ANNUAL REPORT TO
Control Systems Division
Office of Air Programs
Environmental Protection Agency
under
GAP Contract No. EHSD-71-15
CONSOLIDATION COAL CO.
Research Division
Library, Pennsylvania
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DEVELOPMENT OF THE C02 ACCEPTOR PROCESS
DIRECTED TOWARDS LOW-SULFUR BOILER FUEL
ANNUAL REPORT TO
Control Systems Division
Office of Air Programs
Environmental Protection Agency
under
OAP Contract No. EHSD-71-15
September 1, 1970 - November 1, 1971
by
George P. Curran
James T. Clancey
Carl E. Fink
Bedrich Pasek
Melvyn Pell
and
Everett Gorin
CONSOLIDATION COAL COMPANY
Research Division
Library, Pennsylvania 15129
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TABLE OF CONTENTS
I. SUMMARY , 1
II. CONCLUSIONS AND RECOMMENDATIONS . 9
III. FEASIBILITY STUDY 13
A. Introduction 13
B. Feasibility Study - Phase I of Contract 14
1. Process Description - Gasification Section 14
Basic Pressure Case I 14
Alternate Pressure Case II 17
Atmospheric Pressure Case III 17
Feed Coal 20
Preoxidation 20
Gasification 21
Regeneration 22
2. Process Description and Design Basis-
Sulfur Removal and Recovery 22
Sulfur Removal 22
Sulfur Recovery , 23
Case I - Sulfur Rejection from CaS in the
CaC03 Regenerator 23
Recovery of Elemental Sulfur 24
Case II - Sulfur Rejection from CaS in an
External Reactor 26
Recovery of Elemental Sulfur 27
Case III - Sulfur Rejection from CaS 29
Recovery of Elemental Sulfur 29
3. Material and Heat Balances - Gasification Section 29
Basic Pressure Case I 29
Alternate Pressure Case II 32
Atmospheric Pressure Case III 32
4. Economic Evaluation - Gasification Section 48
Basis for Cost Estimation 48
Basic Pressure Case I 48
Alternate Pressure Case II 50
Atmospheric Pressure Case III 5O
5. Integration with Combined Cycle Power Systems 5O
Basis for Evaluation 5O
Supercharged Boiler Cycle 54
Exhaust Gas Cycle with 2800°F Turbine
Inlet Temperature 57
Economics 57
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ii,
TABLE OF CONTENTS (Cont'd.)
Page
6. Low-Sulfur Fuel Gas for Conventional Stations 60
Introduction 60
Basis for Evaluation 63
Economics 63
IV. EQUIPMENT AND PROCEDURE - CONTINUOUS UNIT 65
A. Description of the Unit 65
1. Process Piping and Equipment 65
a. Process Vessels 65
b. Solids Handling 66
c. Gas Flows 67
d. Miscellaneous 69
2. Instrumentation and Control 70
a. Pressure Measurement and Control 70
b. Temperature Measurement and Control 70
c. Control of Solids Flow Rates 71
di Solids Level Control 71
e. Gas Flow Measurements 73
f. Gas Analyzers 73
3. Safety Features 74
B. Materials 90
1. Preoxidizer Feeds 90
2. Gasifier Feeds 9O
3. Regenerator Feeds 90
C. Feedstock Preparation 94
1. Feed Coal Preparation 94
2. Fuel Char Preparation 94
3. Acceptor Sizing 94
D. Operating Procedure for a Typical Two-Vessl
Integrated Run 97
1. Prerun Procedure 97
2. Startup 97
3. Establishing Programmed Conditions 99
4. Routine Run Procedure 99
5. Shutdown Procedure iOO
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iii,
E. Equipment and Procedure for Single Vessel Operation 101
1. Preoxidation 101
2. Gasification 101
F. Sample Calculations 102
1. Preoxidation Runs 102
a. Temperature Profile 102
b. Solids Rate - Run 5P2 102
Feed Rate 102
Preoxidized Coal Product 102
c. Overhead Fines 104
d. Condensate and Tar 104
e. Gas Flows and Composition - Run 5P2 104
Exit Flow 104
f. Bed Conditions - Run 5P2 105
g. Inlet Oxygen Pressure 106
h. Extent of Preoxidation 106
i. Material Balances 106
Carbon Balance 106
Oxygen Balance 106
Coal Balance 107
2. Gasification Runs 108
a. Temperature Profile 108
b. Solids Rates 1O8
c. Gas Rates 108
d. Hydrogen Balance and Concentration 110
e. Outlet Fluidizing Velocity 111
f. Gasification Rate 111
V. RESULTS AND DISCUSSION 112
A. Tabular Chronological History 112
B. Preoxidation Runs 119
1. Introduction and Tabulated Data 119
2. Pittsburgh Seam Coal (Ireland Mine) 126
3. Conclusions - Pittsburgh Seam Coal 130
4. Illinois No. 6 Coal (Hillsboro Mine) 131
5. Conclusions - Illinois No. 6 Coal 132
6 Laboratory Screening for Operability in
the Gasifier 132
7. Probable Cause of Ineffective Preoxidation
at Elevated Pressure 135
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iv.
TABLE OF CONTENTS (Cont'd.)
Page
C. Gasification 136
1. Introduction 136
2. Operability with Respect to Ash Slagging 13?
a. Runs OG2 and OG3 137
b. Runs OG4 and OG5 137
3. Operability with Respect to Caking -
Use of Preoxidized Ireland Mine Coal 137
a. Run 1G3 137
b. Run 2G3 137
c. Run 4G3 140
4. Operability During Gasification
Use of Preoxidized Illinois No. 6 Coal 140
5. Conclusions 140
D. Integrated Operation with Acceptor Circulation 140
1. Introduction 140
2. Run DIB 142
3. The Sulfur Cycle 143
a. Acceptor Sulfur Cycle 143
b. Char Desulfurization in the Gasifier 147
c. Sulfur Balances 148
d. Elemental Sulfur Content of the
Regenerator Offgas 148
4. Acceptor Activity 150
5. Char Combustion in the Regenerator 151
6. Nitrogen Removal from the Gasifier Feedstock 151
7. Kinetics of the Gasification Reactions 152
VI. LABORATORY STUDY OF COAL PREOXIDATION KINETICS 153
A. Introduction 153
B. Experimental 155
C. Rate of Reaction of Oxygen With Coal 157
1. Effect of Temperature 157
2. Effect of Inlet 02 Concentration 157
3. Effect of Extent of Preoxidation 157
4. Effect of Particle Diameter on Rate of
Reaction with 02 16O
D. Particle Density 160
E. Hydrogenation 163
F. Coal Weight Loss 163
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TABLE OF CONTENTS (Cont'd.')
VII. BIBLIOGRAPHY
VIII. APPENDICES
A - Detailed Investment Costs 167
B - Discussion of Gasification Design Basis 174
C - Data and Procedure Used for Sulfur Removal
and Recovery 177
D - Description and Computer Program for Process
Heat and Material Balances 194
E - Literature Review of Preoxidation of Caking Coals 209
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VI.
TABLE OF CONTENTS (Cont'd.)
TABLES Pace
1 Evaluation of Consol C02 Acceptor Process
Pressure vs Atmospheric Operation 2
2 Summary of Power Costs and Investment 5
3 Overall Mass & Heat Balance - Gasification Section - Case I 31
4 Mass and Heat Balance - Preoxidizer - Case I 33
5 Mass and Heat Balance - Gasifier - Case I 34
6 Mass and Heat Balance - Regenerator - Case I 35
7 Mass and Heat Balance - Reductor - Case I 36
8 Mass and Heat Balance - First-Stage Claus Reactor - Case I 37
9 Mass and Heat Balance - Second-Stage Claus Reactor - Case I 38
10 Mass and Heat Balance - Acceptor Sulfur Reactor - Case II 39
11 Mass and Heat Balance - "Wackenroder" Reactor - Case II 4O
12 Mass and Heat Balance - S02 Absorption Tower - Case II 41
13 Overall Mass & Heat Balance - Gasification Section - Case Hi 42
14 Mass and Heat Balance - Preoxidizer - Case III 43
15 Mass and Heat Balance - Gasifier - Case III 44
16 Mass and Heat Balance - Regenerator - Case III 45
17 Mass and Heat Balance - Reductor - Case III 46
18 Mass and Heat Balance - Claus Reactor - Case III 47
19 Investment Summary - Case I 49
20 Direct Operating Cost Summary - Case 1 51
21 Economic Comparison - Cases I and II 52
22 Investment Summary - Case III 52
23 Direct Operating Cost Summary - Case III 53
24 Power Cycles - Case I - Supercharged Boiler 56
25 Power Cycles - Case I - Exhaust Gas Cycle with 2800°F
Turbine Inlet Temperature 59
26 Summarized Power Investment Costs 61
27 Power Costs 62
A Properties of Preoxidizer Feedstocks 91
B Properties of Gasifier Feedstocks 92
C Chemical Analysis of Tymochtee Dolomite 93
D Analysis of Regenerator Fuel Char 93
28 Summary of Preoxidizer Operations 113
29 Summary of Gasifier Operations 116
30 Results of Integrated Gasifier-Regenerator Operations 118
31 Preoxidation Conditions and Results 120
32 Properties of Coal and Products - Preoxidation Runs 121
33 Material Balances for Preoxidation Runs 123
34 Distribution of Oxygen in Products of Preoxidation 124
35 Conditions and Results for Gasification Runs 139
36 Properties of Feeds and Products - Gasification Runs 141
37 Conditions and Results for Demonstration Run - Gasifier 144
38 Analyses of Char Feeds and Products - Run DIB 146
39 Sulfur Balances - Run DIB 149
40 Summary of Results - Laboratory Study of Kinetics of
Coal Preoxidation 154
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vii.
TABLE OF CONTENTS (Cont'd.)
TABLES APPENDIX A
A-l Detailed Plant Investment - Coal Preparation - Case I 168
A-2 Detailed Plant Investment - Gasification - Case I 169
A-3 Detailed Plant Investment - Sulfur Recovery -
Solids Disposal - Case I 170
A-4 Detailed Plant Investment - Coal Preparation - Case III 171
A-5 Detailed Plant Investment - Gasification - Case III 172
A-6 Detailed Plant Investment - Sulfur Recovery -
Solids Disposal - Case III 173
APPENDIX B
B-l Outline of Gasification Design Conditions Affected by
Reaction Kinetics - Case I 175
B-2 Effect of Process Variables on Overall Station Heat Rate -
Case I - With Supercharged Boiler Cycle 176
APPENDIX C
C-l Effect of Temperature and Sulfur Partial Pressure on
Distribution of Sulfur Species 182
C-2 Equilibrium Constants, Gases 183
C-3 Numerical Values of Equilibrium Constants for Table C-2,
Gas Reactions 184
C-4 Heat Capacities at Zero Pressure, Gases 185
C-5 Mean Heat Capacities above 60°F, Gases 186
C-6 Heats of Formation at 25°C, Gas Reactions 187
C-7 Equilibrium Constants for Solids Reaction 188
C-8 Mean Heat Capacities above 60°F, Solids 189
C-9 Heats of Formation at 25°C, Solids 190
APPENDIX D
D-l Program to Calculate Heat and Elemental Balances for 196-208
Production of Low-Sulfur Fuel Gas by the
CO2 Acceptor Process
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viii.
TABLE OF CONTENTS (Cont'd.)
FIGURES
1 Fuel Gas Cost - Effect of Coal Cost 3
2 Power Cost - Effect of Coal Cost 6
3 Schematic Flow Diagram - Boiler Fuel Gas -
C02 Acceptor Process - Case I 15
4 Schematic Flow Diagram - Boiler Fuel Gas -
C02 Acceptor Process - Case II 18
5 Schematic Flow Diagram - Boiler Fuel Gas -
CO2 Acceptor Process - Case III 19
6 Flow Sheet - Sulfur Recovery Section - Case I 25
7 Flow Sheet - Sulfur Removal and Recovery System - Case II 28
8 Flow Diagram - Sulfur Recovery Section - Case III 30
9 Schematic Flow Diagram - Case I - Supercharged Boiler 55
10 Schematic Flow Diagram - Case I - Exhaust Gas Cycle 58
11 Flow Diagram, Revised Continuous C02 Acceptor
Gasification Unit 75
12 Process and Instrumentation Flow Diagram 76
13 Flow Diagram for Continuous Gasification-Preoxidation Unit 77
14 Gasifier Assembly 78
15 Regenerator Assembly 79
16 Details for Gasifier and Regenerator Fittings and Internals 80
17 D-l Gasifier Shell 81
18 D-2 Regenerator Shell 82
19 C-2, C-3 and C-5 Coolers 83
20 B-l Saturator 84
21 G-l Cyclone 85
22 C-l, C-4 and C-6 Coolers 86
23 Rotary Feeder L-4 87
24 Rotary Feeder L-l 88
25 Schematic Diagram - Char-Acceptor Interface Probe Circuit 89
26 Feed Coal Preparation 95
27 Acceptor Sizing 96
28 Preoxidizer Temperature Profile, Run 5P2 103
29 Gasifier Temperature Profile, Run 5G3 109
30 Relationships Among Preoxidation Variables 125
31 Fluidity of Ireland Mine Coal and Preoxidized Products 134
32 Configuration of Entry Points in Gasifier 138
33 Schematic Diagram - Apparatus for Study of Coal
Preoxidation Kinetics 156
34 Reaction Rate Versus Cumulative % Preoxidation 158
35 Cumulative $ Preoxidation Versus Time 159
36 Particle Density After Preoxidation 161
37 Particle Densities for Preoxidized Ireland Mine Coal 162
38 Dehydrogenation Versus $ Preoxidation 164
39 Coal Weight Loss Versus % Preoxidation 164
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IX.
TABLE OF CONTENTS (Cont'd.)
FIGURES APPENDIX C Page
C-l Sulfur Vapor Pressure 191
C-2 Equilibria for Some Reactions Involving CaS - I 192
C-3 Equilibria for Some Reactions Involving CaS - II 193
APPENDIX D
D-l Flow Diagram for Computer Program 195
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I. SUMMARY
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1.
I. SUMMARY
A. Feasibility Study
The feasibility study called for in Phase I of this contract was
directed at determining the economic and technical feasibility of the adapta-
tion of the C02 Acceptor Process to the problem of producing low-sulfur, low-
Btu power plant fuel by the gasification of bituminous coals. Combustion of
fuel gas makes possible a drastic reduction in NOX emissions, compared with
conventional PF stations equipped with stack gas cleanup processes.
The first part of the feasibility study was directed at comparing
atmospheric versus pressure operation of the gasifier for the objective of
preparing low-sulfur boiler fuel gas for an existing boiler. Capital and
operating costs were estimated for two conceptual variations of the C02
Acceptor Process operated under 15 atm pressure, and for one lime bed gasifi-
cation system operated at atmospheric pressure.
All economic figures presented here are based on 1976 operation and
include escalation at 7-1/2$ per year on materials and labor and interest at
7-1/2$ per year during construction.
The product gas is delivered to the boiler in both cases at the
same pressure level of 25.7 psia. The results of this part of the feasibility
study are summarized in Table I, comparing one of the pressurized cases with
the atmospheric case. Figure 1 shows the impact of coal cost on the cost of
low-sulfur boiler fuel for both cases. It is clear that pressure operation
is preferred for this operation.
These calculations indicate that pressure operation not only produces
a cheaper fuel gas but also effects a more efficient overall sulfur removal.
The sulfur removal efficiency can be as high as 96% from a coal containing
4.3$ sulfur. The sulfur in the gas is thus equivalent to that in a O.3$
sulfur fuel oil.
The above comparison was based on Case I the lower cost of the two
pressure cases studied. The two pressure cases differ from each other in the
method of regeneration of the elemental sulfur. Both involve C02 and H2S
acceptor reaction in the gasification section and deliver an equally low-
sulfur pressurized producer gas. The second pressure case (Case II) involves
a more expensive sulfur recovery section which adds about five percent to the
capital and operating cost. New experimental data which have been obtained
since this report was written emphasizes that Case II is technically more
feasible than Case I. Although Case II is slightly more expensive, the differ-
ence in costs is well in precision of our estimate. The final conclusions
remain the same. The second pressure case is also more attractive economically
than the low pressure case. The two pressure cases are described in detail in
Section III.
The second part of the feasibility study was concerned with utiliza-
tion of the high pressure clean producer gas in an improved combined cycle
power plant. Capital and operating costs were estimated for the first
pressurized variation of the C02 Acceptor Process, described above, integrated
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2.
Evaluation of Consol C02 Acceptor Process
Pressure vs. Atmospheric Operation
Generation of Fuel Gas for Conventional Boiler
70$ Plant Operating Factor = 6132 hr/yr
Coal Required
Lb/hr (dry basis)
MM Btu/hr (HHV)
Product Gas
Mols/hr
MM Btu/hr (HHV)
MM Btu/hr (HHV + sensible heat)
Temperature, °F
Pressure, psia
% Sulfur Cleanup
Plant InvestmentC1)
Annual Operating Costs (ex. coal)
Coal at 4O^/MM Btu
Capital Charges at
Sulfur Credit at $15/LT
Power Credit at 9 mills/KWH
Net Expenses
Fuel Cost Delivered to Power Station,
^f/MM Btu (HHV)
Fuel Cost Delivered to Power Station,
Btu (HHV + sensible heat)
Case I
Pressure Operation
15 atm.
Case III
Atmospheric Operation
1,087,600
— 13,812
234,980
10,287
11,466
665
96
$112,700,000
$10,090,OOO/yr
$33,890,OOO/yr
$16,900,OOO/yr
(&L,830,000/yr)
($9,970_.OOO/vr)
$49,080,000/yr
25.7
77.8
69.8
158,833
10,078
11,472
980
86
$141,700,000
$ll,410,000/yr
$33,890,OOO/yr
$21,260,000/yr
($l,650,000/yr)
$64,910,000/yr
105.0
92.2
(i) 1976 operation, includes escalation and interest during construction.
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into two different conceptual power cycles. The first utilizes a super-
charged boiler ahead of the gas turbine and is similar to the conceptual
cycle now under study by Westinghouse Research Laboratories under Contract
CPA-70-9 with the Office of Air Programs on fluidized combustion. This
cycle utilizes currently available gas turbine technology with an inlet
temperature of 1800°F.
The second power cycle uses the one proposed by the United Aircraft
Research Laboratories (UARL) in their contract work with GAP on Advanced
Power Cycles. It uses a high-temperature (2800°F) gas turbine with a heat
recovery boiler on the turbine exhaust. Such a high-temperature turbine is
well beyond the capabilities of current technology. UARL feels that the
turbine technology required could be available for this cycle in the 1980's.
The investment, operating costs, and heat rates for the two com-
bined cycles are compared with each other and with that of a conventional
station equipped with S0a removal and sulfur recovery facilities in Table
2. A 15$ capital charge was used for both conventional and future power
stations.
Figure 2 shows the impact of coal cost on power costs for the
different cases.
The values given for the two combined cycle cases in Table 2 and
Figure 2, represent the base case for utilization of the Consol C02 Acceptor
Process. The base case delivers clean hot producer gas at 1300°F to the
power station. It is assumed that the gas can be cleaned sufficiently well
by use of high pressure drop cyclones to permit its use in the combined
cycle power plants. A more conservative case wherein the gas is cleaned by
wet scrubbing is discussed in Section III.
In drawing conclusions from the figures presented in Table 2 and
Figure 2, it should be recognized that the investment figures are only
approximate. The bases upon which they are derived are likewise given in
Section III. The figures given for overall station heat rates are, however,
relatively precise. It should also be pointed out that economic optimization
of the operation of the gasification system or its integration with the power
cycle was not attempted. It was felt that optimization was not warranted
until more experimental data were obtained at conditions required for the
modified C02 Acceptor Process.
The numbers in Table 2 and Figure 2 illustrate the potential of the
CO2 Acceptor Gasification Process in producing a higher degree of sulfur
removal, i.e., 96$, more efficient use of coal for power generation and power
at a lower cost as compared with a conventional station equipped with stack
gas scrubbing facilities.
The figures indicate a potential improvement in power cost with
both power cycles considered, but the potential gain is much greater with
the more advanced UARL gas turbine cycle.
In Table 2 the net power outputs of the combined cycle cases are
considerably greater than that for the conventional PF station. As shown
in Section III (Feasibility Study), multiple gasification process units are
used. Therefore, the investment/KW would be little affected by reducing the
output of the entire combined cycle power station.
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5.
Summary of Power Costs and Investment
1976 Operation. See page 48 for plant cost basis.
70% Plant Operating Factor
System
Busbar Power, MW
Net Power, MW
$ Sulfur Removal (as elemental
sulfur)
Overall Station Heat Rate,
Btu/Net KWH
# Busbar Power From
Steam Cycle
Gas Cycle
System Investment , $ KW Net Power
Power Station
Sulfur RemovalC3)
Fuel Gas Production Plant
Total Station
Conventional
Coal-Fired
With Stack
Gas Scrubbing
loooC1)
1000
90
8900(2)
100
O
218.4
31.1
Case I - Gasification With
Power Cost «
Coal at 4O^/MM Btu
Gasification
Direct Operating Cost
Capital Charges at 15$ Inv.
