in small tubes (2.5 cm diameter and 150 cm-«long) for
downward propagation of flame. Other data listed in Bulletin.
503 (Table 1) indicate that the H2 limits increase with increases
in vessel size, and that direction of the flame propagation (i.e,,
upward, downward, horizontal) also influences flammability limits.
Further, the strength of the ignition spark also-has an effect on
flammability limits.
As a result of the above discussed evalua-
tion of the potential safety hazard arising from mixing combustible
gases with 02-containing gas, the conclusion reached is that a
hazard may exist but available literature data are insufficient
to clearly define the flammability limits for the present fixed
bed design. Therefore, the present preliminary fixed bed design
will be made on the assumption that no hazard exists, but with
-268-
-------
the reservation that if the fixed bed case is competitive with
either of the other schemes (fluid bed or dispersed phase), ad-
ditional detailed study of the flammability limits will be re-
quired. Further, with the proper reducing gas composition, as
determined by regeneration studies, experimental data on the
flammability limits of the actual process gases would be required
before an actual plant could be designed and built.
3 - Heating and Cooling Gases
The heating gas flow was determined by
calculating the heat required for a 600°F temperature increase
in the sorbent and internal sorber structure, to which was
added the slightly endothermic heat of reaction and the heat
losses to the atmosphere from the sorber structure during maxi-
mum winter design conditions (25°F and 20 MPH wind velocity).
Because the heat transfer in a fixed bed is very rapid, the
solids and gas temperatures at any point are approximately
equal, and therefore the inlet gas temperature was set as
close as possible to 1200°F (regeneration temperature) since:
(1) the allowable design stresses of the sorber structural steel
decrease rapidly as the temperature increases above 1200°F; and
(2) a lower temperature provides less risk of sintering the
sorbent. Because of these points the heating gas temperature
was set at 1225°F. Since the system probably won't reach thermal
equilibrium, due to the one hour cycle time, the average bed
entrance temperature will be close to the desired 1200°F regen-
eration temperature.
The heating and cooling cycles require
a source of 1200°F and 600°F gas respectively, and the oxygen
content should be as low as practical to minimize the flammability
hazard as discussed above. Note that 600°F gas is required for
cooling since this is a fixed bed operation and the solids will
be heated to the entering gas temperature. Since the sorber
-269-
-------
effluent flue gas (i.e., the S02 free flue gas) is at 600°F,
it was decided to use this gas for both the heating and cooling
cycles. The cooling cycle uses the gas directly, but because
the heating cycle requires gas at 1200°F, a recycle flue gas
heater is provided wherein the necessary heat is supplied by
burning natural gas (or sulfur-free fuel oil) and mixing the
combustion gases directly with the recycle flue gas. It is
assumed that the degree of thermal regeneration of the sorbent
during the heating and cooling cycles will be small, so that
the recycle gases can be returned directly to the main flue gas
header without exceeding the maximum S02 limit of the stack gases,
The operating cycle of the fixed bed
system is such that either heating or cooling gas is^required
at all times so that recycle flue gas blower operatesncontinuous-
ly. The flue gas heater is required only one half the:time,
however, which means it has to be designed for frequent start-
up and shutdown. Therefore, the furnace has been provided with
two burners; a small pilot burner running continuously and a
large burner operating intermittently, with each burner having
its own air blower. The pilot burner is designed to supply
enough heat to keep ths combustion chamber refractory at its
operating temperature when the preheater is not in use, i.e.,
during the cooling cycle. A vent line is provided to carry
the pilot burner effluent gases from the flue gas .heater to
the main power plant flue gas header.
4 - Insulation
All ducts and equipment in .-contact with
hot gases on one side and exposed to ambient conditions on the
other will be insulated internally to minimize heat losses and
to lower the design temperature of the materials of construction.
To meet process requirements of good insulating properties and
high abrasion resistance, a castable insulation (e.g,, Grefco,
HL-41) gunnited into place was selected. In selecting the
-270-
-------
insulation thickness required, a before-taxes payout of about
two-three years on incremental investment was used as the
basis. That is the estimated incremental investment required
for increasing insulation thickness must be paid for by in-
creased fuel savings over a two-three year period. This
criterion resulted in a six inch insulation thickness for
all insulated metal surfaces. Since the life of the plant
is greater than two-three years N5-6 years after taxes),
using increased thickness of insulation probably would be
economical but the savings would be too small to have an
appreciable effect on overall costs.
The thickness of insulation required
for the concrete sorber train itself was found to be greater
than the six inches required for the metal surfaces. Brief,
approximate estimates showed that concrete construction would
be the most economical for the sorber vessel and since concrete
must be kept below about 500°F, additional insulation was re-
quired. The concrete acts as an insulator itself so the tem-
perature drop through the wall is much higher than for a metal
wall. Calculations showed that multilayers of insulation would
be the most economical means of reducing the inner wall tempera-
ture to about 400°F for the most severe process conditions that
will be encountered. Consequently, two layers of low conductivity
insulation (e.g., Superex M) four inches thick were used on the
inside walls of the sorbers and one four inch layer of castable
refractory (HL-41) was installed on top of the inner layers to
produce a total thickness of twelve inches. The layer of castable
was needed for abrasion resistance while the better insulating
properties of the inner layers reduced the total thickness re-
quired.
(b) Mechanical
1 - Sorber Units
The initial proposed physical design of
-271-
-------
the sorber is illustrated in Figures J-105 j-106 and J-107
Further study of that design indicated that expansion due to
the thermal cycling, needed for sorption and regeneration, would
pose severe design problems. Expansion along the 40 foot dimen-
sion of the sorber panel would be approximately 6 inches, when
heated from ambient temperature up to the regeneration tempera-
ture of 1200°F, and approximately 3 inches of contraction.and
expansion would occur when cycling between the sorption and
regeneration temperatures. Similar expansion problems,---al-
though not quite as severe, would also be occuring along the
side walls of the sorber panels and in the inlet and*outlet
ducting. To alleviate these problems, modifications of the
proposed sorber configuration were undertaken.. It was:decided
that these modifications should follow the lines ofzfuznace
design where similar expansion problems are encountered.
In the revised coceptual.design the sorber
panels are supported by fixed structural members at the four
corners of each sorber panel,J-108. Fixing-the sorber panel
at the corners keeps the structural shape of the sorber intact
but forces the expansion inward. To allow for this expansion,
slip joints are provided at the center of each of the four sides
of the sorber panels. This arrangement allows the steel frame
of the panel to move with respect to itself but produces very
little movement between the frame and the sorbent, thus minimiz-
ing mechanical forces on the sorbent.
To provide space for further structural
support to the lower sorber panel frame, the inlet and outlet
ducts are placed on either side of the sorber .panels rather
than over and under the panels as in the previous design. The
new design allows the lower sorber panel to rest on an insulating
concrete base which provides both structural support and thermal
insulation. The walls of the ducts are internally insulated so
that thermal expansion of the steel ducts is negligible.
-272-
-------
Another design modification is the eli-
mination of the partitions which ran between the sorber panels
(Figure J- 105} . In the new design (Figure J- 109) SO- rich gas
enters a duct common to two sorber panels; the gas flow then
splits, going through each of the two panels for sorption. The
clean gas then enters the ducts common to the adjoining panel
where it flows into the clean gas header. The use of this design
has a second beneficial effect in that in the old design the
partition would have acted as a heat transfer surface during
the heating and cooling cycles if left uninsulated, thereby
decreasing the thermal efficiency of these cycles. Removal
of this partition and its associated insulation should also
result in a noticable decrease in sorber cost as well as decrease
the problems associated with the fabrication of the sorber sec-
tions. In the present design only the plates forming the gas
seals at the ends of the panels act as heat transfer surfaces
and need be insulated to increase the thermal efficiency of
the heating and cooling cycles.
To form the duct common to two adjoining
sorber panels a top, bottom, and side cover plate are required.
Figure J-110 shows the proposed arrangement for these cover
plates. The two channel or plate and angle sections, which
form the sorber panel, overlap one another and are connected
with slip joints at their side walls. The top plate of the
upper channel has a cut-out to allow for expansion of the lower
panel cover. The panel covers are anchored at the channel ends.
At one channel end a fixed pin is used to force the thermal ex-
pansion vectors outward parallel to the cover's edges. At the
corresponding panel end of the adjoining sorber panel, a slip
pin. is used to allow for expansion in the direction perpendicular
to the length of the cover plate and to prevent rotational move-
ment of the cover plate around the axis of the fixed pin. This
causes the remaining expansion vector to be directed along the
length of the panel cover. Since the panel covers rest on the
-273-
-------
sorber panels and overlap each other at their free ends the
duct is sealed, either admitting flow from the inlet header
or to the outlet header. Figure J-lll shows a detailed view
of the center junction of the sorber panel sides and the
dimensions of the channels and cover plates required to seal
the duct and allow for thermal expansion of the panels and
covers.
Figure J-112 shows the sorber support -
detail. Since the sorber is symmetrical, the detail'shown*
is typical of all four corner supports. The side view shows
the upper and side sorber panel frames-and covers, abutting
at the upper left corner of the sorber, and their connection
to the support beam which joins the sorber panel to the outer
steel frame. Views A-A and B-B show the panel cover connections
to the support beam and sorber panel frames. The support beam
is offset to allow for expansion of one of the panel- covers
while rigidly fixing the next panel cover and sorber panel.
Detail C (Figure J-113) shows the proposed screening arrangement
which will remain essentially the same as was discussed in
section (2) (a) 1 above; i.e., a sandwich arrangement with the
sorbent encased by the grate and screening walls.
Overall sorber panel dimensions, length,
width and bed depth, remain the same as before (Figure j-106).
The distance between sorber panels (A), the width of the sorber
panel covers (B) and the width of a sorber panel (C) along with
the sorbent bed thickness (D) are given as a function of sorption
stream time.and bed pressure drop in Table J-98. It should be
noted that these dimensions do not consider mechanical require-
ments .
The conceptual design of a fixed bed sorber
panel as described above was next evaluated for structural and
fabrication feasibility. These panels were originally designed
-274-
-------
from process considerations such as pressure drop, bed thickness,
etc., with only conceptual designs of key mechanical items such
as sliding joints for thermal expansion, and structural supports.
Having selected a base case as discussed in the process section
above, a mechanical design was undertaken in sufficient detail
to insure fabrication feasibility and to provide a sound basis
for the cost estimate. As a result of this mechanical design
it was found that the flat type sorber panel previously
required an excessive amount of structural steel and con-
sequently produced a very expensive design. A sketch of the
cross-section of the sorber train using flat panels is shown
in Figure J-114. This sketch which is drawn to a scale of 1/8"
.equals one foot, graphically illustrates the quantity of struc-
( tural support required in that the overall dimensions are 72
feet wide by 59 feet, whereas the sorber panel is 40 feet by
20 feet. Further, the panels themselves require trusses between
pairs, with tie rods holding the sides of each panel together,
and additional miscellaneous framing is required. Because of
the complicated and rather massive sorbent-containing vessels
and supporting structure needed, it was decided to seek an alter-
nate design which would offer a simpler and cheaper unit.
A cylindrical configuration for the sorber
units was designed which gives substantial savings in fabrication
costs. Since the major factor influencing the panel structural
requirements is pressure drop through the panel (i.e., 30 inches
of water is equivalent to about 150 pounds per square foot), a
cylindrical shape gives a much lower cost because of its much
higher inherent strength compared to a flat panel. This increased
strength is of particular importance for the present case because
the 1200°F design temperature reduces the allowable metal stresses
to only abouc 4000 psi for the type 321 stainless steel used. As
in the flat panel design, thermal expansion will be handled by
Providing sliding joints, which will be stellited and ground .smooth,
-275-
-------
A sketch of the cylindrical sorber design is shown in Figure J-115.
As indicated by the sketch, this new design essentially comprises
rolling the original flat panel into a cylinder, thereby greatly
reducing structural requirements. Each of the cylinders has a
mean diameter of 3 feet, a bed thickness of 7 inches., and a length
of 22 feet. The length of the sorber units through which the flue
gas flows is 20 feet; the extra two feet are provided to~allow-for
sorbent settling. Gas flow is from the outside to the inside of
the units.
Loading of the sorber units will, be done
by pneumatically conveying the sorbent to the top of the cylinders.
and gravity filling. A hopper truck equipped with a blowerccould
be used for. loading the units. When it is necessary tcr unload
the units, the>sorbent can be gravity dumped and removed by a
front-end loader. Space under the sorber buildings, is required
for heat transfer purposes, i.e., to provide for:adequate,air
circulation to avoid heat build-up in the vessel! floor, and
this space has been made large enough to permit:passage of a
front-end loader.
2 - Sorber Trains
The sorber buiMing (or train) itself
has concrete bottom and sides with a steel roof.. There are
twelve inches of insulation over the concrete secdbions. and
six inches over the steel. A sketch of the crosst section of
a sorber building is shown in Figure J-116. There are four
sorber trains required with 6 rows of 14 sorber-units in each
train to give a total of 336 which, as discussed in Section (a)
-1 above, provides 50% more surface than calculated from the
model. Thus-the superficial gas velocity through the beds is
reduced from the value of 2.7 feet per second shown in-Table 42
to 1.8 ft/sec. A transition piece is used to connect the inlet
-276-
-------
and outlet ducts with the sorber buildings, so that the entire
cross-section of the sorber is used for gas flow. A plan of
the layout of the sorber trains is shown in Figure 56.
3 - Gas Ducts
The gas ducts are sized for 60 feet per
second superficial gas velocity and a maximum pressure of 45
inches of water. Six inches of insulation on the inside are
used to reduce the metal temperature to about ambient conditions
and material of construction is carbon steel. The diameters
of the various lines are: main flue gas ducts — 40 feet; flue
gas inlet and outlet branch ducts -•- 28 feet; heating and cool-
ing gas ducts — 15 feet; and reducing gas ducts — 5 feet. To
minimize duct cost and maximize duct stiffness with minimum weight,
cellular steel floor deck is used for the duct walls with curved
ring trusses around the outside of the duct for support, stiffen-
ing, and tension stresses. The cellular floor deck is a readily
available item and combines high rigidity with light weight, and
also provides a smooth inner surface with maximum stiffness to
support the insulation against duct vibration.
4 - Valves
The large valves required for switching
gas streams posed a problem in that valves of the size needed
were not commercially available in a price range that could be
tolerated. Consequently, various vendors were contacted to see
if they could meet our process requirements with a suitable valve.
The results of this search are included as Table J-99. Grove
Valve and Regulator Company has tentatively designed and estimated,
specifically for the fixed bed case, a retracted disc valve 20
feet in diameter, after conferring with MWK concerning process
specifications. The gas ducts are 28 feet in diameter but a
valve this size does not appear feasible so the 20 feet diameter
values are used. The higher gas velocity through these valves
-277-
-------
FIGURE.
FIXED BED SCHEME- CYLINDRICAL SORBER UNITS
GENERAL ARRANGEMENT
TO STACK
700'±
FROM FD.FANS
mnn MW pm/FR STATION
PREHEATER
-------
GLAUS PLANT
PREHEWER
1000 MW PO^NER STATION
35'0 F.D. HEADER
ID. FANS
280
BY-PASS --
REFORMER
FURNACE
MR FANS
REFORMED GAS
SORBENT
TRAIN
NO.i
nor
EXPANSION
JOINT
-------
BUTTERFLY GKTES
to
^1
00
FUTURE TRNNS
PL/VM
FLUE GAS
MALVES
SORBENT
TRMN
CLEAM 6AS
_75
_50'
5' REFORMED GAS,
ELEVATION
-------
is not considered to be a problem. The initial preliminary
valve design will probably have to be refined to include water
cooling of the valve seats (to avoid the excessive thermal stresses
in the cold duct walls that would result from connecting a hot
and cold mass together^ but this modification can easily be made.
Fixed bed economics includes Grove's estimate for the large valves.
5 - Flue Gas Fans
The additional gas pressure drop caused
by the fixed bed system requires the use of flue gas booster
fans. Since the power plant itself already requires flue gas
fans, use of these existing fans as the booster fans by increas-
ing the discharge pressure was considered. The fixed-bed pres-
sure drop is 30 inches of water and another 5 inches has been
allowed for line losses, etc., so the total incremental pressure
drop due to the fixed-bed system is 35 inches of water. Subse-
quently, it was found that increasing the discharge pressure of
the existing fans would be impractical so an additional set of
fans was included as part of the alkalized alumina process.
6 - Air Preheater Ductwork
The sorption reaction requires a tempera-
ture of about 600°F so the flue gas is withdrawn from the power
plant circuit upstream of the power plant air preheater. The
layout of the alkalized alumina plant is such that the power
plant air preheater must be relocated, thus increasing the
quantity of air ducts required. The cost of these additional
ducts is included in the alkalized alumina giant.investment.
d. Comparison of Processes
A brief comparison of the major design factors and key
process features based on the final design of each process is
given in Tables 44 and 45. It should be noted that although
-279-
-------
TABLE 44
MAJOR PROCESS FACTORS - ALKALIZED ALUMINA
ITEM
FIXED BED
FLUID BED
DISPERSED PHASE
N>
00
o
I
Sorption Temp., °F
Regeneration Temp., °F
Solids Heating Method
Regenerator Type
Solids Transfer Method
Sorbent Density in Reactor, Ib/CF
Reactor Gas Velocity, fps
Flue Gas Press. Drop, in H~O
Sulfur Loading on Sorbent, %:
In
Out
% Sulfur Removal from Flue Gas
Solids Circulation Rate, Ib/hr
Number of Reactor Units
Approximate Size of Reactor Unit
Approximate Plot Size
Approximate Max. Elev. Above Grade
Sorbent Inventory, Ibs.
600
1200
Indirect
Fixed Bed
None
45
1.8
30
4.0
13.0
100
None
336
3'
-------
TABLE 45
RELATIVE COMPARISON OF FIXED, FLUID, AND DISPERSED PHASE SYSTEMS
FIXED BED
FLUID BED
1
10
CO
t
1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
12.
Attrition Losses
Integration into New Power Plant
Integration into Existing Power Plant
Unusual Special Equipment Required
Compressor Required on Glaus Feed Gas
Explosion Potential Exists
Fly Ash Accumulation on Sorbent
Operation Gas Velocity Dependent
One Train Shut Down Required at Half Gas Flow
On-Off Operation Required
By-Pass Included
Space Required-Percent of Powerhouse
Low
Good
Poor
Large Valves
Yes
Yes
Probably High
No
No
Flue Gas Heater
Yes
•v70
Medium
Good
Poor
—
NO
NO
Slight
Yes
No
clone
No
•vSO
DISPERSED BED
High
Good
Poor
No
No
Nil
Yes
Possiblv
One-Train
No
-v.50
-------
some of the ratings listed in Table 45 are subjective in nature,
it is believed that they accurately represent the relative merits
of each process. The space needed for each process and shown as
a percentage of that required by the powerhouse is calculated ,
based on the approximate plot size and the area reported for
TVA's Gallatin power plant. As such, the values reported are
approximate in nature but do serve to point out that the spacc.7. .
requirements of the alkalized alumina processes are large and.^
particularly in areas of high cost real estate, could cause -a -•&
substantial increase in capital investment.
All of the items listed in Table 45 have either been _
discussed in detail in other sections of this report and/or,. .-
are self explanatory. There are two items which could, have a „
major effect on the fixed bed design and probably should, be.«i-
specifically noted, viz., the explosion potential and fly ash
accumulation on the sorbent.
When the regeneration cycle starts, reducing gas will
come into contact with gases containing oxygen (i.e., the
sorber unit contains the flue gas used for heating), thus
creating a potential explosion hazard. Lack of experimental
data for the specific system involved prevents a precise evalua-
tion to b«3 made at this time. If it became necessary to purge
the sorber with an inert gas before and after each regeneration,
the cost would be significantly increased. Experimental data
in this area are essential before a fixed bed unit is built.
The second item noted above, i.e., fly ash accumulation
on the sorbent, is due to the fact that the extent to .which the
fixed bed will act as a filter is presently unknown. ..It is pos-
sible that the sorbent could filter out fly ash such that the -
pressure drop through the bed becomes excessive and/or the sor-
bent becomes coated with fly ash and looses its S02 removal ability.
Again, experimental data are required before a plant is built.
-282-
-------
e. Reformer, Glaus Plant, and Precipitator
The commercial scale designs MWK have developed are based
on presently available experimental data from the several contrac-
tors (e.g., AVCO, USBM, W. R. Grace) working on the alkalized
alumina process with various assumptions required where data are
lacking. Three basic designs have been developed (i.e., fixed
bed, fluid bed, and dispersed phase) with the objective of deter-
mining which of the three, based on both technical feasibility
and process economics, would be the best process for further
development. Since several sections of the processes are common
to all three schemes, these sections were not included in the
preliminary designs. Therefore, in order to obtain a complete
picture of process economics, it was necessary to estimate these
additional sections. The remaining sections which are common to
all three designs are discussed below:
(1) Regeneration Gas Plant
For the base case designs it was decided to use a
steam-reformed natural gas as the reducing gas needed to regenerate
the loaded sorbent. A flow sheet P1374-B, of the reformer plant
is shown in Figure 57. Natural gas feed is desulfurized in the
activated carbon guard drums before entering the reformed furnace
since sulfur will poison the reformer catalyst and therefore must
be removed .from the feed gas upstream of the reformer. In the
reformer furnace natural gas and steam flow through catalyst
packed tubes and react to form H2, CO, and CO^. The heat of
reaction is supplied by burning additional natural gas with air
outside the tubes. Reaction conditions to produce the desired
gas composition for the present design were determined to be
50 psia and 1600°F. The effluent flue gas from the furnace
is at 490°F and this gas is added to the power plant,,flue gas
stream downstream of the sorber, and thus passes through the
air preheater before being vented to the atmosphere.
-283-
-------
FOIt FIXED BFO CITSf
ra/e fixeo PCD ease
ens fxotv is
n*cfen&eo ~ro
13 •* 7-». 71 r+l-MT A
' ' "
iv. H. eon. fit
l1.25J3t.SO)'*'*
BTv/Hf
KSFORMER FUHHRCG
(tent»«rir)PKOCFSS purrs +t- to lie. nl
-------
The product gases from the reformer contain over 50%
water and since a "dry" gas is needed for regeneration, this water
must be condensed. The regenerator temperature has been set at
1400°F which is also the desired temperature for the reducing
gas feed. Thus the heat exchanger train has been arranged such
that the cooled, dried reformed gas is preheated by exchange with
the hot effluent from the reformer furnace. The preheating opera-
tion requires two stages of exchange in series: 1) a regeneration
gas heater where the dry reformed gas is heated from 675° to
1400°F by cooling the reformer effluent from 1600° to 1300°F,
2) a regeneration gas preheater where the dry gas is heated from
100° to 675°F by further cooling the reformer outlet gas from
1300 to 1066°F. The temperature of the wet gas stream effluent
from the regeneration gas preheater is reduced from 1066° to
350°F by generating 50 psig steam in the waste heat boiler.
Finally, the water is condensed by exchange with cooling water
and the resulting gas-liquid mixture separated in the water knock-
out drum. The dry gas, containing some 2% water, is preheated as
described above and sent to the regenerator. The dried, preheated
reformed gas is at about 37 psia which provides an ample allowance
for friction losses and control valves since the maximum pressure
in the sorber is 30 psia.
The flow rates shown on P1374-B are for th« fluid bed
and dispersed phase cases. The reducing gas quantity arequired
for the fixed bed cas is shown on the outlet stream oniy but if
intermediate values are desired they can be prorated using the
ratio of the two product streams shown.
To date, the maximum allowable water content in
the reducing gas to the regenerator has not been determined and
the present designs are based on "dry" gas (i.e., about 2% water).
If later data show that higher water concentrations can be tolerated,
the.p the reformer can be redesigned to produce the regenerator feed
gas directly thus eliminating the cooling and reheating operations.
-------
An evaluation of this alternative reformer design will be made
in the phase 3 work. .,
(2) Glaus Gas Pre-Treatment
The regenerator gas effluent in all three of the
present process designs contains about 18% H2S, and greater than
40% H20. In addition, equilibrium calculations have shown that
there will probably be some elemental sulfur present as vapor.
It was originally planned to condense the water from the regen-
erator exit gas, thereby increasing the H2S concentration to
about 30% and thus reduce the size of the Glaus plant. Further
study showed, however, that because of the sulfur vapor present
in the gas, cooling and condensing the water would also condense
the. sulfur. Further, because of the temperature range and sulfur
concentrations involved, the equipment required became complicated
and expensive; i.e., the sulfur temperature-viscosity relationship
necessitated several steps involving quench towers, scraped sur-
face exchangers (or equivalent type equipment), the cost of which
would more than offset the incremental cost of the Claud plant due
to a dilute H_S feed. Therefore, for the fluid bed and dispersed
phase cases, no pretreatment to the regenerator off gas is used
other than a waste heat boiler for heat economy. The feed to
the Glaus reactor should be about 400°F and since the regenerator
effluent is at 1400°F, the gas is cooled to the desired temperature
by generating 50 psig steam. A credit will be taken for any
steam in excess of that needed for the alkalized alumina plant.
The fixed bed case does require a pretreatment section,
however, because the fixed bed regenerator is at atmospheric pres-
sure and the off-gas must be compressed since the pressure drop
through the Glaus plant is about 10 psi. Limitations on the dis-
charge temperatures of the compressor results in a maximum inlet
-286-
-------
gas temperature of about 250°F for the pressure range involved,
and cooling the gas to this temperature condenses the sulfur
vapor present. A process flow scheme, P1375-B, for this fixed
bed regenerator off gas treatment and compression is shown in
Figure 58.
A brief process description for the fixed-bed
Claus gas pretreatment section is as follows: referring to
Figure 58, the hot gaseous effluent from the regenerator passes
through a waste heat boiler, C-l, where it is cooled to 350°F
by generating 65 psia steaitv. From C-l, the gas flows through
a sulfur condenser, C-2, where most of the sulfur vapor is
condensed to liquid sulfur at 250°F by exchange with the boiler
feed water for C-l. At 250°F liquid sulfur will flow readily
and is pumped to the main sulfur storage pit. The effluent
gas from C-2 is sent to the water-cooled, scraped surface C-3
sulfur crystallizer (or desublimer) and cooled to 200°F thereby
causing the remaining sulfur vapor to be deposited as sulfur
crystals on the cold tubes. The sulfur crystals are continuously
scraped from the tube walls and fall by gravity into the steam-
heated bottom section of C-3 where they are melted and heated
to 250°F and then pumped along with the condensed sulfur from
C-2 to the sulfur storage pit.
The effluent gas from C-3 passes through a guard
filter, L-l, which removes any entrained sulfur particles and
is then compressed to 30 psia in the Claus gas compressor, J-5.
The J-5 discharge stream is cooled to 100°F in the C-4 water
condenser wherein most of the water is condensed, thereby in-
creasing the H2S concentration in the gas stream to about 27%.
The liquid water and gas are separated in F-l knockout drum
and the dried compressed gas is sent to the Claus unit for
conversion to elemental sulfur.
-287-
-------
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-------
It should be noted that since it is necessary to
remove the sulfur, the water can be easily condensed thus up-
grading the Glaus feed gas. The fixed bed case, therefore, has
a more concentrated feed gas for the Glaus plant than the fluid
bed or dilute phase processes.
(3) Glaus Plant and Incinerator
As discussed above, the Glaus plant is identical
for the dispersed phase and fluid bed cases but is somewhat
different for the fixed bed process. The cost of the gas
pretreatment and compression is included as part of the Glaus
plant investment for the fixed bed scheme.
The Glaus plant itself is a typical installation
and comprises a burner, reactors, heat exchangers, and sulfur
separators. One-third of the H2S in the feed gas is converted
to S02 by direct combustion in the burner, the effluent gases
cooled to 400°F by generating 50 psig steam, and the 400°F gas
stream then passes through the reactor where sulfur is formed:
2H2S + S02 -> 3S + 2H20
Two, stages of reactor are used to increase the efficiency of
sulfur recovery. The exit gases are cooled to about 260°F in
the. heat exchangers and then flow into the separator from which
molten sulfur is pumped to storage.
About 90% of the sulfur in the feed gas is recovered
as molten sulfur so the exit gases from the final separator con-
tain about 10% of the entering sulfur as H2S and S02. If released
directly to the atmosphere this gas would be a source of air
pollution, and it may have to be recycled to the flue gas sorber
after the H2S has been converted to S02 in the Glaus incinerator. The
total volume of gas recycled is only 2-3% of the total flue gas
feed so it has been assumed that no adjustments to the present
reactor design are required because of the recycle gas. It has
-289-
-------
been also assumed that the additional sulfur due to the recycle
stream can be sorbed by letting the sorbent loading be increased
by 10%, by increasing the sorbent circulating rate, or some com-
bination of these two procedures. The additional sulfur pickup
by the sorbent owing to the recycle stream would cause the re-
ducing gas consumption in the regenerator to be increased by
about 10%. Consequently the presently assumed reducing gas
utilization of 80% would be increased to 88% for the same size
reformer, or alternatively, a larger reformer would be needed.
However, it has not been established that recycle is^required,
and comparisons would not be affected since all three'processes
would suffer essentially the same penalty, and therefore, because
the effect of recycle on process design and economics is not
significant in any case, it has been neglected in the present
evaluation.
Since HpS will probably not be sorbed by the alkalized
alumina, it is necessary to convert the H-S to SC>2. This is done
by heating the Glaus effluent gas to about 1200°F at which tempera-
ture the H«S reacts with oxygen to form SO,. The gases are heated
to 1200°F in an incinerator by combustion of natural gas and air,
and the excess combustion air supplies the oxygen needed to convert
the H2S to SO2.
(4) Electrostatic Freeipitator
In all three of the present designs a larger electro-
static precipitator is required for the power plant, and the in-
cremental cost of this precipitator is charged to the alkalized
alumina process. In the fixed bed and the fluid bed cases, it
has been assumed that the precipitator is needed upstream 'of the
reactor to prevent excessive fly ash build-up since each of these
systems will act as a filter. The extent of fly ash build-up-is
presently unknown but it seems highly probable that feeding the
untreated flue gas directly into either system would cause blinding
-290-
-------
of the sorbent, hence excessively high pressure drops, in the
fixed bed case, and very high internal circulating loads of fly
•
ash, causing high pressure drops and overloading of the reactor
cyclones, in the fluid bed scheme. Thus, a relocation of the
precipitator to a position upstream of the reactor and therefore
upstream of the air preheater, results in an increased flue gas
volume of about 40% because of its higher temperature, i.e.,
about 600°F instead of about 300°F. With a larger gas volume to
be treated the precipitator size and cost are correspondingly
increased. This increased cost is directly caused by and thus
charged to the alkalized alumina process.