Power Cost for Air Compression
Power Plant
Direct Operating Cost
Capital Charges at 15$ Inv.
Sulfur Credit at $15/LT
Total
249.5
3.56
.77(3)
.76(3)
.48
5.34
-.15
10.76
Supercharged
Boiler
1562
1544
96
8930
67.2
32.8
149.2
Exhaust Gas Cycle
With 28OO°F
Turbine Inlet Temp,
2053
1867
96
7390
35.3
64.7
134.8
60.2
195.0
Mills/KWH Busbar Power
3.54
.94
1.48
2.69
(0 After subtraction of power for auxiliaries.
(2) Heat rate is 8900, based on coal burned in boiler. 640 Btu/KWH as fuel
equivalent to generate reductant for sulfur recovery and power and steam
requirements to drive process equipment.
(3) S02 removal and sulfur recovery plant via Consol's Formate Process. Fuel
cost is included in direct operating cost at 80^/MM Btu as purchased
natural gas. Investment does not include cost of separate scrubber to
remove residual particulate matter prior to S02 removal which may be
required.
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STATION!WITH STACK
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B. Pretreatment Studies
The continuous unit which was used for the experimental study of
the modified C02 Acceptor Process in Phase II-A of the contract was
reconstructed and put into operation.
The main reactor was used initially to study continuous preoxidation.
The preoxidation study on the continuous unit was preceded by a literature
search which is included in Appendix E and by a brief exploratory laboratory
study of the preoxidation kinetics. These two companion studies did not eluci-
date quantitatively all of the features of the preoxidation kinetics. They did
show, however, that the reaction is very rapid and that the preoxidation
kinetics are thus not an important process problem. These findings were con-
firmed by subsequent runs in the continuous unit.
The continuous unit was operated primarily to define the limits of
operability in the preoxidation step and the minimum conditions of severity
of pretreatment required to establish operability in the subsequent gasifica-
tion step. Two coals were used in this study which represent the extremes
in caking properties among the high-sulfur bituminous coals in the eastern
half of the United States. Both coals were used in the 24 x 100 mesh size
range.
The first coal, an Illinois No. 6 coal from the Hillsboro Mine,
represented the more weakly caking coals. It was shown that this coal
required pretreatment to establish freedom from caking in the subsequent
gasification step. Operability was, however, readily established in both
the pretreatment and gasification steps within the framework of preoxidation
severity prescribed in the feasibility study, i.e., 8.5 wt % preoxidation.*
The second coal was a Pittsburgh seam coal from the Ireland Mine
and represents the most highly caking coals. Operability limits were roughed
out for both the pretreatment step and the severity of pretreatment required
to establish operability in subsequent gasification. Preoxidation is operable
at 75O°F but inoperable at 800°F. Two-stage preoxidation is operable, wherein
the first stage is conducted at 750°F and the second stage is conducted at
8OO°F but the product is of very low density and thus not readily suitable for
subsequent gasification.
The minimum severity of preoxidation required in the case of
Ireland Mine coal to establish operability in the subsequent gasification
greatly exceeds the 8.5$ preoxidation level called for in the feasibility
study. The lowest level of preoxidation found to date which satisfies the
above objective is 28 wt %. Further work should reduce the required amount
of preoxidation.
A simple laboratory test has been developed which provides a good
screening procedure for testing the operability in the gasifier of preoxidized
coals.
* Severity of preoxidation is defined as;
Wt. % preoxidation = 100(lb 02 reacted)/lb dry feed coal.
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C. Gasification Studies
Studies were made in the continuous unit of operability in the
gasification step at 1700°F not only with respect to caking but also of
potential ash slagging due to insufficient rate of dissipation of the heat
of combustion of the coal with incoming air. The experimental gasifier was
operated at conditions simulating the full range of conditions in an envis-
ioned full-scale unit. No ash slagging was detected in any of the runs even
at oxygen partial pressures up to 2.4 atmospheres.
A limited amount of data were obtained as a byproduct of the
gasification tests performed here and in the demonstration run discussed
below on the question of gasification kinetics. The data indicate that the
average gasification rate will be roughly comparable to the value assumed
in the feasibility study.
D. Integrated Operation
A single demonstration run was carried out during this contract
period using Disco char, a precarbonized coal char, as feedstock in which
integrated operation of the gasification and regeneration steps were performed.
The run encompassed circulation of acceptor through the two process steps.
The purpose of the run was to demonstrate the main features of the process and
to delineate any special problem areas.
The duration of the run was limited to 13 calcining-recarbonation
cycles. However, the main features of the process were demonstrated. These
include efficient removal of H2S from the gas by the acceptor in the gasifier,
lack of ash slagging in the regenerator at temperatures up to 1860°F> and
rejection of sulfur from the acceptor in the regenerator. Acceptor life as
determined indicates that the equilibrium activity will be higher than that
estimated in the feasibility study. The Case I sulfur recovery system was not,
however, fully demonstrated.
The run showed that a high degree of desulfurization can be achieved
in the product char from the gasifier. The organic sulfur almost completely
removed and the residual sulfur was largely tied up as CaS in the char ash.
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II. CONCLUSIONS &
RECOMMENDATIONS
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II. CONCLUSIONS AND RECOMMENDATIONS
A. Feasibility Study
The CO2 Acceptor Process as modified to produce low-sulfur boiler
fuel has potential in two different types of applications as listed below:
1. Production of low-sulfur, low-Btu fuel gas for
an existing boiler.
2. Production of low-sulfur, low-Btu fuel gas for
a new combined cycle power plant.
An offshoot of the first application is also technically feasible,
i.e., wherein both low-sulfur producer gas and char are produced as boiler
fuel. This case is now being considered as an extension of this contract.
It has been shown in the feasibility study that operation of the
gasification system under pressure is advantageous for both applications as
compared with atmospheric pressure operation.
The major incentive for development of the modified C02 Acceptor
Process resides in the second application. The integrated process has the
potential of producing power at a lower cost, with lower emissions of S02 and
NOX, and with a higher efficiency of utilization of coal than is possible in
a conventional station equipped with stack gas treating facilities.
The higher efficiency of such a power cycle is of great national
importance in view of increasing fuel costs and the need to conserve our
fossil fuel reserves.
The first application noted has less scope in the long run but can
have shorter term use. Full utilization of the second application cannot be
expected until the 198O1s while the first can be put into use within the
decade of the 70's. The first application is limited to situations where for
various reasons it is not desirable or possible to add stack gas clean-up
facilities to existing boilers. The alternative here is the use of low-sulfur
fuel oil (ca. 0.3 wt $) which is not presently available in significant quanti-
ties. The C02 Acceptor Process can under certain circumstances, i.e., at
inland locations close to coal supplies, potentially compete economically with
imported low-sulfur fuel oil. It should also be noted that this application
also has a better potential for NOX control than a conventional station burn-
ing either coal or fuel oil.
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10.
B. Pretreatment Studies
The coal feed to the gasification step must be pretreated to estab-
lish operability in this step, i.e., to eliminate defluidization problems
due to agglomeration and caking. The pretreatment step studied here was via
preoxidation under full system pressure as called for in the feasibility
study. This process was studied because it comprises the most convenient
and economic way of carrying out the pretreatment step.
Operability limitations exist in the pretreatment step itself which
are reflected in the maximum temperature that can be used. The maximum operable
preoxidation temperature in the case of the most highly caking Pittsburgh Seam
coals was found to be in the range of 750-800°F. For the more weakly
caking Illinois No. 6 coals, operability is good at 800°F and the upper limit
was not established.
Operability in both the pretreater and the gasifier is adversely
affected by operation at elevated pressure as is desirable from the economic
point of view. Also, a coal pretreated at atmospheric pressure is more
severely decaked than one pretreated to the same degree of severity under
pressure.
For the Illinois coal, the 8.5$ preoxidation level required for
adiabatic operation of the preoxidizer at 8OO°F is adequate to prevent caking
in the gasifier. However, for the Pittsburgh Seam coal, a much more severe
level of preoxidation (~ 28$) is required to prevent caking.
The effectiveness of the pretreatment process for a given level of
preoxidation in decaking the coal increases with increasing preoxidation
temperature and with decreasing particle size. An upper limit on the preoxi-
dation temperature that may be used, however, is posed not only by operability
in the preoxidizer but also by the properties of the preoxidized product.
The particle density of the preoxidized coal, for example, decreases with
increasing preoxidation temperature and thus reaches a point where it becomes
unsuitable as a gasifier feedstock. The particle density under comparable
conditions is higher the smaller the size of the pretreated coal. It also
increases as the oxygen partial pressure in the preoxidation increases.
The preoxidation rates under process conditions are very rapid and
thus offer no operating problem. Sufficient kinetic data were not obtained to
derive a quantitative correlation of the effect of the variables on rate. The
qualitative trends were in line with expectations from the literature on the
subject. One interesting observation was made, however, in that the rate
tended to become independent of particle size as the preoxidation temperature
was increased.
Further experimental work is required to define conditions where
the severity of pretreatment for the most highly caking coals may be substan-
tially reduced over the levels so far demonstrated. Some particular subjects
which should be studied are listed below;
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11,
1. Effect of more rigorous control of particle size with
particular reference to elimination of coarse sizes.
2. Effect of oxygen partial pressure on degree of decaking
achieved at a given preoxidation level.
3. Study multi-stage preoxidation (one or more stages) with
increasing temperature between stages.
4. Same as above with the first stage carried out at atmos-
pheric pressure.
Process studies should also be carried out in entrained phase pre-
oxidation with a rising temperature regime as well as on other pretreatment
methods which do not involve preoxidation.
C. Gasification Studies
The gasification step of the process must not only be operable
from the point of view of caking but also it must be free from slagging due
to ash fusion. It was concluded from the experimental work carried out under
this contract that the gasification step will be free of operating difficulty
from either cause if a properly pretreated feedstock is used.
The gasification kinetics as determined from the limited work
carried out to date is of the same order as used in the feasibility study.
More comprehensive experimental data are required to delineate the effects
of temperature, carbon burnoff, coal type and gas composition on the gasifi-
cation rate. It is recommended that such data be obtained in future work.
D. Integrated Operation
The main features of the process were demonstrated in a single
demonstration run. The most important conclusions from this run were that
acceptor activity is high enough to efficiently remove H2S from the product
gases in the gasifier, the acceptor activity for C02 absorption and its rate
of decline is within or superior to the values estimated in the feasibility
study, and that ash slagging in the regenerator is not a problem at least at
operating temperatures of 1860°F or below.
Although rejection of sulfur from the acceptor was demonstrated,
all ramifications of sulfur recovery via the Case I process were not proven.
In particular, operation at the CO levels postulated in the feasibility study
was not achieved nor was the recovery of elemental sulfur in the predicted
quantity proven.
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12.
Further work is recommended to obtain more quantitative data on
the following questions:
1. Operating limitations in the regenerator such as
temperature, etc., as dictated by ash slagging.
2. Equilibrium acceptor activity as a function of process
conditions in integrated operation.
3. Limitations imposed on recovery of sulfur as SO2 and
elemental sulfur via the Case I route.
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III. FEASIBILITY
STUDY
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13.
III. FEASIBILITY STUEY
A. Introduction
The process development described herein is a logical extension of
the development of the C02 Acceptor Process now being carried out under the
auspices of the Office of Coal Research. The Office of Coal Research work is
directed towards the production of high-Btu gas from low-rank Western coals.
A pilot plant to demonstrate the process is now nearing completion at Rapid
City, South Dakota. The basic technology, process economics and prior bench-
scale research on the process as applied to low-rank coals are described in
detail in reports to the Office of Coal Research^1)
The modified process now under development under the auspices of the
Office of Air Programs (CAP) is aimed at the production of low-sulfur boiler
fuel rather than high-Btu gas. Another major distinction is that the process
is directed towards the utilization of high-sulfur Eastern bituminous coals
rather than low-sulfur Western coals.
The boiler fuel is either low-sulfur producer gas or a mixed product,
i.e., producer gas and low-sulfur char. These products may be utilized either
in a conventional boiler or in a new combined cycle plant. Generally speaking,
it would be desirable to utilize the producer gas in a combined cycle power
plant. The integration of coal gasification processes with combined power
generation is, of course, an old concept. As a matter of fact, a 17O MW power
plant utilizing this principle and the Lurgi pressurized gasification process
is now under construction in Germany; The plant has been described by Rudolph
in a recent paper.(a) Similarly, United AircraftC3) in a study conducted for
OAP have discussed the use of a slagging entrained phase pressurized gasifica-
tion process to supply gas to a gas turbine based advanced power cycle.
Where low-sulfur char is produced as a co-product, it could be used
as fuel for a fluidized-bed boiler or could be burned along with the gas in a
conventional boiler.
The modified process as originally conceived was described in a
paperC4) presented before the Second International Conference on Fluidized
Combustion. Some changes, of course, have been made in the process concept
since initiation of work under this contract, and they are discussed later,in
this report.
The process differs from other gasification processes previously
proposed in that it removes sulfur within the process proper thus obviating the
need for sulfur removal from the product gas. The Lurgi gasification process
does have the advantage over this as well as over all other gasification
processes proposed in that it is fully developed and is available for commercial
exploitation. The Lurgi process is, however, limited to the use of 4-1/8-inch
size coal and to, coals that are not too strongly coking. The modified C0a
Acceptor Process under development for OAP should be free of these limitations.
The initial phase of the program was a process design study to
determine economic and technical feasibility. The study included a comparison
of atmospheric and pressurized versions of the gasification process. Integration
of the- gasification process with combined cycle power generation was considered
in order to arrive at an understanding of potential merits of the system.
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14.
Simultaneously with the process design study, construction of
experimental equipment to study key features and potential problem areas in
the laboratory has been completed. The major process features involved are
listed below. The order corresponds more or less to the time sequence of the
experimental program.
1. Coal pretreatment and elimination of coking problems
in the gasifier.
2. Methods of introduction of air into the gasifier to
avoid ash fusion problems.
3. Methods of rejecting and recovery of sulfur from the
acceptor.
4. Acceptor life and activity in cycling through process
conditions peculiar to the modified process. As a
corollary, methods of reconstituting spent acceptor
will be studied.
5. Operating limitations imposed by ash fusion in the
regenerator.
6. Kinetics of gasification of Eastern coals under
required process conditions.
7. Kinetics of acceptor reactions with C02 and H2S in
the gasifier.
If the above potential problem areas are successfully resolved in
the laboratory, then it will be possible to consider the next stage in
development of the process. This is demonstration of the process at the
Rapid City pilot plant after required modifications to the plant are made.
The plant as now designed is dedicated to the production of high Btu gas
from low-rank Western coals.
B. Feasibility Study - Phase I of Contract
1. Process Description - Gasification Section
The size of the gasification plant was established to produce
100O MW of electricity by combustion of the cold "washed" product gas in a
conventional boiler at a heat rate of 10,200 Btu/KWH. Since most of the cycles
considered here generate electricity at a much higher efficiency, the net power
output for the different cases varied and was generally much higher than
10OO MW, i.e., up to 1867 MW. Since multiple process units are used, the in-
vestment/KW would be little affected by reducing the size of the gasification
plant.
Basic Pressure Case I (Dwg. XF-3143)f Figure 3
The basic pressure case is shown schematically in Figure 3.
Pittsburgh seam coal (13,884 tons/stream day) is charged to the gasification
section of the plant by a system of belt and distributing conveyors, and is
reduced in size from 4" x 0 to approximately an 8 Tyler mesh x O size consist.
This size reduction is accomplished in two successive stages. Both stages
employ multiple units of double-cage mill units to minimize excess fines
production.
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16.
The prepared coal is then lockhoppered into fluidized, air-
blown preoxidizers which operate at approximately 800°F. Twelve parallel
units are supplied. For this study, it is assumed that a preoxidation level
of about 8.5 wt % oxidation is required with highly-caking Pittsburgh seam
coals to maintain operability in the gasifiers. Experimental work under
this contract will define the required level more precisely.
The total effluent from the preoxidizers is entrained overhead
and passes to twelve parallel fluidized gasifiers together with additional
air, steam, and tail gas from the Claus plant. Calcined acceptor from the
regenerator flows in at the top of the gasifier bed, trickles down through
the bed, and concentrates in a fluidized "boot" at the base. The main
fluidized bed of the gasifier operates at 1700°F and 15 atm pressure. These
operating conditions for the gasifier are dictated by the gasification rate
of the fixed carbon in the coal, and by the kinetic and thermodynamic con-
straints of the C02 acceptor reaction. These constraints are discussed in
greater detail in the section on Design Basis. Most of the sulfur in the
coal reacts with the calcium oxide in the acceptor to form calcium sulfide.
About 65% of the fixed carbon in the coal is gasified.
The remaining char from the gasifier bed is conveyed in a
stream Of inert gas to four parallel regenerators where'lt is burned with a
deficiency of air at about 1880°F. Partially carbonated acceptor from the
"boot" of the gasifier is also conveyed to the regenerator in an air stream.
At the operating conditions of the regenerator, the calcium carbonate in the
acceptor is calcined, and the calcium sulfide in the acceptor reacts to
release sulfur dioxide and sulfur vapor. As noted, the calcined acceptor
then flows back to the top of the gasifier bed completing the acceptor "loop".
Of the heat required to maintain the gasifier at 1700°F, about 20% is supplied
by the sensible heat in the calcined acceptor, about 3O% by the chemical
reaction of carbon dioxide with the calcium oxide in the acceptor, and the
remaining 50% by combustion of a portion of the char with air.
Sulfur-bearing gas from the regenerator is cooled by preheat-
ing the Claus tail gas and by raising a portion of the required gasifier
steam. After passing through electrostatic precipitators, approximately one-
third of the gas flows to four parallel packed beds of alumina catalyst
(called reductors) operated at a temperature of 800°F. In the reductors,
practically all of the S02 in the gas is reduced to S2, H2S, and COS. The
effluent gas from the reductor is then remixed with the by-passed regenerator
offgas. The molal ratio of (H2S + COS)/S02 is then exactly 2.0.
This stream is then fed to a modified two-stage Claus system
where about 96% of the sulfur in the regenerator offgas is recovered as
liquid sulfur. The tail gas from the Claus system is then preheated to 1200°F
(by indirect heat exchange with the regenerator offgas) and charged to the
gasifier. This C02-rich gas provides the partial pressure driving force for
recarbonation of the acceptor and thereby supplies about 3O% of the gasifier
heat requirement, as noted above.
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The product gas from the gasifier is cooled to about 1300°F
by raising most of the required gasifier steam. In the simplest conception
of this process, the gas (234,980 mols/hr of 115.5 Btu/SCF HHV) would then
be delivered directly to the power station at a temperature of about 1300°F
and at a pressure of about 206 psia. If later experience shows that the gas-
fired turbine in the power station is too sensitive to any alkali content in
the product gas, then the gas would readily be water-scrubbed in a venturi-
type scrubber to the gas dew point of about 240°F and then reheated by
indirect heat exchange with itself to a delivery temperature of about 110O°F
and a pressure of about 200 psia.
Unburned char and char ash from the regenerator, and the
rejected fraction of the acceptor are lockhoppered out of the system and
water quenched in a circulating black water system. The solids in the black
water circuit are concentrated by passage through conventional hydroclones,
and stripped with a gas stream rich in carbon dioxide to convert the calcium
sulfide in the acceptor to calcium carbonate. All of the rejected solids
are then disposed of to an ash pond as in conventional wet-bottom, coal-fired
boiler systems.
Alternate Pressure Cape II (Dwg. XF-3144), Figure 4
The alternate pressure case differs primarily in the method
by which the sulfur is removed from the acceptor. In this case, the regenera-
tor is operated with very slight excess air and the sulfur is not released
from the acceptor. One-third of the circulating acceptor leaving the
gasifiers is charged to a fluid bed of acceptor where it reacts at 1300°F
with an equimolar mixture of C02 and water vapor to convert the calcium
sulfide to calcium carbonate. The effluent gas from this reactor is then
passed to a packed column where it mixes in a concurrent flow with a liquid
stream of sulfurous acid. This reaction produces liquid sulfur at a tempera-
ture of 328°F and is a variation of the Wackenroder (5) reaction. Finally, a
portion of the liquid sulfur (about one-third) is burned to produce the
required sulfurous acid via an S02 absorption column.
In all other respects Cases I and II are quite similar.
Atmospheric Pressure Case III (Dwg. XF-3145)f Figure 5
Under the terms of this OAP contract, the C02 Acceptor Process
at pressure is to be compared with a similar case at atmospheric pressure.
Under these conditions, there is no reaction of C02 with calcium oxide and
the acceptor serves only as a sulfur acceptor and as a sensible heat carrier.