The fly ash does not affect the sorbent in the dis-
persed phase design, based on pilot plant results, and therefore,
the electrostatic precipitator does not have to be relocated. Re-
moval of the SO2 from the flue gas upstream of the precipitator
does lower the collection efficiency, however, and the precipi-
tator size and cost are again increased. Values of precipitator
efficiency as a function of the sulfur content in the fuel are
reported in the literature^and an extrapolation of the data indi-
cates a decrease of about 40% for the present process conditions.
Thus the incremental cost of the electrostatic precipitator is
about the same for all three designs and this cost is included in
the total investment for each system.
(5) Power Plant Air Preheater
Standard power plant practice is to install an
air preheater in the exit flue gas stream whereby the entering
combustion air is heated by exchange with the effluent flue gas
whose temperature is reduced from about 650°F to about 300°F.
Reese, J. T., and Greco, J., "Experience with Electrostatic Fly-Ash
Collection Equipment Serving Steam-Electric Generating Plants",
J. AIR POLL. CONT. ASSOC., 18 (8_) , 524 (Aug. 1968).
-291-
-------
As a result of installing an alkalized alumina plant both the
volume and temperature of the effluent flue gas are increased.
That is, the effluent gas from the solids heater and the solids
transport air are added to the flue gas and in addition, the
hot regenerated solids are cooled by the flue gas. Initially,
no incremental cost was added to the air preheater because of
the higher temperature and flow rates, based on phone conversa-
tions with a vendor who stated that this equipment is normally
designed for increases up to 20%. However, subsequent written
inquires did in fact, result in cost increases to the air pre-
heaters. This additional cost is a direct result of and there-
fore charged to the alkalized alumina plant.
^• Final Sorber Design
The original base case designs for both the fluid bed
and dispersed phase schemes included a solids cooler to reduce
the sorbent temperature from the regeneration temperature of
about 1400°F to the sorption temperature of about 600°P. As
an alternative design, it was decided to eliminate the solids
cooler and return the regenerated sorbent directly back to the
reactor without prior cooling. Subsequently, a modified process
design was made for both the fluid bed and dispersed phase schemes
and ultimately, because of favorable economics, the design with-
out the solids cooler was selected as the base case.
Elimination of the solids cooler resulted in 1400°F
solids.being fed directly into the reactor and required that
an evaluation of the effect, if any, on reactor design be made.
In the fluid bed case, no changes to the reactor were necessary
because of the rapid heat transfer obtained in a fluid bed; i.e.,
the hot solids fed into the reactor are cooled to the equilibrium
bed temperature almost instantaneously and no extra height for
heat transfer is needed. The dispersed phase case reactor did
require modification, however, since very few sorption data were
available at high temperatures and to insure an adequate reactor
-292-
-------
design it was assumed that sorption would not occur until the
solids had been cooled to the equilibrium temperature. It was
further assumed that the sorption rates would not change because
of the new equilibrium temperatures which are increased from the
original value of 600°P to 634°F and 658°F for the fluid bed and
dispersed phase cases respectively. The higher temperatures both
cases are a result of cooling the sorbent with the flue gas which
enters the system at 600°F, and the higher dispersed phase equili-
brium temperature is due to the difference in solids throughput.
In addition to the increased height of the dispersed
phase reactor other changes resulting from eliminating the solids
cooler are a smaller air compressor, and deletion of the after-
cooler and water knockout drum. The heat balance around the
solids heater is also affected since the solids enter at a higher
temperature and preheated air is no longer available. The effect
of these reviaions was a net reduction in investment and operating
costs and resulted in the designs without the solids cooler being
selected for the base cases. Investment and operating costs are
discussed in following sections C and D, respectively. The methods
used in calculating the new dispersed phase reactor height are dis-
cussed below. Also presented is a sample calculation for the reac-
tor height.
The increase in reactor height is determined by the time
it takes for the 600°F flue gas and 1400°F sorbent particles to
reach the equilibrium temperature of 658°F. Various methods were
used to calculate the equilibrium time as follows:
• Bird, R. B., Stewart, W. E., Lightfoot, E. N.,
TRANSPORT PHENOMENA, Wiley & Sons, Inc., New
York, 1960, p. 357-360.
This method involves the use of a plot indicating the
variation of fluid temperature with time when a homo-
geneous sphere of solid material, initially at a uni- -
form temperature, is suddenly immersed in a well-stirred
-293-
-------
fluid, at another temperature in an insulated tank.
The temperature distribution in a sphere resulting
from unsteady-state heat conduction can be predicted
since the necessary mathematical equations are solved
and the results presented in graphical form.• This
procedure considers the sphere surface to be suddenly
raised to the fluid temperature at initial time zero
and predicts the temperature distribution then obtained
with increasing time.
• McAdams, W. H., HEAT TRANSMISSION, 3rd ed.,
McGraw-Hill Book Co., Inc., New York, 1954, p. 40
Graphical solutions of the equations for the transient
heating or cooling of a sphere are presented. The
relations between temperature and time for various points
along a radius can be obtained from the Gurney-Lurie
chart for spheres. These relations involve the thermal
conductivity, density, and specific heat of the body, its
shape and size, and the external conditions, including
the temperature of the surroundings and the coefficient
of heat transfer between surroundings and the surface.
• Particle Heat Balance
The heat evolved in decreasing the temperature of the
particle must be equal to that determined from an overall
heat requirement, namely the product of three parameters:
the fluid-particle heat transfer coefficient, the area of
the particle, and the log-mean temperature driving force.
In eliminating the solids cooler in the dispersed phase
case, the optimum condition would be for the solids to cool from
their entering temperature of 1355°F to the equilibrium temperature
of 658°F in a relatively short time so that any additional required
-294-
-------
sorber height would be kept at a minimum. Assuming that the
lighter particles are not recycled, sufficient height must be
provided in the sorber such that they can be cooled down in
one pass since the heavy particles will be recycled. To determine
which particle size would require the greatest height, and thus
set the design, calculations were made for several different
size particles, using the size distribution obtained from an
attrition run reported in the Bruceton Quarterly of 12-31-67
(see Figure 5, this report). A sample calculation of the time
required for equilibrium to be reached in the dispersed phase
sorber for a 1410 micron particle diameter is included as
Appendix K. A summary of the additional sorber height required
for the attainment of the proper equilibrium temperature, for
various particle sizes, is included as Table 46.
Values of thermal conductivity for alkalized alumina
were not available, so alumina (Al-Cu) values were used instead.
Different values of thermal conductivities are reported in the
literature for A1203, however, so two values (0.392 and 0.087
Btu/hr-ft) were used in these calculations. For the 1410 micron
diameter particle, which, as shown in Table 46, sets the addition-
al sorber height required, the results came out quite similar,
namely 36 and 38 feet using the Gurney-Lurie chart for spheres.
These heights are in good agreement with the value of 29 feet
as calculated by a partial heat balance. Both these methods
yielded a considerably greater height than that predicted using
the method of Bird, Stewart and Lightfoot but this is probably
due to the fact that the latter approach assumes the surface of
the sphere to be instantaneously adjusted to the fluid temperature.
As a result of these calculations, it was decided that the dis-
persed phase sorber should be 35 feet higher for the case of
operation without a solids cooler. Therefore, the previously
calculated reaction height of 80 feet was increased by 35 feet
and another 10 feet was added to allow for gas distribution baffles
such that a "reaction zone" height of 125 feet was obtained. A
sketch of the final design of the dispersed phase sorber is shown
in Figure G-93.
-295-
-------
TABLE It
HEIGHT REQUIRED FOR PARTICLE-GAS HEAT TRANSFER
DISPERSED PHASE REACTOR - 1000MW POWER PLANT
SOLIDS: 1330 » 65B°F
GAS: 600 + 658°F
PARTICLE SIZE ,=L1P VELOCITY
(MESH)
8
10
-10+14
60
''gas "set.vel.)
(MICRONS) (FT) (ft/sec)
2362 0.00776
1651 0.00542
1410_ 0.00462,,,
246 0.00080
-9.5
-1
43
+ 22.1
FLUID TO
PARTICLE
HEAT TRANSFER
COEFFICIENT*
(Btu.'hr/ft2F)
22.2
8.2
15.6
102.3
SOLID
THERMAL
CONDUCTIVITY CALCULATED EQUILIBRIUM TIMES
Ks
(Btu/hr/ft F)
0.392**
O.OB7t
0.392
0.087
0.392
0.087
0.392
0.087
Itt
0.9
3.9
I
0.42
1.9
0.31
1.4
0.009
0.04
(SECONDS)
fttt
14
18
30
26
12
12.5
0.32
0.33
3
11.4
--
21.6
9.7
0.29
REQUIRED
1
9.
37.
0.4
1.9
1.
4.
0.2
0.9
EQUILIBRIUM HEIGHT
(FT)
~F~^ 3
133 108
171
30 22
26
36 29
38
7 6.5
7.5
• h calculated - Treybal, R. E., Mass Transfer Operations, McGraw-Hill Book Co., Inc., New York, 1955, p. 54 - Curve 4.
** ks A1,O, from Lange, N. A., Handbook of Chemistry, revised 10th ed., McGraw-Hill Book Co., Inc., New York, 1963.
t ks A12O3 from Thermophysical Properties Research Center, Data Book, Purdue University, Lafayette, Indiana, 1964.
ft Bird, R. B., Stewart, W. E., and Lightfoot, E. N., TRANSPORT PHENOMENA, Wiley * Sons, Inc., New York, 1960, p. 360, Fig. 11.1-4.
ttt McAdams, W. H., HEAT TRANSMISSION, 3rd ed., McGraw-Hill Book Co., Inc., New York, 1954, p. 40-Fig. 3-7.
1 - Bird, Stewart, Lightfoot Method
2 - Gurney-Lurie, Chart for Spheres
3 - Particle Heat Balance
-296-
-------
C. INVESTMENT
Estimates have been completed for the entire battery limits
plant (i.e., excluding offsites) for each of the three processes
and the investments, including the initial charge of sorbent,
are: dispersed phase - $30,720,000; fluid bed - $35,900,000;
and fixed bed - $40,210,000. A breakdown of the investments
into major items is shown in Table 47 including a tabulation
of the material and subcontract costs of the major items.
Accuracy of the investments, based on the current flow sheets,
is in the range of -5% to +25%.
The cost of the sorption-regeneration section in each case
includes that equipment as shown on the individual flow sheets
for each design while the regeneration gas plant investment
shown in Table 47 is the total installed cost based on the flow
sheet described in section B.2.e.(l) (Figure 57).
The total installed cost of the Glaus plant pretreatment
section is $810,000 for the fixed bed design, based on the
discussion and flow sheet described in section B.2.e.(2)
(Figure 58), and is $100,000 for the fluid bed and dispersed
phase designs. The only equipment required for the latter
two designs is a waste heat boiler.
The total erected cost of the Glaus plant and incinerator,
(Section B.2.e.(3)) including the waste heat boiler, sulfur
pumps, etc., is $1,500,000 for the fixed bed design and $1,400,000
for the fluid bed and dispersed phase cases. The slightly higher
cost for the fixed bed unit is due to the increased sulfur re-
moval obtained in the fixed bed sorbers, hence the Glaus plant
is correspondingly larger. The Glaus feed gas from the fixed
bed design has a higher concentration of H2S than the fluid bed
or dispersed phase and this partially offsets the cost due to
the larger size but the net result is an increase in cost for
-297-
-------
TABLE 47
APPROXIMATE DISTRIBUTION OF PROCESS INVESTMENT-BATTERY LIMITS PLANT
DOLLARS
vo
ITEM
1. Sorption-Regeneration*
Converters
Fans & Compressors
Large Valves (Gas Duct)
Ducts & Piping
Other
Sub-total (1)
2. Regeneration Gas Plant
3. Glaus Gas Pretreatment
4. Glaus Plant, including incinerator
5. A Cost Electrostatic Precipitator
6. A Cost Air Preheater
Sub-total(1*5)
7. Sorbent Inventory (@ 25<=/lb)
Total
DISPERSED PHASE
12,200,000
1,300,000
8,300,000
300,000
22,100,000
3,300,000
100,000
1,400,000
2,870,000
400,000
30,170,000
550,000
FLUID BED
13,600,000
2,500,000
10,900,000
630,000
27,630,000
3,300,000
100,000
1,400,000
2,870,000
200,000
35,500,000
400,000
FIXED BED
11,600,000
2,800,000
6,200,000
9,900,000
480,000
30,980,000
**3,300,000
810,000
1,500,000
2,870,000
260,000
39,720,000
490,000
30,720,000
35,900,000 40,210,000
*
**
See next page for material cost breakdown.
Fixed bed requires about 10% higher capacity and lower temperature than other
two cases but investment shown does not include allowance for this difference,
Added cost for higher capacity would be about $250,000.
-------
TABLE 47 (CON'T)
ITEM
MATERIAL & SUBCONTRACT COST, DOLLARS
DISPERSED PHASE FLUID BED FIXED BED
I
to
«3
VO
I
Vessels
Reactors
Reactor Cyclones
Gas-Solids Separators
Solids Heaters
Regener ator s
Sorbent Storage Hopper
Recycle Flue Gas Heater
Fans & Compressors W/Motors
Air Compressors
I.D. Fans (Increment)
'Recycle Gas Blowers
Air Blowers
Glaus Gas Compressors
Piping
Flue Gas Ducts & Expansion Joints
Air Ducts (Incremental)
Other Piping & Valves
Duct Valves
Total Major Material
1,077,200
2,640,000
42,000
187,400
557,400
19,000
395,000
228,000
1,200,000
800,000
900,000
8,046,000
3,693,000
1,730,000
41,600
149,000|
189,800
19,000
265,000
988,000
>4,200,000
125,000
968,000
75,000
15,000
300,000
1,800,000 2,450,000
800,000 800,000
1,340,000 300,000
2,750,000
11,015,400 11,983,000
-------
the fixed bed Glaus plant.
All three alkalized alumina processes result in a larger
electrostatic precipitator being required by the power plant.
The reasons for the increased size were cited earlier (section
B.2.C.(4)) and, very briefly are: increased gas volume because
of a higher temperature for the fixed and fluid bed designs,-:and
lowered efficiency in the dispersed phase case. The incremental
size, hence cost, of the precipitator is; directly caused .by, -and
therefore charged to, the alkalized alumina process. In all three
cases, the total incremental cost including additional piping,
ductwork foundations, etc. is estimated to be $2,670,000. The
basic equipment cost was obtained from a vendor and the total
installed cost was estimated by Kellogg.
The cost of the power plant air preheater is increased by
the alkalized alumina plant owing to the higher temperature
and flow rate of the exit gas. The incremental cost of the
air preheater is included as part of the alkalized alumina
plant investment. The material cost of this item for each
case was obtained from a vendor and the total installed cost
was estimated by Kellogg to be $400,000, $200,000, and $260,000
for the dispersed phase, fluid bed, and fixed bed designs re-
spectively.
It should be noted that a major item in all three designs
is ducts and piping, comprising about 70-80% of the reactor
cost in all three cases. Thus it is apparent that total invest-
ment obtained through the use of normal scale-up factors and
cost of major equipment would likely be substantially lower
than the actual cost because the full effect of the very costly
ductwork would not be obtained. If it is desired to determine
total plant investment from the major equipment cost, it is
essential that the large ductwork be included as a major equip-
ment if reasonably realistic results are to be obtained.
-300-
-------
The approximate quantity of large ductwork and piping
required for each process was obtained from the rough layout
prepared for each design. Since the ductwork and piping comprise
such a large percentage of the total investment (i.e., ^30%), it
is essential to have at least an approximate plant layout from
which the amount of ductwork can be estimated. Without a plant
layout the quantity, hence cost, of the large ductwork can be
substantially in error which in turn would result in a plant
investment that is probably too low.
The major equipment items were designed in sufficient detail
to insure process and mechanical feasibility, and cost estimates
for the individual items were obtained from outside vendors when
ever practical. When vendor quotes could not be obtained, esti-
mates were made by MWK. In the fixed bed case, the large valves
were designed and estimated specifically for this process by the
Grove Valve & Regulator Company, after approximate specifications
were jointly developed with MWK. The present status is a prelimi-
nary design of a retracted disc valve which appears to be suit-
able for the process conditions involved.
-301-
-------
D. ECONOMICS
1. Base Case
a. Gross Costs
Based on the complete plant investments shown in
Section C and a load factor of 60% (5250 hours/year), process
economics have been calculated for each case and are summarized-
in Table 48. A tabulation of the individual items for the
dispersed phase, fluid bed and fixed bed designs is shown in-
Tables 49, 50, and 51 respectively. It should be noted .that
the costs shown in Tables 48 through 51 do not include sorbeat
attrition losses or credit for by-product sulfur.
The values used for the individual items shown-in* ,
Tables 49, 50, and 51, are essentially those listed in. a
report obtained from NAPCA (1). The NAPCA recommended guide
lines agree quite closely with MWK recommendations tiezived from
a study of various sources (2 thru 7).
1. NAPCA, "General Guide-Lines for the Process .
Economic Evaluation of the Nine-Area Surveys -
for SO2 Removal".
2. Office of Saline Water, "A Standardized
Procedure for Estimating Costs of Saline
Water Conversion," March 1956.
. 3. The M. W. Kellogg Company, Report RED-68-1173
to the office of Coal Research, Contract-No.
14-01-0001-380, September 1, 1968, p. 123.
4. The M. W. Kellogg Company, Report Rt)-62-!952
to the U. S. Atomic Energy Commission, -
Contract No. AT (30-1)-3009 (NYO 10,301),
November 30, 1962, p. 167.
5. Aries, R. S. and R. D. Newton, "Chemical -
Engineering Cost Estimation," McGraw Hill ;
Book Company, Inc., New York, 1955, pp. 162-
182. **
6. Peters, M. S. "Plant Design and Economics
for Chemical Engineers," McGraw-Hill Book
Company, Inc., New York, 1968, pp. 108-9.
113. '
-302-
-------
TABLE"48
ANNUAL COSTS - BASE CASE
(EXCLUDING ATTRITION LOSSES AND SULFUR CREDIT)
ITEM
DISPERSED PHASE
DOLLARS/yr
FLUID BED
FIXED BED
I
u>
o
OJ
I
1. Capital Charges
2. Labor, Maintenance, and
Overhead
3. Utilities
TOTAL
3,993,600
2,889,300
991,100
7,874,000
4,667,000
3,396,200
1,271,500
9,334,700
5,227,300
3,767,900
1,433,300
10,428,500
Unit Costs
mills/kwh
C/MM Btu
/ton coal
1.50
17.42
4.53
1.78
20.67
5.38
1.99
23.10
6.01
NOTE
: Coal used in design contains 13,000 Btu/# and costs 25C/MM Btu ($6.50/ton).
-------
TABLE 49
DISPERSED PHASE - PROCESS ECONOMICS
DIRECT COST $/YEAR
1. Operating Labor $-122,600
(4 men/shift @ $3.50/hr)
2. Supervision - 15% of labor -18,400
3. Maintenance 8 5% °F FCI 1,536,000
4. Supplies - 15% of (3) • 230,400
5. Utilities: a. Cooling Water @ 5<:/M gal. -79,100
b.- Natural Gas @ 27C/MM Btu :92i;000
c. 'Power @ 0.6
-------
TABLE 50
FLUID BED - PROCESS ECONOMICS
DIRECT COST $/YEAR
1. Operating Labor
(5 men/shift @ $3.50/hr) 153,300
2. Supervision - 15% of labor 23,000
3, Maintenance @ 15% of FCI 1,795,000
4. .Supplies - 15% of (3) 269,300
5. Utilities: a. Cooling Water @ 5C/M gal. 79,100
b. Natural Gas @ 27C/MM Btu 808,700
c. Power <§ 0.6<:/kwh 686,200
6. Heat (credit) § 25C/MM Btu (201,000)
7. Steam (credit) <§ 25C/M Ibs (101,500)
8. Total Direct Cost $3,512,100
INDIRECT COST
9. Payroll Overhead - 20% of (1 + 2) 35,300
10. Plant Overhead - 50% of (1+2+3+4+5) 1,120,300
11. Total Indirect Cost 1,155,600
FIXED COST
12. Capital Charges @ 13%°F FCI 4,667,000
(Depreciation, insurance, taxes,
cost of capital, etc.)
13. TOTAL OPERATING COST (8+11+12) ....... 9,334,700
UNIT COST
14. a. $/ton coal 5.38
b. C/MM Btu in coal feed 20.67
c. Mills/kwh 1.78
NOTES: 1. No credit taken for by-product sulfur which is
produced at rate of 46,870 short tons/yr.
2. Power plant operates at full capacity 60% of time
or 5250 hours/year.
3. Costs dp_ not include attrition loss.
FCI - Fixed Capital Investment =» $35,900,000
-305-
-------
TABLE 51
FIXED BED - PROCESS ECONOMICS
DIRECT COST $/YEAR
1. Operating Labor
(5 men/shift e $3.50/hr) 153,300
2. Supervision - 15% of labor 23,000
3. Maintenance @ 5% of FCI 2,010,500
4. Supplies - 15% of (3) • -301,600
5. Utilities: a. Cooling Water @ 5C/M gal. 138,200
b. Natural Gas <§ 27C/MM Btu 876,400
c. Power @ 0.6<=/kwh 698,600
6. Heat (credit) $ 25C/MM Btu (169,500)
7. Steam (credit) @ 25C/M Ibs (110,400)
8, Total Direct Cost 3,921,700
INDIRECT COST
9. Payroll Overhead - 20% of (1 + 2) 35,300
10. Plant Overhead - 50% of (1 + 2 + 3 + 4 + 5) . 1,244,200
11. Total Indirect Cost 1,279,500
FIXED COST
12. Capital Charges @ 13% °F FCI 5,227,300
(Depreciation, insurance, taxes,
cost of capital, etc)
13. TOTAL OPERATING COST (8 + 11 + 12) J.0,428,500
UNIT COST
14. a. $/ton coal 6.01
b. $/MM Btu in coal feed 23.10
c. Mills/kwh 1.99
Notes: 1. No credit taken for by-product sulfur, which is produced
at rate of 50,370 short tons/yr.
2. Power plant operates at full icapacity 60% of time
or 5250 hrs/yr.
3. Costs do not include attrition loss.
FCI - Fixed Cost Investment = $40,210,000
-306-
-------
7. Letter from C. W. Carry, Los Angeles County
Sanitation Districts to A. N. Massi,
Federal Water Pollution Control Administration,
October 31, 1968.
; Referring to Tables 49, 50, and 51, the operating
labor includes two men per shift for the reformer and Glaus
plant associated with each process. The number of operators
required for the reformer and Glaus plant was derived from
Kellogg's experience with reformers for ammonia plants and
is believed to be a realistic estimate. However, it should
be noted that the total annual cost of the operating labor
is small compared to the overall cost of the complete plant
so even an error of 100% in the estimate of the number of
operators required will not have a significant effect on
process economics. It is assumed that any additional labor
required for start-up will be available from the power plant
operating staff. Supervision, maintenance, supplies, over-
head, and capital 'Charges are calculated as indicated in
Tables 49, 50, and 51.
The base case economics include an annual maintenance
cost of 5% of fixed capital investment (FCI). It should be
noted that calculating maintenance as a percentage of FCI is
a standard procedure by many companies, especially those which
use contract maintenance. The annual maintenance costs typical-
ly vary with plant life, being higher at the beginning, decreasing,
and then increasing. Two examples of maintenance costs based
on actual plant experience are:
1) The average maintenance charges for the first 10
years of a low-overhead chemical plant are about 4.5% of FCI,
ranging from a high of about 8% to a low of 3.0%,
2) The average maintenance costs of a complete refinery
are about 4% of FCI. It should be noted that these values are
-307-
-------
for well established^ proven processes as opposed to a"new
design such as the alkalized alumina. Since the large-sized
ductwork comprises a significant portion of the alkalized
alumina plant investment, possibly lower maintenance costs
could be used. Because of the different operating condi-
tions (i.e., higher temperature, higher pressure, longer -
runs), of the alkalized alumina plant ductwork compared-to
that normally encountered in a power plant, howeverj:it is
felt that the maintenance costs will be higher than normal.
For example, the maintenance requirements for the larger
expansion joints in the large diameter ducts are not well
defined; painting of the ducts will also be required on
some regular basis and owing to the quantities^involved,
this could be a major item.
Another point to consider is the local conditions'.-
at any given'• site; i.e., contract maintenance may not be *
available, or if.internal maintenance is used, certain-crafts
may be required that are not available from the power plant
maintenance force. Thus, depending on the local laborBunion
situation, it might be necessary to carry several different
crafts on the payroll even though they are not needed-full
time.
To summarize, the base case economics includedannual
maintenance costs of 5% of FCI as compared to 4 to 4-1/2%
for well established processes and because of the various
qualifying factors discussed above, it is felt that-the 5%
charge is a reasonable value. To account for other-possibili-
ties, however, the effect of different maintenance charges
on process economics will be studied on part of the phase 3
evaluation.
The utilities shown are for a battery limits plant,
including heat and steam credits. It is assumed that?all
flue gas streams (e.g., from reformer, Claus incinerator)
-308-
-------
will be mixed with the main flue gas stream at locations
upstream of the power plant air preheater. Using an exit
flue gas temperature of 300°F (i.e., downstream of the air
preheater) the heat available from the various flue gas
streams above 300°F has been calculated and credited to
the alkalized alumina process. Since the total volume of
the gas flowing through the air preheater is increased by
adding these additional streams, an incremental cost has
been added to allow for the increase in the size of the
air preheater.
Credit has also been taken for excess steam
generated by the alkalized alumina process. Most of the
steam can be used internally (e.g., steam-reformer, steam
drive on the Glaus plant air blower) and the excess can
be used either in the power plant circuit or in the alkalized
alumina process itself if steam drives are used on the 'booster
fans. In either case, a credit can be taken for the excess
steam and a value equal to the coal cost has been assumed.
Possibly, a higher steam value could be used but the total
steam credit is small and does not significantly affect pro-
cess economics, and therefore, the value used is considered
adequate for the present evaluation.
The capital charge rate used in the base case
economics is 13% nominally based on NAPCA guide lines com-
prising 10% straight line depreciation and 3% for taxes and
insurance. Since the total capital charge rate, not the in-
dividual items, is the important factor, and since this rate
usually falls between 13 and 16% of original investment for
power plants, the lower value was selected because it agreed
with the NAPCA figure. It is recognized that using a 10 year
straight line depreciation and adding appropriate values for
taxes, insurance, and cost of money would result in a higher
-309-
-------
capital charge rate. In fact, TVA has presented an example1
wherein a 35 year plant life was used and the annual charges
- levelized basis were calculated to be 14.5%. Thus the value
of 13% may be low in view of the fact that the alkalized alumina
plant life would be less than the 35 years expected for the power
plant. However, the lower value was used in keeping with the
general rationale of taking an optimistic approach .whenever
possible.
As was also pointed out by TVA2installing, a sulfur
removal process as part of a power plant leads to complications
because of the differences in the two industries involved; i^e.,
the regulated power company or the free enterprise- chemical
company. Three different investment approaches for air: pol-
lution control facilities are cited by TVA, viz: 1} all-power
company investment; 2) all chemical company investment; 3) mixed
investment venture. Each of these three methods including vari-
ations would require a different method of calculating costs.
Since a detailed analysis of power plant costing methods was
beyond the scope of this study, the simplified approach of
using an average capital charge rate based on original- invest-
ment was selected. It is recognized that the 13% rate selected
is probably low in view of the above discussion, but in the
present case the plant costs are so high that adding additional
costs owing to a higher capital charge rate seemed .-superfluous.
As previously noted, the economics presented herein,
are based on a load factor of 60% (5250 hours/year). The ration-
ale used in arriving at this value is given below..
Normally, as large-size, more efficient plants come
on stream they are base loaded and thus run at a higher stream
efficiency than older, less efficient units. However, since
the trend in power plants appears to be toward larger units,
1 SULFUR OXIDE REMOVAL from power plant stack gas
Prepared by Tennessee Valley Authority for Second Semiannual
Control Process Contractors' Meeting, National Air Pollution
Control Administration, Cincinnati, Ohio, June 11-13, 1969, p. 18
2 Ibid, pp. 29-30.
-310-
-------
and with the installed generating capacity expected to double
about every 10 years, it would seem that at some future time
most of the power plants .will be large; units. Therefore, as
long as there is a large difference between peak and off-peak
power demands, and if most of the units are large-sized, then
some of the large-sized plants will have to be shut down and
their load factor will not be as high as if only a few large
plants existed. It should also be noted that' the net heat
rates for large plants has been leveling off over the past
few years (Federal Power Commission, National Power Survey-
Part I", Washington, D. C., 1964) and thus the difference in
thermal efficiency between large plants built after about
1960 is relatively small. Thus, in the future the load factop
of the older plants will approach that of the newer ones and
both will approach that of the entire system.
TVA-"- have presented a suggested time-load factor
ranging from 80% for the first 10 years down to 17% for the
last 15 years. The overall average for a 35 year power plant
line is 44% based on the TVA numbers, with the first 20 years
averaging about 64%. The life of the plant therefore is an
important factor in setting the load factor.
Information was also obtained from Consolidated
Edison Co., NYC, and Commonwealth Edison, Chicago, and these
sources suggested that a 60% load factor was too high and, in
fact, a value of about 55% would be better. This latter figure
also agrees with the results 'of a survey made for NAPCA by
Battelle Memorial .Institute. However, it was subsequently
decided to use the 60% value to calculate the base case costs.
Finally, it is very unlikely that the identical
stream efficiency will apply to all locations. The present
Ibid,p. 28.
-311-
-------
design is for a general case, however, and an average, value
must be used. Based on the information available and the
rationale presented above, a 60% load factor was selected
as being a reasonable value to use in the economics analysis.
Further, because of the uncertainty as to what the actual
load factor might be at the time and for the location at
which a sulfur oxides removal plant is installed, ,a separate^
study of the effect of load factor on plant economics will
be made in the phase 3 work.
Total annual operating costs and unit costs have
been calculated for each case and are shown in Tables 49, 50,
and 51. Based on the >.values shown it can be seen that the
alkalized alumina process will add about 76% to 92% to the is,
fuel cost when burning a 13,000 Btu/lb coal at 25 cents per
million Btu . (§6.50/ton). These are the gross costs and do
not include attrition losses or sulfur credit, whose effects
are discussed later. Two major reasons for the high costs
are the high capital investment required and the low load
factor of the power plant. Possibly the power plant ..would
operate at a higher load factor during the first 10 years
of its life, the assumed life of the alkalized alumina plant,
which would decrease the unit costs. On the other hand, the
capital charges used (13%) may be too low as discussed above,
and higher rates would of course, increase costs. A study of
the effect of capital charges, load factor, and depreciation
period is planned for the Phase 3 study and, therefore, only
the base case,is used in the present design.
b. Net Costs
The effect of attrition losses and by-product sulfur .
credit is shown in Table 52. An assumed attrition rate is
used for each process as follows: 0.1% of the solids through-
put for the dispersed phase case, one half this rate (0.05%)
-312-
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TABLE 52
PROCESS ECONOMICS INCLUDING ATTRITION LOSSES
u>
I
Basis: Sorbent @ 25C/1L
Sulfur @ $20/ton
Assumed Attrition Loss, Ib/yr
AND
BY-PRODUCT SULFUR CREDIT
DISPERSED PHASE
3,706,500
FLUID BED
997,500
DOLLARS/YEAR
FIXED BED
633,500
ANNUAL COSTS
Cost of Sorbent
Credit, by-product
Gross Cost (Tables
Net Cost
Sulfur
49, 50, 51)
7
7
926
(937
,874
,863
,
/
i
i
600
400)
000
200
249
(937
9,334
8,646
,400
,400)
,700
,700
(1
10
9
158,
,007,
,428,
,579,
400
400)
500
500
UNIT COSTS
C/MM Btu in coal feed
$/ton coal
Mills/kwh
17.42
4.53
1.50
19.15
4.98
1.65
21.22
5.52
1.82
-------
for the fluid bed design, and fixed bed sorbent life of
twice the fluid bed life based on the number of sorption-
regeneration cycles. It is expected that the dispersed phase
attrition will be the highest since the highest gas velocity
is used in this process; similarly, the attrition in the fixed
bed should be the lowest since it uses the lowest gas velocity,
and, in addition, does not require the transfer of solids be-
tween sorber and regenerator. The fluid bed attrition will
probably fall in between these two. The absolute values
used, however, are purely arbitrary since actual attrition
data are not presently available. The unit price of-make-up
sorbent is assumed to be 25 cents per pound.
Credit for by-product sulfur is calculated from .
the annual tonnage produced by each process assuming a net
price of $20 per ton will be obtained. The term "net price"
for sulfur as used here is defined as the sulfur selling price
less all costs associated with producing and selling the sulfur,'
including handling, storage, distribution, sales expenses, etc.
Basis for the selection of a $20/ton selling price-for the by-
product sulfur produced by the alkalized alumina plant is dis-
cussed below.
The competition for sulfur markets would be from
the existing sulfur producers and since the cost of producing
Frasch sulfur is in the order of $10 per ton it seems apparent
that sulfur prices above $20 per ton are unlikely. It should
be noted that historically the sulfur selling price ?has fluctuat-
ed over a wide range as the supply and demand relationship changed.
From the changes in sulfur price that have occurred over the
past few years (i.e., ^$20/ton), it would appear that-the
sulfur producers have the capability of making money, at prices
in the $20/ton range. While the 40,000 to 50,000 tons per
year of sulfur output from a single power plant probably would
not cause any great effect on sulfur price, it does seem most
likely that if enough of these plants were built such that their
-314-
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sulfur output was a significant portion of the market, then
the price of sulfur would probably drop to the $20 per ton
range. As a point of interest, it should be noted that
sulfur recovered from natural gas is presently (first quarter
1970) being sold for less than $15 per short ton.
Since the price for which by-product sulfur can
be sold does not appear to be clearly established at the
present time, the effect of different sulfur prices on pro-
cess economics will be evaluated in our phase 3 work. Sulfur
prices ranging from $0 to $40 per ton will be used and the
effect on process economics determined.
Comparing the results in Table 52, which are
based on a sulfur net selling price of $20/ton, with those
in Tables 49, 50, and 51 shows that a net reduction in cost
is obtained only in the fixed and fluid bed cases, with the
dispersed phase case cost remaining essentially unchanged.
Further, the amount of reduction obtained is only about
9% and therefore does not greatly improve the overall economics,
As can be seen from Table 52, the total annual
revenue from the sale of by-product sulfur is only about
$1,000,000. The sulfur credit is limited by the relatively
low price and the small quantities produced. The sulfur
selling price that can reasonably be expected under normal
competitive conditions will most likely be less than the
list price since the sulfur is a by-product, and the quanti-
ties are limited by the sulfur content of the coal burned
and the low load factor or the power plant. Comparing the
sulfur credit with the sorbent attrition cost shows credit
to cost ratios of about 1:1, 4:1, and 6:1 for the dispersed
phase, fluid bed, and fixed bed cases respectively. Thus,
the assumed attrition losses can be in error by a factor
of 6 for the fixed bed and a factor of 4 for the fluid bed
-315-
-------
without increasing the process costs shown in Tables 50 and
51. Any increase in dispersed phase attrition losses, how-
ever, will cause an increase in the process costs shown in
Table 49. In fact, if the dispersed phase attrition losses
are doubled (0.2% instead of 0.1%, basis solids throughput),
the net annual costs of the dispersed phase and fluid bed
designs become about the same. Since the attrition losses
are assumed, and since i;he dispersed phase losses are'the
highest and probably the most likely to be in error, the
sensitivity of dispersed phase economics to attrition losses
precludes any firm conclusion that the dispersed phase design
offers the lowest costs. Once a suitable sorbent is developed
and long term steady state data are obtained, a more exact
comparison of the three processes can be made.
2. Final Process Design
The original design of the dispersed phase and fluid bed
processes included a solids cooler in both cases, f As an alter-
native design, the solids cooler was eliminated and-the changes
in process investment and operating costs were calculated.
Table 53 summarizes the reduction in total investment for both
the dispersed phase and fluid bed schemes when the solids
cooler is removed. There is a decrease in total investment
of about $1,800,000 and $1,300,000 for the dispersed.phase
and fluid bed.cases respectively. It should be noted that
these values are independent of sorbent inventory, Claus
plant, reducing gas- generator, and the incremental- cost of
the power plant electrostatic precipitator, and therefpre
can be applied directly to the base case plant investment.
The savings in investment and process economics for the
dispersed phase and fluid bed cases owing to deletion of
the solids cooler have been summarized in Tables 54 and 55.
-316-
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TABLE 53
TOTAL INVESTMENT DECREASE FROM REMOVAL OF THE SOLIDS COOLER
(DISPERSED PHASE AND FLUID BED SCHEMES)
DOLLARS
i
to
EQUIPMENT ALTERNATION FROM
ORIGINAL FLOW SHEETS
Compressor Aftercoolers Removed
Absorber Heights Increased
Solids Coolers Removed
Water Knock-out Drums Removed
Inlet Air Compressor Decrease
in Capacity
Change in Process Equipment
Material and Subcontract Cost
Change in Total Investment
DISPERSED PHASE
MATERIAL SUBCONTRACT
FLUID BED
MATERIAL SUBCONTRACT
-80,000
+90,000
140,400
385,000
515,400
1,840,000
+47,200
-364,400
-58,800
_
-376,000
-48,000
-78,800
-32,000
-175,000
-333,000
1,290,000
-260,000
—
-260,000
-------
TABLE 54
CHANGE IN PROCESS ECONOMICS FROM DELETION OF SOLIDS COOLER
DISPERSED PHASE
(EXCLUDES ATTRITION LOSS AND SULFUR CREDIT)*
CHANGE (DECREASE)
Fixed Capital Investment ($): 1,840,000
( -v-5.7%)
Annual Operating Costs ($/yr):
Capital Charges 240,000
Labor, Maintenance, and Overhead 150,000
Utilities 580,000
TOTAL 970,000
(•vll%)
Unit Costs;
$/ton coal 0.56
-------
TABLE 55
CHANGES IN PROCESS ECONOMICS FROM DELETION OF SOLIDS COOLER
FLUID BED
(EXCLUDES ATTRITION LOSS AND SULFUR CREDIT)*
ITEM CHANGE (DECREASE)
Fixed Capital Investment ($) : 1,290,000
Annual Operating Costs ($/yr);
Capital Charges 170,000
Labor, Maintenance, and Overhead 120,000
Utilities 220,000
TOTAL 510,000
(^5.6%)
Unit Costs:
$/ton coal 0.29
«/MM Btu 1.1
Mills/kwh 0.097
*By-product sulfur produced at rate of 46,870 short tons/year.
Power plant at full capacity 60% of time or 5250 hours/year.
Note: Coal used in design contains 13,000 Btu/lb and costs 25C/MM
Btu ($6.50/ton).
-319-
-------
As can be seen from Tables 54 and 55, elimination of the
solids cooler r.esults in an overall reduction in costs for
both cases. The reduction in gross annual cost is about
$970,000 for the dispersed phase plant and about $510,000
for the fluid bed plant, which are equivalent to unit cost
reductions of $0.56 per ton of coal (2.1*/MM Btu, 0.2.mills/
kwh), and $0.29 per ton of coal (1.1C/MM Btu, 0.1 mills/kwh) ,
respectively. Based on these results it was decided* to^use
the process scheme from which the solids cooler had;-been
eliminated as the base case for both the dispersed'phase
and fluid bed designs. Therefore, the investments and operat-
ing costs shown earlier in this report are for the designs' in
which the solids cooler has been eliminated.
If the assumptions upon which these designs and economics
are based are correct, then it appears that the dispersed
phase design offers the best economics of the-.three-processes
studied. The absolute level of cost is still high, however,
since the best case (dispersed phase) shows an increase in
fuel cost of about 70% (i.e., $4.53/ton increase on the base
cost of $6.50/ton). In addition, if some of the assumptions
are changed (e.g., relative attrition rates of the three designs),
the advantage indicated above for the dispersed phase case could
be lost.
-320-
-------
APPENDIX A
ATTRITION TEST METHODS
1. Introduction
A summary is presented of the different attrition test
methods used by various investigators to measure attrition
losses of alkalized alumina. Also included is a copy of
Cyanamids1 test methods which are based on the Standard Oil
Company (Indiana) method used by Forsythe and Hertwig, IEC,
41, 1200(1949). A copy of the latter method is also included.
a. Figure 59 depicts the USBM pneumatic conveyor attriter
which is discussed in Section A. 1. a. (1) of this report,
b. The accelerated air-jet attrition (AAJA) test for
alkalized alumina developed by W. R. Grace and Company.
This test was later modified by AVCO and adopted as the
standard test. Modifications comprised increasing the
opening of the flask to 3 1/2 inches and using a 60 in-
stead of a 14 mesh screen; second degree of severity
conditions only are used.
c. Cyanamids1 test methods for synthetic fluid cracking
catalyst.
d. Albany Metallurgy Research Center - the air-lift and
fluid-bed equipment and test procedure.
e. Standard Oil Company (Indiana) test method of Forsythe
and Hertwig.
-A321-
-------
40
Sight gloss
Orifice
Closure
FIGURE A-59 - Pneumatic Conveying of Alkalized Alumina.
-A322-
-------
ACCELERATED
AIR-JET ATTRITION T2ST
VOR *!.XA1 T7.^n ALW-1TNA
W. R. GRACE & CO.
A. T. Lengade
Tech. Service Lab., Curtis Bay
December 5t 196?
-A323-
-------
Accelerated Air Jet Attrition (A.A.J.A.)
Because of the need for outstanding attrition characteristics
of a sorbent, a new attriti r test was developed for testing alkalized
alumina beads. A lot of significance in this test has been placed by
Oil Industry which they fir.c correlates rather well with attriting
tendency of catalyst in their pneumatic lift uci'ts.
The attached figure (p&ge 2) illustrates the Air Jet apparatus,
Description of the apparatus;
Is Compressed air inlet with pressure = 2? psig
temperature = 65ef. to 75°F.
dew point = -35*F. to -45°F.
2: 1/4" needle valve
*
3: 1/4" eteel nipple
4: 1/4" x 1/2" eteel bushing
5: 1/2" solenoid valve - skinner electric'valve Div. Conn.,
115 volts, 60w, 8 watts, 150 psi. Valve #LC2DB4l50,^normally closed.
6: Universal timer - Dimco Gray Co.
7: 1/2" steel nipple
8: 1/2" x 3/4" eteel bushing
9t Flowmeter - Fischer & Porter Co. #6111 A329 3B1. Precision
bore florator* tube #FP-3/4-2?-G-10
10: 3/4" x 1/2" steel bushing
11: 1/2" iron nipple length = 1"
12: 1/2" Globe valve
13: 1/2" x 3/8" steel bushing
14: Imperial brass hose connector - connecting 3/8" bushinirswith
1/2" plastic tubing.
15: 1/2" I.D. flexible plastic tubing length = 40"
16: 1/4" eteel pipe length = 2"
1?: 0 to 30# pressure gauge
18: 1/4" eteel tee
19: 1/4" x 1/4" half union coupling, Imperial
20: 1/4" copper tubing
21: 1/4" nut, Imperial
-A324-
-------
22: 1/V' x 1/V' half union coupling, Imperial
23: I/1*" x 1/V1 steel coupling
21*: 1/V' x 1/4" ho.se connection, Imperial Brass
25: Rubber stopper >lo. 9 Neojror.r, concave, surface 7/16" deep
26: 1000 ml conical flask witu ]' hole in bottom, Pyrex No.
27: No. I1* nesh support screens (upper - 2 1/2" x 2 3/2", lower
1 3/V' x 1 3/V') bolted into position.
Working Principle:
Dry air is admitted foe a certair period of tine, through a
concave stopper at a certain flow rate. The attrition of the
preweighed seir.ple is done by air. At the end of test, the iraterial
is rescreened on .T!-* mesh and the retained material is weighed.
Attrition values are reported as the weight percent loss.
Established conditions for testing Alkalized
Alumina beads with the following properties;
Shape Spherical
Size 8 to 1** mesh
Average Bulk
Density kO to 50 Ibs./cu.ft.
Total Volatile
at 300°F. 2.0?6
Testing Conditions:
Degree of Severity 1st IInd
Rotameter Reading, % 26 30
Volumetric Flow Rate in
.S.C.F/M. 3.0 3.55 **.25
Time in hours 2.5 1 «5
Screen Mesh 1* 1*» 1*
All attrition losses for I, II, and III degree are reported on an
average time of 1/2 hour.
Standard Procedure for Illrd Degree
Severity for Alkalized Alumina Beads^
A 30 gram sorbent sample (+_ 0.1 gram) previously screened on a No.
1*» U.S. Standard Screen is placed in an inverted one liter-conical
flask (26). The flask has a one-inch hole centered in its botton which
-A325-
-------
is covered by a lk r.esh screen. Dry fiir (-^O0?. D.P.) is adi-
for 1/2 hour through a concave stopper at '».25 S.C.F.". flcv: rat*.
At the end of 1/2 hour, the ir." :*riU is rescrtcnod on a U.S. «1<*
mesh screen at;d th'e retained r.V.teri.il ir> veijhed (X £rnns). Attrition
values are reported an f ollc.'..'.-? :
Let X=grams r.nteriaj hnve been retained or. ..-I1* ci«."~.h after attrition.
Then 30 ^(30-X) x IQQ = ^ attriticB Ip5s
Sane procedure ii> followed for I«t and Ilnd Dt-.yr ?-J Severity except
for time which is 2 1/2 hours for 1st decree i-.\A 1 hoar for: Ilnd,
degree. But the attrition Ibr.i:. ii rercrted or. an average, tir.e of
1/2 hour. Hcoults in a give.-! piece of apparatus hove jer.erally teen
reproducible within a maxirr.ur. doviotion of *_ 10;J of test results.
A typical . result of ;» attrition, loss of alkt.lized alu.-nir.a has
been reported in Table #1.
.TABLE A-56
Alkalized Alumina % Attritior. Loss &
1st Degree of Severity 3 <
Ilnd Degree of Severity 11
Illrd Degree ol Severity 15
Needed Correlation between test results & plant performance.
ERRATA
30 - X = Weight Loss
then [3°-^°-X)] 10Q = 10Q _ z Attritlon Loss
I- To -] 100 - % Attrition Loss
-A326-
-------
FIGURE A-60
ACCELERATED AIR J2T ATTRITION APPARATUS
- -3
-•_. '" . v /-,.'. ' >- •--—•
KAT1T
DO.V?. tor Clvrric^?. Co.
Tech . £ orv . 31 v . , C. rfrv- .
2altiT.ore, !•!:!.
ACC^J.LRAT.ID AIA .-35? A
Dv.n: A. . .
Date:
327-
-------
/
»*<*- - «••
* * f.. .'.
/<*' .
.7__.< - -a
-
**»
i ' •- i r
"| il- f
L. J - • \- /?
A- - A>
1
?r|-^ ac \-2i\\-.
t. t - !• - '«» F J^l « -
•M>-
^'
r
- X
; V..
*,,, -
,{;wT>»xC
u
•r:«*^-^fttj
• •« • - f-?'5'
^.JTTTTl
-- i
ti\
, -
^•v,}//^iF"v:^r^ifj^'^^-'iF'7^^^
—
•*\
•\
DETERM/NAT/ON OF
RESISTANCE TO ATTRITION
CYANAMID TEST METHODS
FOR SYNTHETIC FLUID
CRACKING CATALYST
FORTY-FIVE
ONE-HOUR
HOUR METHOD
METHOD
EFFECT OF ATTRITION ON
PARTICLE SIZE DISTRIBUTION
- _ Apporotui 43
^^ frixtdur. 44
Apporotui 46
46 Procedure 46
Celculolioni 4*
49
-A328-
-------
Fine catalyst dust capable of escaping from the re-
covery equipment of commercial fluid cracking units
b produced in part by the attrition of catalyst par-
ticles as they collide with each other and with the
walls of the equipment. This attrition is accelerated
when such collisions occur at high velocities.
laboratory tests to measure the resistance to at-
trition of fresh or equilibrium catalyst are conducted
under conditions approximating those in the com-
mercial fluid cracking process. At the present time.
experience and observations have not revealed any
clear-cut, simple, quantitative relationship between
laboratory and commercial results. It is important.
nevertheless, to continue such laboratory studies as
• guide for the development of new catalysts and as
a means of advancing the technology of fluid cata-
lytic cracking.
Two tests, the Forty-five Hour and the One-hour,
are employed by the Cyanamid laboratories. In both.
air jets impart high velocities to some of the par-
ticles in a bed of catalyst. The collision of these
particles with others in the system produces attrition.
Of the two tests, the Forty-five Hour .: hod is
believed to give a closer approximation of ;he con-
ditions that exist in commercial 'cracking units. The
One-hour method is a quick, approximate measure of
attrition, useful for comparison of catalyst samples.
FORTY-FIVE HOUR METHOD
The Forty-five Hour method measures attrition at
an air jet velocity of 890 ft./sec. Fine catalyst is re-
moved continually from the attrition zone by elutria-
tion into a flask-thimble assembly, which is weighed
at intervals. These test conditions are similar to those
usually encountered in refinery operations. The at-
trited or overhead catalyst so measured is expressed
as the weight per cent overhead, and is a rough esti-
mate of the particles that would pass out the stack
of a commercial cracking unit.
APPARATUS
Source for compressed air
Pressure regulator
Rotomerer, range 2-20 SCF/H oir o» operating
pressure
Moor* Flow Controller, Type 63 BU
Pressure gauge, 0-50 Ibs.
Perforated plate (See Figure 20). Tungsten carbide
• Th« apparatus b a imxtiftcalton of ihii UKt» by For-
lyllie aiMf Hcrt»<». M. Cxg. dim. 41.
*Sccpa|cM.
draw plater inserts with 0.0150 - 0.0002*
holes from W. Dixon Company, Newark, N. J.
may be inserted in larger holes in the plate
instead of drilling the small holes.
Brass tubing, 27%" x 1 %", 1.0.
Stainless steel sheet, 22" x 5", thin walled, for
upper section of apparatus
Gloss tubing, 9/16" Oiam.
Filler flask. 250-ml.
Extraction thimbles; 43 mm. diam. * 123 mm. long
(Greens #733 from Fisher Scientific Company]
Boll joint clamp for holding Desk assembly in
place (1 part from #35 and 1 part from #28)
Hose clamp, 1 ft" O.D., to hold thimble in place
Rubber stoppers,' two #9, one #10, and one #4
Needle valve, Hoke. one #34} (t * taper)
Needle valves, Hoke, ordinary (or preferably
#450 Hoke toggle valves)
Copper tubing, %"• ond suitable pipe and flare
fittings
Harvard Trip Balance
The apparatus is shown in Figure 20. p. 44.' It b
designed to admit air through the opening of a per-
forated plate at sufficient velocity to cause jets of
catalyst to be blown into the main bed of the cata-
lyst at fairly high velocity.
The large upper section above the catalyst tube
serves as an elutria tor-dust collector. As such, the
particles with diameters over approximately-16 mi-
crons (depending somewhat on density of the par-
ticles and Stokes Law2) fall back and concentrate
on the outer walls of the 5" diameter tube. There
they become agglomerated into large enough aggre-
gates to drop back into the catalyst bed.
Panicles with diameters smaller than 16 microns1
are carried over into a flask-thimble assembly that
can be weighed at intervals to determine the amount
of fines formed as a function of time. Of course, dif-
ferent flow rates and/or upper sections of different
diameter than the 5" tube will cause other diameter
catalyst to pass overhead.
Assemble the apparatus as shown in Figure 20.
p. 44. The 0.0150" holes in the perforated plate
should be drilled to close tolerances, because the at-
trition depends markedly on diameter of the holes.
Sonic erosion of the holes occurs during use, and
the disc must be replaced after several months of
continuous use. For this reason, it is advisable to
run a standard sample weekly and to note the pres-
-5329-
43
-------
. Jlpporofui lot Delttminolion of Attrition
Hose Clomp
9/16" Gloss Tubing —•
Upper Section •
Pctoili of Perforated Plot*
18
Extraction
"Thimble
Pineh Clomp
(Ball 1 SocM Type)
5" Oio. Thin
"WolledS.5.
Weld
Weld
Catalyst Tube
VID erau Pip*
27w
Pressure Reflulotor 75»/in.a / ^ J
/A \t \/ rH » L
V4""Copper Tubing
nrfl Perforated Plate
~sf—
_J IS« Otralltd Sk.tch)
FIGURE A-61
.s..~ ii^*^ .-•»; -i.. a-'i-j* «- • «• -^...w.-,-*^-;.^.- 4.Sw*«~.-~i»»»n.-«t-i-<
drop through the pliite (Gauge 6). This drop
usually decreases when th
-------
Attrition = 100 x
% Overhead — % Initial "Fines"
100 — % Initial "Fines"
The initial amount of "fines" may be determined
by subsieve analyses and knowledge of the particle
size maximum in the overhead. (Section V. p. 24).
Or it may be estimated from the intercept of the
5-hour tangent of the overhead versus time curve
shown in Figure 22, p. 46. The time required to
eliminate 95% of the initial fines probably ranges
from 2-8 hours, depending on the amount present
originally. The elutriation process can be speeded by
tapping or vibrating the upper section.
The compressor at the Stamford laboratories of
American Cyanamid Company compresses the air
to 100 p.s.i.g., so that the air invariably is saturated
after being compressed under prevailing atmospheric
conditions. When it reaches the catalyst, therefore,
h has a relative humidity of 15/115. No attempt is
made to dry the air as electrostatic effects are re-
duced when some moisture is present. The small
gain or loss in weight due to change of moisture con-
tent of the sample can be neglected.
If 10% water content for overhead catalyst (nor-
of 2.50 g./cc. are assumed and Stokes Law is ap-
plied, the particle size maximum of the "fines" de-
pends on pore volume as follows:
rv
0.2
0.4
0.6
0.8
1.0
fortlcl. Sl»
(Micron.)
13.3
15.3
17.1
18.7
20.3
A sample calculation is given "below and results
are plotted in Figure 23, p. 46.
Pore Volume = 0.96
Particle size
maximum of fines = 20 microns
Weight per cent
of initial fines (— 20/0 = 7.0
(12.2 - 7)
Per cent attrition at 5 hr$.= 100 = 6.2
(100 — 7)
(19-7)
Per cent attrition at 20 hrs. =
flftft — 71
100 = 12.9
(23.5 — 7)
Per cent attrition at 45 hrs. = 100 = 17.7
(100 — 7)
TABLE A-57
Attrition Resfttom » Oof a Sheet
FORTY-FIVE HOUR METHOD
Attrition Run. No.
Sample
Unit No.
Heal Treatment 1
Hr». Tempefoture 1100
•F. Somple weight 50
Data
9/14
9/14
9/15
9/16
Sampling
Tim*
10:15
3:15
9:00
9:15
Elop«ed
Hours
_
5%
22y4
47
P
14.7
14.5
14.9
24.4
Rotamitor
Reading
13.7
13.7
13.8
13.7
Total
Weight, g.
57.42
63.56
67.15
68.25
Weight
Ov«r, g.
—
6.1
9J
11.9
Wilghl
Ov.r, %
—
12.2
19.4
23.8
AMritlon
%
—
6.2
12.9
17.7
-A331-
45
-------
30
c
«
U
20
10
Attrition Ri-siiloocc: forty-five Hour Method
Wri'yM Per Cent Orer
Attrition Resistance: forty-five Hour Method
Weight Per Cent Attrition
10 20 30
Time. Hours
FIGURE A-62
50
20 30
Time. Hours
FIGURE- A-63
ONE-HOUR ATTRITION METHOD'
The One-hour method evaluates attrition on the
basis of the change in the minus 20-micron content
of the sample by treatment with an air jet at a linear
velocity of 2670 ft./sec. The attrition in one hour is
onr-mvi'mnrclv ffnunt tn that nhtainprl hv the FortV-
nvc flour mcmou.
The fines produced are not separated from, but
returned to, the main catalyst bed. What specific
effect these fines have on the atlrition mechanism
has not been determined as yet.
APPARATUS (See Figure 24, p. 47)
250-ml. graduated cylinder
4-oz. stoppered bottle
Glass funnel
Tared, itoppcred. fi-o: . bottle
Source for compressec air
Pressure regulator
Rotamelcr, range 2-2C SCF/H air ol opeioting
pressure
Moore flow Controller, Type 63 BU
Pressure gauge (2) 0-100 Ibs.
Perforated plate
Flanged PYRFX glass f ipe, 1" I.D. K 5.0'
Extraction thimbles, 4^: mm. diam. x 1 23 mm.
long (Greens #731!)
Gooch tubing, 1 '/j ". diam. to-hold thimble on
glass pipe
Needle valve, Hake #341, 1 ".taper
Toggle valves. Hake #450
Copper tubing,1/," with suitable fittings
Harvard Trip Balance
Assemble the apparatus as ,shown in Figure 24,
p. 47. The 0.0150" hole in the plate should be drilled
to close tolerance, because the attrition depends
markedly on the diameter of the hole. Some erosion
of the hole occurs during use. A standard sample
should be run weekly, noting the pressure drop across
the plate. \Vhen the hole becomes worn enough to
affect the results, replace .the perforated plate.
PROCEDURE
Weigh £ 250-cc. graduated cylinder.
Fill the graduate with catalyst up to the 80-cc.
level. Tap the graduate- on a firm surface until a
constant, compacted level-is obtained. Continue add-
ing cataly;: and tapping until the compacted level
reaches 80 cc. Weigh the filled graduate and deter-
mine the weight of the catalyst sample.
Transfc- the sample to a 4-oz. bottle. Add water
in an amount equal to 10% of its weight. Breakup
the lumps with a spatula. Stopper the bottle and shake
it to dispc ;c the water evenly."
46
' Thii n.dl.ud ii ba\io!ty the sa > r at thai develops! b> St.inj.utl
Oil Company (InJunj). Tubli htd tnd int. CHim 41 UOO
(1949).
'' Humiilifica on minimi/cs build-up of italic charge and avoids 1C-
cumulation .tt charged particles ...i the apparatus during the lest
-A3 3 2-
-------
Clamp the attrition column securely in a vertical
position and connect the air inlet.
Pour the humidified sample from the bottle onto a
clean piece of hard-surfaced paper. Use a brush to
remove all adhering particles from the bottle. Insert
• funnel into the glass column and pour the catalyst
from the paper into the column being careful to
•void spilling.
Weigh the Nimble and fit it onto the top of the
column and secure tightly by means of Gooch rub-
ber tubing.
Set the air pressure through the calibrated rota-
meter at 70 p.s.i.g. (gauge A, Figure 24).
Open valves (1 and 3), and adjust the air flow
by means of the control needle, valve (2), to give a
rotameler setting equivalent to an air rate of 16.3
CF/H at room conditions.
Record the exact lime of starting the nir flow.
Maintain the air pressure at 70 - 0.5 p.s.i.g.
throughout the run.
The particles in the bed should be moving down-
ward along the glass walls at the bottom of the pipe
so that the bed will be mixed. The length of the
0.0150" hole should be such that pressure drop across
the plate (gauge B) is in the range of 55-60 p.s.i.g.
(«»., 1/24").
After exactly 60 minutes, shut off the air.
Tap the thimble vigorously to return the catalyst
fines to the main part of the sample. Remove the
thimble from the column and weigh.
Unclamp the column and carefully pour the cata-
lyst onto a clean piece of hard-surfaced paper. If a
significant amount of particles cannot be dislodged
from the column, wash it with acetone. Dry the
recovered fines and add them to attrited catalyst
samples. Transfer the total recovered catalyst into a
tared, stoppered, 8-or. bottle. Weigh the total re-
covered catalyst
Prepare the sample for determination of the minus
20-micron content by s^bsieve analysis (See Section
V, p. 26).«
i If icdimentitlon l« med u the uibsfeve method of analytis, the
attrited Mmpk ilurry should be mixed in * Waring blender to
break up the agglomerations of fine particle*. -
FIGURE A-64
Apparalut lor the Determination of Attrition
by the One-hour Method?
oiinct rtATt
V»" Stainless Steel
| ——^4- D
I ~~ J E
Drill 0.01 50" Hole it
xact Center of Plain
• Extraction
Thimble
,, Cooch Tubing
End of Pipe
Wrapped with
Rubber Tape
1" I.D. Flanged
' Pyrex Glass Pipe
• Support Clamp
Pressure Regulator
Rotamcter !/„" Copper Tubing
Pressure Cage \
From 125 #
Air Lint
Flange
Orifice Plate
Rubber Catketi
Support Clamp
Prctiure Cage
Moore Flow Controller
Air Filter
A -A333-
-------
CALCULATIONS
Per cent nonrecoverablc fines. The increase in
weight of the thimble is assumed to represent 0 to 10
micron particles, which are too fine to be dislodged
by tapping. These particles must be included in the
TABLE- A-58
• Sompfe Worksheet vied In the Determination
•f Attrition by th» One-Hour Method
subsicve analysis after attrition and may be deter
mined by means of the following expression:
lab. No.
Car No.
Weight of thimble, g.
Final _ 6.4
Initial
Increate IN)
4.0
Dale
Grade
Weight of lample, g.
final 47.0
Initial 45.0
0.4
•am volume, cc./g.
Minus 30-mlcron fraction. Initial weight %
Rolameter reading
0.10
5.1
14.3
P«r Cent Nonrecoverable Fines
_ 100 X N _ 100 x 0.4
f 47.0
= O.S5
Correction of Minus 20-Micron Analysis for Thimble Fines
WoigM per cent of — 20-micron fraction a> rvn.' a = 16.3
(a -f- 0 100 (14.3 + 0.«5J 100
Corrected per cent liner =
100 4- F
100 + 0.15
= 17.6
(17.0 — 5.1) 100
100 — 5.1
= 12.6
Test Recovery
IOI on Initial iomp4e. %
IOI In recovered torn >le, %
T (100 — V,>
S (100 —
X 100 =
14.2
20.3
47.0(100— 14.2) 100
45.4 (100 — 20.3)
' Sedimentation null.