Again, coal is prepared to 8 Tyler mesh by zero, preoxidized at 8OO°F and
gasified at 1700°F with steam and air. A small amount of gasifier product gas
is recycled to the bottom of the "boot" to prevent oxidation of the acceptor
CaS by the incoming steam. The sensible heat in the gasifier gas is used to
raise the required gasifier steam, and to raise the plant steam required to
run the main air compressors. The final product gas (158,833 mols/hr of
167 Btu/SCF HHV) is then delivered to the power station at about 98O°F and
25.7 psia. Again, if water scrubbing of the gas is required, it could be
reheated to about 750°F.
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20.
The regenerator is operated at 1900°F with stoichiometric air
and the calcium sulfide is converted to calcium oxide with the release of
S02. The regenerator offgas is cooled by means of steam generation and
the gas is passed to a catalyst-packed redactor and a one-stage Claus plant
to convert and recover the sulfur as liquid sulfur. The Claus tail gas,
still containing some H2S, cannot be vented directly but has to be passed
through the power station boiler fire box to convert the H2S to S02.
As this is a feasibility study, optimization of each processing
step by an extended study of alternates was not attempted. The following
sections show the basis for design and the assumptions used for each major
processing step. Some further details are given in Appendix B.
Feed Coal
This study is based on the use of a high-sulfur content, highly
caking Pittsburgh seam coal. The analysis of a typical coal of this seam
is shown below:
Feed Coal Analysis
Moisture, as received 6.0 Wt %
Ultimate Analysis, MF Basis
Hydrogen 4.8 Wt
Carbon 69.8 Wt %
Nitrogen 1.2 Wt %
Oxygen 7.6 Wt %
Sulfur 4.3 Wt #
Ash 12.3 Wt i»
Higher Heating Value, MF Basis 12,700 Btu/lb
The coal is assumed to be delivered to the plant at run-of-mine
size of 4" x 0.
Preoxidation
Highly caking Pittsburgh seam coal requires a pretreatment step
(in this case, preoxidation) to maintain operability in an air-blown
gasifier. The degree of preoxidation, and the temperature required, will
be determined experimentally under Phase Il-A of this contract. For this
feasibility study, it was assumed that a preoxidation level of about 8.5
wt % at a temperature of 8OO°F would be adequate pretreatment. Operating
conditions and yield structure are directly based on prior Consolidation
Coal Company pilot plant data for atmospheric pressure operation.
The pertinent design data incorporated into this study are
summarized below:
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21,
Preoxidizer Design Constraints
Pressure Operation Atmospheric Pressure
Cases I and II _ Case III _
Temperature, °F ^ - 800 -
Outlet Pressure, psia 224 41
Percent PreoxidationC1) 8.6 8.8
Oxygen Utilization, $ ^* - 78
Partial Pressure Oxygen, psia
Inlet 47.7 10.1
Outlet 7.6 1.4
Ln Mean 21.8 4.4
Outlet Fluidizing Velocity, ft/sec « - 1.25
Avg. Solids Residence Time, min. 10 50
(i) 100 (ib 02 consumed)/(lb MF coal).
As is obvious from the above, it was assumed that the rate of oxida-
tion is directly related to the oxygen partial pressure.
Gasification
There are numerous design conditions and constraints that are of
importance in fluidized, air-blown gasification of a caking coal. Consol
has been actively engaged in research in this field, and in the use of
an active acceptor to supply heat to the gasifier for over twenty years. The
following compilation of bases and assumptions was drawn from this background
information.
Gasifier Design Basis
Pressure Operation Atmospheric Pressure
Cases I and II _ Case III _
Temperature, °F «• 17OO
Outlet Pressure, psia 220 35
Outlet Fluidizing Velocity, ft/sec -« 1.09
Acceptor
Type «« Dolomite •
Wt ia Impurities -* 15
ActivityC1) .13
Make-up Rate(2), $ 0.5 0.2
Avg. Residence Time in Bed,
# Fixed Carbon Gasified « 65(3)
% Hydrogen in Residual Char •« O.4
Outlet Gas Composition
Partial Pressure CO2, atm 1.35 O.14
Partial Pressure H20 (v), atm 1.81 O.23
Shift Equilibrium ^ Yes
Equilibrium in CaO + H2S = CaS + H20 « Yes
Steam Conversion, % 37 58
Mean Gasification Rate, Atoms fixed C
Gasified/atom C in bed/min. 75 x 10~4 15 x 10~4
(*) (mols MgO-CaC03 formed in gasifier)/(mols MgO-CaO fed to gasifier).
(2) 1OO (mols/hr MgC03'CaC03 added)/(mols/hr MgO-CaO circulated).
(3) Approximately 8O$ of total carbon fed.
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22.
Regeneration
Regenerator Design Conditions
Pressure Operation Atmospheric Pressure
Cases I and II Case III
Temperature, °F 1883 19OO
Outlet Pressure, psia 220 35
Outlet Fluidizing Velocity, ft/sec 1.6 2.5
$ Carbon Burnout * 90 ft
Avg. Acceptor Residence Time, min. 10 20
Outlet Gas Composition
Partial Pressure CO, atm 0.38 0
Partial Pressure C02, atm 4.4 0.44
Shift Equilibrium < Yes __________
2. Process Description and Design Basis -
Sulfur Removal and Recovery
This section gives a more complete description of the sulfur
removal and recovery features than has been discussed above. The design
basis is also given in more detail.
Sulfur Removal
In Eastern steam coals, sulfur occurs mostly in the forms of
iron pyrites and organic sulfur. In all three Cases included in this study
the coal sulfur undergoes the following reactions in the gasifier, operated
at partial-combustion conditions:
FeS2 = FeS + 1/2 S2
t- 02 = S02
S02 + 2 H2 = 2 H20 + 1/2 S2
FeS + H2 = Fe + H2S
[S] + H2 = H2S
u Jorg 2 z
1/2 S2 + H2 = H2S
All the above reactions are favorable, at the conditions used, with respect
to kinetics and equilibrium, with the result that substantially all the
coal sulfur is converted to H2S. Studies of the reaction [s]org + Ha <= H2S,
have been published by Consol.v6;7)
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23.
In addition to being a C02 acceptor, CaO also is an efficient
H2S acceptor through the reaction,
CaO + H2S = CaS + H20. (l)
Consol has patented processes (U.S. 2,824,047 and 3,117,918) for desulfuri-
zation of carbonaceous materials based on the above reaction. The acceptor
also removes COS by the reaction,
CaO + COS = CaS + C02. (2)
Unpublished data by Consol show that a high degree of COS removal occurs
by reaction (2).
For all three cases, equilibrium calculations show that about
96$ of the coal sulfur is converted to CaS in the gasifier. On each cycle
through the gasifier, about 2 mol % of the CaO in the recirculating acceptor
stream is converted to CaS in the two high pressure cases. In the low
pressure case, the acceptor circulation rate is much greater and conversion
to CaS is about 0.7$.
A compilation of thermodynamic data for pertinent sulfur
reactions and the procedure for equilibrium calculations are given in
Appendix C.
The first step in recovery of the coal sulfur is regeneration
of the CaS formed in the gasifier. Regeneration, and subsequent recovery
of the sulfur, are described individually below for each of the three cases.
Sulfur Recovery
Probably, the best long-term method for disposal of recovered
coal sulfur is conversion to elemental sulfur. The solid sulfur could be
sold or, in the case of poor market conditions, could be stockpiled cheaply
without causing pollution.
In general, technology of the processes to be described are
not fully developed, for the conditions used. Design basis and assumptions
are given for the individual cases.
Case I
Sulfur Rejection from CaS in the CaC03 Regenerator
In the regenerator the CaS content of the incoming acceptor is
oxidized to CaS04 by both air and C02,
CaS ..gg > CaS04
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24.
Then the following sequence takes place: The CaS04 decomposes by the
reaction, CaS04 = CaO + S03. The S03 is unstable and decomposes by the
reaction, S03 = S02 + 1/2 02. The 02 then reacts with unconverted CaS by
the reaction, 1/4 CaS + 1/2 02 = 1/4 CaS04. The sum of these reactions is
3/4 CaS04 + 1/4 CaS = CaO + S02 (3)
At Case I conditions, the mol ratio, incoming CaS/exit gas, is such that
the equilibrium partial pressure of S02 in reaction (3) would be reached
after only about one-third of the CaS had reacted. To allow completion of
the reaction, the regenerator is operated with a slight deficiency of com-
bustion air to provide CO in the upper part of the fluidized bed which then
reacts with S02 via
2 CO + S02 = 1/2 S2 + 2 C02, (4)
thereby decreasing the S02 partial pressure to about 70$ of the equilibrium
value in reaction (3). Details of the calculations which determine the
necessary CO partial pressure are given in Appendix C.
The technical basis for this method of sulfur rejection was
established during our studies for the OCR,(1) and during some preliminary
work in which the C02 Acceptor Gasification Process was operated as a "CO-
maker".(4) However, in those studies in which low-sulfur, low-rank Western
coals were used, the partial pressures of S02 and S2 in the regenerator were
considerably lower than those required for Case I. The possibility exists
that sulfur vapor will react with residual CaS04 in the acceptor or with the
regenerated CaO via,
S2 + CaS04 = CaS + 2 S02, (5)
3/4 S2 + CaO = CaS + 1/2 S02, (6)
with the result that the necessary rejection of sulfur from CaS will not
occur. The kinetics of reactions (5) and (6) must be slow relative to the
other sulfur reactions for successful use of the method. Study of the fore-
going reactions at Case 1 conditions will be a major part of Phase II-B of
the GAP contract.
Recovery of Elemental Sulfur
A flow diagram of the sulfur recovery section is shown in Figure 6.
Mass and heat balances for all three cases are given in Section III-B-3.
The regenerator exit gas, after removal of most of the coal ash and acceptor
fines by cyclones, passes through heat exchangers which reduce the temperature
to 472°F. Final clean up of particulates is by an electrostatic precipitator.
About 34% of the total gas, after heat exchange, is passed over an alumina
catalyst at 800°F in the reductor where the following reactions occur;
CO + H20 = C02 + H2
H2 + 1/2 S2 = H2S
CO + 1/2 S2 = COS
2 CO + S02 = 1/2 S2 + 2 C02
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GLAUS
TO
TO GASIFIES
STEAM SUPERHEATER
FIGURE 6
FLOW SHEET-SULFUR RECOVERY SECTION
CASE I
NOMINAL SYSTEM PRESSURE I95PSIA
Stream numbers refer to
Tables 7-9, pages 36-38.
LIFT GAS
2 CONDENSER
Q=6OMM BTU/HR
SULFUR
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26,
After heat exchange; the redactor gas is combined with the remainder of the
clean regenerator exit gas.
Discarded acceptor and overhead fines from the gasifier and re-
generator contain CaS. These materials cannot be dumped because the CaS
would hydrolyze on weathering, with release of H2S to the atmosphere. These
materials in all three cases, are slurried with water and contacted with a
C02~rich gas in a stripping column where the following reaction occurs:
CaS + C02 + H2°/1) - CaC03 + H2S.
The stripper gas then is combined with the regenerator gas.
The mol ratio, (H2S + COS)/S02 in the combined gas streams now
is 2.O, the optimum ratio for the Claus reactors which produce elemental
sulfur by the reactions,
2 H2S + S02 = 3 S + 2 H20
2 COS + S02 = 3 S + 2 C02
The tail gas from the No. 2 sulfur condenser is recompressed, reheated to
1220°F, and sent to the bottom of the gasifier.
Very high conversions to elemental sulfur are not of paramount
importance in this method for sulfur recovery. Residual S02, H2S, and COS
in the Claus tail gas would be converted to CaS in the gasifier with the
result that slightly more CO would be required in the regenerator to reject
the incremental sulfur in the acceptor. In the heat and material balance
calculations, the sulfur compounds in the Claus tail gas were ignored.
Unpublished work by Consol has shown that the reductor reactions
very nearly approach equilibrium over an active alumina catalyst at 8OO°F at
a space velocity of 4000 SCFH gas/CF catalyst bed. This work, done at
atmospheric pressure, also showed that reactions involving CO and H2 become
very slow at temperatures below about 600°F. For calculations at Claus
reactor temperatures, CO was assumed to behave as an inert gas. Claus
reactor technology, also at atmospheric pressure, is well established. Since
the mechanisms of the reductor and Claus reactions are not known, the effect
of operation at elevated pressures could not be determined. Therefore, the
space velocities used in Case I were those based on atmospheric pressure
data.
Details of the reductor and Claus reactor calculations are given
in Appendix C.
Case II
Sulfur Rejection from CaS in an External Reactor
Although Case I embodies the preferred method of sulfur rejection
from the standpoint of cost and simplicity, an alternate method was desired
in the event that forthcoming experimental studies show that the back reaction
of sulfur vapor is an insurmountable problem.
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27,
In Case II, a very slight excess of air is used in the regenerator
to prevent any substantial rejection of sulfur to the regenerator offgas.
Most of the acceptor CaS is oxidized to CaS04. The regenerator gas is
recompressed, without cooling, by a steam ejector and enters the bottom of
the gasifier as shown in Figure 4. The sulfur is rejected from the acceptor
by withdrawing one-third of the recirculating acceptor stream from the gasi-
fier into an external fluidized bed reactor where the following reaction
occurs:(18)
CaS + H20 + C02 = CaC03 + H2S.
This acceptor stream, now free of sulfur, is then sent to the regenerator.
Thus, the recirculating acceptor stream contains a sulfur "flywheel" in which
the acceptor leaving the regenerator contains about 4 mol $ of the total
calcium in the form of CaS04 and the acceptor leaving the gasifier contains
about 6 mol $ of the total calcium in the form of CaS. In the gasifier the
CaS04 in the recirculating acceptor is reduced to CaS.
A flow diagram for the Case II process is shown in Figure 7. In
the external reactor, the acceptor is fluidized by a recirculating stream
of steam and C02 in equimolar proportion at 1300°F. The ratio of inlet gas
to acceptor is controlled to give a 90$ approach to equilibrium in reaction
(?) in the exit gas. Nominal retention time is one hour in the fluidized bed.
One mol of C02 for each mol of CaS is supplied from a hot potassium carbonate
unit which processes about 6.5% of the regenerator offgas, as shown in
Figure 4.
The reactor offgas, after removal of most of the acceptor fines by
cyclones, passes through heat exchangers which reduce the temperature to
480°F. Final cleanup is by an electrostatic precipitator.
Kinetic and equilibrium restrictions in reaction (7) limit the H2S
content in the reactor offgas to about 3.7 mol $. Lower reactor temperatures
would give a higher concentration of H2S, but unpublished work by Consol has
shown that the kinetics of reaction (7) are poor at temperatures below about
12CO°F.
The low concentration of H2S in the reactor offgas, coupled with the
high concentration of steam, provides an undesirable feedstock for sulfur
recovery via the Claus process. For Case II another method of sulfur recovery
was chosen.
Recovery of Elemental Sulfur
The reactor offgas is passed concurrently with a stream of aqueous
H2S03 through a packed tower where the Wackenroder reactiom5) takes place;
2 H2S + H2S03 = 3 S/j\ + 3 H20. (8)
To avoid the very large heat duty involved in the condensation and re-evapora-
tion of steam in the recirculating steam-C02 stream, reaction (8) is carried
out at the saturation pressure. The reaction products pass to a separator
where liquid water and liquid sulfur are removed. The steam-C02 stream is
recompressed and sent back to the acceptor sulfur reactor. The water stream,
after cooling, is sent to a tower where aqueous H2S03 is formed by absorption
of SO2 produced by combustion of about one-third of the liquid sulfur produced
from the Wackenroder reactor.
-------
Stream numbers refer to
Tables 10-12, pages 39-41.
FIGURE 7
FLOW SHEET - SULFUR REMOVAL 8- RECOVERY SECTION
CASE U
NOMINAL SYSTEM PRESSURE 206 PS IA
480°F m _
90°F
ELECT.
PRECIP.
S02
ABSORBER
WACKENRCOER
REA
Q=I89-2MM
BTU/HR
BFW
7&7MM
BTU/HR
MME 1VHR
ACCEPTOR
SULFUR
REACTOR
SEPARATOR
0=77MM
BTU/MA
SULFUR
BURNER
PRODUCT
SULFUR
TO REGENERATOR
^^^f
ACCEPTOR
CO2
BFW
Q«79.8MM
BTU/HR.
-------
29.
The kinetics of reaction (8) are not well established at the
conditions used. However, as in Case I, very high conversion to liquid
sulfur is not of paramount importance. By operating with a deficiency of
H2S03, the unconverted sulfur in the form of H2S would be recirculated
with the steam-C02 stream to the acceptor sulfur reactor which then would
require a higher ratio of (H20 + C02)/CaS than is shown in the mass and
heat balances .
Case III (Low Pressure Case)
Sulfur Rejection from CaS
At the low system pressure in Case III, the recirculating acceptor
stream serves only as an H2S acceptor and heat carrier, i.e., no CaC03 is
formed. The regenerator is operated with stoichiometric air and sulfur is
rejected by the reaction sequence:
3/4 CaS + 3/2 02 = 3/4 CaS04
3/4 CaS04 + 1/4 CaS = CaO + SO2 (3)
CaS + 3/2 02 = CaO + S02
as in Case I. However, no equilibrium restriction on S02 partial pressure
exists because of the low system pressure.
Recovery of Elemental Sulfur
A flow diagram for Case III is shown in Figure 8. Regenerator
of f gas is cleaned and cooled as in Case I . The regenerator gas is mixed
with 3.35# of the total product fuel gas and is passed over an active
alumina catalyst at 80O°F in the reductor. This amount of fuel gas provides
sufficient CO and H2 to give a mol ratio of (H2S + COS)/S02 = 2.0 in the
reductor product gas .
Elemental sulfur is formed in a single-stage Claus reactor. Calcu-
lations showed that a second Claus stage would give only a small increment
in overall sulfur recovery .
The Claus tail gas must be sent to the boiler for incineration of
the residual sulfur compounds. Because the Claus tail gas is not recycled
to the process as in Case I, the overall sulfur clean-up level drops from
to
3. Material and Heat Balances ~ Gasification Section
B asic Pressure Case I
Complete mass, elemental, and heat balances have been
computed for all the major processing steps incorporated into the basis
pressure case flow sheet as described previously. The overall mass and heat
balance is summarized in Table III. The cold gas efficiency is:
10.286.6 MM Btu/hr /,_«i
13,812.4 MM Btu/hr (l00) =
-------
FIGURE 8
FLOW DIAGRAM - SULFUR RECOVERY SECTION
CASE HI
NOMINAL SYSTEM PRESSURE 20 PSIA
FROM
SLURRY
STRIPPING
FROM
REGENERATOR 695T
AFTER RAISING (ie)
PLANT STEAM ^
^^^
V
ELECTRO
PRECIP.
<
S>
iao°F
< \
PRODUCT GAS
FROM GASIFIER
2)
I2O°F
I
DUST
BFW
CLAUS
REACTOR
REDUCTOR
800° F
0=265 MAflBTU/HR
•402eF
410° F
Stream numbers refer to
Tables 17-18, pages 46-47.
TAIL GAS TO BOjLER
FOR INCINERATION
TO SLURRY
28OeF
STRIPPING
310° F
CONDENSER
CW
26O°F
Q-9OMMBTU/HR
280°F
PRODUCT SULFUR
u
o
-------
Basis: 1 Hour Datum: 60°F H2O,jx
TABLE 3
Overall Mass and Heat Balance
Gasification Section-Case I
Input
Feed coal (6% moisture)
Air
Boiler feed water
Make-up Acceptor
Totals
Output
Product Gas
CO
C02
NH3
H2S
H20
(v)
Sub- total
Unburned Char and Ash
Product Sulfur/1A \.
Spent Acceptor
Net Heat of Reaction and
Sensible Heat Losses
Totals
Pounds
1,157,000
3,949,2OO
731,3OO
77,3OO
5,914,8OO
5,914,2OO
Mols
137,095
Mol #
—
—
—
Temp.
°F
60
398
60
6O
AH
MM Btu
X
335.8
X
X
335.8
AHc
HHV
MM Btu
13,812.4
X
X
X
13,812.4
AH + AHc
MM Btu
13,812.4
335.8
X
X
14,148.2
48,4OO
71,8OO
1,069,500
932,700
3,026,900
8,700
1,8OO
510,900
5,67O,7OO
152,800
44,800
45.9OO
3,017
35,624
38,182
21,194
108,041
5O9
54
28,359
234,980
—
—
__
1.3
15.2
16.2
9.0
46.0
0.2
X
12.1
1OO. O>
—
—
— —
13OO
13OO
1300
13OO
1300
1300
1300
1300
12O
28O
120
48.1
310.7
348.6
296.7
976.5
6.9
O.6
849.7
2,837.8
1.1
2_8
0.1
1,154.8
4,388.0
4,646.8
X
X
83.9
13.1
X
10,286.6
271.7
178.9
X
1,202.9
4,698.7
4,995.4
296.7
976.5
90.8
13.7
849.7
13,124.4
272.8
181.7
O.I
2,841.8
10,737.2
569.2
14,148.2
CO
-------
32.