-A334-
-------
where F = Nonrccoverablc fines, %
N = Increase in weight of thimble, g.
T = Total weight of catalyst recovered
from attrition test, g.
Correction of Minu* 20-Micron Analysis for
Thlmbl* Fines
^'article size = 20 microns
Per cent finer as run = *
(a + F) 100
Corrected per cent finer =
100 + F
Role of Attrition. Relative rates of attrition for
catalysts are calculated as the increase in fines con-
tent caused by the standardized treatment for attri-
tion. The rate of attrition is based on the difference
in the quantity of 0 — 20-micron fraction present,
as shown by micron analysis before and after attri-
tion, and is determined by the following expression:
Attrition rate =
A X 100
B
where
A — Increase in weight per cent of
0 — 20-micron fraction present
B = Weight per cent of plus 20-micron
material in charge
Test Recovery
% Recovery =
T(100—Vr)
S (100 — V.)
X 100
where
T = Total weight of recovered
catalyst, g.
S = Weight of original sample before
humidificalion, g.
V, = Volatile content of recovered
catalyst, %
V, = Volatile content of original
sample, %
ON r-ARTICU SIZE
Attrition effects appreciable change in the particle
size distribution of fluid cracking catalyst. The degree
to which this change occurs can be measured by de-
termining the particle size distribution of the catalyst
before and after attrition.
In the following example,-particle size distribution
was determined by sieve sedimentation (see Section
V, p. 24), and attrition by the One-hour method.
However, sieve-CAE Roller and/or the Forty-five
Hour attrition method can be used.
Particle Size Distribution
TABLE A-59
FortUU 5!i»
IMlcrentl
(50
100
80
74
60
50
40
30
20
10
f»r Ctnl Finer
»*fort Allrillon)
100
92
80
74
56
41
26
13
5
1
Mr Cent Finer
(AfUr Attrition)
100
95
•7
•3
71
61
44
31
17
6
-A335-
49
-------
ALBANY METALLURGY RESEARCH CENTER
ATTRITION TEST METHODS
Fluidiz».:J-Bed Attritioning Equipment
The fluidizcd-bed attritioncr consists of a "Roots" low
pressure blower (60ft /min), a 2-inch manifold with 1/4-inch outlets
to each fluidized bed, a laboratory-built electric furnace (hot-zone
12" x 12" x 60") and 12 "Vicor" fluidizing tubes. The fluidizing
tubes are 2 inches in diameter and 16 inches long, with a 5-inch
diameter by 6-inch boil on the upper end to reduce air velocity and
thus retain attritioned material. The feed ends of the fluidizing
tubes are tapered at about 60°, from 2 inches dam to just under 3/8
inch, to retain 3/S-inch steel balls used to disperse the.air as it
enters the bed. The air is not preheated before entering the fluidizsd
beds; therefore, there is a temperature gradient across the bed of
from room temperature to 300° C.
Attritioning-Testing Procedure
1. Representative 50-gram samples of each product to be
treated are weighed and sized to obtain the size distribution cf the
material.
2. Samples are then charged into the top of the fluidizing
tubes and the air adjusted to produce a turbulent bed. Gas velocity
through the bed is about 1.8 ft/sec.
3. Every 24 hours the cliurgc is weighed and additional material
is added to maintain a 50-gram charge.
-A336-
-------
4. The 50-gr;un charge was screened to determine the change
in size distribution of the material after coaipletion of the 15-day
attritioning test period.
5. Air-lift attritionny sample size was 10 grains ant! the
gas velocity about 25 ft/sec.
6. 80 mesh screen on top.
-A337-
-------
FIGURE A-65
AIR
LIFT
diom
\-
i
j?**— 3" d.io m H
Gloss
guides
2" diam-
l" diam-
Gluvi-^
supports
12"
I,
1
k
15"
FLUID
BED
i
diom-4-
-A338-
-------
Attrition
Fluid Cracking ©aialy
Catalyst attrition, or particle breakdown, In a fluid-
catalytic cracking unit Is an Important factor In operating
costs, as It directly affects the losses of catalyst from a unit.
A simplified laboratory accelerated-attrition test has
provided useful data for evaluating attrition resistance.
The catalyst is subjected to the action of a high-velocity
•Ir Jet; the extent of particle breakdown is determined by
screen and Roller analyses. A comparison of test data on
commercial grades of fresh natural, silica-alumina, and
silica-magnesia ground catalysts shows that all are within
the same range of attrition resistance. Fresh micro-
spheroidal catalysts show better resistance to breakdown
than ground catalysts except where there are excessive
proportions of agglomerates. The effect on attrition
Laboratory Studies
resistance of heat pretreatment at 1100* r". varies with
different catalysts. Silica-magnesia catalyst showed the
greatest Increase In attrition resistance among the
ground catalysts, while mlcrospheroidal silica-alumina
samples showed little change. Attrition data on catalyst
samples from the start of'runs in a commercial fluid-
cracking unit with fresh synthetic ground catalysts show
that the attrition resistance is greatly Improved during a
short period of operation. A comparison of commercial
and pilot plant data indicates that the greater part of thli
Increase In attrition resistance results from mechanical
action on the particles. The smaller effect of subjection
to cracking operations is similar to the laboratory observa-
tion on heat-treated catalysts.
W. L. FORSYTHE, JR., AND W. R. HERTWIG
•TAMOAffO OIL. COMPANY (INDIANA). VWHITINO. IIMO.
FLUID catalytic cracking (5) has become one of the principal
processes in petroleum refining. The most important feature
at this process is the maintenance in a fluidiied condition of a
powdered catalyst which ranges in particle size from less than 1 to
over 100 microns. Appreciable quantities of this relatively ex-
pensive catalyst are lost from a fluid-cracking unit because of the
incomplete recovery of fine catalyst particles from the regenerator
flue gases by cyclones and CottreU precipitators. Furthermore, in
some recently designed cracking units CottreU precipitators have
been omitted, and all the particles finer than about 20 microns are
rapidly lost through the cyclone separators. In addition to the 0-
to-20-microo particles present La the fresh catalyst charged to the
unit, fines in this size range are produced during operation by the
attrition of the coarser particles. Consequently, to minimize
tones of catalyst fines, the catalyst must have good resistance to
•ttrition.
Information on attrition resistance is therefore an important
factor in the laboratory evaluation of a new cracking catalyst,
which involves a comparison with the attrition resistance of a
commercially acceptable catalyst. In order to provide data for
these comparisons, a laboratory accelerated-attrition test was de-
veloped to determine the relative resistance to particle break-
down of various catalysts. This paper describes the test, evalu-
ates some of the factors influencing catalyst attrition, and pre-
sents attrition data on a number of different cracking catalysts.
Attrition of catalyst in a commercial fluid-cracking unit is
caused by several factors, including collisions between catalyst
particles and impact and abrasion of the particles upon the walls
of vessels tod catalyst carrier lines. In the design of a suitable
laboratory test for comparing the attrition resistance of various
catalysts, it was desirable to reproduce as nearly as possible the
•We mechanism of particle breakdown, so that the results
would be representative of commercial operation. Although it
wa u«t possible to reproduce exactly the plant conditions on a
laboratory scale, it developed that a high-velocity air jet im-
pinging on a bed of catalyst particles was a readily controlled
tteans for attrition of catalyst and gave the desired accelerated
breakdown rate.
In the present test, the charge of catalyst is subjected to the
action of a single air jet for a fixed period of time. The break-
1200
down shown by the change in particle-size distribution is com-
pared with that for other catalysts under the same conditions.
Although the comparisons of catalysts are relative and do not
permit direct quantitative predictions of actual plant losses, the
test results have been useful in laboratory evaluation of new
catalysts for commercial operations. In addition, information hat
been obtained on the trend of catalyst attrition resistance with
length of time of operation in a commercial unit.
In contrast with the conditions in a commercial unit, the fines
produced by attrition are retained in the test apparatus, so that
the severity of attrition on the remaining coarse particles is some-
what reduced. As a second deviation from plant operation, the
attrition in the test is presumably caused by the collisions of
particles rapidly accelerated by the air jet with slower moving
particles; hence there is no effect of particles hitting pipe walls or
obstructions. The advantages inherent in this test method-
small catalyst charge, simplified apparatus and operating proce-
dure, good reproducibility, and short time requirement—are be-
lieved to offset the advantages of possibly closer approaches to
plant attrition conditions that could be obtained with a more
complicated method.
From the standpoint of attrition resistance, the catalyst* com-
merically available for use in Quid-cracking operations may tw»
divided into two principal types: ground catalysts and micro-
spheroidal (MS) catalysts. Fresh-ground catalysts have gen-
erally irregular shapes with sharp edges and protruding corners.
Microspheroidal catalysts, as the name indicates, have spherics!
shapes and smooth surfaces. The spherical shape was sought in
order to reduce attrition losses from cracking units by eliminating
the catalyst surface irregularities which give rise to high initial
breakdown rates.
ATTRITION APPARATUS
A diagram of the attrition apparatus is presented in Figure 1.
The catalyst at room temperature is subjected to the attrition
action of a high-velocity air jet issuing from a V«-inch orifice at
the center of a Vi-inch stainless-steel plUe. This orifice plate a
bolted between flange* at the bottom of a vertical 5-toot section of
1-inch-inside-diameter Fyrcx pipe. The high-pressure air supply
is reduced to 70 pounds per square inch guge by means of the
-A339-
-------
NOT REPRODUCIBLE
I')t9
KIGUILE- A-.6-6
ROTAM£IERj
INDUSTRIAL AND ENGINEEI1ING CHEMISTRT
1201
PRESSURE
CAGE
(no Pa icy
^
PRESSURE
GAGE
9J
I
_r~
L,
3 CANVAS
: JUTtR
*
i ' i. a
GLASS PIPE
ORIFICE
, PLATE
rJ*-(l/«4' HOLE)
Attrition Apparatus
pressure reducing valve. The air rate is measured with a rotam-
t\fr am controlled by a needle valve downstream of the
rotnioctt r. A pressure gage upstream of the orifice plat.? indicates
the pressure drop across the orifice. A canvas filter bag is
clamped securely at the top of the glass pipe in order to prevent
loss of catalyst entrained in the exit air stream.
TEST PROCEDURE
In this laboratory accelerated-attrition test a weighed catalyst
sample .is subjected to the jet attrition action for 1 hour. The re-
lulting ctiange in distribution of catalyst particle size is deter-
mined by screen and Roller analyses. The more pertinent details
of the attrition procedure and particle-size analyses are described
below!
Attrition Procedure. A representative 50-gram sample of the
catalyst is charged into the top of the glass pipe, apd the weighed
ulv^. wttt. u oi&itiuru HI in.Mt r. -*n ntr CMIM IM il ,',•> vranfjrtr'1 ?Mfy<*
iixic, per minute is men maintained through the ontice for 1 hour.
This air rate corresponds to a superficial velocity of about 0.8 foot
per second in the glass pipe, which thoroughly fluidizcs tlie
catalyst bed and provides good mixing of the particles. The pres-
sure upstream of the orifice plate is 55 to 58 pounds per square
inch gripe under these conditions, and the air velocity in the jet
approaches the speed of sound. Catalyst particles in its path are
very rnpidly accelerated and collide with the slower moving
particles in the aerated bed. Tbe?e high-speed collisions cause the
attrition of the catalyst particles in this test.
During the test, the g!.is.« pipe and filter bag are periodically
lapped in order to return to the bed any catalyst thnt has stuck to
the wall or canvas filter. Following the test period, the catalyst
ia recovered as completely as possible from the pipe and the filter.
Acetone is used to wash free any catalyst particles I hat have stuck
Io the pipe walls. The canvas bag is weighed to determine the
quantity of catalyst fines not directly recoverable. These fines
are included in the particle-size analyses in the finest size fraction'.
Including the fines in the be,r. the catalyst recovery is seldom less
lhan 97% of the charge. To correct the weight of recovered
catalyst for any change in moisture content during the test, vola-
tile content at 2200* f. is determined on the catalyst before and
after attrition.
Screen Analyses. The screen-analysis data presented in this
paper are based on the Tyler standard scries from SO thrpugh
326 mesh using a combination wet- and dry-screening method. A
representative 25.0-crara sample of catalyst is placed on a 325-
mesh serein, and tiie finest particles Are washed through with
water. The catalyst remaining on the screen is then washed with
acetone to remove excess water and dried for I hour in an oven at
220'F.
This wet method eliminates the passage of large amounts of fine
particles thruuph the whole scries of screens and thus greatly
inmimi/c air-jet attrition in changing the
particle-size distribution of the coarser fractions of cntsJyst. Par-
ticularly in evaluating a new cat.-ilyst, it is significant to show the
manner of breakdown in the coar ;er size range which results in the
observed increase in — 325-mesh fines.
During the screening procedi re the moisture content of the
catalyst changes to an extent d< pending on the level of catalyst
activity, previous heat treatments, and exposure to atmospheric
humidity. It is therefore neces-ury to correct the weight of the
sample and the weights of the screen fractions to a constant
moisture basis before calculation of the — 325-mesh fraction by
difference. Total volatile content at 2200* F. is determined on
the sample before screen analysis and on the combined screen
fractions.
Rcproducibility of screen analyses ia checked periodically to
avoid varying precision of result::. Frequent checks with the set
of screens currently used are made on a standard catalyst sample
to detect deterioration of the wreens or errors in procedure
When it becomes necessary to replace worn screens, results with
the new set are compared with previous results on the standard
sample and with analyses with 2 set of reference screens. Varia-
tions between duplicate analyse;; with a single set of screens iel-
dom exceed =* 1 weight %. Because measurement of attrition ia
based on the difference between ' wo analyses with the same set of
screcus, small variation in absolute values among different Sets of
screens has a minor effect on the calculated attrition rate.
N'o significant catalj-st attrition to form —325-mesh fines has
been found during the normal Ro-Tap screening procedure.
Therefore, no corrections are needed to obtain the true particle-
size distribution by screen analysis.
Roller Analyses. The present discussion is concerned only with
measures taken to imorove the gr.rnrsrv nf Hf>t« nht.iinpd Kv th*
effects shown by Matheson (S), the air supply is humidified by
bubbling through a 3S weight % sulfuric acid solution. A con-
stant air rate of 9.6 liters per minute is used in all analyses, based
on a particle density of 1.35 grams per cc. In the absence of con-
sistent data, on particle density, no correction for the lower par-
ticle densities of the fre-sh catalysts has been made. Experience
indicates, however, that such a correction would amount to less
than 1 weight % decrease in each of the three finest size fractions.
In order to .oinimize attrition of catalyst particles during the
Roller analysis, a modified aeration tube similar to that described
by Matheson (S) has been employed.
DISCUSSION Or TEST
Information on the test variables that affect the proper inter-
pretation of test results leads to a better understanding of the re-
lation of the laboratory test to attrition conditions in a comrner-
TABLE A-60
Effect of Catalyst Fines Content on Attrition
(Fmb-crouod wliei-tlumin* eatalyit B)
Ch»rRe io A tuition Teat
A* Reerivrd Finn-Free
Screen
An*\yii»
4- 80 mnh. %
-80 + 100 mnh. %
-100 + 150 nw.h. %
-150 4- 200 n.T,h. 7,
-IOC + 270 m^h. %
-270 + Mi mwh. %
-324 m»»h. %
Aj
received
o.e
2.5
19 1
22. J
12.1
«.»
36. B
100 0
irrcMe IB % of -325 mob
After
attrition
0.2
0.4
7.3
14.5
10.8
4 A
C2.2
100.0
as. e
At
received
O.I
4.*
31.3
33.*
17.9
(.4
2.»
100.0
A(l*r
•tuition
0.1
O.ft
10.0
19.0
11.4
1.0
53.1
100.0
M.ft
are not recovered. The small loss occurring during the
•lr>-MTtiiiing is proportioned among all the fractions.
Attrition r»t«. fi
In.-rraai'in '",. of -325 mf»h X 100
mnh in
| 40.4
•I.*
-A340-
-------
1202
INDUSTRIAL AND ENGINEERING CHEMISTRY
Vol. 41. No. 6
FIGURE A-67
FRESH GRQUHD
SILICA-ALUMINA
TIME,
40
MINUTES
Increase in Fines Content during
-Hour Attrition Test Period
cial unit and provides knowledge of the actual attrition mecha-
nism.
Screen-analysis data show that the attrition obtained in the
laboratory test in nearly all cases results in a net increase only in
the — 325-mesh fraction. Attrition-test data in Table I for an as-
rc-eived fresh-ground silica-alumina catalyst sb"w that the in-
crease in the — 325-mesh fraction from 30.6 to 62.2% resulted
from a net decrease in all the coarser fractions. On the screen-
analysis basis, therefore, the extent of attrition of —325-mcsh
particles already present in the sample to be tested is not meas-
ured. Particles finer than 325-mcsh are termed fiuts in subsequent
discussions.
Fines Content of Catalyst. The presence of fine particles in the
catalyst tested.is one of the most significant factors influencing
•' •••••--'•- TV j 1....tf.-,....r-.,..*,»....:,.
01 tne coarser parucics in two ways: oy a cusnioning eiieci
which limit* the force of collision impact between coarse particles,
and by dilution which reduces the number of coarse particles
available for attrition. The significance of these effects is illus-
trated by Table I. An attrition test with fresh-ground silica-
alumina catalyst resulted in an increase in — 325-mesh fines from
36.6 to 62.2%. When most of the fines had been screened from
the catalyst before testing, there was an increase in —325-mesh
from 2.8 to 52.8%.
It has been found possible to compcrisste for the dilution effect
in an empirical manner by d-viding the increase in — 325-mcsh
material by the percentage o:' courser than —325-mesh particles
present in the test sample. Siiiilur corrections may be applied to
Roller-analysis data. When the data in Table I are corrected o
this fines-free charge basis, th-s as-received sample shows an in-
crease of only 40.4%, compaied to 51.5% for the sample which
was initially nearly free of firts. The calculation to a fines-free
charge basis therefore does not completely correct for the total
effect of the fines on attrition rate. It is believed that this differ-
ence is largely attributable to the cushioning effect of the fines.
Because it does not seem U ;ely that any simple additional ad-
justment for fines would properly correct for the cushioning
effect, which may vary from one type of catalyst to another, no
further correction has been jndertakcn. However, as long as
different catalysts are compaied on the ba^is of similar fines con-
tents, this deviation should not interfere with the usefulness of the
test results nor the validity of such conip.-mron.;. In all cases, the
previous correction of the breakdown to a fines-free sample basis
hiubcen employed. This corrected increase in -325-mcsh is im-
plied in thU paper when the term nltntion rate U used.
The fines formed during tfe attrition test ter 1 to reduce the
breakdown rate in the same manner a? the fines present in the
original charge, through an at vUtion.-il cushioning effect and a fur-
ther decrenxe in the proportion of coarse particles available for
•trritiou. In any given commercial unit th« cushioning effect w
kept reasonably constant by the loss of fine* to some equilibrium
value. The inventory of coarse piirticlfi-s is maintained by addi-
tions of fresh catalyst. No attempt har, been mae increase
in attrition resistance is shown below for operations in commer-
cial units.
Precision of Test Data. The range of variation in th-; data from
differer t attrition tests on the same catalyst is not f.pprecubly
TABLE A-61 NOT REPRODUCIBLE
Precision of Attrition Test Data
(Microipheroidal jpray-dried iilic»-iluraini cmlyit •<)
CtUlyjt
u
Scre'a »a»ly>U Received
4 80 mi-sh. % 0.4
- JO 4 100 n.'ih. % I.J
-100 + 150 mp.ih. % 11.9
-150 4 200 nicjh. % 20.2
-200 4 270 mesh. % 13.8
-270 4 3J-> mcah. % 7.8
-32ittojh. % 44.3
ipo.o
Increase io % of —325 mesh
Attrition rite, %/hour
Test coi lilion.i
Air ntt. >t.in
-------
INDUSTRIAL AND ENG1HEEHING CHEMISTRY
1203
Attrition
Silica-Alumina A
A. Ali'r
gcrfo An«lyti» received attrition
4- 80 mesh. % 0.4
_ 80 + 100 mesh. % 2.2
-100 + loO mebh. % 18.8
-l» + 200 mc^h. % 27.5
-JOO + 270 rn'sh. % 13.8
-270 + 32i mesh, % «.4
-125me>h. % 30.8
100.0
Roller Analyaia
0-10 micron, % 4.1
10-20 micron. % 4.5
20-40 micron. % 11.9
40-80 micron, % 30.3
80-r micron, % 48.0
100.0
Attrition Breakdown
1. Incre»e in % of -325 mesh 30.0
Attrition rate. Si/hour 43.4
11. Incie»e in % o( 0-AO micron 25.7
Inrreaie in % ol 0-40 micron X 100 -, „
% 40 + micron ia chttrge
III. Increase in % of 0-20 micron 20.8
Increase in % of 0-20 micron X 103 „
% 20+ micron in charge
0.0
0.2
4.9
14.8
11.9
6.4
60.8
100.0
19.4
11.2
16.8
27.4
24.2
100.0
of Fresh Ground Cr.icking C.italysts
At
received
0.2
2.9
2«.4
18.9
10.3
3.9
37.4
100.0
2.8
13 .'o
20.9
46.5
100.0
22.*
35.
13.
10
17.
19
TABLE
c-in C
After
attrition
0.2
10
13.4
13.4
7.2
5.2
59. S
foo.o
15.8
11 3
19.3
20.0
33.8
100.0
A-62
C.lAlX.t
E
A> Alter
received attrition
4.7
3.6
11.9
13.6
9.8
6.4
49.8
foo.o
4.0
8.4
16.6
30 9
40.1
100.0
14.5
48.1
12.0
45.1
11 .9
16.4
3.1
2.3
o\2
4.5
3.4
74.3
100.0
20.8
23.5
16.7
17.0
22.0
100.0
Catalyst-*—
C«tnl)"it
F
received • attrition
4.S
3.7
12.1
14.4
11.5
5.2
48.3
100.0
3.8
7.3
15.5
29.8
43.8
100.0
11.8
36. 0
19.1
Z6.1
18.9
21.1
1:1
7.0
9.1
7.1
4.1
66.9
100.0
19.1
10.7
15.9
21.8
32.3
100.0
greater than the variation to be expected from the screen analyses.
Data from a scries of tests on a microspheroidal silica-alumina
catalyst (Table II) show the typical range of deviation. The
increase in -325 mesh varied from 15.6 to 13.7%, for a maximum
variation in attrition rate from 2S.O to 33.6%. This range reprc-
KIIU - »vi.I—Iv-*. _ -.-„ ' ''•''• fJOoY TKnf^t f-nnrlitinns
have been included in Table 11 to snow mcu uuima. .............
FIGURE A-68
*
•'an
o-*°
U
•a AD
o
ce
0- SO
5
SEO
i
n
Rio
t
i 0
^
FRCSH
SILICA-
V
\
<
JBOUNO SYN
ILUMINA CA
TH£TIC
fAuYST
I 2
TIME. HOURS
Rate of Attrition during Successive
Hour Periods
The amount of breakdown obtained in the laboratory test ia
»ot significantly a/Tccted by variations in orifice pressure drop
caused by rc.isonable variations in the accuracy of drilling the
V«-inch hole. Results obtained with plates giving pressure
drops as high as 65 pounds per square inch have not shown any
significant difference from the data 'or 55 to 5S pounds per square
Inch in Table II. Hence, test rcsul * from different sets of ap-
paratus may be expected to be in i;ood agreement.
ATTRITION nCSISTANCC Or FRESH CATALYSTS
Ground Catalysts. Samples of fio three principal commorml
types of frc.- to
40-micron fractions. For the silica-alumina catalyst this amounts
to increases cf 15.1, 5.7, and 4.9%, respectively, for these three
fractions.
The material finer than 20 microns produced in the laboratory
attrition test is of particular significance from the standpoi it of
catalyst lossc; from a commercial unit having no Cottrcll pruipi-
tator; all particles fuicr than about'20 microns are rapid!) lost
from such a unit. A comparison of the ground catalysts or this
basis shows t' 325-
mc«h screen, which was shown by microscopic exanfm ilion.
With its higher initial fines content, tlie scrern analysis ••insc-
qucnlly pro/idcs a greater correction when calculated to a lines-
tree charge basis. However, th; deviation of screen-analysis ilnta
from the IMkr data docs not limit the usefulness of r. lativft
-A342-
-------
1J04
INDUSTRIAL AND ENGINEERING CHEMISTRY
Vol. 41. No. C
Attrition of Frevli Microsphcroidat Cracking
Catalysts i
JiSilira- Silira-
XXIurnin" Ma"ttr«iii
~>I» J»
,
Attrition t.if. %/h'.,.r
InrrejiC in % of 0-80 ntirru
In«e»ie_in %^or_£-S
*" -
II.
_ ^___
% 804- niirron in charge™ "
n|. lucres^ in % of 0-40 micron
Inrrta^in % ot 0 -40 rnicrun X 100
~~
% 404- micron in chanif
IV Inerewc in % of 0-iO micron
lncr>.ye in J^o/ 0 -JO micron X_100
% 204- micron ia charge "
13 5
27.0
.. •
13.4
18 8
11. a
12.1
90
17.7
..
2.7
2.8
1.5
1.4
3«.9
60.9
..
31.4
40.1
14. S
14 9
18.3
20.4
22.8
SI. 8
IS. 7
17.0
• .2
9.4
• Si>r>y dried ?MMy.l.
t Oil-dropped catalyst
comparisons among screen analyses as employed in the present
method
Hicrospheroidal Catalysts. Samples of some microsphcroid.il
catalysts have shown considerably better attrition resistance than
ground catalysts in the laboratory attrition test. Data in Table
IV on spray-dried silica-alumin.', microsphcroidal catalyst sumplc
Gshcw an attrition rate of 27.0%, compared to the 43.4% shown
in Table III for ground silica-alumina catalyst A. There is a cor-
rected increase "' 0-1°20-micron material of 12.7% for this micro-
api--roidal catalyst, while the ground silica-alumi!)-1 shows 23.1%.
An oil-dropped microsphcroid;d silica-alumina catalyst, sample
H, givss an attrition rale of 17.7%; on the Holler-analysis basis
ftc breakdown is even less, with an increase of only 1.5% in the
0 to 20-niieron fraction.
The large difference in the amount of breakdown shown by
scrcrn and Holler data for sample H illustrates an effect of orig-
inal particle-size distribution upon the measurement of attrition
rule. If a large proportion of particle* of a catalyst is close to the
.. ' '•••-• '•' -1--. ' ••-•••- ••- «--^,l ;.c_
j«_mcsn--a relatively sinau cnange in size 01 inesc particles
could lie reflected in a markedly higher attrition rate than would
be the case for an originally coarser sample. Over 60% of
thepxrticlcs of sample H were in the 40- to SO-nucron range;
therefore the 325-meih screen which has a cut point within this
wogc, will show a greater amount of attrition than indicated on
the Holler basis. Other catalysts had more uniform and mutu-
ally similnr initial particle-size distributions, so that this effect
was not an important factor.
Excessive amounts of agglomerated particles in a. micro-
ipheroidal catalyst can haves highly detrimental effect on its at-
trition resistance. The agglomerates break apart under attrition
conditions much more readily t'i«n the spheres themselves; hence,
if the agglomerates are composed of fine particles, the over-all at-
trition resistance can decrease into the range for ground catalysts.
Aammplc of this condition is the case of sample J. (Table IV).
This shows an attrition rate of 60.9%, which is greater than
tint for any of the ground catalysts in Table III. The in-
crease in 0- to 20-micron aiati rial la not so great as for a ground
catalyst :ii proportion to the increase in —325 mcih, as hnlf of the
particles result'in;.; from breakdown of agglomerates were in the
20- to 4.0-micron ran;;o. The corrected increase of 11.9% 0 to
20 microns neverthrlcss approaches the 19.2% shown for
catalyst C.
Microscopic examination of the original ciitatyst showed indi-
vidual particles to be spherical in shape, hut practically all the
particles in the SO- to 200-mrih range arid over 50% in the 20')- to
325-mcsh range were agglomrMtes. During the attrition ie.it,
these agglomerates were largely broken down; the resulting finer
fractions were composed practically completely of individual
sphere*. This showed that the hi;;h proj>ot tion of fine particles in
the agglomerates making up this batch of microsphcroidM cata-
lyst did not permit it to have the desired advantages of attrition
resistance over a ground catalyst.
The data on sample K (Table IV") show better attrition resist-
ance than for sample J. The attrition rate decreased to 20.4%
because of the presence of fewer agglomerates. Although there
were considerable proportions of agglomerates in the coarser frac-
tions, these were made up of larger particles. The data in Table
IV show increases of 9.2% in 0 to 20 microns, 15.7% in 0 to 40
microns, and 22.5% in 0 to SO microns. It follows that the fine
material foruic-d "during attrition consisted of about 40% 0 to 20
microns, 30% 20 to 40 microns, and 30% 40 to SO microns.
Attrition Resistance of Heat-Treated Fresh Catalysts. The
effect on attrition resistance of heating catalyst to operating tem-
perature of the cracking unit varies, with different catalysts and
with the extent of previous calcinations. In Table V are pre-
sented data on samples of various types of catalyst which, before
testing, were heated at 1100* K. for 4 hours in a muffle furnace.
A comparison of the data for ground silica-alumina catalyst A
before and after healing shows a decrease in attrition rate from
43.4 to 42.1%. A similar comparison for the more attrition-re-
sistant of the two samples of natural catalyst also indicates the
hcatinc to rive a decrease in the attrition rate, from 36.0 to 31.7%.
ground *;'ici-magi>c.>ia catalyst D and resulted in an attrition rate
of 29.1% .Muxrwrcd to 30.4% for the untreated catalyst.
With b-'ih the spray-dried and oil-dropped microspheroidal
silica-a!-: .-iui catalyst samples, however, no decrease in attrition
r»to occurred as a result of the heat treatment. Sample G had
an •ttri::c:i rate of 27.1% after heating, compared to the origi-
nal valu'; of 27.0%. The oil-dropped catalyst showed an in-
crease m attrition rate from 17.7 to 20.4%.
On the basis of the Roller increase in 0- to 20-microu material
from attrition, the breakdown of ground silica-alumina decreased
from 23 1 to 1S.1% as a result of the beating; this u a greater
effect than was shown by screen-analysis data. The Kaller dati
also show that the decrease in attrition rate from heading silira-
magncsi.i sample D was great enough to give it a considerably
lower rite than microjptieroidal silica-alumina sample G, which
bad shown little change.
Such data on the attrition resistance of catalysts after beat
treatment provide a clearer picture of possible effects occurring
TABLE A-64
Attrition of Fresh Catalysts after Heating at 1100* F. for 4 Hours
Attrition Breakdown
t bureau in % of -3Ii
Attrition rat", %/hour
1. IncrMie ia % of 0—10 mirr >n
iBcrea-iejn % o( 0- 10 mccr i
% 404- micron in eh..
"t Inere«.«c ia T, of 0-20 .niefun
Incm.t in 7C of 0-?0 micr
ir>_X IUO
ise
.n X 100
received
30.0
43.4
25.7
32.8
20.8
23.1
Healed
27.9
43.1
19.0
24. S
18.3
'8.1
received
24.6
36.4
15.8
21.1
12. S
14.2
gneu'a D Si lunl E Natural F
Heated
17. 7
29.1
8.r
11.7
6.0
6.8
Aa
received
24..'
48. t
32.0
IS.)
31.1
36.-.
\j
Heated received
18.6
36.0
19.3
36.3
18.1
21.3
lieatej
(6.7
31.7
10.6
15. i
11. 1
12.9
Sili
Microcphfroidrl
Spray- Dric.i. G
A,
received
13.5
27.0
.13.4
18.8
11.6
12.7
HeiteJ
13 3
27.1
11.0
IS. 3
10. S
11.6
*i— — .
O.I-f'.opccJ. H
^3
received
17>
»./
2.11
l.i
i.;
Hi-J«e4
20 '.4
3.*
3.6
».»
2.8
-A343-
-------
NOT REPRODUCIBLE
!»»« '
rf iO
X
. «0
* 30
z
i
S 20
s»
1
I £
,1 1 V
• J
V"
-^0
f-rnitup A_£Q
A-ALUMlNj
FftCSI
SILICA-MAC
i
S»UCA-
o
A
l
NESIA '
KACNESIA
/
"L *
StUCA-AL
ilN UN
T
o—
[) K) 20 30 40 50 60
INEERING CHEMISTRY
1205
Decrease In Attrition Rate of Ground Catalysts during
Operation of Commercial Fluid Unit
ra & cracking unit, even though the much more rapid heating
which occurs when fresh catalyst is added to a unit may have a
somewhat different vffcct. On tiie other hand, inasmuch as sig-
nificant . ttrition of a fresh catalyst may occur before !• paling to
cracking temperature during the charging of fresh catalyst to a
commercial unit, attrition characteristics of catalyst samples both
before and after healing should be considered in the evaluation of
* new catalyst.
ATTRITION HESISTANCE OF USCO CATALYSTS
The attrition of a fresh-ground cracking catalyst decreases
rapidly during the initial period of operation in a commercial
«.— .— e .....0—_.—_,.- .. — -.L .
tests on samples ol ground silica-alumina catalyst withdrawn
during -JL commercial run. Before being charged to the unit, the
fro.-.h catalyst had an attrition rate of -10%, whereas the used
catalyst from the unit showed a breakdown averaging only 12%
over the first 2 months of the run.
Similar data in Figure 4 from a run with a ground silica-
m.itnrfia catalyst sbuw a large decrease in attrition rale during
transfer of catalyst from the storage bin to the regenerator and be-
fore circulation of catalyst from the regenerator to the reactor
l«a«J lirgun. The fresh, catalyst from the bin thowed an attrition
rate of 36%, whereas that in the ref^neratcd-catalyst standpipe
nt the start of transfer of catalyst to '.he reactor showed a rute of
only about 20%. There appear to 'ie three factors contributing
lo the rapid increase in attrition resistance before the start of nor-
inftl cracking operations: pas-age »f catalyst through the line
from the storage bin to the unit, ac. ation in the unit, and heat.
The attrition rate of this catalyst decreased further (luring the
first 2 weeks of operation unlil Inter samples showed a breakdown
of loss than 10% under the attrition-test conditions. Differences
iu operating conditions between tin: commorcial-unit runs with
these two catalysts do not permit an exact comparison of their
attrition resistance at equilibrium ic a unit.
The larger part of the decrease in Attrition rate shown by these
ground catalysts during commorcial-unit operation results from
elimination of weak particles and tlie rounding-off of the rough,
ifr.-'ijuliir particle surfaces by mechanical wearing. During oper-
ation in a fluid-cracking pilot plant" in which erosion conditions
were much less severe, the decrease i t attrition rate of the ground
»ilic.vMmimi.\nnd filica-magncyia r: lalyfts was much hss. The
»iiiaUi-r c\leiit of attrition in the pilot phut rvsultcti from oper-
ation «i:h lower tranter-line vi-lovi ios and U:-e of filu/s instead
of cyclones for catalyst recovery. n Figure- 5 are shown data
fiom pilot -plant runs with fresh bate' K-S of the Mine catalysts that
were used in the enmmiTri.il unit. With silica-alumina catalyst*
following an initial decrease in nltiition rate from -10% for the
fre>h catalyst to about 3C% at (|1C .start of the run, which rep-
f s the off. ot of temperature, th re was a further decrease to
only about 29% lifter 2 weeks of operation. Be-
rausc the pilot plant was operating under the fame
cracking process conditions as the commercial unitj
which showed a decrease in attrition rate to 12%,
it is concluded that e.xposure to heat and alternating
reaction and regeneration cycles has relatively little
effect on the attrition rc.-:istance of this catalyst.
The decrease in attrition rate from about 36 to 12%
in commercial-unit operation as a result of mechani-
cal action is similar to the rapid decrease which was
shown in Figure 3 to occur io the laboratory test.
The ground silica-magnesia catalyst showed a
greater initial decrease in attrition rate in the pilot
plant than did the silica-alumina, but considerably
less than the decrease in the commercial operation.
The laboratory attrition,rate fort-ilica-magnwia.sam-
ples after 2 weeks of operation in the pilot plant v/ag
19%, compared to about 7% for catalyst from the
commercial unit. The greater effect of temperature
on the silica-magnesia catalyst is the explanation.for its greater
initial decrease in attrition rate. It can bc-cstimatod from extra-
polation of the silica-magnesia curve in Figure 5 that the initial
decrease in attrition rate from 35 to 2i% corresponds to the
initial effect of temperature. The decrease in attrition rat* in
the commercial unit from 25 to 7% wouM therefore represent
largely the effect of mechanical action...
FIGURE A-70
1 '
1 FRESH
1 ^^.^ — SILICA- ALUMINA ^
f
SILICA -ALUMINA
5 10 15
DAYS OPE RATIOS
20
Decrease In Attrition Rate during Opcrat'on
of Fluid-Cracking Pilot Plant
CONCLUSIONS-
In a laboratory accelerated-attrition test for me..ifuring the
relative re.-is.ance '° breakdown of .fluid-cracking cataly.-'t:! the
catalyst is subjected to the act ion-of a high-velocity aii jet, and
the amount of breakdown is determined by the change in par- icle-
t>izc distribui »n. Such datn are useful in)lhe estimation of rela-
tive loss rat*-; in commeniiil operations.
The at I ri l on reiiil.inccs of fresh-around commercial f lira-
alumina, silici-iiiagnctia, and natural clay cracking catalysts are
in the same r mge.
Microsphe^uidnt catalysts *how a considerably belter attrition
resistance th:in tround catalysts, because of their improved par-
ticle shape. The presence of agglomerates formed during iranu-
fncture can greatly reduce this advantage.
Heat treat.ncnt of frc.>h catalysts prior to attrition t»lin<;h;i»
a varjing eff.-ct for different caVilysts. The attrition roifti:«nce
of the ground catalysts increased after heating at 1100* F.,
whcicns that of the micro«pheroidal catalysts showed liltl*
change. Af rition ilata on heat-treated Mniples thrrvfore pritide
a more ri-ali. lie evaluation of nmv catal.vst*.
The attiil on resistance
-------
120}
INDUSTRIAL AND ENGINEERING CHEMISTRY
Vol. 41. No. 6
ton \ilcd off ami weak particles "liininatcd. Tlic attrition causes
a icutuliiig oil of sli:irp edges of particles similar to that occurring
in 8 commercial fluid catalytic Bracking unit.
Die attrition resistance of ground catalysts increases rapidly
during operation in a commercial unit. Mechanical erosion of
par .iclcs u the primary cause of this increase. AYith a ground
silica-magnesia catalyst the effect of process temperatures is sig-
nificant.
The presence of fine particles in a catalyst reduces the attrition
rate of the coarser particles, and a correction to a fines-free charge
bas,s helps to compensate for this effect. The most valid com-
parisons of attrition resistance arc obtained when the catalyst
gan pies have comparable fines contents.
1 he particle size of the fine material produced during the attri-
tion of a catalyst in the laboratory test is predominantly less
thati 10 microns with the exception of some agglomerated cata-
lysts. The particle-size distribution resulting from the attrition
of f.n agglomerated catalyst depends on the size of the individ-
ual particles in the agglomerates.
ACKNOWLEDGMENT
It is desired to acknowledge the development of this
test method by J. T. Clap;> anil J. 1J. Gray. Appreciation U
also expressed for many helpful suggestions bv A. L. Conn ami
J. 0. Howe.
LITERATURE CITED
(1) Inncs. W. B., and Ashle,-, K. D., Proc. Am. Petroleum Iiut.,
27, 1III1.9-17 (1917).
(2) Mathe»oi>. G. I,., /lid., pp. 18-22.
(3) Murphree, E. V., Brown, C. L., Gohr, E. J.. Jnhnig. C. E.. Martin.
II. Z., and Tys>on, C. W.. Tram. Am. Inst. Chcm. Kngrt.. 41,
19-33(1915).
(4) Roller, P. S.. J. Am. Ctram. Soc.. 20, 167 (1937).
(5) Holler, P. S.. Tram. Am. Soc. Testing Malcrio.lt. 32. 607 (1932).
(6) Roller, P. S.. U. S. Bur. Mines, Tech. Paper 490, 40 (1931).
(7) Webb. G. M.. Petroleum frofti-iiny. 2 (7). 497 (1917).
RECEIVED January 9, 1049.
-A345-
-------
APPENDIX B
EFFECTS OF OXYGEN AND WATER VAPOR ON SO? SORPTION
The following discussion, which is taken from the AVCO
Final Report "Removal of S02 From Flue Gas", Contract No.
PH86-67-51, dated November 1, 1967, appears to adequately
describe the subject effects.
Effect of Moisture..and Oxygen in the Flue Gas
A large amount of data vas taken early in the progran on the mistaken assump-
tion that sorptlon rates were unaffected by the presence of water vapor In
the flue gas. The assumption was based on a series of experiments which were
apparently in error. All recent data, and the majority reported on here were
taken with a flue gas containing 5-7 percent moisture by volume.
The effect of water vapor was demonstrated by sorptions S-51 through S-55
which were run at 300 c and about 1 percent SO.. A 216 tng sample of. Peter
Spence sorbent was used. Run S-51 used dry flue gas and showed a weight gain
of 15 mg after 100 ninutes. Run S-52 used flue gas with 5.2 percent moiature
and gained 96 mg in the same period. In S-53 the dry results of S-51 were
replicated at a high oxygen concentration and slightly higher rates were
-B346-
-------
TABLE A-65
SUMMAR'' DF SORPTION DATA
I
W
Co
Run No.
S-47
S-49
S-51
S-52
S-53
S-54
S-55
S-62
S-63
S-64
S-65
S-67
S-68
S-69
S-70
S-79
S-80
S-81
S-82
S-83
S-85
S-86
Sample
5
'5
6
6
6
6
6
9
9
9
9
9
9
9
9
12
12
13
13
13
14
14
T(°C)
300
300
300
300
300
300
300
300
300
300
350
200
250
150
350
300
300
300
300
300
300
300
Caa
Velocity
(ft/sec)
.5
.5
.5
.64
.5
.66
.64
7.3
1.9
1.9
1.9
1.9
1.9
1.9
1.9
7.6
7.6
7.6
7.6
7.6
7.6
7.6
Vol. Percent
S02 '
1.03
eat 1.0
1.03
.83
esc 1.0
0
.78
.90
.89
.87
.95
.68
.87
.91
.87
.81
.96
.85
.92
.81
.96
1.02
Vol. Percei
H20"
0
0
0
5.2
0
7.2
5.2
3.5
7.2
7.2
7.2
7.2
7.2
7.2
7.2
.7.2
7.2
7.2
7.2
7.2
7.2
7.2
Vol. Percent
02* »
15.5
3.6
3.6
16.6
21.0
0
2.85
3.27
3.145
3.145
3.145
3.145
3.145
3.145
3.145
3.15
3.15
3.15
3.15
3.15
3.15
3.15
Particle "
Diameter
(in.)
.120
.120
.102
.102
.102
.102
.102
.102
.102
.102
.102
.102
.066
.066
.066
.079
.079
Theoretical
kc
(ft/hr)
505
505
1050
1050
1460
1460
1880
1880
1880
1*70
1670
fcc frou
data
(ft/hr)
261
161
117
105
103
151
179
179
171
148
252
145
391
498
K
from data
(hr-1)
139
118
332
174
197
312
442
D
from data
(cmZ/eec)
.012
.024
.024
.014
.014
.0099
.012
.0?"
.013
.0062
.0071
.0085
.012
.010
•tfrom pressure change on addition of 02 to feed cylinder
t by chromatography
>t by sieving
** from humidifier temperature
-------
observed. In S-54 the sorbent was exposed to a mixture containing 7.2 per-
cent HO and 92.8 percent N2 by volume. The weight gain was only 1 rag. The
results of S-52 with moist flue gas were replicated in run S-55 at a low
oxygen concentration. Sorption curves for each of these runs are given in
Figures 25 and 26. Note in Figure 26 that raising the oxygen concentration
from 2.8 percent to 16.6 percent had no effect on the sorption in moist flue
gas.
Thus, increased concentration of 02 gave slightly higher sorption rates in
the absence of moisture in the gas, but there was no effect of oxygen content
with moist gas.
An equal increase in rate over previous dry sorptions was observed with the
W. R. Grace material. The effect of water vapor on W. R. Grace sorbent can
be seen by comparing runs S-49 (dry, 300° C) with S-79 (7.2 percent moisture,
300° C). The sorption curves are presented in Figure 27 and clearly show the
increase in sorption rate in the presence of moisture.
The mechanism by which water vapor accelerates the sorption is not understood.
The rate of the surface sorption reaction may be catalyzed by the presence
of water. Alternately, capillary condensation of the moisture might allow
rapid sorption of SO. in the condensed water. A third possibility is that- a
surface diffusion mechanism might be activated by absorbed water.
The possibility of capillary condensation may be checked with the Kelvin
equation which relates vapor pressure lowering to the radius of curvature of
a curved surface.*
i. !l . 2-i JL
p rp RT
where
pg » vapor pressure of water at ~f
p - partial pressure of H-0 in gas phase
y • surface tension of water at T
p = density of condensed liquid
R - gas law constant
M m molecular weight of condensed liquid
T m temperature
r * pore radius
GUsstont, Textbook of Physical Chfmincy, v»n No«tf»nd, Princeton, N.J. (1946) p. 49J.
-B348-
-------
d,
OJ
*h
vo
•4
0»
I
O S-51 Opct H20,3.6ocf Oz,fpct S02
X s-53 Opct H20, 21 per O2,1 per SO2
A S-54 T.2 per H20 IN NITROGEN
NOTE'
W = TOTAL WEIGHT GAIN ASSUMED TO BE S03
40 60
t, minutes
IOO
FIGDRE B-71 SORPTION DATA FOR SAMPLE NO. 6
-------
-4
I
I
(0
U)
en
o
I
KU
too
r"
: *°
>
j
40
20
0
0
S7-8C
£
o
O
/
XJ
)O
K>
S>
>
rfc
0
f
ox
0^
p
x°~~
0 x
,)0-^X
c
)
. X—
=»x •
-X— X
) S -52, 16.6 pet 02 IN FLUE
C S -55. 2.8 pet Oa IN FLUE
NOTE:
W = TOTAL
-X—
GAS
GAS
WEIGHT GAIN ASSUMED TO 8
i— X—
i S03
ZO 4O 60 80 100 120 140 160 160 20O
ra t, minutes
FIGURE B-72 - SORPTION DATA FOR SAMPLE NO. 6
-------
tx)
OB
1
120
too
r80
E
! 60
c
2
£
4O
2O
O
•7-8
j
0°
O
O
O
)
I
-x— x—
r
X — X
3"-1"
x — -x-
r-o--
i
-0— <
>—
O S- 79 , SAMPLE NO. 12 , 277 mg SORBE
MOIST FLUE GAS
X S-49. SAMPLE NO. 5, 225 mg SORBE
DRY FLUE GAS
YOTE :
W = TOTAL
-x
•X
x-
WEIGHT (
X—
JAIN ASSU
X
MED TO B
— X
:NT,
NT
E SO3
0 2O 4O 60 8O IOO I2O I4O I6O 180 200
BZO t, minutes
FIGURE B-73 - SORPTION DATA FOR WET AND DRY FLUE GAS
-------
The maximum pore radius which could support capillary condensation of mois-
ture was computed using typical sorption conditions.
TABLE B-66
TfK)
423 (150° C)
573 (300° C)
Po (atm)
4.70
84.8
p (atm)
.072
.072
/(gm)\
<>[ ; cc )
1.0
1.0
Advne)^
>\ cm /
47
20
o
r( A)
1.2
0.21
The concept of capillary condensation in pores of one molecular diameter is
not sensible. As a result capillary condensation is probably not the expla-
nation of the high sorption rate in the presence of.moist flue gas.
Alternately, sorbed moisture might provide a path for surface-diffusion of
SO- on the pore walls. If this were an important factor, it could be checked
by stopping the sorption periodically by sweeping tha"sorption chamber with
N£. In the presence of N? the surface diffusion would be expected to con-
tinue causing a levelling of the sorbate concentration-profile. Restoring
the sorption should result in a different rate than.previous to'standing in
N.'due to the redistribution of the sorbate within-the particle.
All sorptions from S-78 on have been run on an intermittent basis. Sorption
was carried out for a given time, then the sample-charaher*swept with nitrogen
and the sample weighed in static nitrocen. Sorption was then resumed in
been continuous. There are no evident differences in the sorption curves, or
diffusivitics obtained in the diffusion controlled region (high loading
region) of the sorption. This observation suggests that surface diffusion
is not a large contributor to the total diffusion process.
By elimination, the implication is that moisture catalyzes the sorption step
at the sorbent surface but this conclusion is by no means based on adequate
testing.
The'rate of sorption curve, given in Figure 28, indicates that there is a
large jffect of moisture on the sorption rate at loadings greater than two
percent of saturation. However, there was no increase in the initial i.ite
when noisture was added. In fact, the data show a decrease, but this my
not bt significant in light of the poor resolution of initial rates by these
experiments.
The mclsture levels used, 5.2 percent and 7.2 percent, are similar to those
found in actual flue gas. In the regi-n where pore diffusion is contm-linp,,
higher moisture content should not affect the sorption rate. The mecfai.iisn
of tht- Initial rate is not known, nor is the effect of water content on this
rate Vnown exactly.
Enrimnicd fro.. v»lucs given m trmpemuret up to 100°C.
-80-
-B 352-
-------
O S-47
X S-49 I GAS,300*C
0.0001
0.001
S7-80ZI
0.01
w
FIGURE B-74 - SORPTION RATE DATA, W.R. GRACE SORBENT
1.0
-79-
-B 353-
-------
APPENDIX C
NITROGEN OXIDE CHEMISTRY:- THERMODYNAMIC CALCULATIONS'* .>
The following discussion of nitrogen oxide chemistry was-.
presented by the W. R. Grace Co. in the September 1968 monthly
report, Contract No. PH-86-67-129. It is included here since -
1) it shows some of the reactions that might occur during sorption: r
of SOp by alkalized alumina, and 2) Grace developed these-cal-
culations in support of their theory of absorption. Although •
the exact mechanism by which sorption of S02 from stack gases", »
occurs has not been agreed upon by all investigators,>Grace's
theory (i.e., oxidation of S02 to SO., is substantially accelerated
by N0x and the resulting S03 reacts rapidly with the basic,surface
sites on the alkalized alumina) is certainly a distinct possibility,
-C354-
-------
W. R. Grace Discussion
The stack gases from which sulfur dioxide is absorbed by
alkalized alumina contain oxides of nitrogen. In order to
substantiate the postulated theory of absorption, it became
necessary to identify the state of oxidation of nitrogen.
The data summarize the thermodynamic equilibrium considerations,
and a discussion of the thermodynamic feasibility of some of
the probable reactions is presented.
The reactions considered are:
(1) NO + 1/2 02 ~»- N02
(2) 2N02 + 02 y-g. N204
(3) S02 + 1/2 02 *•* S03
(4) NO + SO- —-^ SO, + 1/2 N_
f. J £•
(5) N02+ S02 -—*. S03 + NO
The standard free energies of reaction for the temperature
range 300°K(80°F) to 700°K(800°F) for these five reactions are
given in the attached data sheets, one through five respectively.
Data sheet six shows the concentrations of components when reac-
tion #(1) is superimposed on reaction #(5) at equilibrium at
600°K (620°F).
The assumptions of ideal components and ideal mixtures were
made in the calculations. The flue gas at one atmosphere was
assumed to contain 5% 02/ 0.05% mole equivalent of nitric oxide
(NO) and 0.3% sulfur dioxide (S02), the balance being inert gases.
-------
The thermodynamic considerations predicted that when the
stack gases contain 5% oxygen and 0.05% nitric oxide (NO), it
is possible that equilibrium exists between nitric oxide and
nitrogen dioxide (NO2). At 700°K (800°F) 0.015% nitric oxide
can come to equilibrium with a maximim 0.035% nitrogen dioxide.
Dimerization of nitrogen dioxide to dinitrogen-tetroxide
is thermodynamically not feasible between 400°K (260°F) and
700°K (800°F). However, it can take place below 325°K (125°F).
-C356-
-------
' in the temperature range considered, reactions # (^) and # (5)
arc feasible. The high negative value for the standard free energy
of reaction of '# (k) as compared to # (5) does not necessarily predict
that'. ,f (l>] is favored to take place. it merely indicates chat the
reaction # (l\) at equilibrium is probably more complete than reaction
ft (5) at the same temperature and equivalent concentrations.
':Based on the above information the following sequence of steps
for the absorption of S02 was considered.
NO + -|02 - N02 (1)
N02 + SOo ss NO + £03 (5)
S03 .H;-2N,aA.10'2 -. Na2S04 + A1203 (6)
Attempts were made to superimpose the two gas phase reactions
involved in the above sequence of steps. It was assumed that at
600°K (620°F), reaction -// (1) attains equilibrium and the concentrat-
ions of NO 0.00375$ and of N02 0.0'l63^ were assumed to be the initial
concentrations of the respective gases . for reaction £ (5). Sulfur
dioxide concentration of 0.3/S was assumed. Equilibrium analysis of
reaction # (5)' under these conditions at 600°K (620°?) indicates that
the forward reaction is 99. 9£ complete. The equilibrium concentrat-
ions of the components of the reaction mixture assuming ideal gases
and ideal mixture are shown below.
Mole £ Vol 55
] = O.Olg .005
[S02] = 72.6$ .2& .
[NO] «= 14.2£ .050
[S03] = 13.2g -0^6
It must be noted that the above calculations are based on
thermodynamics only. The actual occurance of any of these reactions
be governed by the kinetics of the competing or series reactions
-C 357-
-------
Data sneet
NO + 1/2 02
Nitric
Oxide
Nitrogen „
dioxide
Assumptions:
Reference:
(1) Ideal gas and ideal mixture
(2) Stack gas containing 0.05% mole; equivalent of
NO and 5% oxygen at 1 atmospherer,
JANAF Thermochemical Tables
Tgmp.
300
A 00
son
600
700
^mP-
80
260
MO
620
800
Std._.Frec Energy
of Reaction
AF kcals/mole
-8.4
-6.6
-4.8
-3.0
-1.2
Equilibrium
Const-ant
^ K ^
4300
12.56
0.37
X 7,
100
93
70
Conclusions
(1) The forward reaction is .feasible in the-temperature
range.
(2) Thcrmodynamieally at. 300°K, it is possible that the
state of oxidation of nitrogen is completely nitrogen
dioxide (N02) and at 700°K, equilibrium may exist for
the stack gases-to contain 0.0357, nitrogen dioxide
(N02) and 0.015% nitric oxide (NO),
-C358-
-------
Data Sheet (2)
2NO,
Nitrogen
dioxide
Dinitrogen
tetroxide
Assumption:
Reference:
Ideal gas and ideal mixture
JANAF Thermochemical Tables
Tgmp.
Is.
300
400
500
600
.700
80
260
440
620
800
Std. Free Energy
of Reaction
A_F kcals/cnole .
+3.1
+7.2
+11.3
+15.4
Conclusions
By interpolation, the forward reaction is feasible
at temperatures less than 325°K cr 125°F.
-C359-
-------
Data Sheet (3)
1/2
SO,
Reference: International Critical Tables Vol. 7, page 236
AF° = Standard Free Energy of Reaction, kcals/inole
= -22.6 + 0.02136 T at T °K.
Temp.
°K
300
400
500
600
/uu
Temp .
Op
80
260
440
620
ouu
Std. Free Energy
of Reaction
AF° kcals/mole
-16.2
-12.1
-11.9
- 9.8
- v.6
Std. Free Energy t of
Formation of S02*
'AF0 kcals/mole
-71.7
-71.9
-71.9
-71.8
-/J..6
Std. Free Energy
Formation of SO,"
AF° kcals/mole u
• -87.9
--84.0 '
-83.8
-81.6
-/y.i
^Reference - JANAF Thermochemical Tables.
?°.
so.
AF
Reaction
Conclusion: The forward reaction is feasible in the temperature range..
-C 360-
-------
tictua i>ncuu
NO •»- SO,
Nitric Sulfur
oxide dioxide
1/2 N,
Nitrogen
so3
Sulfur
trioxide
Assumption:
Reference:
Ideal gas and ideal mixture
(1) JANAF Thermochemical Tables
(2) Data Sheet (3)
Std.. Free Energy
of Reaction
AF kcals/mole
-.36.9
-32.5
-32.0
-29-6
-07.1
Temp.
°K
300
400
500
-600
•TAA
ToemP.
80
260
440
620
800
Conclusions
The forward reaction is feasible up to almost
completion in the temperature range considered,
-C 361-
-------
Data Sheet (5)
S0 ^=±: NO - S0
Nitrogen Sulfur Nitric Sulfur
dioxide . dioxide oxide trioxide
Assumptions:
Reference:
Ideal gas and ideal mixture.
(1) JANAF Thermochemical Tables
(2) Data Sheet (3)
Temp.
£K _
300
400
600
700
Temp.
°F
80
260
WO
620
800
Std. Free Energy
of Reaction
AF° kcals/mole
-7.8
-6.8
-6.5
Conclusions:
The forward reaction is feasible in the
temperature range considered.
-C 362-
-------
Data Sheet. (6)
NO -V 1/2 0 ^rr=±^ NO,
N0
S0
NO + SO,
(1)
(5)
Assumption:
Reference:
(1) Ideal gas and ideal mixture
(2) Stack gases containing 0.05% mole equivalent
of NO, 5% oxygen and 0.3% S02 at 1 atmosphere,
(3) Reaction //(I) at equilibrium at 600°K
Data Sheets (1) and (5)
NO + 1/2 0
Initial Concentration 0.05%
1 mole
2 -v-
5%
100 moles
NO,
at 600°K
At Equilibrium 0.07 moles 99.54 moles 0.93 moles
Concentration Vol.# 0.0037% 0:0463%
N09 +
Vol.g 2
Initial Concentration 0.0463%
1 mole
SO
2 ^-~
0.3%
6.5 moles
NO +
0.0037%
0.08 moles
SO,
at 6(
At Equilibrium
Concentration Mole
0.001 moles 5.501 moles
0.013% 72.6%
0.251$
1.08 moles 0.999 moles
14.2% 13.2%
-C 363-
-------
APPENDIX D
DERIVATION OF FIXED BED SORPTION MODEL
The following derivation of a fixed bed sorp
-------
KINETIC STUDIES OF SQs ABSORPTION BY ALKALIZED ALUMINA19
.Introduction
The purpose of this investigation was to develop a mathematical model
which would accurately describe the removal of S02 from flue gas by alkalized
alumina in fixed- and falling-bed reactors. Rate constants associated with
the reaction were determined from data obtained in bench-scale tests. Using
the model, these data were analyzed with the intention of gaining some insight
into the factors which affect the reaction under various operating conditions.
It is felt that such an understanding will be of great help in predicting and
analyzing the behavior of large-scale systems operating under similar
conditions.
Mathematical Analysis for Fixed-Bed Reactor
Nomenclature
x = distance into column, ft
t = time, hr
y = SOa concentration in flue gas, by weight, dimensionless
p, = S0g concentration on absorbent, by weight, dimensionless
pB = gas density, lb/ft3
p, = solid bulk density, lb/ft3
F = void fraction of bed, dimensionless
v = gas velocity, ft/hr
A = cross sectional area of column, ft3
G = pG'vF = flow rate of gas, lb/hr(ft)3
S = feed rate of solids, Ib/hr (ft3)
K = reaction rate constant
A gaseous mixture containing S03 is passed continuously upward through a
uniformly packed column of granular material. It is assumed that the concen-
tration of SOa in the gas at the column inlet is constant for the time the
column is in operation, and further that the gas flow rate (G) also remains
constant. The sorbent material is initially assumed to be uncontaminated by
SOa, and finally we assume a constant temperature throughout the system. For
such a situation, a material balance on SOa over a horizontal cross section of
height &x gives
(Rate of accumulation of SOa) [(Input-Output)/unit time]
(on solid) (in void)
~ [A(Ax)ps(J, + A(Ax)pGyF] = ~ tpGyAvF](Ax) (1)
19Prepared by M. J. Lempel, mathematician, Pittsburgh, Coal Research Center,
Bureau of Mines, Pittsburgh, Pa.
-D365-
-------
This equation reduces to
1 ay = _ jpj. au. (2>
v dt G dt
It can be shown that the term 1 Q. may be neglected when the void volume of
v dt
the fixed bed is negligible compared to the total volume of gas mixture passed
through the bed. Therefore, equation 2 becomes
ax G at
A second equation is obtained by postulating the rate at which'SOa-is absorbed
by, or reacts with, the alkalized alumina. We assume that this -rate is pro-
portional to the product of concentration of S02 in the gas with the .unreacted
Traction of the solid. This is expressed by
at
where p,, = total capacity of r.orbcnt (Ib/lb) . The constant K appearing in
equation 4 is not a true constant r.incc it will depend upon such, things as
Equations 3 and 4 arc to be solved for y(x,t) and u(x,t) subject to the
boundary conditions,
y(o, t) = y0 n(x,o) = o (5)
where y0 = SOa concentration in gas at column inlet. Omitting the details,
the solution is stated as follows:
At
M, « Mo (l -
where A = ——, V, = —i-L.