A more realistic evaluation of the thermal efficiency of the process
is obtained by relating the total heat content of the product gas and of the
product sulfur to the total heat input:
13,306.1 MM Btu/hr
14,148.2 MM Btu/hr
Detailed mass, elemental, and heat balances for the following opera-
tions are shown as follows:
Preoxidation - Table 4
Gasification - Table 5
Regeneration - Table 6
Reductor - Table 7
First-Stage Claus - Table 8
Second-Stage Claus - Table 9
Alternate Pressure Case II
As mentioned previously, the main difference between Cases I
and II is the method by which the sulfur is released from the acceptor. Since
the heat and material balance calculations showed that the energy contained by
the product gas in Case II was within 1% of that in Case I, only the sulfur
removal and recovery details are presented here. Detailed mass and heat
balances for the novel acceptor processing cycle were then computed, and are
shown as follows:
Acceptor Sulfur Reactor - Table 1O
"Wackenroder" Reactor - Table 11
S02 Absorption Tower - Table 12
Atmospheric Pressure Case III
Again, complete mass, elemental, and heat balances have been
computed for all the major processing steps incorporated into the basic flow
sheet as described previously. The overall mass and heat balance is summarized
in Table 13.
For this case, the cold gas efficiency is;
10.078.5 MM Btu/hr /,^^x „„ ^y
13;812.4 MM Btu/hr (10°) = 73'°*
The thermal efficiency is expressed by the ratio of the total heat content of
the product gas and of the product sulfur to the total heat input is:
11,627.4 MM Btu/hr
14,082.6 MM Btu/hr
Detailed mass, elemental, and heat balances for the following
operations are shown as follows:
Preoxidation - Table 14
Gasification - Table 15
Regeneration - Table 16
Reductor - Table 17
Claus Stage - Table 18
-------
TABU 4
Mass fc Heat Balance
Preoxidizer
Case I
Basis: 1 Hour Datum: 6O°F. Hg>f ^
Pounds Mo Is
Input
ID Feed Coal
MF Coal
Wa • e r/ 1)
Sub-Total
© Alr
02
S2
H20 (v)
Sub-Total
1,087,600
69,400
1, 157,000
119,800
394,700
1,400
515,900
-
-
-
3,744
14,087
77
17,908
Mol *
-
-
-
20.9
78.7
0.4
100.0
- See Figure 3
Elemental Balance, Pounds
H C N
52,200 759,200 13,000
7 . 700 x x
59,900 759,200 13,000
XXX
x x 394,700
200 x x
200 x 394,700
0 S
82,600 46,800
61,700 x
144,300 46,800
119,800 X
x x
1.200 x
121,000 X
Ash
133,800
x
133,800
x
x
X
X
Heat Release by Oxidation
Totals
Output
fsl Preoxldlzed Ceal
(3} Preoxtdizer Vapor
Tar » Light Oil
CH4
H2
CO
CO2
°2
it's
H20 (v)
Sub Total
Heal Loss
1,672,900
1 , 009 , 000
(v) 28,300
1,600
800
900
10,200
90.000
• 26.400
395.000
1.300
109,400
663 , 900
(2921 mo Is
-
-
189
98
26
461
364
2,046
824
14.099
39
6,068
24,21-4
02 x 185
-
-
0.8
0.4
0.1
1.9
1.5
8.4
3.4
58.2
0.2
25. 1
100.0
,OOO Btu/mol)
60, 1OO 759,200 407,700
44.100 705,200 12,400
2,100 23,300 300
400 1 , 200 x
200 600 X
900 x X
x 4,400 X
X 24,500 x
X XX
x x 395,000
100 x x
12,300 x x
1 6 , OOO 54,000 395, 300
265,300 46,800
68,400 45,100
2 , 100 500
X X
X X
X X
5,800 X
65,500 x
26,400 x
X X
x 1 , 200
97,100 X
196,900 1,700
133,800
133,800
x
X
X
X
X
X
X
X
X
X
X
Temp.
Of
60
398
800
800
AH or Cp
2365
Btu/mol
1210.8 Btu/lb 1.7
43.9
540.4
584.3
266 Btu/lb
268.4
Totals
1.672,900
60,100 759,200 407,700 265,300 46,800 133,800
268.7
47.2
584.3
Heat of
Combustion
HHV
MM Btu
13.812.4
13,812.4
12,653.9
533
8,192
13,569
5,168
5,288
7,784
5,494
5.251
6,525
1402.
Btu/lb
Btu/mol
"
"
"
"
"
"
"
5 Btu/lb
15
0
0
2
1
15
4
74
0
153
.1
.8
.4
.4
.9
.9
.5
.0
.3
.4
452.4
37. S
17.6
56.8
44.3
x
X
X
9.S
X
618.1
CO
00
-------
TABLE 5
Has a b Heat Balance
GaslMer
Case I
~ See Figure 3
Basis: 1 Hour Datua: 60°* H20(i) a * •
Pounds Mols Hoi < Elemental Balance, Pounds
Input
(S) Preoxldlitd Coal
1,009,000
Q Preoxldlter Vapor 663,900 24,214
(fi) Air
02
N2
H20 (V)
Sub- Tot. 1
0 Steam
(7) Claus Tall Gas
CO
C02
H2
>*2
SOj
H2S
H20 (V)
Sub-Total
(5) Acceptor
\\SO- caO
Inert
Sub- Total
Heat of Reaction
402.600 12,581 20.9
1,326.600 47.351 78.7
4.700 259 0.4
1,733.000 60,191 1OO.O
731,300 40,595
32. BOO 1.171 1.7
931.600 21.167 30.4
100 24 X
1.300,900 46.433 66.8
300 5 x
300 10 x
13.900 771 1.1
2.279.900 69.5H1 100.0
7,399,800 76,761
1,305.800
8.705.600
UgO-CaO to MfO-CaCO3 (9,699 mols
MgO-CoO to JlKO'CnS (1,416 mols
B C N
44,100 70S, 200 12,400
16, OOO 54,000 395,300
XXX
x X 1,326,600
500 X x
500 x 1,326,600
81,800 x x
x 14 , 100 x
x 254,200 x
100 x x
X X 1,300,900
XXX
XXX
1,500 x x
1.600 268,300 1,300,900
XXX
XXX
XXX
x 76,2OO Btu/mol)
x 45, SCO Dtu/mol 1
O S
68,400 45,100
196,900 1,700
402,600 X
X X
4.200 x
406,800 x
649,500 x
18 , 700 x
677,400 x
X X
X X
100 200
x 300
12,400 x
708 , 600 500
1,228,200 x
x x
1,228,200 x
Ash MgO-Ca
133,800 x
x x
X X
X X
X X
X X
X X
X X
X X
X X
X X
X X
X X
X X
X X
X 6,171,600
1,305, 80O x
1,305,800 6,171,600
Heat of Combustion (Reactants - Products ex Acceptor)
Totals
Output
(7) Char
(*) Char Loss
M-M Acceptor
^^ MgO* CaCO3
ItgO- CaS
MgO-CaO
Inert
Sub- Total
(9) Gas
CH4
H2
CO
C02
N2
H2S
HjO(v)
Sub-Total
15.123.600
293 .700
3 , 600
1,361.800 9.699
159. 20O 1.416
6.328.300 65.646
1,305.800
9,155,100
48,400 3.O17 1.3
71.800 35.624 IS. 2
1.0U9.500 38.182 16.2
932.700 21.194 9.O
3,026.900 108,041 46.0
8.700 509 0.2
1.800 54 x
SlO.iMlO 28.359 12. 1
5.670.700 234.980 1OO.O
144,000 1,027,500 3,035,200
1,200 159,200 x
100 2,500 x
x 116,500 x
XXX
XXX
XXX
x 116,500 x
12,200 36,200 x
71,800 x x
x 458.600 x
x 254.500 X
x x 3,026. 9OO
1.500 x 7.2OO
100 x x
57,200 x x
142,800 749,300 3,034,100
3,258,400 47,300
X X
300 200
465,500 x
x 45.4OO
1,050,300 x
X X
1,515.800 45,400
X X
X X
610.900 x
678 , 200 x
X X
X X
X 1,700
453,700 x
1,742,800 1,700
1,439,600 6,171,600
133,300 x
500 x
x 779,800
X 113,800
x 5,278,000
1 , 305 , 8OO x
1.305.8OO 6,171,600
\ x
X X
X X
X X
. X X
X X
X X
X X
X X
Heat of Reaction
Heat Loss
Heat of comuuallon (1. 116 uol»
1I2S tu CnS)
Te»p.
Of
800
800
398
120O
1883
1700
1700
1700
1700
AH or Cp
266 Btu/lb
(Table 4)
2365 Btu^"°l
1210.8 Btu/lb
8,344 Btu/mol
12,715
8,004
8.263
13.107
10,574
16O7.8 Btu/lb
23.2 Btu/nol-»K
0.255 Btu/lb-of
545 Btu/lb
545 Btu/lb
268.4
268.7
141.7
5.7
IT7TT
1607.8 Btu/lb 1,175.9
Beat of
CoBbustlon
RHV
MM Btu
12,653.9
618.1
142.5
x
3.0
x
X
2.4
x
147.9
9.8
269.1
0.2
383.7
0.1
0.1
22.3
685.3
3,246.5
607.0
3.853.5
739. 1
64.4
738.3
7,941.0
16O.1 2,349.9
1.9 4S.1
38.7 Btu/mol-cr 615.6
23.0 " S3.4
23.0 " 2,476.1
0.251 Btu/lt>-°r 537.5
3,682.6
23,134 Btu/mol
11,641
12,347 "
19.29O
12.213
19.203
18.O94 "
1884.9 Btu/lb
69.8
414.7
471.4
408.8
.319.5
9.8
0.9
963.0
1,154.8
4,388.0
4.646.8
83.9
13. I
X
Totala
15. 123. 100
111.100 1,027.500 3.03-1.100 3,iC,B.UOO 47.3UO 1.439.6OO 6.171,600
3,657.9 10.286.6
342.6
95.9
7,941.0
CO
-------
TABLE 6
Mass L Heat Balance
Regenerator
Case I
Baals: 1 Hour Datu. : 60°F. H2O/i)
Pounds Mola Hoi
Input
© Char
lo) Lift Caa
CO
C02
H2
N-
H2O (?)
Sub-Total
i»7) Alr
^ 02
H20 (»)
Sub-Total
Tw Acceptor
MgO- CaCO3
MgO- CaS
MfiO- CaO **
Inert
Sub- Total
13) Make-up Acceptor
^"^ Mtf*CO3' CaCO3
Inert
Sub-Total
Heat of Reaction
293. TOO
2, 100
60.700
x
84,700
900
148. 100
394 , 700
1,300. 10O
4,600
1.699,400
1,354,900
158,600
6.296.600
1. 299.300
9, 109.400
70.800
6,500
77.3OO
UgO- CaS to MeO-
-
76 1.7
1,378 30.4
1 X
3.022 66.8
51 1. 1
4,528 100.0
12,336 20.9
46. -IOC 78.7
254 0.4
58,996 100.0
9.650
1,410
65,317
-
.
384
-
-
CaO (141O mols x
Heat of Combustion (Reactants -
Totals
Output
5*) ch'r
ffi\ Acceptor
UcO-CaO
Inert
Sub-Total
(Tfi) Gas
^"^^ CO
co2
H2
X2
SO-
s2
HjS
COS
H-o <»)
Sub-Tots 1
11,328.200
149.200
7.399.800
1 . 305 . 800
8 . 70S . £00
52,300
964 . 900
100
1.385.500
20.800
34.400
200
300
1 1 . TOO
2,473,200
-
76,761
-
-
1,867 2.5
21,925 29.2
38 O. 1
49.455 66.0
325 0.4
536 0.7
5 X
6 x
HI 1 1.1
74.971 1OO.O
- See Figure 3
£ Elemental Balance. Pounds
H C N
1,200 159,200 x
x 900 x
X 16,600 x
XXX
x x 84.700
100 x x
100 17,500 84.700
XXX
X X 1,300,100
500 x x
500 x 1,300,100
X 115,900 x
XXX
XXX
XXX
x 1 15 , 900 x
x 9 , 200 x
XXX
x 9,200 x
196, GOO Dtu/mol)
Products ex Acceptor)
1,800 301,800 1,384,800
X 15,900 x
X XX
X XX
X XX
x 22,400 x
x 263,300 x
1OO x x
x x 1.38S.SOO
X XX
X XX
X XX
x 100 x
1 , GOO x x
1,700 285,800 1.385,500
O
x
1.2OO
44, 100
X
X
8OO
46, 100
394,700
x
4, 1OO
398,800
463,200
x
1,045, 1OO
x
1,508,300
30.7OO
x
30,70O
1,983,900
x
1,228,2OO
x
1.228,200
29,900
7O1.6OO
x
X
10,400
X
X
100
13.100
755 , 100
S Ash
x 133,300
x x
X X
X X
X x
x x
x x
x x
x x
X X
X X
X X
45,200 x
X X
x 1,299,300
45.200 1,299,300
x x
x 6,500
X 6.5OO
45,200 1,439,100
x 133.30O
x x
x 1,305,800
x 1,305,800
x
x
X
X
10,400
34,400
200
100
x
45,100
MgO-Cs.
x
x
X
X
X
X
X
X
X
X
X
775,800
113,400
5,251.500
x
6,140,700
30,900
x
30.900
6,171,600
X
6,171,600
X
6, 171, 6OO
Heat of Reaction
UsO
UgO
Heat Loss
Totala
CaCOj to McO-CnO (0. liio ruolu »
J3'CaCO3 to llgO
11.328.0UO
CaO (384 mols x
-
7C,2OO IHu/mol)
127, 4OO Btu/raoll
1.700 301,700 1,385,500
1,983,300
45. 1OO 1,439,100:-
6.171.60O
TCnp.
°F
17OO
280
"
**
"
*•
398
"
11
1700
'•
'*
"
60
"
1883
1883
"
1883
"
"
„
"
"
M
"
£H or Cp
945 Btu/lb
1537 Btu/mol
2072
1528
1534
1158.4 Btu/lb
< 2365
<
1210.8 Btu/lb
AH
mi Btu
ISO. 1
0.1
2.8
x
4.6
l.O
8.5
| 138.9
(
5.6
144.5
Beat of
Combustion
HHV
KM Btu
2.349.9
9.2
X
0.1
X
X
9.3
x
X
X
X
38.7 Btu/mol-°F 612.5 z
. 23.0
23.0 "
O.251 Btu/lb-*F
—
490 Btu/lb
23.2 Btu/mol-»F
O.255 Btu/lb-°F
13,852 Btu/nol
21,780
13.0O7
13.7OO
22,021
15,684
18,215
23.051 "
1987.4 Btu/lb
53.2
2,463.7
534.8
3,664.2
x
x
X
277.2
1,731.4
5,985.9
73.1
3,246.5
607.0
3.8S3.5
25.9
477.5
O.5
677. S
7.1
8.4
O. 1
0. 1
29. 1
1,226.2
735.3
48.9
48.9
5.985.9
X
X
X
X
x •
X
X '
226.6
X
X
X
227.2
x
4.7
x
X
166.7
1.2
1.4
x
401.2
u
01
-------
36,
TABLE 7
Mass it Heat Balance
Reductor
Case I - See Figures 3 and 6
Basis: 1 hour Datum: 60°F H20'^\
Input
fl^\ Regenerator Gas
V^ CO
C02
H2
N2
so2
S2
H2S
COS
H20 (v)
Sub-total
Heat of Reaction:
Totals
Output
f^\ Reductor Gas
*J CO
C02
H2
N2
S02
*S2
H2S
COS
H20 (v)
Sub-total
Heat Loss
Pounds
17,530
323,400
25
464,500
6,976
11,370
61
132
4,916
828,910
Mols
625 . 8
7349.3
12.7
16,577
108.9
177.3
1.8
2.2
272.8
25, 128
Heat of Combustion (Reactants
828,910
148
343,100
1
464,500
2
1,551
8,288
10,540
780
828,910
25,128
5.3
7,795.9
0.3
16,577
0.03
24.19
243.5
175.7
43.6
24,866
Temp AH
op
678 4385
6356
4320
4366
6668
5113
5415
6996
1344.3
- Products)
800 5288
7784
5168
5251
8123
6144
6525
8510
1404.
or Cp
Btu/mol
11
"
M
"
if
it
it
Btu/lb
Btu/mol
ff
it
M
(I
It
tf
ft
7 Btu/i;
Totals
828,910
24,866
AH
MM Btu
2.74
46.71
.05
72.38
.73
.91
.01
.02
6.61
130.16
24.97
155.13
.03
60.68
x
87.04
x
0. 15
1.59
1.50
1.10
152.09
3.04
155.13
Shown as S2. Actual composition is 3.875 atoms S/mol
-------
Basis: 1 hour Datum:
Input
Feed Gas
CO
C02
H2
N2
S02
* S8
H2S
COS
H20 (v)
Sub-total
20Gas from Slurry
Stripping
CO
C0
H2S
Mass & Heat Balance
First-Stage Claus Reactor
Case I - See Figures 3 and 6
37,
H20
(v)
Sub-total
60°F ^0/jv
Pounds
34,910
984,600
52
1,385,700
13,840
24 , 100
8,420
10,810
10,540
2,472,970
110
2600
4000
200
40
6950
Mols
1,246.5
22,372
25.6
49,455
216.1
93.97
247.1
180.0
584.8
74,421
4
59
143
6
2
214
Temp
op
535
120
AH or Cp
3348 Btu/mol
4753
3309
3334 "
4996
18,887 "
4051 "
5238
1276.8 Btu/lb.
AH
MM Btu
4.17
106.33
.08
164.88
1.08
1.77
1.00
0.94
13.45
293.70
X
542 Btu/mol
417
X
X
X
.03
.06
X
X
709
Heat of Reaction
S02 + 2 COS = 3/2 S2 + 2 C02
(89.40 mols S02 x 8510 Btu/mol)
.76
S02 + 2 H2S = 3/2 S2 + 2
(98.65 mols SO2 x 17,780 Btu/mol)
4 S2 = Sg
(71.59 mols S8 x 178,000 Btu/mol)
Totals
2,479,920
74,635
1.75
12.78
309.08
Output
(21) Product Gas
^^ CO
co2
H2
S02
«* Sg
H2S
COS
H20 (v)
Sub-total
Heat Loss
Totals
1
2
35,
995,
,389,
1,
42,
1,
14,
,479,
020
100
52
700
634
460
690
72
190
920
1,
22,
49,
74,
250
610
25
598
25
165
49
1
788
514
.5 535
.6
.51
.56
.80
.20
.3
3348
4753
3309
3334
4996
18,887
4051
5138
1276.
Btu/mol
it
ti
ii
it
ii
it
it
8 Btu/lb
4.
107.
,
165.
.
3.
.
.
18.
298.
10.
19
48
08
36
13
13
20
01
14
72
36
2,479,920
74,514
Shown as Sg. Actual composition is approx. 7 atoms S/mol
Shown as Sg. Actual composition is 7.294 atoms S/mol
309.08
-------
Mass & Heat Balance
Second-Stage Claus Reactor
Case I - See Figures 3 and 6
38.
Basis: 1 hour
Datum:
60°F H20(i)
Input
12J Feed Gas
CO
C02
H2
N2
so2
H2S
COS
H20 (v)
Pounds
35 , 020
995,100
52
1,385,700
1,634
1,690
72
14,190
Mols Temp.
OF
1,250.5 390
22,610
25.6
49,598
25.51
49.80
1.20
788.3
AH or Cp
2312
3193
2295
2306
3372
2765
x
1208.
Btu/mol
ii
it
ii
ii
ii
8 Btu/lb.
AH
MM Btu
2.89
72.21
.06
114.37
.09
.14
x
17.17
Sub-total
2,433,460
74,349
Heat of Reaction
S02 + 2 COS = 3/2 S2 + 2 C02
(0.57 mols SO2 x 8510 Btu/mol)
S02 + 2 H2S = 3/2 S2 = 2 H2OM)
(19.97 mol S02 x 17,780 Btu/mol)
4 S2 =
I
Totals
Output
--v
(23) Product Gas
CO
C02
N2
so2
* S8
H2S
COS
H20 (v)
Sub-total
Heat Loss
(7.69 mols Sg x 178,500 Btu/mol)
35,020
995,100
52
1,385,700
318
1,973
336
4
14,930
2,433,430
1,250.5
22,611
25.6
49,598
4.96
7.69
9.86
.06
828.4
74,336
380
2242 Btu/mol
3090
2225
2236
3306 "
12,508
2684
x
1204.2 Btu/lb
Totals 2,433,430 74,336
Shown as Sg. Actual composition is 7.518 atoms S/mol
206.93
.01
.36
1.37
208.67
2.80
69.88
.06
HO. 90
.02
.10
.03
x
17.98
201.77
6.90
208.67
-------
39.