-D366-
-------
APPENDIX E
DERIVATION OF DISPERSED PHASE SORPTION MODEL
The following derivation of a dispersed phase bed sorption
model was contained in the Air Pollution Research Progress
Report, Pittsburgh Coal Research Center, to Public Health
Service for the Quarter Ended March 31, 1968 (i.e., USBM-
Bruceton). The model shown was used in the present design
of the dispersed phase process.
-E367-
-------
Q » quantity of solids conta.ct.ed, Ib
r « sorption rate, hr~*
t « time,-hr
V - flue gas flow rate, SCFH
v « volume of column, ft3
w • S02 concentration on sorbent, weight fraction
x «= sulfur concentration on sorbent, weight fraction
y - sulfur oxide concentration in flue gas, mole fraction
Z - sorbent condition parameter
p »» average sorbcnt density, Ib/ft
Calculation of•sorption rate
Hie average sorption rate (r) observed within the column is
expressed as the weight of S02 removed per weight of sorbent per hour.
r •= -f (1)
Sorbent weight is defined as the total weight of solid rather than that
of sulfur-free alkalized alumina.
A material balance on S02 over the column height gives
Ib S02 sorted = 0.178 (yo-yf)Vt (2)
uhere the subscripts o and f indicate the column inlet and outlet,
respectively, and t refers to gas contact time. Equation (2).neglects
the change in gas volume due to sorption of S02 and assumes.' plug gas
flow in a constant "diameter tower. The quantity of solids contacted
IB given by
Q - Pv (3)
and if pressure loss due to wall friction and sorbcnt momentum is
neglected
P " 0.1924h '
Comparison of measured solid densities and those calculated by equation
(4) is shown in figure 1. Deviations are believed tg^be primarily due
to the following factors:
a. Expcrimc-nt.il error in measurement of dilute solid density.
b. Measurement of pressure loss over a portion of the column
height, assuming pressure loss to be proportional to height,
an assumption valid only for uniform solid density.
-E368-
-------
UJ
UJ
u.
0
in
2 2.O
.
z
UJ
m
cc
O i 5
Q
P x
UJ JQ
0 •£ |.o
V)
O
— i
Ul
3
frt
CO
UJ
o:
a.
I'll,
.*
/
O ^
— x __
o y
/
*
s
s
O CD '
— OOCDQ, —
O X O O
X
/
O / •
xX L_p-o,924 h
x
XCD
x
x
— x
X
x
0
X
X
xX 1 1 1 1
0.05 0.10 0.15 0.20
SORBENT DENSITY, Ib per cu ft
0.25
FIGURE E-75 -
Measured and predicted sarber presisure
loss.
3-15-68
L-IO477
-E369-
-------
Substitution of (A) into (3) and use of the column cross -sectional
area in place of — gives
h
Q = 5.198pa. (5)
Since w is defined as (2)/(5), insertion into (1) feives the expression
for the average sorption rate:
_ 0.0342(y -yf)V
r -- _ - . (6)
Determination of Sorption Rate Constant
It is assumed that the sorption rate can be adequately repre-
sented- as proportional to the product of two terms, one dependent only
on sorbent condition and the second dependent only on flue gas con-
centration. Thus
dw
r - d7 - Kf(Z,y). (7)
If it is further assumed that sorbent is evenly distributed and- contains
a uniform quantity of sulfur throughout the column, Z becoaaes a constant
dependent on operating conditions and sorbent quality. Equation (7)
can thus be integrated and expressed as
KZy (8)
o
where y is the mean concentration in the flue gas, 'defined as
_ I f
y - t J
(9)
For plug gas flow at constant temperature and pressure, t may be re-'
placed by h. If in addition f(y) can be approximated by yn such that
>
equation (9) may. be replaced by
**f
y '• h J Ly0n~l + (n-i) chj1'11 dh.
. —„ . . - (11)
ho
-E370-
-------
dy
Equation (10) predicts that a plot of logiodl versus ^°Sioy wil1 ***
a straight line of slope n. When averaged data of y versus h, pre-
sented in last quarter's report and reproduced here as figure 2, was
graphically differentiated and tested in this manner, the result shown
in figure 3 was obtained.
That a straight line was observed, indicates that the mean gas concen-
tration predicted by equation (11) adequately describes sorber operation.
The significance of the value found for n—approximately 2/3--io, however,
not clear. Laboratory data indicate n - 1. Differences may be due to
the inaccuracy inherent in the constant sorbent condition assumption or
may be related to a change in reaction mechanism in the presence of
nitrogen oxides. Further investigation is indicated.
Integration of (11) for n = 2/3 and insertion into (8) gives
where c is evaluated using the integrated form of (10)
.-*k1/V/Sl 03,
The sorbent condition parameter, Z, is expected to be a function of
particle size and sulfur concentration distributions within the column.
For the tests considered, however, this information is not available.
As an approximation, Z is assumed to be a function of sulfur concen-
tration alone. In addition the sorbent contacted is assumed to consist
of particles having a linear loading distribution from the inlet to
exit concentration. Thus
(14)
If Z can be approximated by f 2£p£ J , equation (8) may be written
10*10 • 10810K
¥
...x -x ^
and a plot o£ !OBIO ~ versus log1Q ( -fi—J-J will be a straight line
of slope m. The result of such a plot is shown in figure 4. The da£a,
at least for test series E and F, may be approximated as a straight line
of slope = -1 and in = -1 is assumed for all tests.
-E371-
-------
O 10 20 30 40 50
SORBER HEIGHT, feet
FIGURE E-76.-Voriotion in sulfur oxides concentration
with sorber height.
L7
o
o
o>
CL
8
u
o
o
E
O
eT
3
1.6
1.5
1.4
1.3
I
1
I
I
I
L
_ 2
Slope = ^
2.8
2.9 3.0
JL.
3.1
3.2 3.3
_6
3.4
LOG,0y, mole froction x 10
FIGURE E-77 -Test for predicting n in the equation
*1
dh
•=cy
3-15-68
L-10478
-E3*72-
-------
3.3
3
O
g. 3.2
•'c
Q>
JD
° 3.1
JC
.S?
*C>
£ 3.0
CM
O
£ 2.9
o«
o
O
2.8
2.7
0.8
O
o
0.9
1.0
1.1
oTest series C and 0
OTest series E and F
1.2
1.3
1.4
LOG,0 X x 10 , weight fraction
1.5
-f r Xo + X f|
FIGURE K-78 -Test for predicting m in the equation •— = k ' —
y *• €. J
m
3-15-68
L-IQ479
-E373-
-------
The overall rate expression, then, is
- . 2Kr .-. . „„
L Vt
Values of F and K are tabulated in table 3 for the tests considered.
For series C and D, K equaled 9.92 and for series E and F,, 2 1.63.
Standard deviations were 1.97 and 3.92, respectively.
The difference in calculated values of K is significant.., The assumption
of a linear sorbent loading distribution is extremely optimistic.
Actual sulfur content of sorbent within the column is higher than pre-
dicted , by this assumption. Results indicate that this. ideal condition
was more closely approximated in test scries E and F where -the total
overhead stream was recycled than in C and D where a portion was ••removed
after a single cycle. During the past quarter, the overall composition
of sorbent in-. con tact with the gas was determined over a wide range. of
operating conditions . The data should result in a more realistic ex-
pression for the sorbent condition term.
Prediction of Column Performance
The expressions developed should, over the range of ^conditions con-
sidered, predict column performance within about 20 percent. Sulfur
oxides removal, pressure loss, flue gas flow rate, and-, initial and
final sorbent loading arc related by equation (16)
, i-,
»Eyo"1/2cy
where F and c are defined by equations (6) and (13), respectively
_ 0.0342(yo-yf)V
pa
Finnl Borbont In.idinK mny ho rcplncrd by .sorbent food rnto if Borpl,(on
in fiHiiiiiiiril Lo occur tin oulfftlu, uiuco fi'om « SOy ma to rial balance
x m x0LfHO.Q890(yo-yf)V
£ Lf40.2225(y0-yf)V
-E374-
-------
TABLE - E-67 - Calculated values of sorption rate
Test No.
C401
D*04
D115
D120
D121
D122
D123
D124
E101
E102
E103
E104
E105
E106
F103
F104
F105
F106
F112
and sorption
r
1.34
1.66
0.94
1.38
1.48
1.36
1.17
1.25
1.24
1.16
1.36
1.15
1.16
1.05
1.66
1.60
0.95
1.28
1.34
rate constant
K
8.51
12.20
7.06
7.49
10.64
11.87
9.39
12.21
22.86
18.60
23.72
20.20
16.95
15.73
22.66
23.61
18.27
26.03
29.34
-E375-
-------
APPENDIX F
REGENERATION DATA
A summary is presented of the available data on regenera-
tion of alkalized alumina. Sources of the data are indicated
on the individual sheets. For a complete discussion of the
data the-original reports should be consulted.
-F376-
-------
TABLE F-68
RATE OF CONSUMPTION OF REDUCING GAS
SAMPLE
NO.
-4
-J
1
-
26.
28.
32.
33.
34.
38.
40.
46.
49.
54.
43.
Note
Ref :
REDUCING
GAS
Pilot plant spent
Pilot plant spent
Pilot plant spent
Pilot plant spent
Pilot Plant spent
Pilot plant spent
Pilot plant spent
Bench-Scale spent
Bench-Scale spent
Pilot plant fines
Pilot plant fines
: Initial charge was 4
USBM-Pittsburgh-Sec .
H2
H2
H2
H2
CO
CO
CO
H2
CO
H2
CO
grains .
II
SULFUR,
BEFORE
2.5
2.5
2.5
2.5
2.5
2.5
2.5
6.5
6.5
8.4
8.4
PERCENT
AFTER
0.7
0.6
0.9
1.4
2.0
1.7
1.8
1.6
4.0
2.1
7.0
RATE OF CONSUMPTION (CC/MIN)
600°C 620°C
1
1
2
1
3 7
2 8
4 9
2 4
5 14
3 6
9
640°C
6
7
6
7
11
12
13
11
24
11
16
eeo'c
—
10
9
11
21
22
31
34
28
25
680°C
—
—
19
20
34
35
50
57
41
42
3/31/68 Quarterly Report - Table 1
-------
APPENDIX F-69
00
I
REGENERATION OF
Test
NO.
82
81
83
79
80
86
85
84
74
73
90
91
92
Reducing
Gas
H2
H2
H2
H2
H,2
CO
CO
CO
CO
CO
(H2 + )
(CO )
(H,+)
(CO )
(c3 )
GRACE NO. 1 SPENT SORBENT FROM PILOT PLANT MIXED PARTICLE
Temp.
620
640
660
660
660
640
660
680
660
640
640
640
660
Flow
(cc/min. )
57
55
53
52
52
51
9\
93
61
62
(38 + )
(19 )
(19+)
(40 )
(26+)
(28 )
Space
Vel.,
(hr'1)
850
825
790
780
780
770
1360
1390
920
930
850
860
800
Max.
Consump.
(cc/min. )
8
17
28
39
29
24
59
75
—
28
23
24
43
Total
Consump.
(cc/hr . )
350
810
805
730
755
600
708
765
720
755
615
750
Percent
Final
3.2
1.8
1.4
1.4
1.7
2.2
2.6
3.3
3.3
2.1
1.7
1.65
1.65
SIZES,
S
Removed
45.2
71.4
77.6
77.1
72.4
63.5
58.6
46.5
45.5
65.4
72.7
72.9
73.2
5.3 PERCENT S*
Reac-
tion
Time
1 hr.
1 hr.
1 hr.
45 min.
1 hr.
1 hr.
1 hr.
1 hr.
80 min.
70 min.
1 hr.
1 hr.
1 hr.
% Wt.
Loss
9.2
15.4
15.2
13.5
13.9
12.2
14.0
14.0
13.9
12.8
15.0
13.0
13.9
* USBM-Pittsburgh-Sec. II-6/30/68 Quarterly Report - Table 1.
Test Conditions: Differential Reactor - 4.25 gm-charge
8 mm bed height
4 cm 3 volume
-------
TABLE F-70
HALF HOUR REGENERATION
TESTS-ORIGINAL GRACE NO. 1 SORBENT*
(9.8 PERCENT SULFUR AND
Test Reducing
No.
109
105
107
i
2 108
2 110**
102
101
103
104
Gas
H2
H2
H
2.
H2
H2
CO
CO
CO
CO
Temp.
Flow
Space
Vel.
(°C) °F (cc/min) (hr-x
620
640
660
680
680
620
640
660
680
1148
1184
1220
1256
1256
1148
1184
1220
1256
53.2
61.5
58.9
73.3
37.0
58.2
58.2
58.3
72.7
709
820
785
977
1000
776
776
111
969
Max.
Consump,
1.0 PERCENT CARBON)
Total
Comsump. Percent S
) (cc/min) (cc/1/2
17.7
31.5
40.4
61.5
31.2
24.5
46.0
52.3
66.1
445
910
1100
1200
675
375
855
1135
1400
Percent
Weight Percent
hr) Final Removed Loss
8.6
5.9
3.4
2.60
2.8
8.8
8.7
9.2
9.6
25.9
53.2
74.5
81.8
80.8
23.2
29.0
30.2
29.3
15.6
22.3
26.5
31.2
32.6
14.5
20.3
25.7
27.4
Carbon
0.4
0.1
0.1
0.2
0.3
0.6
0.3
.0.2
0.1
* USBM-Pittsburgh - Sec. II - 9/30/68 Quarterly Report - Table 1
** 2 gm Sample, Differential Reactor. All others contained 4.25 gms.
-------
TABLE F-71
REGENERATION OF GRACE NO. 1 SORBENT-MODIFIED THIN BED TEST UNIT*
00
o
I
(SAMPLES LOADED IN DEEP FIXED-BED)**
REGENERATION TIME, (MINUTES)
TEMP. (°C)
Sample loaded to
600
650
680
700
700
730
Sample loaded to
600
650
680
700
730
Sample loaded to
600
650
680
700
730
* USBM- Albany -
04 d 16
30 60 120
SULFUR, (PERCENT)
8.8% S.
8.66 8.72 8.50 8.46
9.33 9.26 8.92 8.94
8.10 8.87 8.05 7.23
8.89 7.00 5.73 3.09
8.80 6.36 4.49 2.37
8.47 2.79 1.78 0.72
5.87% S.
5.96 5.77 5.45 5.92
5.59 5.23 5.82 6.07
5.86 5.46 5.34 3.56
5.67 4.52 3.10 1.97
5.70 2.24 0.87 0.61
2.29% S.
1.95 1.93 2.16 2.26
2.33 2.05 1.92 2.03
2.48 1.61 1.62 0.92
2.18 0.91 0.73 0.42
2.40 0.54 0.44 0.25
January 1969 Progress Report -
Test Conditions: 1.5" diameter sample holder
-
8.45 8.75 9.00
8.54 7.42 5.32
5.57 2.70 1.22
1.56 0.99 0.39
1.29 0.85 0.23
0.58 0.28 0.21
4.97 5.23 5.28
5.82 4.80 3.47
3,07 1.73 0.62
1.54 0.91 0.21
0.33 0.25 0.14
2.23 2.22 2.32
1.92 0.87 0.62
0.60 0.24 0.19
0.29 0.15 0.12
0.13 0.16 0.13
Table 1
: Simulated Reformed Reducer Gas-50%N^,
20%H9/ 5%H00; Gas Flow Rate - 0.12 ftj/min @ room temperature, das
** Pellet loads
analyses show
10 f t/min ; Space Velocity -
limited to a single pellet
for the three samples are shown
Rate
Constant,
(Min-1)
.001
.005
.017
.059
.064
.195
.001
.004
.021
.044
.235
.001
.005
.058
_
—
io%co2,
Velocity-
57,600 hr~! @ room temperature with the bed depth
layer (1/8" deep) .
above the zero minute clry head analysis.
the spread in analyses from run to run, probably due to wide variation
The zero
in load on
the individual pellets.
-------
U)
CO
REGENERATION OF
SPACE VELOCITY
REDUCING
TEST
NO.
162
176
177
129
155
157
149
135
138
169
170
154
156
160
136
137
142
141
140
159
GAS Ex
SORBENT
Kaiser PR-15
Kaiser DN-112-F
Kaiser Dd-113-F
Kais.er (B) +
Kaiser DN-105+
Peter Spence +
USBM X-3-F
USBM 423-X-4E
Kaiser PR-15
Kaiser DN-112-F
Kaiser DN-113-F
Kaiser DN-105 +
Peter Spence +
Kaiser PR-15
USBM 423-X-4E
M
Kaiser PR-15
II
ll
II
TEMP(°C)
H-
2
H2
H2
1:2
H2:CO
1:2
H9:CO
1:1
H7:CO
1:1
660
680
640
620
600
660
660
680
640
1000
TABLE
SPENT
hr-1;
F — 72
ALKALIZED ALUMINA*
REACTION TIME = 1 hr
REDUCING MAX. CONSUME.
GAS FLOW REDUCING GAS % S
(cc/min)
61
64
61
63
62
63
63
65
64
64
60
65
61
60
63
62
64
65
64
62
.8
.5
.9
.4
.5
.5
.5
.8
.2
.7
.4
.6
.6
(cc/min)
31.
48.
45
14.
10
8.
42.
51
51.
53.
51.
16.
15
15.
23.
13
38.
34.
SO,
23
6
2
7
3
1
1
5
3
9
7
2
1
3
INITIAL
11.7
8.8
9.1
8.6
10.4
10
6.7
6.2
11.7
8.8
9.1
10'.4
10
11.7
6.2
6.2
11.7
11.7
11.7
11.7
4
REMOVED
WEIGHT
FINAL (WT LOSS BASIS) LOSS
3.7
2.3
1.9
6.2
9.6
10.2
1,5
2.5
2.3
2.1
1.7
8.5
8.9
9.1
2.8
3.3
6.2
5.5
5.9
8.2
77.7
79.5
83.7
30.4
13.6
4.7
81.7
66.3
86.7
81.3
85.3
25.3
19.4
33.3
61
52.9
61.5
66.3
64.8
42.8
29.8
21.7
21.9
10.3
6.0
6.6
18.0
15.7
32.5
21.7
21.5
8 .5
9.5
14.2
13.7
11.6
27.3
28.3
30.3
ld.4
JSBM-Pittsburgh - Sec. II - uec. z'L, 1963 Quarterly ^
Taoie
Test Conditions:
1" diameter Vycor Reactor
4.25 gm charge
7-8 ram bed height
4 cm^ volume
Samples containing no iron. All. other sorbents tabulated contained iron.
-------
TABLE F-73
TWO-STEP
REGENERATION OF SPENT ALKALIZED ALUMINA*
SPACE VELOCITY
NO . SORBENT
171 Kaiser DN-113-F
173
174
Lit
175
178
187 Kaiser DN-112-F
186
179
180
181
166 Peter S pence +
167
165
164
150 USBM X-3-P
151
148
161
153
143 USBM 423-X-4E
144
145
146 Kaiser PR- 15
163
•
168 Grace (1) B7M1
REDUCING
GAS
CO
CO then CO, •
CO then CO.
CO then COj
(34H-0)
CO then CO,
(3*H,0)
CO 2
CO then CO,
CO then CO,
CO then CO,
-------
TABLE F-74
REGENERATION OF SORBENTS USED FOR SO, REMOVAL*
SPACE
VELOCITY
lOOOhr"1
MAXIMUM OVERALL
TEST
NO.
182
183
184
185
188
189
190
191
192
199
198
200
193
194
211
212
201
202
203
206
207
208
SORBENT
BuM No.
BuM No.
Bu-4 No.
BuM No.
BUM NO.
BUM NO.
BUM NO.
BUM NO.
BUM NO.
BUM NO.
BUM NO.
BUM NO.
UOP-2
UOP-2
UOP-V-1
UOP-V-1
Kaiser
Kaiser
Kaiser
Kaiser
Kaiser
Kaiser
466* 17%Na
467-22%Na
470-22%Na
473-ll»Na
474-17»Na
474-17»Na
474-17«Na
474-17»Na
473-ll%Na
473-ll%Na
473-lltNa
PR-18-1
PR-18-1
PR-18-1
PR-18-1
PR-18-1
PR-18-1
REDUCING TEMPERATURE RATE RATE
GAS (°C) (cc/gra S/roin)
H, 660
2
H, then CO, 660 300
• 660 600
660 100
CO 680
CO then CO, 680 300
* 680 600
H, 660
H, 640
CO 680
CO then COj 680 300
H, 680
Hj 660
H2 640
CO 640
CO:H2-5:1 640
C3|H2-2:1 640
27
136
135
17
163
145
143
140
144
187
180
183
266
245
180
176
134
94
63
102
102
ya
16
54
52
8
47
53
53
51
53
48
48
48
52
51
50
46
56
52
40
45
50
48
» SULFUR
INITIAL
8.2
9.2
8.8
9.2
6.5
8.1
8.1
8.1
8.1
6.5
6.5
6.5
4'. 9
4.9
6.7
6.7
9.3
9.3
9.3
9.3
9.3
9.3
FINAL
7.0
2.1
2.8
9.1
2.1
2.8
1.7
1.9
2.6
4.7
2.2
4.1
2.2
2.5
5.1
1.4
2.5
2.5
4.7
5.5
4.1
4.1
REMOVED
20.7
82.5
74.8
6.7
73.2
72.7
53.1
81.3
74.6
39.2
70.1
44.9
61.3
55.9
36.8
81.8
79.5
79.8
58.6
52.5
65.6
65.5
HARDNESS (LBS)
SPENT
14.0
•12.0
14.8
14.2
13.6
14.6
14.6
14.6
14.6
13.6
13.6
13.6
10.3
10.3
2.4
2.4
_
_
-
-
_
REGENERATED
15
5
7
15
10
11
11
11
10
11
12
S
9
10
4
4
4
3
3
4.
j.
3 .
.0
.6
• *
.0
.8
. X
.2
.0
.1
.3
.7
.1
4
4
2
7
7
5
8
9
o
6
• USB.'4-Pittaburga-Sec. II-March 31, 1969 Quarterly Report-Table 1.
Test Conditions; 4.25 gm charge with 7-8 mm bed height
60 cc/min reducing qas flow rate.
1.17 sec. superficial contact time 9 660'C
Reduction reaction using Hf, CO and these mixtures was usually terminated after
1 hour, even though at the higher temperatures consumption of the reducing gas
was completed before that time.
+ Samples containing no iron. All other sorbents tabulated contained iron.
-F383-
-------
TABLE F-75
REGENERATION OF GRACE 1 AND GRACE 2 SATURATED WITH SO,, SPACE VELOCITY -1000 hr'l*
TEST
NO.
209
210
217
218
227
220
225
224
221
222
SAMPLE
PREP.
Grace 1
Grace 2
Grace 1
Grace 2
Grace 1
Grace 2
Grace 1
Grace 2
Grace 2
Grace 2
REGENERA-
TION GAS
CO
then
CO,
• co
then
CO,
H
H2
H,
H2
H2
»2
1:1, H,:CO
then
C02(5%H20)
1:1, H :CO
thin
TEMP.
CO
680
MAXIMUM
(cc/gm
Room
RATE
S/min)
101
OVERALL RATE
(cc/gra S/min)
53
680-»Tenp.
680
Room
107
53
680-»Tenp.
700
700
680
680
640
640
680
250
680
(2 hrs.)
(2 hrs.)
(0.5 hr.)
106
112
102
99
44
56
104
110
54
54
53
53
25
24
54
54
SULFUR,
INITIAL
12.5
li.9
12.5
11.9
12.5
11.9
12.5
11.9
11.9 '
PERCENT
FINAL
' 7.2
6.8
2.9
3.0
3.2
3.9
3.4
•3.7
2.8
REMOVED
52.2
52.2
84.4
82.6
82.7
77.4
81.6
78.5
82.4
HARDNH
SPENT
12.4
6.9
12.4
6.9
12.4
6.9
12.4
6.9
6.9
;SS, (LBS)
~R£GENE"KATET
4.4
3.0
3.0
0.8
2.4
1.5
1.8
0.9
2.1
C02(5*H20)
250 (1 hr.)
11.9
1.2
92.4
6.9
1.8
USBM-Pittsburgh-Sec.
Test Conditions:
II-March 31, 1969 Quarterly Report - Table 3.
4.25 gm charge with 7-8 rra bed height
60cc/min reducing gas flow rate
1.17 sec. superficial contact time 9 660°C
Reduction reaction using H,, CO, and tneir mixtures was usually
terminated after 1 hr., even though at the higner temperatures
consumption of tli« reducing gaa was completed before that time.
-F384-
-------
TABLE F-76
EFFECT OF H2 PARTIAL PRESSURE ON REGENERATION TIME
25
20
15
10
5
Sorbent
Loading Given at
Each Point - % SO2
_ e 3.6
— T 0 4. 3
c
H
a
04.2
_* Q 4.1 05.4
— in
r-
— o
4J
_ VOLUME PI
e 5.5
]RCENT H0
2| 1
0 10 20 30 40 50 60 70
RUN TEMP. REGENERATION GAS
COMPOSITION
(°C) (VOL. %)
R-3 650 H0 - 100
R-5 " «2 - "
R-6 " " - "
R-7 " H2-N2 - 50/50
R-8 " " - 12.2/87.8
R-9 " " - 26/74
R-10 " " - 33.1/66.9
R-13 " " - 61.2/38.8
R-21 " " - 61.2/38.8
INITIAL SORBENT
LOADING
(% S02)*
4.2
4.3
4.2
5.4
3.6
4.3
4.1
5.5
4.2
0 i, o
9^1 4
^ 4B 3
1 1
80 90 100
REGENERATION TIME
(75% COMPLETION)**
7.7
7.2
8.9
13.
20.8
17
13
8.5
15
* Sulfur free basis.
** 75% of weight gained during sorption.
Ref: AVCO FINAL REPORT
TABLE 8 - p. 106
NOVEMBER 1, 1967
-F385-
-------
TABLE F-77
EFFECT OF SORBENT LOADING ON REGENERATION TIME.
RUN
P-5B
P-6A
R-3
R-5
R-6
R-56
R-71
R-72
R-57
R-58
R-62
R-63
R-13
R-16
R-17
R-19
R-20
R-21
TEMP
(°C)
650
II
II
n
it
n
il
ii
700
II
II
II
650
II
II
II
It
II
REG. GAS COMP.
(VOL %)
H2 -- 100
n
"
"
n
ii
n
11
H2 ~ 100
II
II
II
H2-N2-61.2/38.8
11
n
M
ii
ii
INITIAL SORBENT
LOADING (% SO2)*
2.8
9.4
4.2
4.3
4.2
37. **
7.2 **
8.6 **
33. **
41. **
31. **
28.5 **
5.5
2.4
2.2
6.1
2.5
4.2
REG. TIME
(75% COMP. **
3.
17.
7.7
7.2
8.9
95.
11.
14.5
22.
31.
8.
9.
8.5
4.
4.9
13.
5.3
15.
GRAPH 1 SYMBOL"
**
X
II
II
II
II
A
A
A
0
II
It
M
O
II
II
II
II
11
*Sulfur free basis.
**Sorption with wet flue gas
***75% of weight gained during sorption
Ref: AVCO FINAL REPORT
TABLE 8 - p. 106
NOVEMBER 1, 1967
-F386-
-------
TABLE F-78
EFFECT OF CARBON MONOXIDE ON REGENERATION TIME
RUN
R-24
R-27
R-28
R-3
R-5
R-6
TEMP
(°C)
650
11
II
II
II
II
REG. GAS. COMP.
(VOL %.) '
CO - 100
H--CO; 75/25
tn-CO; 25.5/74.5
H^-100
f-n
ii
INITIAL SORBENT
LOADING (% SO-))*
4.9
4.1
4.7
4.2
4.3
4.2
REG. TIME
(75% COMP.)
5.2**
5. **
3.2**
7.7
7.2
8.9
*Sulfur free basis
**In the presence of CO, regeneration time is defined as the
time required to remove 75% of the sorbate that can be
removed before reaching equilibrium as opposed to 75% of the
weight gained during sorption.
10
9
8
7
6
5
4
3
2
1
0
Z
s
— o
— w
o
l-S
REF: AVCO FINAL REPORT
NOVEMBER 1, 1967
TABLE 8 - p- 106
10
% CO IN REGENERATING GAS
20 30 40 50 60 70 80 90 100
-F387-
-------
30
or
£
20
10
EFFECT OF SORPTION TIME ON THERMAL REGENERATION
TABLE F-79
••
•
'
o
X
o
)
X
>
X
REGEN PERCENT SO, IN GAS
TEMP DURING SORPTICN
O 650'C 1.0 ~
• 700 1.0
X 650 .06
O 6SO .15
X
V
o
X
X •
O
Q
O
20 40 60 60 100 120
SORPTION TIME (MIN)
140 160
REF: AVCO FINAL REPORT
PP. 104-106
NOVEMBER 1, 1967
ISO
REGENERATION DATA FOR S02 SORPTION (Tsorplion = 300° C)
Run
R-56
R-57
R-58
R-59
R-60
R-62
R-63
R-71
R-72
ISR-1
tSR-2
lS.R-3
tSR-4
ISR-5
ISR-6
1SR-7
ISR-11
ISR-21
ISR-31
ISR-51
T(°C)
650
700
700
700
700
700
700
650
650
650
650
650
650
650
630
650
650
650
650
650
Regeneration gas
Composition
(volume percent)
H2-100
Hj-100
Hj-100
H2-69.7, H2S-7.2
H20-22i8
H2S + H2 -
H2S meter broken
H2-100
H2-100
H2-100
H2-100
H2-100
H2-100
H2-100
H2-100
H2-100
H2-100
H2-100
H2-100
H2-100
Hj-100
ii., -ion
Initial*
Loading
(mg/mg)
.37
.33
.41
.367
.372
.31
.285
.072
.086
Regeneration
Time (rain)
95
22
31
8
9
11
14.5
Sorptton
time
(mln)
95
91
162
155
142
60
56
14
16
45
50
85
87
79
60
70
50
86
75
95.2
Weight loss upon
heating In N2,
percent of gain
In sorptlon
7.0
12.6
6.2
6.8
5.5
21.8
23.8
22.6
34.8
31.8
20.6
22.9
19.4
12.2
28.2
5.9
21.0
6.9
18.0
8.7
*con>i>uc«d as .8 X (weight pain (luring sorptlon?
weight of abrbenc
tor P'5B throu«h R-28 «>>senc for R-56
-------
TABLE F-80
EFFECT OF TEMPERATURE. ON REGENERATION TIME
RUN
R-23
R-3
R-5
4, R-56
cj
oo R-57
1 R-58
TEMP.
550
650
650
650
700
700
REG. GAS COMP.
(VOL %)
H2 - 100
II
tl
It
II
II
INITIAL SORBENT
LOADING (% SO2)*
4.3
4.2
4.3
37.0
33.0
41.0
SORPTION CONDITIONS
Dry Flue Gas
„
Wet Flue Gas
n n n
n if n
REG. TIME
(75% COMP.)**
36.0
7.7
7.2
95.0
22.0
31.0
*Sulfur free basis
**75% of weight gained during sorption
Ref: AVCO FINAL REPORT
TABLE 8, p. 106
NOVEMBER 1, 1967
-------
TABLE F-81
TWO-STEP REGENERATION SUMMARY-GRACE NO. 2 SORBENT*
I
Ul
V£>
0
1
Run #
5650-243
5650-244
5650-245*
5650-246
5650-247
5650-248
Wt % S
Absorbed
5.07
4.86
4.83
4.92
4.92
5.18
Time
(Min)
20
75
75
45
60
60
REDUCTION
CO Flow
Temperature Rate
op
1270
1200
1200
1200
1250
1200
VHSV
2625
2625
1500
1500
1500
2625
REGENERATION
Flow
Time Temperature Rate
(Min)
75
40
50
20
50
40
(°F)
500
340
340
340
500
500
VHSV
1211
750
750
750
750
750
^
80
0
0
0
0
0
, %co
0.0
50
1
50
75
75
87
Final
Analysis
%H2O
20
50
50
25
25
13
(Wt %)
1.28
1.32
3.12
0.61
0.52
0.63
* Sample to be rerun to confirm data.
+ W. R. Grace & Co.- Jan. 1969 Monthly Report - Table III
-------
10.0
3
v»—
"5
c
(D
O
UJ
REF: USBM-ALBANY —|
JANUARY 1969
PROGRESS' REPORT!
FIG. 1
Sorbed sulfur
Residual sulfur
100
40 60 80
TIME, minutes
FIGURE F-79 -Effect of Time and Temperature
on Sulfur Removed (Grace No. 1-8.5 pet S).
120
-F391-
69-29
-------
8.0
&
"V
-"H.
?—
a —
==»*
1
I
1
;^__
•MM^M
I^MIMB
BB|H|HH
•MlBMHI
MMM^
|
1
600° C ~~
4.0
o
Q.
Ld
CO
<
2
<
REF: USBM-ALBANY _
JANUARY 1969
PROGRESS REPORT!
FIG. 2
0
40 60 80
TIME, minutes
FIGURE F-SO -Affect of Time and Temperature
on Sulfur Remove4J£race No. 1-5.5 pet S)
~jt jyA" ' ''
-------
C
O
CO
LJ
CO
<
z
<
REF:
600° C
USBM-ALBANY JANUARY 1969
PROGRESS REPORT - FIG. 3
Sorbed sulfur
Residual sulfur
650° C
0
20
' 100
40 60 80
TIME, minutes
FIGURE F-8i -Effect of Time and Temperature
on Sulfur Removed (Grace No. 1-2.3 pet S).
69-31
-F393-
-------
100 pet regeneration
of sorbed sulfur
US1M-ALBANY
JANUARY 1969
PROGRESS REPORT
FIG. 4
0
20
FIGURE F-82
Grace No.
40 60 80
TIME, minutes
-Percent Regenerated on
with 2.3 pet Sulfur Lo.ad.
69-32
-F394-
-------
100
100 pet regeneration
of sorbed sulfur
USBM-ALBANY
JANUARY 1969
PROGRESS REPORT
FIG. 5
0
20
FIGURE F-83
Grace No.
40 60 80 100 120
TIME, minutes
-Recent Regenerated on
with 5.5 pet Sulfur Load.
69-33
-F395-
-------
100
c
o>
o
O
CC
UJ
-z.
UJ
o
UJ
a:
40
30
20
10
REF: USBM-ALBANY
JANUARY 1969 "
PROGRESS REPORT
FIG. 6
650° C
,600° C
3 20 ' 40 60 80 100 ' 120
TIME, minutes
FIGURE F-84 -Percent Regenerated on
Grace No. I with 8.5 pet Sulfur Load.
69-34
-F396-
-------
Grace No 2
Kaiser PR-I5
Kaiser ON II3F
USBM-PITTSBURGH - SEC. II
DEC. 31, 1968 QUARTERLY REPORT
FIG. 2
24
TIME, minutes
40
48
56
FIGURE F-85 H2 utilization at 6GO°C and 1000 hr"1 space velocity
for commercial sorbents containing iron.
-F397-
-------
i 1 r i r
423X-4E
660°C
620"C
600°C
580°C,
REF: USBM-PITTSBURGH-SEC. II
DEC. 31, 1968 QUARTERLY;REPORT
FIG. 3
8
16
24 32 40
TIME,minutes
-1
48 56
FIGURE F-86 H2 utilisation at 1000 hr~x space velocity of Bureau's
H23-X-ME at four diff-rent reaction temperatures.
-F398-
-------
T
z
o
p
c.
s
D
to
Z
O
o
ca
z
o
5
a:
tu
z
UJ
o
UJ
a:
REF: USBM-PITTSBURGH-SEC, II
DEC. 31, .1968 QUARTERLY REPORT
FIG. 4
T
1:1 H2 to CO of 660 °C
H2 o> 660 °C
I'l H2 to CO at 640 °C
H2 at 640 °C
\
\
•
32 40
TIME, minutes
56
FIGURE F-87 Utilization of Regeneration G-is for Grace (1) at 660° and
6*»0°C and 1000 hr~]- st-ace velocity.
-F399-
-------
100
90 —
USBM-PITTSBURGH - SEC. II
MARCH 31, 1969 QUARTERLY REPORT
FIG. 1
UOP-2
—680°C
•-640°C
8
16 24 32 40
TIME, minutes
48
56
64
FIGURE F-88 ~ ^2 Re8enerati-on °f the UOP-2 Scrbent Impregnated With Sodium
Aluminate and Ferric Nitrate at 1,000 hr"1 Space Velocity.
-F400-
-------
100
Kaiser PR-IS-i
680°C
660°C
640°C
fREF: USBM-PITTSBURGH - SEC. II
MARCH 31, 1969 QUARTERLY REPO
TIME, minutes
FIGURE F-89 ' H2 Regeneration of Kaiser TR-18-1 Alkalized Alumina at Three
Different Temperatures and 1,000 hr'l Space Velocity.
-F401-
-------
7O
60
c
Ol
o
v 50
QL
a.
5
:>
CO
z
o
o
V)
<
o
o
?
o
40
30
? 20
UJ
a:
10
I
I'
Koiser PR-I8-I
,CO
H2
^0% CO~20%H2
65 % CO-35% H2
USBM-PITTSBURGH - SEC. II
MARCH 31, 1969 QUARTERLY REPORT
FIG. 3
8
16 24 32 40
TIME, minufes
48
56
S4
FIGURE F-90 - Regeneration of Kaiser PR-18-1 Using Various Combinations of
Reducing Gases at 640° C.
L-II09I
-F402-
-------
FIGURE F-91
EFFECT OF INITIAL SORBENT LOADING ON REGENERATION TIME
• i •
9
8
7
6
5
-- I
9
8
7
6
5
UJ
z
2
UJ
2
z
o
z
UJ
o
O 61.2% H2 IN N2, 650°C, DRY FLUE GAS
X 100% H2, 650CC, DRY FLUE GAS
A 100% H2, 650°c, WET FLUE GAS
100% H2, 700°c, WET FLUE GAS
0
61.2% H2
650°C
100% H2
700°C
WET FLUE GAS
REFERENCE:
Avco Final Report, Table 8, p. 106,
and Figure 44, p. 107 (November
I, 1967).
INITIAL SORBENT LOADING, % S02"
-"-I
4 567891
X I
2 3
— x 10 —
456
SULFUR FREE BASIS
-F403-
-------
APPENDIX G
PROCESS CALCULATIONS AND DISPERSED PHASE VESSEL SKETCHES
Material balance sample calculations are shown in Table G82
and G83 for the dispersed phase regenerator.
Sample calculations are presented for the various methods
which were used to design the solids heater. Note that the
examples presented do not necessarily represent the final design
but do show the various methods that were evaluated in selecting
the final design. After the examples were typed several process
changes were made and it was not considered worthwhile to change
the sample calculations to reflect the final design numbers since
the methods remain the same.
The calculations for the dispersed phase sorber material
balance and vessel dimensions are shown in Tables G-89, G-90
and G-91. Note that the examples shown for sorber height are
for the design which includes a solids cooler. The cooler was
later deleted and extra reactor height added for heat transfer
as described in the main body of this report.
Vessel sketches are shown for the major pieces of equipment
required in the dispersed phase process.
-G404-
-------
TABLE c-_82.
REGENERATOR MATERIAL BALANCE CALCULATIONS
(1) AVCO* Regenerated Solids Product
moles Ibs
NaAl02 1.912 156.73
A1203 1.479 150.80
Na2SO^ 0.044 6.25
3.435 313.78
0.044(32.06)(100) ^_
313778 - °'45 Wt % S
(2) Regenerated Solids Product for Design Condition of 0.5 wt % S
moles Ibs
NaAl02 1.912 156.73
Al-0, 1.479 150.80
Na^SO^ 0.049 6.97 (1.573 Ibs S)
3.440 314.50
(3) Amount of Na2S(\ Formed in Absorber for Design Condition of
3 wt % S in Regenerator Feed
SO, + ^O, -1- 2 NaAlO? ->• Na2SO^ -' M2O,
fL £. *~ *~ ~ £- ^
Let x = moles Na2SOlt formed in absorber
x = moles A1203 formed in absorber
2x = moles NaAlO2 consumed in the reaction
wt.sulfur formed + wt. entering sulfur = Q 03
"' total wt. solids leaving
*AVCO 16th Monthly Summary, April 29, 1968, Table VIII, p. 50
continued
-G405-
-------
or 32 .Mx + 1.573 _
' [314.50 + 142.04x + 101.96x] -2x (81.97) ~ U>
x = 0.265 moles Na2SO4 formed in absorber and
,0.'265 moles Al 0_ formed in absorber
(4) AVCO* Regeneration .Gas Adjusted for the Above^Deai^n4Canditions
AVCO MWK
moles Factor moles
H2 3-42 0.265 0.9063
C02 0.58 0.265 a."1537
CO 0.444 0.265 0.'1177
-G406-
-------
TABI.K G-83
REGENERATION REACTION
(3 wt % S •> 0.5 wt % S)
(1) Balanced Reaction
MPLS
Reactants: 0.265 Na2SO1+
0.265 A12O3
0.9063 H2
0.1537 CO,
0.1177 CO
Products: 0.53 NaAlO2
0.0022 CO
0.0046 COS
0.265 C02
0.0202 H2
0-6491 H20
0.2370 H2S
0.0170 S02
0.0032 S2
(2) Flow Rates for 80% Utilization of Regeneration Gas (1000MW Plant)
* S removed to off-gas = 619.3 moles (0.90) = 557.4 moles S
hr hr
Equation Scale Factor = 557.4/0.265 = 2103.4
Table G-89
continued.
-G407-
-------
Regeneration Gas:
mole
moles* fraction scale factor hr
Component
H2
CO;
CO
1.472 1.000
*includes 25% excess over theoretical
equation mol«s/
1.1329
0.1921
0.1471
0.769
0.131
0.100
2103.4
2103.4
2103.4
Ibs/hr
2382.9 4,813
404<1 17,784
309»4 8,666
3096. A 31,263
Regeneration Off-Gas:
Component
CO
COS
CO2
H2
H2O
H2S
SO 2
S2
moles
0.0316
0.0046
0.3034
0.2468
0.6491
0.237
0.0170
0.0032
mole
fraction
0.0212
0.0031
0.2033
0.1653
0.4348
0.1588
0-0114
0.0021
1.0000
equation
scale factor
moles/
hr Ibs/hr
2103.4
2103.4
2103.4
2103.4
2103.4
2103.4
2103.4
2103.4
66.47
9.68
638.17
519.12
1365.32
498.50
35.76
6.73
3139.75
1,862
581
28,079
1,049
24,603
16,989
2,291
432
75,886
S in Off Gas =
3139.75
(100) = 17.75% mole
-G408-
-------
TABLE G-8 4
SAMPLE CALCULATION - SOLIDS HEATER DIAMETER
FOR TWO
TEMPERATURE APPROACHES
1000 MW PLANT
Conditions:
2 psig inlet pressure
Assume 1 psig pressure drop
0.5 ft/sec average superficial gas velocity
75.6MM Btu/hr duty for each of 2 trains
Assume cylindrical vessels for purposes of calculation
50° APPROACH
Solids:
Gas:
600
650
1271°F
1500°F
Amount of air needed:
75,600,000 = x (7.75). (1500-650)
x = 11,476 moles/hr/train or
332,804 lbs/hr/train
ACFM
out
ACFM.
in
ACFM
= 11,476(10.73) (460 + 650)
60(15.7)
= 145,098
= 11.476(10.73) (460 +1500)
60(16.7)
= 240,868
c
(0 . b;
= 192,983
2
= 6433 ft cross section
:. Diameter = -90.5 ft for each of
2 solids heaters
10° APPROACH
Solids: 600 -»• 1271°F
Gas: 610 «- 1500°F
Amount of air needed:
75,600,000 = x (7.75) (1500-610)
x = 10,960 moles/hr/train or
317,840 Ibs/hr/train
ACFM = 10,960(10.73)(460 + 610)
out 60(15.7)
- 133,581
ACFM. = 10,960(10.73) (460 + 1500)
ln 60(16.7)
= 230,037
ACFMav = 181,809
181/809 = 6060 ft2 cross section
60(0.5)
.•. Diameter = -88 ft for each of 2
solids heaters
-G409-
-------
TABLE G-85
SAMPLE CALCULATION - SOLIDS HEATER BED HEIGHT
BRINN* CORRELATION
lOQQ-MW PLANT
Conditions: 75.6 MM Btu/hr duty for each of 2 trains.
0.5 ft/sec average superficial gas velocityv
Assume 1 psig pressure drop
50° approach—
Solids: 600 -»- 1271°F
Gas: 650 «- 1500°F
1) Calculation of Total Tube Length
Feed to Regenerator = 705,792** = 352,896 Ibs/hr/train
2
Mean heat capacity of solids =0.29 Btu
lb°F
Z = (WCD) solids _ 352,896(0.29) = .
;wrp. 7 7q (Data from .Table G-84)
Pgas 332,804^)
t2~ fcl 1271 - 600 n
..,.„ -.,. — ,. =: —,-.,--- I, - • — n
T2- t± 1500 - 600 u
Assume 2" ID tubes, triangular pitch
Assume U f coefficient of heat transfer from jacket fluid to
inside tube surface ~
= 9 Btu/hr/ft^ °F (Brinn Article)
* Brinn, M. S. et.al,, IEC 40, 1050 (1948)
** Table G-89.
continued
-G410-
-------
kAl O @ Tavq = 4'62 Btu/hr ft °F (Thermophysical Properties
23 Research Center, DATA BOOK,
Purdue Univ., Lafayette,
Indiana (1964) )
ft _ k _ 4.62(12)
- —
U R
From the design curves (Brinn Article, Figs. 14-15-16 extrapolated)
we
iTr— =0.6 It can be seen that for appropriate
values of the temperature ratio,
there is a narrow spread in values
for we /kL for 3 = 2 and 6 =20.
352,896(0.29) _ n ,
4.62(L) °'6
• L = 36,919 ft total tube length for each of 2 solids
heaters.
2) Calculation of Number of Tubes and Bed Height
where C = constant =0.86 for triangular pitch
P = tube spacing = 1.25(D) + 1/2 = 3 inches
L = outer tube limit, inches
(Floating Head Exchanger = Shell ID - 1-1/2")
N = number of tubes
* Nelson, W. L., PETROLEUM REFINERY ENGINEERING, McGraw-Hill Book
Company, New York (1958).
continued.
-G411-
-------
TABLE G-85 (continued)
For 16.7 psia inlet pressure and 90.5 ft diameter vessels (2)
N = 0.86
90.5(12)-!.5
N = 112,387 tubes-for each of 2 solids heaters
Bed Height = 36,919 ft total tube length
112,387 tubes
= 0.33 ft for each of 2 solids heaters.
-G412-
-------
TABLE G-86
SAMPLE CALCULATION - SOLIDS HEATER BED HEIGHT
FURNAS* METHOD
1000 MW PLANT
Assumptions: 1) Negligible resistance to heat transfer by
conduction in the solid.
2) Countercurrent heat exchange between a fluid
and solid without heat generation within the
solid.
3) Area for heat exchange is the total external
area of the particles in the bed at one time.
4) Solid consists of uniform spheres.
5) The exchanger is adiabatic, the only heat
transferred being that between the solid and
fluid.
6) Heat conduction in the fluid in the direction
of flow is negligible.
* Munro, W. P., and Amundson, N. R. , IEC, 42_, 1481 (1950)
Furnas, C. C. , IEC, 22_, 26 (1930)
Furnas, C. C., TRANS AIChE, 24, 1942 (1930)
Furnas, C. C., U. S. BUR. MINES, BULL., 361 (1932)
continued.
-G413-
-------
Conditions: 75.6 MM Btu/hr duty for each of 2 trains.
0.5 ft/sec average superficial gas velocity
Assume 1 psig pressure drop
50° Approach—
^Solids: 600 -»• 1271°F
Gas: 650 «- 1500°F
16.7 psia inlet pressure
90.5 ft diameter vessels (2)
T " to 1500- 600
T - t 650- 600
o • o
C* C!
(Data ^om,,Tables G-84
ana G—85)
Based on Grace #2 Catalyst, Bruceton Quarterly, June 30, 1968:
After 3 regenerations - 50% size = 1300u- Further cycling will
reduce this. Therefore, in these calculations, assume a lOOOp
particle size.
• R, radius of sphere = 2(2 54?{l2) = °*00164 ft
hf, heat transfer coefficient at solid-fluid interface:
REF: Leva, FLUIDIZATION, McGraw-Hill Book Company; New York
(1959)
hD
= 0.0063
D G
M
(Eq. 8-53)
continued.
-G414-
-------
avg
= 0.0325 Btu/hr ft °F
= 0.0955 Ibs/hr ft
Gair = 332,804/6433 = 51.73 Ibs/hr ft2 (Data from Table 1)
h = 0.0325(0.0063)
0.00328
0.00328(51.73)
0.0955
1.8
h = 0.18 Btu/hr ft °F
Furnas Approximate Equation*:
T -
fr exP
3(0-l)yx
6-1
where E = k /h_R
s i
Y = k8d-f>
for
GsCs
= (l-f)hfR
GsCs
Let fraction voids = 0.60 = f (from MWK Progress Report No. 7, p-14)
Gg - 352,896 =
6433
lbs/hr
(Data from Tables G-84
and G-85)
» I « 0.40(0.18) (0.00164)
e 54.86 (0.29)
* Munro, W. D., and Amundsen, N. R., IEC, 42, 1481(1950)
continued.
-G415-
-------
18 =
18 =
x =
1 1.15
1-1.15 1.15-1 6XP
3(1.15-1)(0.00000742)x
(0.00164)*
-6.67 + 7.67 exp fl.2414x1
-1.3 ft bed height for each of 2 solids heaters
-G416-
-------
TABLE G-87
SAMPLE CALCULATION - TEMPERATURE HISTORY CURVE -
HEATING OF SOLIDS HEATER BED FOR 0.075 HRS HOLD-UP
MWK CORRELATION - 1000 MW PLANT
Conditions: 75.6 MM Btu/hr duty for each of 2 trains
0.5 ft/sec average superficial gas velocity
Assume 1 psig pressure drop
10°F approach—
Solids: 600 -*• 1271°F
Gas: 610 «- 1500°F
16.7 psia inlet pressure
88 ft. diameter vessels (2)
1) Calculation of Convection Coefficient
D°'9
where C , mean specific heat of the gas, = 7.75/29 = 0.267 Btu/lb °F
G, gas mass velocity, = 317,840/6060 = 52.4 Ibs/hr ft
Z . eras viscosity at average temperature, = 0.0955/2.42 *
9 0.039 cp
f, fraction voids in bed, =0.60
D, particle diameter, = lOOOu - 0.00328 ft
(The above data are from Tables G-84 and G-86)
• hA = 17.5(0.267)(52.4)0-7(0.039)0-3(10-°-2736)
(0.00328)0*9
hA = 2593.4 Btu/hr ft °F
continued.
-G417-
-------
2) Time - Temperature Relationship
C = mean heat capacity of solid =0.29 Btu/lb °F
s 3
p = packed solids density =45 Ibs/ft
S
t = hold-up time, hrs
x = bed height, ft
T = temperature of gas at any point at any time,. i.°F
T = temperature of solid at any point at any time,...?F
S
T = initial temperature of gas, °F
T = initial temperature of solid, °F
SO
z =
Y =
hAt _ 2593.4t
Co ~ 0.29(45)
= 198.7t
hAx 2593.4x
GC 52.4(0.267)
= 185.4x
T -
y
T -
o
_
s
m _
O
T
SO
T
so
T
so
T
so
T - 600
g
1500- 600
T - 600
s
1500- 600
T - 600
. g
900
T - 600
s
900
The following data,are plotted on Figure 49:
t = 0.075 hrs ; Z = 14.94
x(ft)
Bed Ht.
0-1
0.125
0.15
0.025
0.05
0.075
Y =
185. 4x
18.5
23.2
27.8
4.6
9.3
13.9
g
J 0.29
0.11
0
1.0
0.895
0.63
Trr '
g
900(1 ) + 600
861
699
600
1500
1406
1167
s
0.23
0.08
0
1.0
0.855
0.535
T= =
s
900(1 ) + 600
S
807
672
600
1500
1370
1082
continued.
-G418-
-------
For 0.075 hrs hold-up:
352,896 lb_s(0.075 hrs)(ft3/45 Ibs)(1/6060 ft2) = 0.097 ft bed ht.
hr
The area under the solids curve from 840°F (corresponding to
0.097 ft bed ht.) to 1500°F = 45.19 ft °F (from planimeter)
*
.-. Average solids temperature = I.1;?J:t;;°F + 840°F = 1306°F
u.uy/it /, ,.
(above the
required 1271°F)
Also shown in Figure 49 are the temperature history curves for
other holdup times, viz., 0.150, 0.225 and 0.300 hours.
Figure G-92 shows the temperature history curves when the
pressure is increased to 116 psia which changes the vessel
diameter to 33 feet. Other conditions remain the same as in
Figure 49.
-G419-
-------
FIGURE G92
TEMPERATURE - HISTORY CURVES FOR SOLIDS HEATER
MWK METHOD
o.
2 TRAINS REQUIRED
33 FT DIAMETER VESSEL
116 psio INLET
0.5 FT/SEC GAS VELOCITY
I psig PRESSURE DROP ASSUMED
0.225 HRS HOLD-UP I0» APPROACH
SOLIDS - 60O -*I27IF
GAS - 610 —-I5OOF
_ 75.6 mm BTU/HR/TRAIN
TAV - I346F
^
SOLID AND GAS
TEMPERATURE
0.150 HRS HOLD-UP
0.6_9_
BED HT 0.075 HRS HOLD-UP
I I I I
600 700 800 »00 1.000 I.IOO 1.200 1.900 1.400 1.500
-------
TABLE Q-88
SAMPLE CALCULATION - PRESSURE DROP THROUGH SOLIDS HEATER MOVING BED
HAPPEL* CORRELATION
1000 MW PLANT
Conditions: 75.6 MM Btu/hr for each of 2 trains
0.5 ft/sec average superficial gas velocity
Initially 1 psig pressure drop assumed
50 °F approach—
Solids: 600 •*• 1271°F
Gas: 650 + 1500°F
16.7 psia inlet pressure
90.5 diameter vessels (2)
The correlating equation is applicable to gravity flow of solids
when under countercurrent or cocurrent vapor flow, as long as the
solids velocity relative to the wall is not in excess of 20% of
the vapor velocity. For moving beds with solids linear velocities
exceeding 20% of the superficial gas velocity, it is suggested
that the correlation be modified by substituting fluid velocity
relative to the moving particles, rather than relative to the
vessel wall.
Check to see if Vg > 0.20Vg
2
V = 54.86 Ibs/hr ft _ 0>00034 ft/sec (Data from Tables G-86
3 45 Ibs/ft (3600 sec/hr) .
V- = 0.5 ft/sec
ij
• V i 0.20V- and correlation is applicable without modification.
s G
* Happel, J-, IEC, 4_1, 1161 (1949)
Leva, M., FLUIDIZATION, McGraw-Hill Book Company, New York (1959)
continued.
-G421-
-------
Modified Re = D G(l-e)
where D , composite particle diameter = lOOOp = 0.0394 in.
P (Data from Table G-86)
2
G, air mass velocity = 51.73 Ibs/hr/ft
e, " bed voidage = 0.60
y, air viscosity = 0.0955 Ibs/hr ft
• Modified Re = 0.0394(51.73)(0.4) .
0.0955(12) U
AP = fM2L uo2 pf(l-e)3 (Leva*, p!64, Eq. 6-20)
q D
yc p
where fM, modified friction factor = 2500 (Leva*, p!64, Fig. 6-14)
M
u , superficial gas velocity = 0.5 ft/sec
Pf , fluid density = 16.2(29) a Q 02g Iba/ft3
avg , 10.73(460 + 1075) O'°^y •Lbs/tt
L, bed height; Use 3 ft since maximum calculated bed
height was 3 ft.
=
2500
(2)
'
(3)
32
(0.
.2
5)
(0
2
•
(0.029)
0394)
(0
.4)
3
12
144
AP = 0.46 psig (O.K. since less than 1 psig assumed)
op. cit.
-------
TABLE G-89
DISPERSED PHASE ABSORBER MATERIAL BALANCE
1000 MW PLANT
1) Feed Rate to Absorber
xQLf + 0.0893 (yQ- yf)V (Ref. USfiM Quarterly
f Lf + 0.2230 (y - y_)V Report - Section I -
1 or March 31, 1968 - Eq. 17)
where (y - yf) = mole fraction S02 in flue gas removed
MPH S02 in flue gas initially* = 0.002146(288,583) = 619.3 MPH
•. for 90% removal of SO-,
=
V = SCFH flue gas
V = 62,012 ftjx3600 sec x 492°R =
sec hr 1060°R
x = weight fraction S in sorbent** feed = 0.005
o
xf = weight fraction S in sorbent leaving absorber =0.03
Lf = sorbent feed rate, Ibs/hr
• Substituting into the above equation gives
0-03Lf + 0.03(0.2230) (0.001931) (103,618,542) =
0.005Lf + 0.0893(0.001931)(103,618,542)
and Lf = 661,169 Ibs/hr
-G423-
-------
2) Total Feed to Regenerator
Sulfur removed in Absorber:
619.3 moles x 32.06 Ibs x 0.90 = 17,869 Ibs S/hr
hr mole
Sorbent weight gain as SO.,:
S02 + 1/2 02 + 2NaAlO2 -*• Na2O.S03 + AljO-j
17,869 Ibs S x 80.06 Ibs SO /mole = 4 lfas /hr
hr 32.06 Ibs S/mole J
Sulfur in Absorber Fresh Feed:
0.005 (661,169) = 3306 Ibs S/hr
Sulfur in Feed to Regenerator:
3306 + 17,869 = 21,175 Ibs S/hr
Total Feed to Regenerator:
661,169 + 44,623 = 705,792 Ibs/hr
* Table H-96.
** Sorbent weight is defined as the total weight of solid,
including the absorbed sulfur oxides.
-G424-
-------
TABLE G-9Q
EXPRESSION FOR DISPERSED PHASE ABSORBER HEIGHT
C = 3/h (yo1/3 - yf1/3) (Eq. 13)
f . o - f
P = 0.1924 hp (Eq. 4)
Combining the above equations gives, for 90% SO? removal:
h =
xoLf + °-°804 V0V
Lf + 0.2007 yQV + x(
5.562 Kap
where h
x
V
a
p
K
= absorber height, ft
= weight fraction sulfur in sorbent feed
= sorbent feed rate, Ibs/hr
= mole fraction SO- in flue gas to absorber
*
= flue gas flow rate, SCFH
o
= absorber cross-sectional area, ft
= average sorbent density, lbs/ft3
= rate constant, hours
Ref: USBM Quarterly
Report - Section I -
March 31, 1968
-
-------
TABLE G-91
SAMPLE CALCULATION - ABSORBER DIMENSIONS
1000 MW PLANT
Absorber inlet flue gas
(600°F, 14.7 psia)
= 62,012 ft
sec
(Table H-96)
Desired superficial,gas velocity = 25 ft/sec
.-. Absorber cross sectional area = 62,012 _ 2491 ft^
2
Initially, assume a maximum vessel diameter of 40 ft; a = L256 ft
therefore, 2-40 ft diameter absorbers are required.
Calculate absorber height from equation in Table G-90 for 90% S02
removal:
h =
xoLf
Lf +
+ 0.0804y V
0.2007yQV Ko
V
5.562 Kap
x = 0.005 (wt fract S in sorbent feed)
y = 0.002146 (mol fract S02 in entering flue gas)
p =0.15 Ibs/ft (average sorbent density)
_ i
K = 22 hrs rate constant at 10/1 recycle ratio calculated from
USBM linear sorbent loading model, USBM~-Bruceton Quarterly
Report, March 31, 1968).
Since 2-40 ft diameter absorbers are required, use 1/2" the flow
rates from Table
a = 1256 ft cross sectional
Lf = 661,169/2 = 330,585 Ibs/hr
V = 103,618,542/2 = 51,809,271 SCFH flue gas
continued
-G426-
-------
Substituting these design conditions into the above equation
gives:
h - 0-Q05 (330,585) + 0.0804 (0.002146)(51,809,271)
'330,585 + 0.2007(0.002146) (51,809,271)
5J.,809,271
(5.562) (1256) (22) (0.15)
h = (0.035) (2247.36) = 78.66
.-. h = 80 ft
2 absorbers are needed, each 40 ft diameter x 80 ft high.
-G427-
-------
TABLE G-92
SAMPLE CALCULATION FOR SOLIDS HEATER
1000 T1W PLANT
2. 2-Stage Direct Combustion Fluid Bed Solids Heater
The solids heater duty was calculated to be 90.1MM,Btu/hr/
train based on the following conditions:
• Regenerator feed is heated from 600 -> 1400°F.
• One exchanger train required for each regenerator,
or a total of two trains. Quantities shown are for
each of two trains.
• Heat loss based on 100°F AT between the vessel and
ambient air, wind velocity of 20 miles/hr, and an
estimated solids cooler size,is essentially, negligible.
• The duty shown above includes a 10% excess over
theoretical for process uncertainties.
• The two stages are assumed to be equilibrium stages;
therefore the temperature of the gas and solids
leaving each stage are identical.
By trial and error, the exit gas temperature is found to
be 1000°F from which the natural gas fuel and combustion air
quantities can be calculated.