TABLE 10
Mass & Heat Balance
Acceptor Sulfur Reactor
Case II
Basis: 1 hour Datum
Input
M/) Acceptor
^-^ MgO- CaC03
MgO. CaS
MgO'CaO
Inert
Sub-total
(2) Gas
C02
H20 (v)
Sub-total
Heat of Reaction
MgO'CaS + C02
Totals
Output
fsj Acceptor
MgO- CaCX>3
MgO- CaO
Inert
Sub-total
(T) Gas
W co2
H2S
H20 (v)
Sub-total
Heat Loss
-
: 60°F H2°(l
Pounds
370,550
154,700
1,627,400
355,500
2,508,150
852,400
348,950
1,201,350
+ H20(l) =
(1375.5 mols
3,709,500
563,700
1,627,400
355,500
2,546,600
791,850
46,900
324, 150
1,162,900
See Figures
)
Mols
2,639.
1,375.5
16,882.
-
—
19,368
19,368
38,736
MgO-CaC03 +
4 & 7
Temp
op
1700
If
M
ft
600
It
H2S
x 30,620 Btu/mol)
-
4,014.5
16,882.
-
-
17,992.5
1,375.5
17,992.5
37,360.5
1300
II
If
1300
If
M
AH or Cp
38.7 Btu/mol-°F
23.0
23.0 "
0.251 Btu/lb -°F
5473 Btu/mol
1299.8 Btu/lb.
36.9 Btu/mol-°F
22.6
0.242 Btu/lb-°F
13,999 Btu/mol
11,644
1661.8 Btu/lb
AH
MM Btu
167.4
51.9
636.8
146.4
1,002.5
42.1
1,604.2
183.9
472.4
106.9
763.2
Totals
3,709,500
806
34.4
1,604.2
-------
40.
TABLE 11
Mass & Heat Balance
'Wackenroder" Reactoi
Case II
- See Figures 4 & 7
Basis: 1 hour Datum: 60°F H20(i)
Input
®Gas
C02
H2S
H20 (v)
Sub-total
(jp\ Liquid
H2SOg
H20(l)
Sub-total
Heat of Reaction
2 H2S + H2S03 =
(1375.5
Totals
Output
QT) Gas
C02
H20 (v)
Sub-total
Liquid
ff\ Sulfur
/§N H20/j\
Sub-total
Heat Loss
Totals
Pounds
791,850
46,900
324, 150
1,162,900
56,450
726,450
782,900
3 S(l) -f
mols H2S x
1,945,800
791,850
348,800
1,140,650
66,150
739,000
805,150
1,945,800
Mols Temp. AH or Cp
op
17,992.5 480 4152 Btu/mol
1,375.5 " 3557
17,992.5 " 1240.9 Btu/lb
37,360.5
fiSS 99Q 1 ^
I 10 Btu/lb- °F i
40,324 ) S
41,012 I
3 H20(i)
43,900 Btu/mol)
-
17,992.5 328 2556 Btu/mol
19,361 " 1159.1 Btu/lb
37,353.5
2,063 328 0.2 Btu/lb-°F
41,019 " 270.5 Btu/lb
43,082
-
AH
MM Btu
74.7
4.9
402.2
481.8
132.3
132.3
60.4
674.5
46.0
404.3
450.3
3.5
199.0
202.5
21.7
674.5
-------
41
TABLE 12
Basis: 1 hour Datum:
60°F H.
t
Mass & Heat Balance
S02 Absorption Tower
Case II
- See Figures 4 & 7
vQf i N
6 \ i. /
Pounds Mols Temp. AH or Cp
Input
fgj Gas
^ so2
N2
H2O (v)
Sub-total
g) Water(i)
Heat of Reaction
S02 + H20(i)
(688
Totals
Output
!ll) Gas
S02
N2
H20 (v)
Sub-total
O Liquid
^ H2S03
Sub-total
44,500
73,250
500
118,250
738,500
= H2S03
mols x
856,750
450
73,250
150
73,850
56,450
726,450
782,900
op
695 400 3480 Btu/mol
2,614.5 " 2376
29 " 1213.5 Btu/lb
3,338.5
90 29.93 Btu/lb
(aq)
13,400 Btu/mol)
7 90 284 Btu/mol
2,614.5 " 208
8.5 " 1072.8 Btu/lb
2,630.
688 10,1 ? 1.0 Btu/lb-°F
40,324 " )
41,012
Heat Rejected to Cooling Water
Totals
856,750
-
AH
MM Btu
2.4
6.2
0.6
9.2
22.1
9.2
40.5
X
0.55
0.15
0.7
^ 32.1
32.1
7.7
40.5
-------
Overall Mass and Heat Balance
Gasification Section-Case III
Basis; 1 hour Datum; 6O°F H2O/j\
Input
Feed coal (6% moisture)
Air
Boiler feed water
Make-up Acceptor
Process water
Totals
Output
Product Gas
CH^
H2
CO
CO 2
N2
NH3
H2S
H20 (v)
Sub- total
In-Process Fuel Gas
Unburned Char-Ash
Product Sulfur/
Glaus Plant Tail Gas
CH4
CO2
H2S
S02
S8
H20
Sub- total
Spent Acceptor
Net Heat of Reaction and
Sensible Heat Losses
Totals
}
Pounds
1,157,OOO
4,O54,8OO
614,700
84,800
200
5,911,500
3,7OO
88,500
1,O34,3OO
4O3,5OO
1,469,3OO
8,1OO
1,1OO
281,700
3,290,200
35,200
164,500
38,3OO
100
654,200
1,614,1OO
3,3OO
3,100
1,7OO
57,8OO
- See Figure 5
Mols Mol %
140,996
— —
—
—
14Oj996
232 O. 2
43,916 27.6
36,928 23.3
9,167 5.8
52,445 33.0
479 0. 3
30
15^636 9.8
158,833 100.0
1,7OO
—
"~— ™ —
7
14,865
57,615
97
49
6
3,213
Temp.
°F
60
3OO-315
60
60
60
980
980
980
980
980
980
980
98O
12O
280
280
280
280
280
280
280
2,334,300
47,9OO
5,910,400
75,852
Temp.
°F
AH
MM Btu
AHc
HHV
MM Btu
AH + AHc
MM Btu
120
270.2
27O. 2
2.5
282.8
245.4
91.3
345.6
4.6
0.2
421.2
1,393.6
14.9
1.2
2.4
13,812 4
13,812.4
88.6
5,409.5
4,494.3
78.9
7.2
1O,O78.5
1O8.O
440.6
152.9
2.7
1,599.6
10,812.7
13,812.4
270.2
14,082.6
91.1
5,692.3
4,739.7
91.3
345.6
83.5
7.4
421.2
11,472.1
122.9
441.8
155.3
2.7
30.8
88.4
23.6
0.1
6.7
67.0
219.3
0.1
1,671.1
14,082.6
u-
to
-------
TABLE 14
Haas b Heat Balance
Preoxidlzer
Case III
Bail*: 1 hour Da tun; 60°
Input
T) Feed Coal
«F Coal 1
WaterM)
Sub- total 1
5) *"•
02
X2
H20 (v)
Sub-total
Heat Release by Oxidation
Totala 1
Output
@ PreoxldUed Coal 1
(3) Preoxldlzer Vapor
Tor t Light Oil (v)
C!I4
C2"6
"'2
CO
C02
°2
II2S
H20 (v)
Sub-total
Heat Loss
F H2O/.%
Pounds
,087,600
69, 40O
, 157,000
122,500.
403,600
2,900
529,000
(2987 mola
.686.000
,007,800
28,300
1,600
800
900
10,200
92,500
26,900
404,000
1,400
111, 500
678,100
Ho la
-
-
-
3,829
14 , 40S
161
18,395
O2 x 185
-
189
98
26
463
364
2, 102
842
14,418
41
6, 188
24,731
M01.4
-
-
-
20.8
78.3
0.9
100.0
- See Figure 5
Elemental Balance, Pounds
H C
52,200 759,200
7,700 x
59,900 759,200
X X
X X
300 x
300 X
N 0
13,000 82,600
x 61 , 700
13,000 144,300
x 122,500
403 , 600 X
x 2,600
403,600 125,100
S
46,800
x
46,800
x
X
X
X
Ash
133,800
x
133,800
x '
X
X
X
,000 Btu/mol)
-
0.8
0.4
0.1
1.9
1.5
8.5
3.4
58.3
0.1
25.0
100.0
60,200 759,200
44 , 100 704 , 300
2, 1OO 23, 300
400 1 , 200
200 600
900 x
X 4,400
X 25 , 200
X X
X X
100 x
12 , 50O x
16,200 54,700
416,600 269,400
12,300 68,300
300 2,100
x x
X X
X X
x 5 , 80O
X 67,300
x 26,900
404,000 x
X X
x 99,000
404,300 201,100
46,800
45,000
500
x
X
X
X
' X -
X
X
1,300
X
1,800
133,800
133,800
x
X
X
X
X
X
X
X
X
X
X
Temp.
op
60
315
800
800
AH or Cp
1816 Btu/nol
1779
1174.S Btu/lb
266 Btu/lb.
AH
MM Btu
7.0
25.6
3.4
36.0
5S2.6
588.6
268.1
533 Btu/lb.
8,192 Btu/nol
13,569 "
5,168 "
5,288 "
7,784 "
5,494 "
5,251 "
6,525 "
1404.7 Btu/lb 156.6
274.2
46.3
Heat of
Combustion
HHV
UM Btu
13,812.4
x
13,812.4
12,641.1
15.1
0.8
O.4
2.4
1.9
16.4
4.6
75.7
0.3
156.6
452. 4
37.5
17.6
57.0
44.3
x
X
X
9.9
X
618.7
Total*
1, 685, 900
60,300 759,000 416,600 269,400 46,800 133,800
588.6
•fe.
CO
-------
TABLE IS
Baals: 1 hour Da tun;
Input
*s\ Prvoxldlced Coal L
T) Air
^ 0,
•V2 l
H20 (v)
Sub-total 1
^ Steam
fis} Acceptor
V4.tr CaO 2O
Inert 3
Sub-total 23
Heat of Reaction
MgO- CaO
Heat of
Totala 27
Output
Q Char
(Q Char laia
[M Acceptor
-S v.ffl. CaS
5*0' CaO 20
Inert 3
Sub-total 23
ff) Xet Ca>
^^ Cllj
H2
CO 1
C02
X, 1
NH3
H S
H20 (v)
Sub-total 3
Heat of Reaction
Heat of
60° F HzOf
rounds
.007.800
678 100
342,000
,126,500
8.1OO
,476,600
614.700
, 2G8.-IOO
,576,800
.8-15.200
to MgO-Cas
Com bus t Ion
,622,400
298,800
16,300
157,900
.133,100
.570.800
.867,800
3.9OO
92 , 5OO
.081.200
421,800
.535,900
8,500
1, 100
294. 5OO
,439,400
Combust Ion
Recycle Caa Cooling (I5UO mole
Heat Lota
Totala 27
.622.300
n
Wo Is Uol }
a c
44,100 704,300
24 731 - 16 20O 54 7OO
10,689 20.8 x x
40.21O 78.3 X x
448 O.9 9OO x
51,347 100. 0 9OO x
34,119 - 68.80O x
210,253 - x x
- - x x
• - x x
(1,404 mole x 45,5OO Btu/moll
(React ants - Products ex Acceptor)
130,000 759,000
1.2OO 166,000
700 11,400
1,404 X x
208,849 x x
x x
X X
242 0.2 l.OOO 2,900
45,907 27.6 92,500 x
38,602 23.3 X 463,600
9,583 5.8 x 115, 1OO
54.822 33. 0 x x
501 0.3 1.500 x
31 x 100 x
16,345 9.8 33,000 x
166.033 1OO.O 128, 10O 581, 6OO
(1,404 mo Is H2S to CaS)
1700«F to 1600°F)
130. OOO 759. OOO
Maaa l> Heat Balance
Oaalfler
Case HI
- See Figure 5
Elemental Balance, Pounds Temp.
N OS Aah UgO-c* °F
12,300 68,300 45, OOO 133, 80O x 800
4O4 3OO 201 100 1 BOO x x 800
x 342, OOO xx x 50O
1,126,500 xx x x "
x 7 , 200 x x x "
1,126,500 349,200 XX x
x 545.9OO xx x 1200
x 3, 364, OOO x x 16,904,400 19OO
x xx 3.576.8OO x
x 3, 364, OOO x 3,576,800 16,904,400
1,543,100 4,528,500 46.8OO 3,710,600 16,904,400
x x X 131,600 x 1700
20O 1,100 TOO 2,200 x 17OO
x x 45.0OO x 112,900 1700
x 3.341,600 x x 16,791,500
x xx 3,576,800 x
x 3,341,600 45.OOO 3,576,800 16,904,400
x xxx x 170O
X xxx X "
X 617, 6OO XX x "
x 3O6.7OO xx x "
1,535,900 xxx x "
7.OOO xxx x "
X X l.OOO X X "
x 261. SOO xx x "
1,542,900 1,185, 800 l.OOO x x
1,543,100 4,528.500 46.7OO 3,710,600 16,904,400
All or Cp
LH
MM Btu
266 Btu/lb
(Table 14)
3183 Btu/mol
3084
1260.2 Btu/lb.
1609.2 Btu/lb.
23.3 Btu/mol-'F
0.255 Btu/lb-°F
542 Btu/lb
542 Btu/lb
23.0 Btu/mol-°F
23.0
O.2SI Btu/lo-"F
23, 134 Btu/mol
11,641
12,347
19,290
12,213
19,203
16,094 "
1885.2 Btu/lb
268.1
34. 0
124. 0
10.2
168.2
989.2
9,014.0
1,678.2
10,692.2
63.9
72.8
12,528.6
161.9
8.8
53.0
7,877.8
1,472.3
9,403.1
5.6
534.4
476.6
184.8
669.5
9.6
0.5
555. 1
2,436.1
339.7
1.4
177.6
12,528.6
Heat of
Cox bullion
HHV
MM Btu
12,641.1
x
X
X
X
z
X
X
X
2,447.3
204.4
X
X
X
X
92. 6
5,654.7
4,698.0
x
x
82.5
7.5
x
10,535.3
-------
TABLE 16
Uaaa fc He»t Balance
Regenerator
Case 111
&••!•: 1 hour Dfttun
Inp
©
©
,12)
T5)
ut
Char
Lift Caa
CO2
X2
H2S
H^ (v)
Sub-total
Air
02
N2
II20 (v)
Sub-total
Acceptor
' MsO'CaS
MgO- CaO
Inert
£ub-total
V-fiCOs" CaCO3
Inert
Sub- total
: 6O«F HjO( i
Pound*
298,800
54 , 90O
136,400
200
5,300
196,000
474.700
1,563.300
1 1 , -00
2,0-19,200
157,500
20.092,800
3,569,600
23.81'J.'.IOO
77.600
7,200
84,800
)
Ho la
-
1,248
4,870
6
296
6. -120
14.833
55,800
C^ 1
71.254
1.401
208.431
-
421
-
- See Figure 5
Uol % Elemental Balunce, rounda
H C
1.2OO 166,000
19.4 x 15.OOO
75.9 x X
O.I x x
4.6 600 x
100. O 600 15,000
20. 8 x x
78.3 x x
0.9 1.3OO x
100.0 1,300 x
- X X
- X X
- X X
- X X
x 10, 100
X X
x 10, 10O
N 08 Ash MgO-Ca
X xx 131,600 x
x 39,900 xx x
136,400 xxx X
x x 200 x X
x 4, TOO XX x
136,400 44,600 200 x x
x 474,700 x x x
1,563,300 xxx x
x 9.9OO XX x
1,563,300 484,600 x x x
x x 44.9OO x 112,600
x 3,334,900 x x 16,757,900
x xx 3,569,600 x
x 3,334,900 44,900 3,569,600 16,870,500
x 33,700 x x 33,800
x xx 7,200 x
x 33,700 x 7,200 33,800
Temp. AH or Co £H
°P KM Btu
1700 542 Btu/lb 161.9
28O 2O72 Btu/mol 2.6
1534 " 7.5
" x
1158.4 Btu/lb 6.2
16.3
315 1816 Btu/mol 26.9
1779 " 99.3
1174.5 Dtu/lb 13. 1
139.3
17OO 23. O Btu/mol-«F 52.8
" 23.0 " 7,862.0
O.251 Btu/lb-°F 1,469.4
9,384.2
60 - x
" X
X
Beat of
Combustion
HHV
SIM Btu
2,447.3
x
X
1.3
X
1.3
x
X
X
X
X
X
X
X
X
X
X
Heat of Reaction
Out
S)
(*)
^"^
©
MKO- CaS
Meat of
Totala
put
Char
Acceptor
'•&>• CaO
Inert
Sub-total
Ca>
C02
"2
II20 (v)
Sub-total
M«COS
Heat Losa
Totala
to MgO- CaO ( 1
,4O1 mola x
Combustion (Reactants -
26,449,500
148,200
20,268,400
3.576.800
23.845.200
639.300
1.699.700
90. 100
J 7 . L'OO
2.456.300
Cn-Oj lo Mfc-0
2li,-l-l;>. 7lhl
-
210,253
-
-
14,525
6O.67O
L.407
1,510
78, 112
•CnO (421 t>
-
196,600 Btu/ooO
Products ex Acceptor)
3,100 191,100
x 16,600
X X
- X X
X X
18.6 x 174,500
77.7 x x
1.8 x x
1.9 3.10O x
1OO.O 3,100 174,500
oln « 127,4OO lllu/Bol)
3. I CIO 101, IOO
1,699,700 3,897,800 45,100 3,708,400 16,904,300
x xx '131,600 x
x 3,364,000 x x 16,904,400
x xx 3^576^800 x
x 3,364,000 x 3,576,800 16,904,400
x 464,800 x x
1,699,700 xx x
x 45,000 45,100 x
x 24, IOO x x
1,699,700 533,900 45, IOO x
1,G'.I!).7()0 3,H!)7.!I()0 45 , IOO 3, 7OH . 40O 1G,9O1,4OO
275.4
2,212.4
12,189.5
19OO 496 Btu/lb. 73.6
19OO 23.3 Btu/mol-°F 9,014.0
0.255 Btu/lb-°F 1,678.2
10,692.2
19OO 22,012 Btu/nol 319.7
" 13,837 " 839.5
" 22,247 " 31.3
" 1997. 1 Btu/lb. 54.3
1,244.8
03. 6
125.3
12, IH't.S
236.2
X
X
X
X
X
X
X
X
.u
CD
-------
46.
Basis; 1 hour Datum
Input
16) Regenerator Gas
C02
N2
SO 2
H20 (v)
Sub- total
|l7) Gasifier Gas
CH4
H2
CO
CO 2
N2
H20 (v)
Sub-total
(is) Stripping Gas
C02
N2
H2S
H20 (v)
Sub- total
Heat of Reaction; Heat
Totals
Output
(14) Reductor Gas
CH4
CO 2
N2
H2S
COS
SO 2
S2 (g)
H20 (v)
Sub- total
Heat Loss
Totals
Mass
- See
; 60° F H20(
Pounds
639,300
1,699,700
90,100
27,2OO
2,456,300
100
3,100
35,900
14,000
51,100
9,800
114,000
2,900
7,400
100
4OO
10,800
& Heat Balance
Reductor
Case III
Figures 5 & 8
1)
Mols
14,525
60,670
1,407
1^510
78,112
8
1,526
1,283
318
1,822
543
5,500
65
265
3
24
357
of Combustion (Reactants -
2,581,100
100
711,800
1,758,200
21,200
1,1OO
20,600
14,300
53,700
2,581,000
2,581,000
83,969
8
16,173
62,757
623
18
322
224
2,982
83,107
83,107
695
695
695
695
128
128
128
128
128
128
- Products)
800
800
800
800
8OO
800
800
800
AH or Cp
6556 Btu/mol
4485 Btu/mol
6859 Btu/mol
1353.2 Btu/lb
470 Btu/mol
473 Btu/mol
616 Btu/mol
473 Btu/mol
1088.9 Btu/lb
120
120
120
12O 1085.6 Btu/lb
417 Btu/mol
8192
7784
5251
6525
8510
8123
6144
1404. 7
Btu/mol
Btu/mol
Btu/mol
Btu/mol
Btu/mol
Btu/mol
Btu/mol
Btu/lb
AH
MM Btu
0.1
0.5
0.6
120.1
547.5
8.4
"54775
-------
47.
TABLE 18
Mass & Heat Balance
Claus Reactor
Case III
Basis: 1 hour Datum
Input
M9\Reductor Gas
^-^ CH4
C02
N2
H2S
COS
so2
^8
H20 (v)
Sub-total
Heat of Reaction
268 S02 + 18
ICO. 5 s
Totals
Output
'20} Claus Gas
^ CH4
C02
N2
H2S
S02
Sg
H20 (v)
Sub-total
Heat Loss
- See
: 60°F H20^^
Pounds
100
711,800
1,758,200
21,200
1,100
20,600
14,300
53,700
2,581,000
COS +518 H2S
8+18 C02 + 518
2,581,000
100
712,600
1,758,200
3,600
3,400
40,100
63,000
2,581,000
Figures 5 & 8
Mo Is Temp.
OF
8 402
16, 173 "
62,757
623 "
18
322
56
2,982 "
82,939
-20(1)
82,939
8 410
16,191
62,757 "
105
53
156.5
3,500
82,770
AH or Cp
3,265 Btu/mol
3,317
2,390
2,868
3,682
3,502
13,400 "
1214.6 Btu/lb.
3,355 Btu/mol
3,403
2,447 "
2,939
3,592
13,729
1218.3 Btu/lb
AH
MM Btu
Totals
2,581,000
82,770
-------
48,
4. Economic Evaluation - Gasification Section
Basis for Cost Estimate
All of the cost estimates in this study are based on earlier
Consol studies on the C02 Acceptor Process carried out for the Office of Coal
Research!1£v updated to present-day cost indices. While the cost of individual
pieces of equipment may be estimated by this method with reasonable accuracy,
the estimate of indirect construction costs (e.g., piping, instrumentation,
etc.) is of necessity quite imprecise. Therefore, the absolute accuracy of
these cost estimates is uncertain. Differences in costs between the various
cases are probably reasonably accurate since the same method was used in all
instances.
To determine the total investment costs, the following assump-
tions were made, and are common to all the gasification cases as well as to the
combined cycle power system described in Section III-B-5, page 50.
1. Minemouth site.
2. January 1971 costs.*
3. Begin construction July 1971.
4. Operation January 1976.
5. Escalation at 7-1/2$ average of labor and material.
6. Interest during construction at 7-1/2$ simple
interest on cash flow.
7. Cooling towers not included.
Costs to be quoted later for the United Aircraft Research Laboratories
power cycle are as of January, 1970.
Escalation of direct operating costs in all cases was taken at
the basic rate of 7-1/2$ per year average for labor and material, subject to
•• the construction schedule outlined above.
The coal costs used in this study represent Consol1s estimate
for the minemouth cost of coal from a new mine in West Virginia to be opened
prior to operation of the power plant.
Basic Pressure Case I
Based on the process design for this case described previously,
and the detailed material and heat balances reported in Tables 4 through 9, a
cost estimate for each section of the plant was evolved as shown in Appendix A.
Tables A-l to A-3 in Appendix A show the ISBL (inside battery limits) cost
estimates for coal preparation, gasification, and sulfur recovery-solids
disposal, respectively. To these costs must be added off-site facilities,
utility costs, electrical substations, cooling water towers, distribution
piping, and boiler feed water treating'.
The installed plant cost is $81,000,000, not including interest
and escalation, and a breakdown by section is shown in Table 19. This is the
installed cost for the basic process scheme wherein the product gas is avail-
able under pressure at 130O°F without water scrubbing. The investment cost
-------
49.
Investment Summary
Case I
Section
Utility Requirements
Electricity, KW
Cooling Water, GPM
Boiler Feed Water, GPM
Investment, $100O
Erected Cost (iSBL)
Off-sites & Utilities
Off-sites
Electrical
Cooling Water
Boiler Feed Water
Sub-Total (OSBL)
Installed Plant Cost
Coal
Preparation
4,2OO
X
X
5,000
400
500
X
X
900
5,900
Gasification
12,40O
17,3OO
1,463
56,OOO
4,300
1,500
6OO
6OO
7,000
63,000
Interest During
Sulfur
Recovery-
Solids
Disposal
1,000
3,700
X
11,OOO
900
100
100
X
1,100
12,100
Escalation
j Construction
Totals
17,600
21,000
1,463
72,OOO
5,600
2,1OO
7OO
6OO
9,000
81,OOO
15,20O
96,2OO
16.2OO
Total
$112,400
-------
50.
includes $13 MM (installed cost, including accessory electrical equipment) for
motor-driven compressors to supply the process air. The compressors are more
properly charged to the power cycle in a case of a combined cycle power station
in order to give the "base" cost of the gasification system only. This base
cost then becomes $68 MM ($94 MM including escalation and interest during
construction). Costs associated with supply of process air for various modes
of integration with a power station are shown in Sections III-B-5 and 6. If
water scrubbing to eliminate alkalis and entrained dust is required, and the
product gas is reheated to about 11OO°F, the installed base plant cost is
increased to $94 MM. Most of the $13 MM ($18 MM with escalation and interest
during construction) incremental cost is associated with the very large gas-gas
heat exchangers which are required.
Direct operating costs for the basic flow scheme were estimated
as shown in Table 20.
Alternate Pressure Case II
The alternate pressure case yields almost the same quantity of
product gas at the same conditions of 1300°F and 206 psia. The higher invest-
ment cost ($4 MM) shown in Table 21 for Case II is largely the result of
increased vessel costs in the sulfur recovery section and the need for stainless
steel-lined equipment to cope with the corrosive H2S03 stream. The somewhat
higher direct operating costs shown in Table 21 stem from increased operating
labor and power costs for compression of the air and gas streams in the sulfur
recovery section.
Atmospheric Pressure Case III
To fulfill the requirements of Phase I under the contract, the
same economic evaluation of this case was prepared as discussed for Case I.
The detailed ISBL estimates are shown in Tables A-4 to A-6 in
Appendix A. The summarized cost for the installed plant is given in Table 22,
and the direct operating costs in Table 23.
5. Integration with Combined Cycle Power Systems
Basis for Evaluation
The cost of producing power from low-sulfur boiler fuel will be
relatively high when the C02 Acceptor Process is used simply to produce clean
fuel gas for an existing boiler. Improved economics and station heat rates can
be achieved by integration of the gasification section with combined cycle power
systems. This section gives the design basis used for evaluation of such an
integrated system.
two studies, sponsored by OAP, were made by Westinghouse Research
Laboratories (OAP Contract No. CPA-70-9) and by United Aircraft Research
Laboratories^19) in which designs and cost estimates for advanced power cycles
were performed. In Germany, STEAG is building a 17O MW combined cycle power
station which uses the Lurgi pressure gasification process to provide low-Btu
fuel gas.
-------
51.
Direct Operating Cost Summary,
Excluding Coal Cost
Case I
Basis: 70$ Plant Operating Factor
Section
Direct Operating Labor, Men/Shift
Direct Operating Cost
1. Operating Labor at
$40,OOO/man/shift/yr
2. Maintenance Labor at
1.6$ Investment
3. Direct Supervision
15$ of 1. + 2.
4. Indirect Overhead
50$ of 1. + 2. -f 3.
5. Payroll Overhead
15$ of 1. + 2. + 3. + 4.
6. Maintenance Material at
2.4$ Investment
7. Miscellaneous Supplies
15$ of Maintenance
Material
8. Utilities
Electricity at 7.5 mills/KWH
Cooling Water at 2^/10OO gal
BFW at 25^/1000 gal
Chemicals & Catalyst
TOTALS
Coal Gasification Sulfur Totals
Preparation Recovery-
Solids
Disposal
3 10 6 19
&10OO/Year
120
94
32
123
55
142
21
194
400
1,008
211
8O9
364
1,512
227
570(0
128
135
830
240
194
65
25O
113
290
44
46
27
781
6,194 1,340
Escalation
Total
760
1,296
308
1,182
532
1,944
292
810
155
135
901
8,315
1.560(2)
9,875
Ex. main air compressor duty.
Direct operating costs shown here are investment-sensitive (items 2
through 7}. Therefore, these costs are escalated in the same manner
used for the plant investment shown in Table 19, i.e., the average
escalation over the 5-year period, 1971-1976, at a rate of 7.5$/yr
for labor and materials.
-------
TABLE 21
Economic Comparison - Cases I and II
52.
Case I
Case II
Installed Plant Cost $81,000,000
Annual Direct Operating Costs
(ex Coal)
$85,000,000
$ 8,315,OOO/yr $ 8,776,000/yr
TABLE 22
Investment Summary
Case III
Section
Utility Requirements
Electricity, KW
Cooling Water, GPM
Boiler Feed Water, GPM
Investment, $1OOO
Erected Cost (ISBL)
Offsites & Utilities
Offsites
Electrical
Cooling Water
Boiler Feed Water
Sub-total (OSBL)
Installed Plant Cost
Coal
Preparation
Gasification
4750
90
4,500
400
600
1,OOO
5,500
6200
45,680
1,310
76,500
6,1OO
7OO
1,600
6OO
9,OOO
85,500
Sulfur
Recovery-
Solids
Disposal
135O
1160
Totals
10,000
800
200
1,OOO
11,OOO
Escalation
Interest During Construction
Total
12,3OO
46,930
1,310
91,OOO
7,300
1,500
1,6OO
6OO
11,OOO
102,OOO
19.200
121,200
2O.500
141,700
-------
53.
TABLE 23
Direct Operating Cost Summary,
Excluding Coal Cost
Case III
Basis: 70$ Plant Operating Factor
Section
Direct Operating Labor, Men/Shift
Direct Operating Costs
1. Operating Labor at
$40,000/man/shift/yr
2. Maintenance Labor at
1.6$ Investment
3. Direct Supervision
15$ of 1. + 2.
4. Indirect Overhead
50$ of 1. + 2. + 3.
5. Payroll Overhead
15$ of 1.+2.+3.+4.
6. Maintenanct Material at
2.4$ Investment
7. Miscellaneous Supplies
15$ of Maintenance
Material
8. Utilities
Electricity at 7.5 mills/KWH
Cooling Water at 2^/1000 gal
BFW at 25^/1000 gal
Chemicals & Catalyst
TOTALS
Coal Gasification Sulfur
Preparation Recovery-
Solids
Disposal
Totals
3
120
88
31
12O
54
132
20
219
1
785
10
$1OOO/Y
400
1,368
265
1,016
457
2,052
308
285
336
121
910
7,518
6
ear
240
176
63
240
1O8
264
39
63
9
105
1,307
Escalation
19
760
1,632
359
1,376
619
2,448
367
567
346
121
1,O15
9,610
1.800
Total 11,410
-------
94.
In this section, integration of the Case I fuel gas with two
advanced power cycles is considered. It is emphasized that no detailed designs
and cost estimates for power station components were made. Results of the
Westinghouse and UARL studies, information obtained from Lurgi,(8) and costs for
conventional coal-fired stations presented by United Engineers and
Constructors(20) provided the bases for estimation.
For each power cycle, two versions of the Case I fuel gas product
were used. In one, the assumption was made that particulate matter and alkali
in the product gas, after cooling to 1300°F, could be reduced sufficiently by
high pressure drop cyclones to allow sustained operation of gas turbine engines
and expanders. In the other, the assumption was made that wet scrubbing would
be required for adequate cleanup. In this instance the gas is wet scrubbed at
the saturation temperature (~240°F) corresponding to its steam content. The
gas then is reheated in a heat exchanger to 11OO°F. Also, ammonia formed from
the coal nitrogen in the gasifier was assumed to be removed from the fuel gas
during scrubbing.
Supercharged Boiler Cycle
In this power cycle fuel gas from the C02 acceptor gasifier is
burned at elevated pressures in a boiler which provides steam for power genera-
tion by a conventional steam cycle. Hot gas from the boiler is then passed
through an expander which drives the compressor for combustion air and a
generator which produces additional electrical power. The expander exhaust gas
is used to preheat the boiler feed water before final discharge to the stack.
This cycle is used in the Westinghouse study (high-pressure fluidized bed
boiler) and in STEAG's KeHermann station (high-pressure Lurgi fuel gas).
Figure 9 is a schematic diagram showing integration of Consol's
Case I gasification process with the supercharged boiler cycle. Flows and
power outputs are summarized in Table 24. For the steam cycle, steam conditions
of 3500 psia/1050°F/1050°F were chosen on the basis of high efficiency.
Detailed design and cost estimation would be required to show whether this cycle
is justified, compared with a lower cost subcritical cycle. Gross output of
the expander was taken as 91$ of the isentropic output for the particular work-
ing fluid used. Air compression duties were calculated at 89$ of the polytropic
efficiency. Thermodynamic properties of dry air were taken from the NBS
tables.(*)
On the basis of available technology, using cooled bleed air for
blade cooling, the expander inlet gas temperature was taken at 1800°F. No
specification is made as to the unit sizes of the expanders and steam turbines.
By the nature of the gasification process (12 gasifiers and 4 regenerators, as
shown in Figure 3), multiple units are envisioned.
The assumption was made that complete combustion of the fuel gas
can be achieved in the supercharged boiler with 5$ excess air. At this level
of excess air and with the use of two-stage combustion(21) of the low-Btu fuel
gas, formation of nitrogen oxides should be drastically reduced from the levels
associated with conventional PF boilers.
-------
FIGURE 9
SCHEMATIC FLOW DIAGRAM
CASE I GASIFICATION SYSTEM
SUPERCHARGED BOILER
55.
AIR
897°
15.3 P
EXPANDER
COAL
1800*
202P
436°
422<
2O6P
(412°)
228 P
(398°)
GASIFICATION
I30O'
FUEL GAS
206 P
TO STACK
STACK GAS
COOLER
300°
14.7 P
Heat Source For Feed Water Heaters;
T) From Air Compressor Intercooter
BOILER Fl
PUMP
CONDENSATE
PUMP
Multistage, By Steam Extracted From LP Turbine.
MM N II *M H
0 Fahrenheit
P • PSIA
Multistage,
IP
-------
TAULE 24
Power Cycles
Case I Gaslfier
Superelm rued Holier Power Cvcle
56.
Method of Fuel Cos Cleanup
(partlculate and nlkali)
Expanders
Inlet Gas
MolF/hr
Temperature, °F
Pressure, psia
Exhaust Gas
Temperature, °F
Pressure, psia
Air Compressors
Combustion Air
Mols/lir
Exit temperature, °F(I'
Exit pressure, psia
Process Air
Mols/hr
Exit temperature, °F
Exit pressure, psia
Bleed air for cooling (150°F),
$ of total airflow
Busbar Power Output, Mw(a/
Gas Cycle
414,9OO
18OO
2O2
897
15.3
216,900
422
206
136,500
436
228
4.O
513
412,4OO
1800
196
9O5
15.3
215,100
418
200
136,500
436
228
4.O
494
Steam Cycle
Steam Turbines
Steam Flow, MM Ib/hr
Main steam
Reheat steam
Extracted for FW heating
Throttle temperature, °F
Throttle pressure, psia
Reheat temperature, °F
Reheat pressure, psia
Condenser temperature, °F
Condenser pressure, inches Hg
Boilers
Fuel Gas
Mols/hr
Temperature, "F
Pressure, psia
Combustion Air
Mols/hr
Temperature, °F
Pressure, psia
Products of Combustion (5$ excess air)
Mols/hr
Temperature, °F
Pressure, psia
Feedwater
MM Ib/hr
Temperature, "F
Pressure, psia
Stack Gas
Mols/hr
Temperature, °F
Pressure, psia
Busbar Power Output, Mwt3'
6.O17
S.95O
1.464
1O5O
3500
1O50
62O
1O1
2.O
234,980
13OO
206
216,9OO
412
2O6
414,9OO
18OO
202
6.017
650
40OO
429,600
300
14.7
1O49
5.578
5.52O
1.222
1050
3500
1O5O
62O
101
2.0
234,470
110O
200
215,1OO
4O8
200
412,400
180O
190
5.578
650
4OOO
427,000
300
14.7
983
(i) One stuge of interceding to 12O"F.
( i) 98$ of expander power output minus total comprissor duty.
(») After allowance of 3.9$ for auxiliaries, heat and generator losses.
-------
57
Exhaust Gas Cycle with 280O°F Turbine Inlet Temperature
This cycle is used in the UARL study.(3>19) Fuel gas is burned
with sufficient excess air in the combustor of a gas turbine engine to provide
inlet gas to the compressor turbine at 28OO°F. The power turbine drives an
electric generator. The hot exhaust gas from the power turbine provides steam
in a heat recovery boiler for power generation by a conventional steam cycle.
No fuel is burned in the boiler and the feed water is not preheated.
Figure 10 is a schematic diagram showing integration of Consol's
Case I with the exhaust gas cycle. Flows and power outputs are shown in
Table 25.
Since the fuel gas composition and sensible heat content are
appreciably different from those used in the UARL study, some modifications
were required to adapt the UARL estimate for use in this study. The gross
power outputs of the compressor and power turbines were calculated, using
of isentropic efficiency for the particular working fluid used in Case I.
Air compression duties, steam conditions, and steam cycle efficiencies are on
the same basis as in the UARL study.
In Case I, the fuel gas is produced at somewhat higher pressure
than the 190 psia level used by UARL. The schematic diagram shows an expander,
which produces additional power, and which brings the fuel gas to the same
pressure as used by UARL. In practice, the gas turbine engine would be modi-
fied to accommodate the higher inlet pressure. However, to keep the same
pressure ratios in the gas turbine engine as used by UARL, the expander is
necessary.
Economics
The capital costs for the gasification system are taken from
Table 19. The costs for the power stations, unit costs ($/KW of installed
capacity) were obtained from four sources, as outlined below;
Power Cycle
Category . Supercharged Boiler Exhaust Gas
Structures and Improvements UE U
Boiler Plant L U
Steam Turbine - Generator UE U
Gas Turbine - Generator W2 U
Stack Gas Cooler Wl
Switchyard and Station Equipment UE UE
Accessory Electrical Equipment
and Miscellaneous UE U
Indirect Construction Costs UE* )
Land and Contingency UE** )
Kev:
Wl Westinghouse, Reference 10. L Lurgi, Reference 8.
W2 Westinghouse, Reference lOa. U UARL, Reference 3, 19.
UE United Engineers and Constructors, Reference 2O.
* Indirect construction costs taken as 7.4$ of direct cost, as shown in
Reference 20.
** Contingency taken as 6.4$ of direct + indirect costs, as shown in
Reference 20.
-------
FIGURE IO
58,
SCHEMATIC FLOW DIAGRAM
CASE I GASIFICATION SYSTEM - EXHAUST GAS CYCLE
2800°F TURBINE INLET TEMPERATURE
COAL
1
AIR
GASIFICATION
(398°) 436°
228P
EXPANDER
AIR
1300°
206 P
5-
MOTOR
694*
176 P
1270°
I90P
COMBUSTOR
2800
1587°
HEAT
RECOVERY
BOILER
40 H20
TO STACK
285°
1000°
STEAM
TURBINE
2415 P
14.7 P
0 Fahrenheit
P = PSIA
-------
59.
jfowcy Cycles
Cose I - Gasification System
Method of Fuel Gas Cleanup
(paniculate and alkali)
Gas Cycle
Engines
Compressor Pressure Ratio
Inlet Gas
Mols/hr
Temperature, °F
Exhaust Gas
Mols/hr
Temperature, °F
Bleed air for cooling (200°F),
% of total airflow
Compressor Airflow, mols/hr
Combustion air
Bleed air
Compressor outlet temperature, °F
Fuel Gas Expanders
Inlet pressure, psia
Inlet temperature, °F
Outlet pressure, psia
Outlet temperature, °F
Busbar Power Output, MW
Engines (') .
Expanders
Total
Steam Cycle
Steam Turbines( 2)
Throttle temperature, °F
Throttle pressure, psig
Throttle flow, MM Ib/hr
Reheat temperature, °F
Reheat pressure, psig
Reheat flow, MM Ib/hr
Heat Recovery Boilers *
Inlet Gas
Heat content, MM Btu/hr
Temperature, °F
Pressure, psia
Stack Gas i
Heat content, MM Btu/hr
Temperature, °F
Pressure, psia
Net Boiler Duty, MM Btu/hr
Busbar Power Output, Mw(*'
Hot
12
565,100
2800
599,400
1587
8.54
367,100
34,3OO
694
2O6
1300
19O
127O
1310
17
1327
1OOO
240O
4272
1OOO
569
4218
721O
1587
16.17
971
285
14.7
6239
726
Wet
Scrubbing
12
540,000
2800
572,000
1590
8.54
343,000
32,1OO
694
2OO
1100
190
1081
1266
1O
1276
1OOO
24OO
3944
1000
569
3894
6690
1590
16. 17
93O
285
14.7
576O
670
(i) 98$ of gross power output minus compressor duty.
(a) Same steam conditions as used by United Aircraft Research Laboratories, Reference 3.
(») At same ratio of power output/net boiler duty as used by United Aircraft Research La
Laboratories.
-------
60.
Supercharged boiler costs obtained from Lurgi for the Kellermann plant were
arbitrarily increased by 50$ to compensate for addition of reheater surface
and for the higher steam pressure. The Kellerman plant uses a modest steam
cycle; 1900 psi/1000°F/no reheat.