Solids: Q = 90.1MM Btu/hr
T = 600 -*• 1400°F
Gas: Air @ 1000°F + natural gas @ 1003F @ inlet
Flue Gas @ 680°F @ exit
continued
-G428-
-------
The assumed composition of natural gas and air and the cal-
culated flue gas composition are:
NATURAL
GAS
% MOLE
CHH
C2H6
C3H8
Ci^Hj o
C5H12
CeHm
N2
95.85
2.50
0.43
0.32
0.20
0.10
0.60
100.00
AIR
COMPONENT % MOLE
02 20.105
N2 76.095
H20 3.80
100.00
MOLES
MOLE N.G.
2.072
7.842
0.392
10.306
M.W. = 28.41 Ibs/mole
FLUE GAS
COMPONENT
C02
N2
H20
M.W. =
4 MOLE MOLES
% MOLE MQLE N>
9.26 1.050
69.25 7.848
21.49 2.435
100.00 11.333
27.34 Ibs/mole
Using the above and total molal enthalpies, the heat
available from combustion of the natural gas is calculated to
be 307,000 Btu/mol with a first stage exit temperature of 1400°]
Balance around bottom stage (by trial & error) is 73.5MM
Btu/hr.
•. Nat. gas rate = 73,500,000/307,000 = 240 mols/hr
Air req'd = 10.306 (240) = 2473 mols/hr
4 psig pressure drop assumed through solids heater.
5 ft/sec average superficial gas velocity chosen.
For a 30 psia inlet pressure:
ACFS. = 2473 MPH Air
in
1000°F + 240 MPH N.G.
10.7311473(1460) + 240 (560)1 . 372
100°F
3600(30)
Flue Gas Leaving = 11.333(240) = 2720 moles/hr @ 680°F
ACFS
out
2720(10.73) (1140)
3600(26)
355
Cross Sectional Area = (372 +^355)/2 = ?2>? ffc2 , 1Q ffc ^
Vessel Height (Table 1) =5+ 10 +5+1+ 3.5(10) = 56
A 2-stage Fluid Bed Direct Combustion Solids Heater,
10U x 56', is required at 30 psia.
-G429-
-------
TABLE G-93
FIRST STAGE TEMPERATURE PROFILE FOR 3-STAGE
FLUID BED SOLIDS HEATER - 1000 MW PLANT
Conditions: 90.1MM Btu/hr duty for each of 2 trains
Solids: 600 -»- 1400°F
Gas: 990 «- 1500°F; 825 Ibs/hr ft2 (airv+ flue gas)
5 ft/sec average superficial gas velocity
1 psig>pressure drop assumed
30 psig inlet pressure
31 ft. diameter vessels
lOOOy (0.00328 ft) diameter spherical particles
assumed.
REF: Frantz, J. F., Chem. Eng. Prog., 57 (7_) , 35 (1961)
in (t»- Vi - -h
- Vo
where t . = point gas temperature, °F
t . = point solids temperature, °F
5 1
t = inlet gas temperature, °F
tgo = outlet solids temperature for countercurrent case, °F
a = surface area/unit volume, ft2/ft-*
h = heat transfer coefficient at solid-fluid interface,
Btu/hr ft2 °F
C = gas heat capacity, Btu/lb °F
G = gas mass flow rate, Ibs/hr ft2
L =» bed height, ft
continued.
-G430-
-------
k D G
h = 0.0063 - (JP_) (Leva, FLUIDIZATION, Mc-Graw Hill Book
p u Co., New York (1959))
where k = 0.0356 Btu/hr ft °F
y = 0.0983 Ibs/hr ft
h = 0.0063(0.0356)
0.00328
0.00328(825)
0.0983
1.8
26.7 Btu/hr ft2 °F
for sphere = *Dp*/Particle _ _ 6 _
^? Dp ~ 0.00328 * 1829'27 ft
—7—/particle
Cp = 0.29 Btu/lb F
ah 1829.27 (26.7) -n. ,. , -1
" CpG 0.29(825) ^u«.J.« rt
(t - ts)Q = 1500 - 1400 = 100F° (1st stage)
The 1st stage bed temperature will be assumed equal to the 1st
stage exit gas temperature .-. tg _ 1400oF
tg - 1400
= 204.14L
Figure 51 is a plot of gas temperature vs. bed height for the
first stage, the temperature profile predicted by this equation,
A bed height of less than 0-3 inch per stage is required to heat
the solids to 1400°F using 1500°F inlet gas.
-G431-
-------
-G432-
-------
IKEUQCCJ
FIGURE G-94
JOB No
PAGE
DATE
J?€V-I
/>«*
<&X/T~ G-AJ 4 &/rafr#&
{
*^
u
^
1
%
VI
1^
N
i
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v
Vt
7
v
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r[«B|
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(If
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=1
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t
Lou
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off
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i.
71 r^
£. 'a. ftf.
-G433-
to<4»/*«-
7900
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I
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-------
Pee TEA*/ » TBmt o/* 2.
FIGURE G-95
M/9X. 6/^/4" f
-G434-
NOT REPRODUCIBLE
-------
,HfiIULf
vw
FIGURE G-96
/I
AX
*1C4*4/£
MATBtrtl.
-G435-
-------
KEUOGCj
G-97
TOTAL <>F I "2.
MCSH Soi-iOi
3J8t 5*00 ACFM
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50UICXS to AD/ A/ 6-
?i
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-------
Lor
.—/5i*\—,
[KEILOCC]
FIGURE G-98
Jon No
^
—
i
1
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—
'7t
—
—
i
s:
—
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Al
^
/i
i '
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9-.
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-G437-
-------
APPENDIX H
COMPOSITION OF-COAL AND FLUE GASES
A discussion .is presented of the methods used to iderive
the volume and composition of the flue gases from a ooat-.fj.red
power station needed for the design of the alkalized Alumina
processes.
-H438-
-------
APPENDIX H
COAL PROPERTIES AND FLUE GAS COMPOSITIONS
Typical flue gas volumes and compositions resulting from the
burning of coal containing 1, 3, and 5% sulfur have been calcu-
lated; coal properties were estimated from analyses covering a
wide range of coals. The calculations are based on a 1000 MW coal-
fired power plant with an assumed heat rate of 8,600 Btu/kwh
(39.7% thermal efficiency). The sources of the coal analyses are
as follows:
1. MWK Progress Report #1 Contract #14-01-0001-380
2. MWK Progress Report #8 Contract #14-01-0001-380
3. Babcock & Wilcox Co., Steam-its Generation & Use, 1955,p. 4-A5
4. USBM (Illinois),
5- USBM (New Mexico),
6. USBM (Wyoming),
7. USBM (Utah),
8. USBM (Ohio),
The proximate and ultimate analyses taken from these references as
well as that used by S. Katell, uSBM-Morgantown, in his economic
study, Report #64-9, 1965 (9), are shown in Table H-94.
Based upon the analyses in Table H-94, the following analyses
were taken as typical of 1, 3, and 5% sulfur coals:
RI
RI
RI
RI
RI
5270,
5270,
5270,
5270,
5270,
1956,
1956,
1956,
1956,
1956,
pp.
PP-
pp.
P-
pp.
9-14
22-23
48-49
36
24-25
% S IN COAL
1.0
3.0
5.0
H.O ASH
2.0 9.0
2.0 9.0
2.0 9.0
1.0
3.0
5.0
H
5.2
5.2
5.2
74.4
72.4
70-4
N
1.4
1.4
1-4
7.0
7.0
7.0
Btu
13,000
13,000
13,000
Using these analyses, the calculation of a theoretical flue gas
resulting from the burning of these coals was obtained. The calcu-
lations assumed that the coals were burned using 20% excess air,
the air was at 80°F and 60% relative humidity, all carbon was son-
verted to CO,, all hydrogen to H2O, all sulfur to SO,, and the air
was 21% mole O9. A typical calculation is outlined In Table H-95.
The results of these calculations are given in Table H-96 showing
the physical properties of the flue gas generated by the combustion
of these three grades of sulfur bearing coals.
-H439-
-------
TABLE H-94
TYPICAL COAL ANALYSES
REFERENCE
PROXIMATE
H00 VOLATILE F
I
a:
0
1
1
2
3
4
5
6
7
8
9
2.
1.
8.
12.
13.
13.
3.
- 3.
1.
5
31
0
2
2
1
8
0
3
39.
37.
__
40.
41.
42.
43.
43.
--
0
3
3
5
7
3
1
53
57
-
48
47
51
48
49
-
.C.
.4
.59
-
.8
.2
.2
.8
.5
-
ULTIMATE
ASH
5.1
3.8
-
11.0
11.3
6.2
8.0
7.4
12.7
S
1.3
0.66
1.6
3.4
1.0
0.7
0.4
4.2
3.0
H
5.2
5.14
4.4
5.0
5.2
5.3
5.6
5.2
4.9
C
79.4
83.6
72.0
70.4
70.3
75.6
75.4
74.3
70.1
N
1.4
1.48
1.4
1.4
1.3
1.4
1.4
1.3
1.4
0
7.5
5.27
3.6
9.3
12.6
13.3
10.0
7.6
6.6
B
AS -
ASH REC ' D
5.2 13,990
3.85
9.0 12,800
11,130
10,780
11,210
12,800
13,260
__
tu
DRY
14,350
—
--
12,720
12,410
12,800
13,300
13,680
_ _
-------
TABLE H-95
SAMPLE CALCULATION FOR A 1% SULFUR COAL
M. W.
C 12.01
H2 2.02
S 32.06
02 32.00
N2 28.00
H20 18.02
ASH
TOTAL
# MOLS O0 FACTOR
74.4 6.195 1.0
5.2 2.574 0-5
1.0 0.0312 1.0
7-0 0.219 -1.0
1.4 0.050 0
2.0 0.111 0
9.0
100.0
REQUIRED O,,
• — e.
6.195
1.287
0.0312
-0.219
0
0
-
7.2942
@
0.79 mol N,
0.21 mol O,
= 3.762
mol N,
i
mol 0,
& 20% Excess Air
0.20 x 7.2942 mol O2 - 1.459 mol O2 Excess
Total O^ = 7.294 mols + 1.459 mols
2
Total N2 = 3.762 mols
x 8.753 mol
= 8.753 mols 0,
= 32.929 mols
mol 0,
Total mols dry air » 41.682 mols
continued
-H441-
-------
@ 80°F & 60% R.H.: 0.0212 mol
(B&W Steam p. 4-A5)
mol dry air
HO in Air =
H2O in Air
41.682 mols dry air x 0.0212 mol H
0.884
mol dry air
Total H2O = H2O in Air + H2O in Coal
Total H2O = 0.884 mols + 0.111 mols
Total H2O = 0.995 mols.H2O
Total N2 = N2 in Air + N2 in Coal
Total N2 = 32.929 mols + 0.050 mols
Total
= 32.979 mols
continued
-------
FLUE GAS
co2 so? 0.
* *
C 6.195
H2
S - 0.0312
°2 ~ ~ 1-459
N2
H2° - - -
TOTAL 6.195 0.0312 1-459
N. H50
— £. — £. —
-
2.574
-
-
32.979
0.995
32.979 3.569
CO TOTAL
6.195
2.574
0.0312
1.459
~ 32.979
0.995
~ 44.233
Coal Required = 8,600 Btu x 1,000,000 KW x
hr K.W.
Coal Required = 661,538 I Coal
hr
13,000 Btu
Flue Gas Evolved = 44.233 mol flue gas x 661,538 # coal
100 # coal hr
Flue Gas Evolved = 292,618 mol flue gas
hr
-H443-
-------
TABLE H-96
FLUE GAS PROPERTIES FOR VARIOUS SULFUR BEARING COALS
% S IN COAL PPM S02 MPH CFS M. W. GAS DENSITY
@ 600°F @ 600°F
1 705 292,618 62,879 29.60 0.0383
3 2146 288,583 62,012 29.61 0.0383
5 3627 284,567 61,149 29.62 0.0383
-H444-
-------
APPENDIX I
FLUID BED PROCESS VESSEL SKETCHES
A sample calculation of the minimum fluidization velocity
in the fluid bed sorber is presented in Table 1-97. Vessel
sketches of the major pieces of equipment in the fluid bed
process are shown.
-1445-
-------
TABLE 1-97
SAMPLE CALCULATION OF THE MINIMUM FLUIDIZATION VELOCITY
BY THE METHOD OF LEVA/ ET. AL.
GMF = 688 • Dp 1'82 • [pf (pa-Pf)] °'94
0.88
M
where:
Gj,ip = Mass velocity @ minimum fluidization, Ib/hr ft2
Dp = Particle diameter, in. = 0.0359 in.
p-f = Fluid density, lb/ft3 = 0.0409 lb/ft3
ps = Particle density, lb/ft3 = 70 lb/ft3
p = Fluid viscosity, cp = 0.0286 cp
T => 600°F
p =1 psig
GiVlF 688 • (0.0359)1-82 • [0.0409 (70-0.0409)] °-94
(0.0286)0'88
GMF = 98.9 Ib/hr ft2
Since:
UMF = GMF
PF • 3600
UMTT = 98.9
0.0409-3600
= 0.671 ft
sec
1. Leva, M., McGraw-Hill Book Company, Inc., New York, 1959
-1446-
-------
16. I-??
cyuone ser*K*ro/t
VESSEL PMHFttfl 40 '-o" Oft
SEDHtltKT
STATIC 6R>
GAS DISTIIBUTOK
ft SOLIDS TO SOLIDS
THE M. W. KELLOGG CO
FLUID BED
DRAWING NO
-------
FOR
LOCATION
SUB.EC, FLU/P &££?
JOB No
DATE __.
Dv v
M-H
,-M-
40
/As.
TO
PJVfi-
-1448-
vu.
P \ACKO
7 ?
rue
/A/
-------
ron /Y/T PC/)
LOCATION
SUBJECT
; T- 99 r—/Si\_
rcofer^p- f«llo«i
JOB No.
A/
o T£,S
•v
DATE
BV_.
1 1
T'";
l£*irfflft/£k£
i '
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FIG. I 99
fco/vr'p)
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AT LEAST 4 VANES
'iTS T-c? 8£"
RLQ'-"*! E.J)
; A/73
DRAWN: C? A7
CHECKED:
APPROVED:
DATED: /3 •
ISSUED FOR
FABRICATION
THE M. W. KELLOGG CO.
NO.
DESCRIPTION
DATE BY CHK
11
REVISIONS
DCFIANCC NO. 1A7H
ISSUED FOR
CONSTRUCTION
-1450
cuAsa
a ITEM
AREA
604,5"
JOB NO.
DRAWING NO.
-------
lKF.iinr.cl
~
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fiEV. NO.. BY-
APPROVED: DATED- 4-17-6)
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1 /8"!
1 1 -
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TABLE OF PROCESS CONNECTIONS
SYM. SIZE
_A__ 20"
a' 30"
C JO"
.D to"
MH
NOTES :
1. VESSEL
SERVICE REMARKS
(rfts /AjLirr Stf A/c-f Q
^AS 6vTter 5«f No-US' 1
6ouD5 INLET SfiT A/ST£ <<>'
^5oi(B5-C>iy7i.eT ^ffV^ffTC 7
M/!A/HOL£5 /)^ P?£"Qi).
t-ECHANlCAL TO SPECIFY MATERIALS
OF CONSTRUCT lOli. SEE THE MATERIALS OF
CONSTRUCTION SPECIFICATION Ai;D/0?,
DESIGN
DATA SHEETS ISSUED FCR THIS J03
FOR CORROSION ALLa-.'ANCo , ETC.
2. .MANHOLE SIZES (IF STATED) CORRESPOND
TO
VESSEL
REOyTREMElvTS. IF NOT STATED,
MECHANICAL TO SPECIFY MANHOLE
SIZES REQUIRED BY JOB SPECIFICATIONS
AND/OR
DESIGN DATA SHEETS.
3. V£$$et. S^FLc tf^p" L,lHEf> W/f>" G-VtllTED
C.mSTfi0L£ >f>SUL»T!OK
4. G-*i> Pl5r/P ,v /• ;: Rare IS A FLAT fL/t^r ^/
8C>0 Hour (2 Xe'V AND 4','C-C SPACiuG.
MAVftt.
r> 5i»7'C Lapp o»* PLATE IS I&8.O4O
L C-A5 AP/fC/roi'S PL/»r(? AS 0.2E PSI Mrt/,
LlNEP w/ EROSJOAf' J?£$l??Tfl*fT ft*T*L-
DlP LEG- W/ FLAPPEK TVPF C'5C^^WElt- A,
lOf-^.Le C /s 4ff" HiLLS>D5. INTSITMIL
p;fy/v^ lc"4S*LRe
7 SHSLL A
ioziLe D conuecj-5 TO /o"/vTe«MAL
57/JWC "fPt V»// e'-0."f COWE TOP. 5" A*£6?S .
fi A/o??if
/) HAS /» BfifFLC ryf't i*4Tetrr-i*L
2/srfiifiWK. ' Stft" Awe- /l/l-A"B-/A2.
9 THROUGH INTERNALS PS. REFff
cr\ /AJLF." Bfirric AA-FB-^T-AZ
TITLE " OV.G KO.
(E:NCE DRAWINGS
-------
VERTICAL
OF VESSEL
SEE NOZZLE SYMBOL
FIGURE 1-100
(con't)
2"(MIN.)
.2" SPACE
'(TYPICAL)
D - SPACES
ARE REQUIRED
(TYPICAL)
U NOTE 1
NOTES
1. NUMBER, SIZE AND PLACEMENT OP VANES TO BE SPECIFIED BY VESSEL!MECHANICAL.
AT LEAST 4 VANES ARE REQUIRED.
TABLE OP DIMENSIONS
NOZZLE
SYMBOL
A
DETAIL
A
C
?8"
D
$
DRAWNi
THE M. W. KELLOGG COMPANY
«» PUUMAN |NCORK>*ATfO
BAFFLE
HKVKIOM DMCMIPTION
DAT! »V CMK.
APPNOVIDl
IOUCO FOR
FABRICATION
IISIUKD FOM
CONSTMUCTION
6045*
JOB NO.
OHAWINO NO.
-------
>CAM.JN
JL &
-------
JOB N
PAGF
DATT
-l a.y'A -
(
- 707 AL Of
-------
APPENDIX J
FIXED BED SORBER DESIGNS
Table J-98 shows sorber dimensions for various sorption
times based on the rectangular sorption panel design shown
in Figure J-lll. These are included to illustrate the varia-
tion in bed thickness, sorbent volume (which can be calculated
from the data shown), and pressure drop as sorption time changes.
Sorbent volume, bed thickness, and pressure drop are the same for
other panel configurations (e.g., cylindrical).
Table J-99 summarizes the results of the search for a vendor
to supply the large diameter (20 feet) valves needed in the
fixed bed design.
Two different types of gas-solids contactors are shown in
Figures J-108 and J-109. These have possible applications in
the fixed bed design but were not evaluated in the present study
since no large economic savings or other advantages were apparent.
Figures J-110 through J-119 show details of the rectangular
sorber design which was first considered for the fixed bed process,
It was later abandoned in favor of the cylindrical design shown
in Figures J-120 and J-121 but is included here since it may have
application in some other process where the temperature level and
variation is less than for the alkalized alumina process.
-J455-
-------
TABLE J-98
FIXED-BED SORBER DIMENSIONS
AP
A
B
D
24
24
12
12
6
6
4
4
30
2
s 30
2
30
2
30
2 '
12"
4"
18"
5"
21"
7"
2'-l"
8"
15"
7"
21"
8"
2'-0"
10"
2'-4"
11"
21-1/2"
8"
16"
6-1/2"
12"
5-1/2"
10"
5"
18-1/2"
5"
13"
3-1/2"
9"
2-1/2"
7"
2"
A = Minimum required distance between sorber panels (may be
increased if necessary for mechanical design reasons).
B = Width of sorber panel covers.
C = Width of sorber panels.
D = Sorbent bed thickness.
T = Sorption stream time, hours.
AP = Pressure drop through sorber panel, inches of water.
-J456-
-------
TABLE J-99
FIXED BED SCHEME - VENDOR SEARCH FOR 20 FT. DIAMETER VALVES
VENDOR
Crane Company
Chapman Div. Crane Co.
Keystone Valve Corp.
Henry Pratt Co.
W. G. Rovang & Assoc.
Darling Valve Co.
Wm. Powell Co.
Continental Equip. Co.
Valcon Equip. Co.
Fluid Dynamics Inc.
Babcock & Wilcox
Equipment Div.
S. Morgan Smith Div.
of Allis-Chalmers
W. S. Rockwell Div. of
Rockwell Mfg.
W. S. Rockwell Div. of
Rockwell Mfg.
(When we lowered temp.)
Requested copies of design
MWK developed for 14' x 16'
louver vane valve some years
ago.
Sun Shipbuilding & Drydock Co.
Golden Anderson Valve
Specialty Co.
Hammel-Dahl Div. of ITT
Grove Valve & Regulator Co.
RESULTS OP CONTACT
Nothing to offer.
Nothing to offer.
Not interested.
Not interested.
Not interested.
Not interested.
Not interested.
Beyond their scope.
Beyond their scope.
Beyond their scope.
Will not build their damper type
valves for others.
Have design & production capabilities
but price is very high.
Will not consider valves for
temperatures over 900eF.
Withdrawing from specialty field and
will consider only that work which
will lead to mass production.
No longer in their files.
Will build anything we design and
take full responsibility for.
Think they can provide these valves
economically but have not sub-
mitted anything of value, but
are currently active.
Expect to submit a realistic butter-
fly type design which they will
have built by someone like Sun
Ship. But H-D will be responsible.
Original offering was completely
unrealistic. They have redesigned
and have a retracted disc valve
which can be reasonably priced.
The preliminary design will have
to be refined to water cool or
otherwise take care of the temp-
erature difference between the
valve and duct walls.
-J457-
-------
FIGURE J-103
COUNTER-CURRENT GAS-SOLID CONTACTING PANEL
Arthur M. Squires
treating, solid (16-30 mesh)
TREATED
CAS
sealing solid
{8-12 mesh)
solid
1
I- UNTREATED
GAS
££&>. A
solid
solid-discharge aprons
(struck from beneath) '
-J458-.
-------
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t •;•'•:•'•." •'. • '* •:•"?• ,'• ".
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vx.s-
^fL^^.i:
••*."• 'V '•»'•''•"•'•''-" '..»•'.•••
**
-------
F I6U
RE J-105
CROSS-SE-CTIONAL V/FW OF SORBER FLOU
T/?EATED FLUE GAS
\
ry •
T
-I
\
y
.A.." ..
—f
UNTREATED FLUE GAS
GAS AFTER 90 % OF THE WLET
flf£) THICKveSS WHICH IS A
OP Ap /(
REMOVED
DRAWN:
CHICKED:
APPROVED:
ISSUED FOR
FABRICATION
THE M. W. KELLOGG CO.
3CRBER FLOVPA7TEM
DESCRIPTION
DATE
BY
CHK.
REVISIONS
DEFIANCE MO. ICTH •«!••
(••UEO FOR
CONSTRUCTION
-J460-
ITEM
AREA
JOB NO.
DRAWING NO.
-------
FIGURE J-106
SORBER PANEL
ME5W
suew/qy
PANEL V//UL
D* BCDTWICKN^S VH/CH 15 A FUNCTION
OF A
-J461-
-------
T-/07
FRONT W£V
SOR8ER
40-c,
TRFATFP Ft Of G-AS
PANEL
UNTRFATCD FlUF G-A5
z'-o
'~-°- GRADE
NO.
DESCRIPTION
OATC BY CHK
M C V I S I O N S
DRAWN;
CHICKED.
APPROVED:
OATCO: II'
IISUCD POM
FABRICATION
I*BUKO FOR
CONSTRUCTION
THE M. W. KELLOGG CO.
SORBFR FKOMT V/EW
CLASS
• ITCM
AREA
JOB NO.
DRAWING NO.
-J462-
-------
FIXED BED
FIGURE J-108
SQRBER FRONT VIEW
-J463-
^PRODUCIBLE
-------
FIXED BED
PL
AN
FIGURE J-1Q9
STEEL
-J464-
-------
FIXED BED
FIGURE J-110
TaP € BOTTOM SoRBER P*N£L
IL
-J465-
-------
FIXED BED
BER
FIGURE J-lll
w
X
y
T
op
PANELS
fe'
79
'
I •/
//-854
zo-S"
•19-9
"
20^-4
-J466-
-------
FIXED BED SORBER
FIGURE J-112
r
2L£/£
SUPPORT.
-J467-
-------
FIXED BED SORBER
FIGURE J-113
TOP
PAVEL.
I I
-------
EXTERNAL FRAMING - FLAT PANEL DESIGN
FIGURE J-13J
: _ .*Tr~(- r •• i—rr-*~
-------
FIQUR! .J.r
SORB3R UgljT -
QR
E3IGM*&
n
-J470-
-------
FIGURE J-116
GENERAL CROSS SECTION - SORBENT TRAIN
-J471-
-------
APPENDIX K
SAMPLE CALCULATION: HEAT TRANSFER OF ALKALIZED ALUMINA
IN GAS STREAM
When the solids cooler was deleted and the hot regenerated!.i
solids returned directly to the reactor, it was assumed that the
sorption reaction would not take place until the solids had
reached the equilibrium temperature. The time (and consequently
reactor height), needed;for heat transfer was calculated, using-
several different-methods and sample calculations for these
methods are shown, in this Appendix. For these examples the .
temperature of the regenerated solids entering the reactor is:.
1330°F, which was the preliminary design figure,.-instead of
the final design value of 1355°F. The equilibruum temperature
of 658°F is unchanged. However, since 'the' calculated time
and reactor height varies from method to method, and since
the sorption reaction will likely start at some temperature
above 658°F, no allowances were made for the higher solids
inlet temperature. ,
-K472-
-------
SAMPLE CALCULATION; EQUILIBRIUM HEAT TRANSFER TIME
FOR ALKALIZED ALUMINA PARTICLES IN FLUE GAS STREAM
BASIS: Dispersed Phase Reactor, superficial gas velocity
25 ft/sec.
TEMPERATURE: Solids : 1330 -»• 658°F
Flue Gas: 658 •*• 600°F
AVERAGE PARTICLE SIZE: 1410 microns (0.00462 ft)
PARTICLE DENSITY: 70 lbs/ft3
THEORETICAL SETTLING VELOCITY: 22 ft/sec
REF: Bird, R.B., Stewart, W.E., and Lightfoot, E.N.,
TRANSPORT PHENOMENA, Wiley & Sons, Inc.,
New York, 1960, pp. 357-360.
continued
-K473-
-------
where:
Wg = flue gas flow rate, Ibs/hr.
W = solids flow rate, Ibs/hr
3
C = flue gas heat capacity, Btu/lb °F
C = solid heat capacity, Btu/lb °F
ps \
4 274 341 (7 69/29 fi21 (Data from
.-. B = ' -dTJ A - H-58 Flowsheet
/A ?Q -
(0.29) ,, -p3ig7_D)
'T~ - T where: T = initial fluid
r
T, = initial soli4*rfeeanperature
Tf = fluid temperature @ time t
M . „, i Tf " To 1 _ 17 cB / 658-600
Ad + B) 1 _ T I = 12.58 (1330,600J
*• o i
ast
From Fi«W»ire 11.1-4 (extrapolated): - = 0.31
where: a.
continued
-K474-
-------
A „, „ 0.31 R2 p C (3600)
i t = °'31 R2 (3600) _ s Ps
where:
ks = solid thermal conductivity, Btu/hr-ft-°P
Pg = solid particle density, lbs/ft3
R = particle radius, ft
t = equilibrium time, sec
. t (gec) = 0.31 (0.00231)2 (70) (0.29) (3600) = 0.12
ks ks
Thermal Conductivity of
Reference
Lange ' s Handbook
TPRC Tables
V
of Chemistry
Btu
hr -f t-op
0.392
0.087
•h — 0 1 ? /V
t(sec) °-i2/ks
1.4
. Gurney-Lurie Chart for Spheres
REF: McAdmans, W.H., HEAT TRANSMISSION, 3rd ed.,
McGraw-Hill Book Co., Inc., New York, 1954, p. 40
t - t where: t = temperature of surroundings
3i cL :
^a ~ b t = temperature @ time e
t, = original temperature of solid
t = flue gas average temperature = 629 F
a
cont.^.n -ed
-K475-
-------
629-1330
n = ^— where: r = sphere radius
m
r = radius from center to particular
point
The entire sphere is to be cooled down to 658°F
so r = o, the center point.
.*. n « o
k where: k = solid thermal conductivity,
m = ^- s Btu/hr-ft-°F
h = heat transfer coefficient between
surroundings at ta and surface
(fluid to particle heat transfer
coefficient), Btu/hr-ft-°F
REF: Treybal, R.E., MASS TRANSFER OPERATIONS,
McGraw-Hill Book Co., Inc., New York, 1955
From page 55, Curve #4 for flow of gases past single spheres:
CIO-3) = 0.00058 [Ref(Prf)2^~| °'52
k
where: D = particle diameter, ft
kf = flue gas thermal conductivity, Btu/hr-ft-°F
continued
-------
and Ref = D (vSLIp) Pf(3600)
where: VSLIP = Slip Velocity = sorber superficial gas
velocity-particle settling velocity,
ft/sec
pf = flue gas density, lbs/ft3
P = flue gas viscosity, Ibs/ft-hr
Re
, = 0.00462 (3) (0.04) (3600) OQ .,
f - - - - 28.51
~ f
Pr - Pf - 0.26 (0.07) _ o 758
prf - TE— - O ~ °'758
Treybal Equation becomes :
h = 15.6 Btu/hr-ft2-°F
ks
continued
-K477-
-------
where: 9 = equilibrium timer sec
X = value from Gurney-Lurie Chart,
Figure 3-7, McAdams
.9 (sec) = X(70) (0.2:9)<0.00231) 2 (3600)
ks
0.39X
s
*s
(Btu/hr-ft-"F)
0.392
0.087
h
(Ritu/fer-ft2-0?)
15.6
15.6
ks
m = ^
10.9
2.4
X (Fig. 3-7,
extrapolated &
interpolated)
12.1
2.78
6^= 9.39X/kD
9
(sec)
12
12.5
Particle Heat Balance
Solids : 1330
Flue Gas: 658
(AT)
658°F
600°F
'LM
250.6°F
Q =
= PPVPCPAT
LM
where: p = particle density, lbs/ft3
V = volume of one particle, ft3
continued
-K478-
-------
C » particle heat capacity, Btu/lb-°P
AT = particle temperature drop, °F
(AT) = log mean temperature driving force, °F
A = surface area of one particle, ft2
PpVpCpAT = 70(V) (0.29) (672) - 13*641.6V Btu
hA (AT)^ = 15.6 A (250.6) = 3909.4 A Btu/hr
V * U/6) (0.00462)3 ft3 (particle assumed to
p be spherical)
A- n(0.00462)2 ft2 (particle assumed to be
p spherical)
0.00462 = 0.00077 ft
TV .-nw . m- 13,641.6(0.00077) (3600) Q 7 C0f,
Equilibrium Time = —l ^9^ j = 9.7 sec
-K479-
------- |