Table 26 shows the summarized power station investment costs
and includes a comparison with a modern conventional PF station equipped with
Consol's Formate Process!17) for stack gas scrubbing. Investment and operating
costs for stack gas scrubbing are from unpublished work by Consol. The Formate
Process for stack gas scrubbing was chosen to put the sulfur recovery systems
on a comparable basis, i.e., sulfur is recovered in all cases as elemental
sulfur. The major item in the operating cost of the Formate Process is the
cost of fuel gas required to provide reducing gas for the regeneration section.
OriginallyC17) natural gas was used as fuel at 40^/MM Btu. Since that time a
severe shortage of natural gas has developed, and it was necessary for this
study to increase the fuel costs for the regeneration section from 40^ to
80^/MM Btu.
Total costs for a conventional PF station were presented by
United Engineers and Constructors in reference 20. Breakdown into the in-
dividual Federal Power Commission account numbers was on the basis of a
previous study by UEC which was made for Westinghouse in 1965.I11)
In the supercharged boiler systems, process air for the gasifi-
cation plant is supplied by compressors driven by the power station expanders.
Thus, the investment associated with process air are included in the entries
for "Gas Turbine - Generator" in Table 26. In the exhaust gas systems, the
process air is supplied by motor-driven compressors, as in the UARL study.
Thus, the gasification plant investment corresponds exactly to that shown in
Table 19.
Table 27 shows the summarized power costs for the same stations
which are in Table 26. Direct operating costs for the power plants shown in
Table 27 attempt to reflect the higher maintenance costs which probably will
be associated with advanced cycle gas turbines and expanders and to reflect
the elimination of coal and ash handling costs in the combined cycle stations.
The direct operating costs for the conventional coal-fired station were based
on annual Federal Power Commission data for modern, efficient stations.
As Table 27 shows, there is considerable incentive to develop
a hot gas clean-up process which operates at about 1300°F. The increased
costs due to wet gas scrubbing can be reduced somewhat over those shown in the
table if the heat rejected to cooling water in the scrubber circuit (~450 MM
Btu/hr in Case l) could be utilized in the power station. This possibility
was not investigated.
6. Low-Sulfur Fuel Gas for Conventional Stations
Introduction
Integration of the low-sulfur gas process with advanced cycle
power stations appears to be attractive for long-term requirements for power
generation. >
-------
Summarized Power System Investment Costs
1976 Operation. See page 43 for Plant Cost Basis
System
Fuel Gas Cleanup
(paxtlculate and alkali)
Busbar Power, MW
Air Compression, MW
Miscellaneous Drives, MW
Net Power, !.IW
Sulfur Removal, %
(as elemental sulfur)
Overall Station Heat Rate,
Btu/Xet KK1I
Busbar Power, $ from
Stean Cycle
Gas Cycle
System Investment
Power Station. $/KV'' Busbar Power
Year for Cost Basis
Structures and Improvements
Boiler Plant
Stean Turbine - Generator
Gas Turbine - Generator
Stack Gas Cooler
Switchyard and Station Equipment
Ace. Elect. Equip, and Misc.
Land
Indirect Construction Costs
Contingency
Sulfur Removal
Total Capitalized Cost
Escalation
Interest During Construction
Total
$/KW Net Pov.-er
Cis if icat i-jr. Plant. M.t$ Inv.
BJSC Plant
Air Compression
Wet Scrubbing
Total
$/K1V Net Power
Total Sv.ste--. Cos:. ; fry: N'et Power
(l) After sub; rac: lor. of po*er for auxiliaries.
(l) Via Consol's Forr.ate Process. Reference 17
natter prior to S02 removal, if required.'
(3) In gas:ficu:ion plant.
Conventional Coal
Fired With
Stack Gas Scrubbing
iooo(0
—
—
10OO
on
so
8300
100
0
1971
13.2
75.3
30.3
—
' —
9.7
8.8
.4
10.2
9.5
22.4(0
179.8
33.7
213.5
36.0
249.5
249. 5
—
—
—
~
249.5
Can
Suoerchareed Boiler
Hot
1562
—
18
1544
f
893O
67.2
32.8
1971
9.5
16.8
20.3
18.5
9.O
9.7
8.8
.4
6.9
6.4
(•)
106.3
JL9.9
126.2
21.3
147.5
149.2
94.4
—
—
94.4
61.1
210.3
Wet
Scrubbing
1477
.
18
1459
9450
66.5
33.5
1971
9.5
16.6
20.1
19.3
9.0
9.7
8.8
.4
6.9
6.4
(>)
106.7
20.0
126.7
21.4
148.1
150. 0
94.4
—
^8.O
112.4
77.0
227. O
le I - Gasification With
Exhaust Gas
Cycle With
28OOP F Turbine Inlet Ternerat'jre
Hot
2O53
168
18
1867
7390
3S.3
64.7
197O
7.7
13.2
13. 5
19.8
—
9.7
7.2
.2
\ 10.6
(0
81.9
23. 0
1O4.9
17.7
122.6
134.8
94. 4
18. O
—
112.4
60.2
195.0
Wet
Scrubbing
1946
168
18
1760
7840
34.4
65.6
1970
7.7
12.8
13.2
19.9
--
9.7
7.2
.2
\ 10.6
(»)
81.3
22.9
104.2
17.6
121.8
134.7
94.4
18.0
18. 0
130.5
74. 1
2O8.8
o>
- Investment docs not include separate scrubber for removal of residual particulate
-------
TABLE 27
Power Costs
1976 Operation. See pace 4S for Plant Cost Basis
Basis;
Coal
at 4Of?/lM Btu 15$ Capital Charges
4.3$ Sulfur Content 6132 hrs/yr = 7O$ plant factor
12,7OO Btu/lb $15/LT Sulfur Credit
Conventional Coal
Case I - Gasification With
System
Fuel Gas Cleanup
(Partlculate and alkali)
Heat Rate. Btu/KVH
Busbar Power
Net Power
Investment, £/KW Net Power
Gasification
Power Plant
Total Station
Coal
Gasification
Direct Operating Cost
Capital Charges
Power Cost for Air Compression
Sub Total Gasification
PQWCT Plant
Direct Operating Cost
Capital Charges
Sulfur Credit
Total
Fired With
Stack Gas Scrubbing
esoof1*2'
8900( *)
Sl.l(')
218.4
249.5
3.56
.77(3}
.76l3)
1.53(>)
.48
5.34
10^76
Superchareed Boiler
Hot
8838
893O
61.1
149.2
21O.3
3.54
.94
1.48
2.42
.48
3.61
9.88
Wet
Scrubbing -
9347
9450
77.0
15O.O
227.0
Mills /KWH Busbar Power
3.74
1.09
1.86
2.95
.48
3.62
_-,18
10.61
Exhaust Gas Cycle With
28OO°F Turbine Inlet Ter.perature
Hot
6727
7390
60.2
134.8
195. 0
2.69
.78
1.34
.74
2.86
.59
3.00
-.13
9.01
Wet
Scrubbing
7099
784O
74.1
134.7
2O8.8
2.84
.90
1.64
.78
3.32
.59
2.98
-. 14
9.59
(i) After subtraction of power for auxiliaries.
(a) Based on coal burned in boiler.
(>) Includes Investment for sulfur removal via Consol's Formate Process. Reference 17 - Investment for separate scrubber for removal of residual particulate, which
r.»y be required, is not included. 640 Btu/KWI fuel equivalent is required to supply rcductant for sulfur recovery, fuel for fired heaters and steam and power
to drive fans and pumps in SO2 and sulfur recovery sections. Major purchased fuel is natural gas which is included in direct operating cost at 8Oj?/MM Btu.
Cl
to
-------
63.
Existing stations, now faced with restrictions on sulfur
emissions and with future restrictions on nitrogen oxide emissions require a
shorter-term approach. Successful development of the CO2 Acceptor Process
through Phases II and III of the GAP contract would provide a practical
alternative for sulfur cleanup in the event that a stack gas treatment
process would not be feasible or practical.
Use of fuel gas produced by the C02 Acceptor Process in a con-
ventional boiler also potentially offers the following advantages over stack
gas clean-up processes;
a. Lower nitrogen oxide emissions because the amount
of excess air can be reduced below that used in
coal or oil-fired stations, and because two-stage
combustion of the fuel gas can reduce greatly the
formation of the nitrogen oxides.
b. Higher levels of sulfur cleanup - potentially as
high as 96$.
c. The process directly produces elemental sulfur,
which can be stockpiled cheaply without danger of
pollution. Not all of the stack gas treatment
processes now under development allow recovery of
elemental sulfur.
d. The fuel gas sulfur content is dictated by in situ
reactions in the gasifier. Therefore, the sulfur
recovery section of the plant need not be highly
efficient.
A limitation of the C02 Acceptor Process is that its use would
be restricted to stations which have adequate land available for installation
of the fuel gas plant.
Basis for Evaluation
For evaluation of the cost of low-sulfur fuel gas, Case I was
chosen. The product gas, at 13OO°F and 2O6 psig, is passed through expanders
which reduce the pressure to 25.7 psia. The expanded gas is delivered at
665°F to the burners at the boiler. The expanders drive compressors for the
process air and produce an additional 170 MW of power.
The cost of the expander-compressor-generator system was esti-
mated at $13.2 MM which, added to the "base" plant cost of $68 MM (page 50),
gives the installed plant cost as $81.2 MM. Direct operating cost was
increased over that shown in Table 20 to allow for increased maintenance.
Economics
The cost evaluation is summarized in Table I in the Summary.
The fuel gas cost is shown on the bases both of the heat of combustion of
the, gas and of the heat of combustion plus the sensible heat content. The
latter cost is more realistic since the sensible heat represents available
heat to the boiler.
-------
64.
The impact of coal cost on fuel gas cost is shown in Figure 1.
From the standpoint of sulfur emissions, the C02 Acceptor
Process converts coal having 4.3$ sulfur content to a fuel gas which has a
sulfur content equivalent to that in 0.3$ sulfur content fuel oil, a material
which is not now available in large quantity at any price.
In areas which impose severe restrictions of sulfur emissions,
the costs shown in Figure 1 indicate that fuel gas from the C02 Acceptor Process
would be competitive with low-sulfur imported fuel oil, with the added advantage
of the lower NOX emissions. Further, production of power from coal assures
continuity of supply and reasonably predictable costs in contrast with
politically-caused uncertainties involved with use of foreign oil.
Low-sulfur fuel oil now is being delivered to barges at East
Coast ports at costs between 70-85^/MM Btu, depending upon sulfur content and
heating value.!12) Residual fuel oil with a sulfur level equivalent to that
potentially possible from the processes considered here (ca. O.3 wt % sulfur)
is not available today. It would be necessary to use low-sulfur distillate
fuels, at least in part, to meet these sulfur specifications.
-------
IV. EQUIPMENT &
PROCEDURE
-------
65.
IV. EQUIPMENT AND PROCEDURE
A. Description of the Continuous Unit
1. Process Piping and Equipment
A simplified flow diagram for the continuous gasification unit is
shown in Figure 11. A more detailed process and instrumentation flow diagram
is shown in Figure 12. These figures show the unit as arranged for "integrated"
operation where the acceptor is recirculated continuously between the gasifier
and regenerator.
Figure 13 shows a simplified flow diagram of the unit as arranged for
single-vessel operation where the reactor (vessel D-l in Figure 12) is used
either as a preoxidizer or as a gasifier. The regenerator is not used since
acceptor circulation is not required.
The following description applies to operation of the integrated
system where the feedstock to the gasifier is a precarbonized char. In future
operation of the integrated system, the char feedstock will be produced in a
separate operation by partial gasification of preoxidized coal in the single-
vessel system shown in Figure 13.
For single-vessel operation, the equipment and procedure in general
are the same as described here except that all processing steps involving the
acceptor and regenerator are eliminated. This mode of operation is described
separately in Section IV-D, page 97.
a. Process Vessels
The continuous unit was designed on the basis of 2-10 Ib/hr of
char feed to the gasifier. The corresponding acceptor circulation rates
required to maintain heat balance were in the range of 5-20 Ib/hr (raw acceptor
basis). These feed rates and the anticipated kinetics for the steam-carbon
reaction and acceptor calcining reaction dictated the sizes of the process
vessels.
The gasifier vessel D-l shown in Figure 14 is 4" in diameter.
The char bed height, fixed by an overflow weir is about 38", and the height of
the acceptor bed directly below is about 7-1/2" .
The regenerator vessel D-2 shown in Figure 15 is 3" in diameter.
The bed height is maintained at 18" by an overflow weir. Details of the in-
ternals and fittings of both vessels are given in Figure 16.
Heat was supplied through the walls of the process vessels. The
gasifier heater contained three separate circuits which corresponded to the
bottom, middle and top zones of the reactor. The regenerator was equipped with
four heater circuits spaced along the length of the reactor. All the.gasifier
heaters and the top zone regenerator heater were controlled by 270 volt Power-
stats supplying 5.1 KW of electrical heat. The three lower heaters on the
regenerator were controlled by 230 volt Powerstats supplying 3.8 KW.
-------
66.
The process vessels and their heaters were contained in pressure
shells which were pressurized with C02, so that the pressure drops across the
walls of the process vessels are zero. Figures 17 and 18 show the gasifier and
regenerator pressure shells, respectively.
Johns-Manville Sil-0-Cel C-3 insulation was used to fill the
annular space between the heaters and the wall of the regenerator pressure shell,
The regenerator pressure shell wall temperature was held at 950°F. The annular
space between the heaters and wall of the gasifier pressure shell was filled
with Carborundum "Fiberfrax" insulation. The gasifier pressure shell wall
temperature was held below 1100°F.
b. Solids Handling
Seven solids streams were handled in the continuous gasification
unit;
(l) Char feed to the gasifier.
(2) Char product from the gasifier.
* (3) Fuel char feed to the regenerator.
(4) Recarbonated acceptor to the regenerator.
(5) Calcined acceptor from the regenerator to
the gasifier.
(6) Recarbonated acceptor from the gasifier.
(?) Overhead ash material from the regenerator.
* Fuel char is ungasified product char from the gasifier
which is burned in the regenerator.
The residual fuel char was stripped from the regenerator bed and
carried out overhead in the regenerator outlet gas stream. The overhead solids
were removed from the gas by an external cyclone G-l.
The acceptor entering the gasifier in Stream 5 recarbonated as
it showered down through the gasifier char bed, and then segregated at the
bottom forming a sharply defined interface with the char bed. This was stabi-
lized by means of a 2-1/2" diameter boot on the bottom of the char bed.
Solids streams 1, 3, and 4, were conveyed in electrically heated
pneumatic transfer lines.
The recarbonated acceptor and fuel char were introduced at the
bottom of the regenerator bed through a single axial tube so that the sensible
heat of the acceptor and the heat of reaction of the endothermic calcining
reaction would provide a heat sink for the heat released by the combustion of
fuel char.
Streams 2} 5, and 6 were conveyed by gravity flow in purged
standlegs. These standlegs formed the seals which prevent cross-contamination
between the H2-bsaring gasifier gas and 02~bearing regenerator gas. The solids
flow rates in Streams 2 and 5 were controlled to maintain the standlegs full
of solids and the solids flow rate in Stream 6 was controlled to maintain a
constant char-acceptor interface level. To protect the Teflon seats in the
valves used to control the solids flow rates, water-jacketed coolers were in-
stalled in each of the standlegs. These coolers are specified in Figure 19.
Sight glasses (Jerguson 12-T-20) were also installed in each standleg to allow
visual confirmation of the amount and nature of solids flow.
-------
67.
The gasifier char overflow weir was placed inside a compartment
which occupied 30$ of the vessel cross-section to prevent contamination of the
product char by the acceptor which entered the gasifier.
The calcined acceptor entered the top flange of the gasifier
and passed through a preheater purged with N2 which was introduced into the
acceptor return line above the vessel. The preheater was a stainless steel
tube, 7/8" OD x 3/4" ID x 10" long having nine baffles installed at 60° angles.
In the preheater the acceptor was heated to about 1200°F before falling into
the char bed.
Feed hoppers and withdrawal hoppers were provided so that the
solids flow rates in Streams 1, 3, and 4 could be controlled independently of
the inventory requirements of the fluidized beds and standlegs.
The gasifier char feedstock was supplied from one of two char
feed hoppers, F-1A and F-1B, each having a capacity of about 100 pounds.
Continuity of feed was maintained by feeding char alternately from the two
feed hoppers.
Gasifier product char was withdrawn into a lockhopper F-2 at a
rate dictated by the overflow in the weir. The capacity of the lockhopper was
approximately ten pounds.
Recarbonated acceptor from the gasifier was withdrawn into one
of two lockhoppers, F-6A and F-6B, at a rate dictated by the level of the gasi-
fier char acceptor interface. The acceptor was removed from the off-stream
lockhopper and charged back to one of two acceptor feed hoppers, F-4A and F-4B.
Continuity of acceptor circulation was maintained by feeding acceptor to the
regenerator alternately from the two feed hoppers. A sight glass was installed
below the feed hoppers to show when the on-stream feed hopper was empty. The
feed hoppers and withdrawal hoppers each have a capacity of six pounds of raw
acceptor.
A lockhopper F-5 was installed on the acceptor return line
between the regenerator and the gasifier to allow addition of acceptor to the
gasifier. The capacity of the acceptor charge pot was 1.1 pounds of raw
acceptor.
A sampling tee was inserted in the L-5 acceptor return line
through which a sample of calcined acceptor could be withdrawn while the unit
is operating.
c. Gas Flows
The gasifier inlet gas consisted of recycled product gas with
H2S, H2, C02, and steam flows added to give the desired gas composition.
The metered dry gasifier fluidizing gas was passed through the
steam generator which is shown in Figure 20. The water temperature in the
steam generator was controlled (±0.5°F) to give the desired steam partial pres-
sure. Calibrations showed that at any given water temperature, in the range
of 295 to 370°F, the steam partial pressure was equal to the equilibrium vapor
pressure of water. The steam generator feed water was passed through a Betz
Laboratory Demineralizer.
-------
68,
The fluidizing gas passed through an electrically heated line
from the steam generator to the top of the gasifier. The wall temperature in
the line was kept above 400°F to prevent condensation of the steam. The gas
inside the gasifier vessel passed through a helically-wound preheat coil and
an axial dip tube which extended into a cone at the bottom of the vessel. The
gas passed downward, reversed direction, and fluidized the bed.
The gasifier outlet gas passed through one of two parallel
filters where solids carried over from the gasifier bed were removed using
F-porosity Micro-Metallic filters. Both of the filters and the upstream piping
were heated electrically to prevent condensation of the unreacted steam in the
outlet gas. Separate heating circuits were provided for each of the filters
so that they could be cooled before being removed for cleaning and heated
rapidly after being reinstalled.
The unreacted steam was condensed in the C-l cooler. The con-
densed steam was collected in the condensate receiver F-8 which was mounted
below C-l. A liquid level controller actuated a control valve C-7 to drain
condensate and maintain a constant liquid level in the receiver.
The gas was further cooled in the C-4 cooler using water supplied
at 4O°F from a water chilling unit. Additional condensate was disengaged in
knock-out pot F-9 and collected in the sight glass below F-9.
The dry solids-free product gas was vented to atmospheric pres-
sure through the gasifier back pressure control valve PCV-1.
The regenerator inlet gas consisted of recycled regenerator
product gas with air, C02, and N2 added to give the desired concentrations of
CO2 and CO or air in the product gas.
About two-thirds of the inlet gas entered the top of the regen-
erator and passed through a helically-wound preheat coil and an axial dip tube
which extended into the cone at the bottom of the vessel. The gas passed down-
ward, reversed direction, and fluidized the bed.
The remainder of the regenerator inlet gas consisted of the
acceptor and fuel char carrier gas which entered the top of the regenerator and
passed through the axial solids inlet tube to the bottom of the vessel.
The regenerator outlet gas passed first through an external
cyclone G-l, shown in Figure 21, which removed about 98$ of the fuel char
residue. The remainder of the overhead solids were removed in one of two
parallel filters which used F-porosity Micro-Metallic filter elements.
-------
69,
The regenerator gas was cooled in a cooler, C-6, which was
supplied with water at 40°F from the water chilling unit. The steam formed
by combustion of the hydrogen in the fuel char was condensed and collected in
a sight glass below the cooler. The gas coolers, C-l, C-4, and C-6, are
specified in Figure 22.
The dry, solids-free regenerator product gas was vented to
atmospheric pressure through the back pressure control valve DPCV-1.
Small metered flows of N2 were used to purge the pressure taps
at the tops of both vessels as well as the L-2, L-6, and L-5 standlegs, the
acceptor return line, and the sheath which contained the leadwire for the
char-acceptor interface probe. A small flow of N2 was added to the on-stream
gasifier char feed hopper and a small flow of N2 added to the fuel char feed
hopper to compensate for the solids fed. All of the purge flows were as
small as possible relative to the fluidizing gas flows.
The H2, C02, and N2 used as process gas, pressure shell balance
gas, and lockhopper pressurizing gas were supplied from cylinders. Regulators
downstream of the cylinder manifolds fixed the supply pressures at 300 psig.
Air was supplied from a compressor. Water and oil were removed by a knock-out
drum followed by a bed of silica gel.
Identical single-stage diaphragm compressors (Pressure Products
Industries, Model 1073) were used to recycle gasifier and regenerator product
gas.
d. Miscellaneous
Wherever possible, connections to the process piping were high
pressure coned fittings (Autoclave Engineers or Pressure Products) which were
used because of their ruggedness, ease of disassembly, and freedom from leaks.
Special adapter fittings were required to connect the solids
outlet tubes (3/8 x .035" wall) to the standleg piping (9/16" OD x 5/16" ID).
The solids outlet tubes from the vessels slipped inside the adapter fittings
and were sealed using a lava sealant and packing gland. A standard high-
pressure connection was made between the adapter fitting and the standleg.
In order to minimize the pressure drop across the gas outlet
piping, 3/8 x .028" wall tubing was used downstream of C-l and C-6 coolers in
the gasifier and regenerator, respectively.
The steam condensate contained dissolved NH3, S02, H2S, and C02.
Type 304 stainless steel gave excellent service in piping and equipment hand-
ling process gases at temperatures from ambient to 40O°F. Previous experience
had shown that condensate and wet process gases are extremely corrosive to
copper, brass and plain carbon steel. These materials were not used, even in
instrumentation piping.
Conax packing glands and lava sealants were used where the heater
leadwires, solids outlet tubes, and the vessel skin thermocouples passed through
the pressure shells.
-------
70.
2. Instrumentation and Control
a. Pressure Measurement and Control
The system pressure was controlled at the top of the gasifier
vessel using a Foxboro pressure cell and recorder-controller which operated
the back pressure control valve PCV-1. The pressure at the top of the
regenerator was maintained at a fixed differential (usually in the range of
zero to five inches H20) from that at the top of the gasifier using a Foxboro
d/P cell, and a recorder-controller which operated the regenerator back pres-
sure control valve DPCV-1
The discharge pressures of the two recycle compressors were
controlled using Foxboro indicator controllers which operated control valves
PCV-2 and PCV-3, which vented to the suction side of the compressors.
Identical balance gas systems were used to maintain a zero
pressure differential across the walls of both process vessels. C02 was
supplied to the pressure shell at 300 psig. Barton d/P cells and recorder-
controllers operated control valves DPCV-11A and 11B and DPCV-13A and 13B which
either loaded the shell of vented CO2 to the atmosphere to maintain a zero
pressure differential between the pressure shell and the process vessel.
The unit was adequately instrumented to monitor pressure drops
across the various transfer lines, standlegs, and fluidized beds. These
pressure drops were measured with respect to the pressures at the tops of the
gasifier or regenerator vessels, and were continuously recorded.
The pressure drops across the solids filters in both the gasi-
fier and regenerator outlet gas piping were continuously recorded. The filters
were switched when the pressure drop across the on-stream filter reached 75"
of H20.
Panel mounted pressure gauges were used to indicate the pres-
sures at the top of both vessels, the pressure shell pressures, the compressor
discharge pressures, the char and fuel char feed hopper pressures, and the
pressures immediately upstream of the back pressure control valves. Locally
mounted pressure gauges were used to indicate the pressures in the steam
generator and the various lockhoppers.
b. Temperature Measurement and Control
The bed temperatures in both the gasifier and regenerator were
measured using chromel-alumel thermocouples contained in wells which were
immersed in the fluidized beds. Two thermocouples were used in the gasifier
thermowell and one in the regenerator.
The bed temperatures were controlled manually by adjusting the
Powerstats which controlled the electrical heat input from the heaters.
In each vessel, the wall temperatures corresponding to each of
the heating zones were measured using thermocouples which were inserted through
the side of the pressure shell and through slots ground in the heater ceramics.
-------
71.
The water temperature in the steam generator was monitored using
an Iron-Constantan thermocouple contained in an axial thermowell. The thermo-
couple signal was recorded using a Foxboro temperature recorder-controller. A
high or low temperature activated relays which supplied power to the steam
generator bayonet heater from the appropriate one of two 7.5 KW Powerstats
which were set by trial to supply slightly more and slightly less than the
required steady-state input.
A total of 50 temperatures were monitored of which 31 were con-
tinuously recorded. The other temperatures included the wall temperatures of
the pressure shells, the transfer lines, the inlet and outlet gas piping, the
steam generator, the gasifier solids filters, the regenerator cyclone, and the
ash receiver. Thermocouples inserted in the sight glasses in the various
standlegs monitored the temperatures of the solids flowing from the water-
jacketed coolers.
c. Control of Solids Flow Rates
The solids flow rates were controlled using rotary feeders which
consisted of a tapered Teflon plug rotating inside a stainless steel body.
Drawings of the feeders used in the continuous unit are shown in Figures 23,
24, and 25. The number, diameter, and depth of the pockets were chosen on the
basis of the bulk density of the solid and the desired range of solids flow
rates. Variation of the solids flow rates over the desired range was accom-
plished by means of a Graham variable speed drive. Slots were milled in the
feeder body around the inlet and outlet ports to allow the pockets to begin to
fill and discharge before reaching top and bottom dead center. Since Teflon
was used, no lubrication was required between the feeder body and rotor.
The solids feeders supplied an essentially continuous flow of
solids over the desired range of feed rates. The solids feed rate was linear
with rotor rpm up to the maximum speed of 12 rpm.
A rotary feeder could not be used to feed calcined acceptor.
Due to lack of physical strength, this material would be severely attrited by
the shearing action of the rotor turning in the stainless steel body. The
flow of calcined acceptor was controlled using a Hoke ball valve having a
5/16" bore.
d. Solids Level Control
To control the level of the gasifier char-acceptor interface, a
curved type 310 stainless steel plate, 2" long x 1-1/2" wide x 1/16" thick, was
positioned as shown in Figure 14. The plate was supported by a 1/16" diameter
310 stainless steel leadwire which entered the top of the gasifier vessel. The
leadwire was insulated from the char bed by a 1/4" OD x .028" wall type 31O
stainless steel sheath, which was purged with N2. Alumina beads, .068" ID x
.170" long, were spaced at 2" intervals along the leadwire to prevent its short-
ing against the sheath.
-------
72.
The dc circuit shown in Figure 25 was connected across the
plate and the vessel walls and internals. The circuit was designed to give
a 50 millivolt signal when the plate was covered completely by the electri-
cally conducting char, and essentially zero signal when the plate was
covered by the non-conducting acceptor. The L-6 valve, which fixed the rate
of acceptor withdrawal from the acceptor layer, was operated at one of three
settings to control the interface level. The medium setting was chosen to
correspond to the nominal acceptor circulation rate. A higher setting was
used if the interface level was increasing and a lower setting was used if
the interface level was decreasing. The best control occurred when the lower
setting was zero.
Two rotax switches were attached to the pen arm of the con-
troller, which recorded the circuit signal. They were activated at pen
positions corresponding to 15 and 35 millivolt signals. Upon being activated,
a switch would open and close (or close and open) the appropriate two of
three solenoid valves which supplied air at a preset pressure in the 3 to 15
psig range to the Conoflow Cylinder Operator used to position the control
shaft of the valve.
To detect the level of calcined acceptor in the standleg above
the L-5 valve, a thermocouple was installed in the adapter section between
the bottom of the regenerator pressure shell and the water-jacketed standleg.
The thermocouple extended 1/16" into the 5/16" bore of the adapter. When hot
acceptor, which had been discharged from the regenerator, immersed the thermo-
couple; a temperature increase occurred. The temperature controller which
read out the thermocouple signal opened the L-5 valve, allowed acceptor to
flow into the gasifier and thereby lowered the acceptor level in the standleg.
When the thermocouple was uncovered, the temperature fell off and the tempera-
ture controller reduced the L-5 valve opening to stop the flow of acceptor
until the thermocouple again was immersed in the hot acceptor which overflowed
the regenerator weir.
The temperature controller opened and closed (or closed and
opened) two solenoid valves which supplied air at preset pressures, in the
3 to 15 psig range, to the Conoflow Cylinder Operator used to position the L-5
valve. When the temperature was below the control index, the L-5 valve opening
was set so that purge gas would flow through the valve but acceptor would not.
When temperature was above the index, the L-5 valve was opened to allow acceptor
to flow through the valve at a rate somewhat higher than the nominal acceptor
circulation rate.
The operation of the L-5 valve was monitored continuously by
recording the air pressure supply to the Conoflow Cylinder Operator.
Since acceptor in this standleg acts as a seal between the gasi-
fier and regenerator vessels, an auxiliary level control system was used to
over-ride the normal control system and close L-5 valve should the level drop
to the level of the sight glass. A photocell circuit was used to detect the
interface level. A rotax switch was attached to the pen arm of the controller
which recorded the output of the circuit. When the circuit signal reached 35
millivolts, the switch activated closing the open solenoid valve in the normal
control system and opened a third solenoid valve which vented the air supply to
the Conoflow Cylinder Operator and 'closed the L-5 valve.
-------
73.
e . Gas Flow Measurements
The inlet gas flows were metered using calibrated rotameters
which were calibrated with either a Rockwell dry gas meter or a wet test meter
depending on the range of the rotameter. The gas meters were calibrated by
passing known weights of C02 through them.
Rotameters using recycle gas were calibrated using two gases
having different molecular weights. The flow rates were correlated with the
rotameter float heights using the following general equation;
Where: Q = flow rate, SCFH .
k = constant for a given float height.
Mr = molecular weight of the recycle gas .
n = constant for a given rotameter at a given
supply pressure.
The constant n varied between .455 and .500 for five rotameters
calibrated at supply pressures of 400 and 250 psig.
A plot of k versus float height was made for each of the rota-
meters at both supply pressures .
The dry product gas flow rates from both vessels were measured
using calibrated Rockwell dry gas meters located downstream of the back pres-
sure control valves .
f . Gas Analyzers
The product gas streams from both vessels were analyzed using
a gas chromatograph. During a run, low levels of CO, C02, and SO2 were con-
tinuously monitored with infrared analyzers. High levels of CO2 were monitored
with a thermal conductivity cell with reference to nitrogen. A paramagnetic
susceptibility analyzer was used to measure the concentration of oxygen. The
outputs of the CO, C02, S02, 02 analyzers and of the thermal conductivity cell
all were continuously recorded. Analyses were made on the gas chromatograph
when the above analyzers indicated that the gas composition was steady.
All of the gas analysis instruments were connected to three-way
valves so that offgas from either the regenerator or gasifier could be fed to
each. For integrated operation the chromatograph analyzed both the regenerator
and the gasifier of f gases . The infrared C02 analyzer monitored the gasifier,
and the remaining instruments checked the regenerator offgas. In addition, gas
samples were taken into a stainless steel cylinder for intensive chromato-
graphic analysis by the analytical laboratory.
-------
74,
3. Safety Features
The unit was equipped to give an audible alarm whenever a potentially
dangerous upset occurred. A light indicating the source of the upset accompa-
nied the alarm. In some situations, relays were activated which closed the
emergency solenoid valve to isolate the gasifier and regenerator vessels and
shut off appropriate solids feeders.
The possibility of an explosion exists if cross-contamination of
the gasifier and regenerator process gases occurred. During normal operations,
cross-contamination was prevented by maintaining the N2 purged standleg above
the L-5 valve full of acceptor and by controlling the pressure differential
between the tops of the gasifier and regenerator vessels such that the pressure
drop across the L-5 valve was zero. The presence of solids in the L-5 standleg
was guaranteed by the photocell circuit in the standleg level control system
which closed the L-5 valve when an acceptor interface appeared in the sight
glass. If an upset occurred such that the pressure differential between the
vessels deviated ± 10" H20 from the set valve or the pressure drop across the
L-5 valve deviated ± 5" H20 from zero, an alarm sounded and the emergency
solenoid valve closed, thereby isolating the gasifier and regenerator vessels.
Recycle gas, which was used to transport the gasifier char feedstock
and the acceptor to the regenerator, was a major constituent of the fluidizing
gases of both vessels, and was vital for operation of the unit. Pressure
switches were installed to sound an alarm, close the emergency solenoid valve,
and shut off the L-4 feeder if the discharge pressure of either recycle gas
compressor fell below a preset valve. The L-l feeder was also shut off if the
gasifier recycle compressor discharge pressure was low.
The supply of process air and feed of fuel char and acceptor, which
were vital to the control of the regenerator temperature, were monitored using
alarms. High and low AP alarms were installed to detect plugging or starvation
in either the acceptor or fuel char transfer lines. Other alarms warned of low
process air pressure and high oxygen concentration in the regenerator product
gas.
The unit was equipped with a high AP alarm to detect plugging in the
gasifier inlet piping. Another alarm sounded in the event of malfunctioning
of the gasifier back pressure valve or of plugging in the gasifier outlet
piping, as indicated by a high pressure at the top of the vessel. Other alarms
warned of malfunction of the liquid discharge valve below the condensate
receiver, as indicated by a high or low condensate level.
Failure of the balance gas system to control pressure differential
across the walls of either process vessel could have resulted in collapse or
rupture of the vessel. An alarm was provided to indicate any deviation in the
pressure differential between either process vessel and its pressure shell
greater than 25" of water and another alarm was used to indicate a low
pressure in the C02 cylinder manifold which supplied the balance gas.
Nitrogen was available to supply pressure to the instruments if the
supply of instrument.air were lost, as shown by an alarm indicating low instru-
ment air supply pressure. Cooling water was available from a cooling tower if
the city water supply were cut off, as indicated by a low pressure alarm.
-------
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101,
E. Equipment and Procedure for Single-Vessel Operation
Only the reactor, vessel D-l in Figure 12, was used for single-
vessel operation. A simplified flow diagram for this system is shown in
Figure 13. The system was used for either preoxidation of coal or for gasi-
fication of the preoxidized coal product. A bed of noncaking start-up char
was fed to the reactor in both instances.
1. Preoxidation
Coal was fed through the L-6 acceptor outlet line. A thermocouple
was installed in the unused overhead char inlet line, and the tar receivers
were connected to the system. All of the analyzers were switched to the
gasifier.
Preoxidizer outlet gas generally contained some heavy tar and pitch
which would rapidly plug the solids filters and the outlet piping. This was
alleviated by installing two parallel tar receivers or knock-out pots to
collect the heavy material as shown in Figure 13. The upstream piping was
heated to prevent condensation of the tar and pitch. The tar receivers were
water cooled to ensure condensation of heavy tars; some water was inevitably
also condensed. The tar receivers were removed for gasification runs.
To begin preoxidation the bed was brought to temperature with only
nitrogen being fed. The coal feed and air were both begun at a reduced rate.
The nitrogen flow was diminished in accordance with the increasing air feed.
At the same time the electrical heat input was reduced and the temperature
was stabilized at the programmed level. Since only a fraction of the final
air was being fed, severe temperature excursions were avoided. The final
feed rates of coal and air were approached in several steps.
The run was shut down by first turning off the coal feed. The air
flow was kept at half rate for 15 minutes while the bed was cooling. This
was to prevent possible agglomeration from occurring in the hot bed in the
absence of air. Air was then shut off and the bed was drained under nitrogen.
2. Gasification
Preparation of the vessel for gasification was the same as for
preoxidation except that the tar receivers were not used. All of the single-
vessel runs were used to test the resistance of preoxidized coal to agglomera-
tion and ash slagging, so preoxidized coal was fed from the bottom in an air-
containing stream in order to simulate the bottom of the process vessel where
the threat of agglomeration or ash slagging would be the greatest.
Gasification conditions were such that the start-up bed completely
reacted with any air being fed. Consequently, the programmed air and gas
flows were used while the char bed was brought up to temperature. When the
start-up char feed hopper ran empty, the programmed feed was switched in and
the run began.
Sampling and shutdown were the same as for gasification in the
integrated operation.
-------
102.
F. Sample Calculations
1. Preoxidation Runs
The procedure is described below using Run 5P2 as a typical example.
a. Temperature Profile
The preoxidizer bed temperature was continuously measured by
three thermocouples at positions corresponding to 10", 16" and 33" above the
bottom of the thermowell. Axial bed traverses were taken hourly, and typical
traverses are shown in Figure 28. The nomincal bed temperature was taken as
the flat zone temperature. In Run 5P2 this was 800°F.
The observed profile is not ideal, but the heat of reaction is
of the order of 220,000 Btu/lb-mol, and the fluid bed did not completely
smooth out this considerable heat release.
b. Solids Rate - Run 5P2
Feed Rate
In Run 5P2, 52.34 Ib of coal were fed over a 6.0 hour
period. The coal contained 6.8O$ moisture, so the average feed
rate was:
. • Dry coal 8.13 Ib/hr
Moisture .592 Ib/hr
Total 8.72 Ib/hr
The feed rates of the various components of the coal were determined
using the ultimate analysis of the coal (dry basis):
Ultimate Analysis Input Rate
Component Wt % Dry Basis Ib/hr
Carbon 6O.44 4.91
Hydrogen 4.19 .341
Nitrogen 1.17 .095
Sulfur 4.38 .356
Ash 20.93 1.70
Oxygen (by diff.) 8.88 .722
Volatile Matter 37.21
Preoxidized Coal Product
In a five-hour period 34.07 Ib of product preoxidized coal
were collected. The hourly rate was 6.81 Ib/hr and the product was
dry when removed and weighed. The feed rates of the various compo-
nents were determined using the ultimate analysis of the product
(dry basis):
Ultimate Analysis Exit Rate
Component Wt H> Dry Basis Ib/hr
Carbon 67.03 4.56
Hydrogen 3.29 .224
Nitrogen 1.44 .098
Sulfur 3.72 .253
Ash 16.55 1.13
Oxygen (by diff.) 7.98 .543
Volatile Matter 21.55
-------
-------
104,
c. Overhead Fines
Overhead fines were collected in two places; the tar receiver,
where it stuck to the trapped tarry material in the stainless steel mesh, and
in the solids filters. Very little tar, on a weight basis, was actually
observed in the tar receiver, and the non-aqueous material was counted as all
overhead fines. In Run 5P2, 0.267 Ib of solids were removed from the tar
receiver and 0.055 Ib from the solids filter at the end of the run. This
was equivalent to 0.249 Ib during the 6 hour balance period or 0.041 Ib/hr.
The overhead fines were assumed to have the same ultimate analysis as the
product coal.
d. Condensate and Tar
Both tar and condensate were collected from each cooler. Over
a 6.47 hour balance period the following results were obtained:
Tar + Water 9.51 Ib
Tar l.OO Ib = 0.155 Ib/hr
Water 8.51 Ib = 1.315 Ib/hr
Of the 1.315 Ib of total condensate, 0.593 were fed in with the coal as moisture
and 0.722 were the products of oxidation.
The components exiting with the tar were;
Ultimate Analysis Exit Rate
Component wt jo Ib/hr
Carbon 78.42 0.121
Hydrogen 9.25 0.0143
Nitrogen 0.25
-------
105.
3.36
.46
.38
.14
.02
.06
.01
.05
.18
95.34
.6267
.0546
.0258
.0178
.0023
.0112
.0017
.0127
.0489
.1710
.0234
.0193
.0142
.0020
.0092
.0015
.0024
—
Component Mol % Ib/hr C, Ib/hr 0. Ib/hr
C02 3.36 .6267 .1710 .4557
CO .46 .0546 .0234 .0312
CH4
C2H4
C3H8
C3H6
COS .05 .O127 .0024 .0034
S02 .18 .0489 — .0244
N2 (by diff.)
The nitrogen purges above the bed were:
.9534 x 164.4 - 146.7 = 156.7 - 146.7 = 10 SCFH
The total purges were, therefore; = 15 SCFH
f. Bed Conditions - Run 5P2
The total flow just above the bed consists of the recycle flow,
the exit gas flow, and the product steam flow less the purges injected above
the bed. For run 5P2 this was:
380 + 164
The general equation for the fluidizing velocity is as follows:
_ Q 144 1. T
V ~ 3600 X A X TT X 530
or
V = 7.55 x Kf5 ^— (10)
where:
V = fluidizing velocity, ft/sec.
Q = gas flow rate, SCFH.
A = cross-sectional area of vessel, in.2
T = bed temperature, °R.
TT = system pressure, atm.
The cross-sectional area of the reactor was 12.20 in.2. Thus,
V.T.».IO-. iXrrff • °-29"/—
-------
106,
g. Inlet Oxygen Pressure
This is equal to the fraction of oxygen in the inlet gas times
the pressure. The recycle flow is included in the inlet gas flow. The inlet
oxygen pressure in Run 5P2 was:
13-34 X 15 = 0.32 atm.
146.7 + 11.34 + 380
h. Extent of Preoxidation
The extent of preoxidation is expressed as pounds of oxygen
reacted per pound of dry coal fed times 1OO. For Run 5P2,
% Preoxidation = 0.938 x 100/8.13 = 11.5$.
i. Material Balances
Carbon Balance
Carbon in on coal .6O44 x 8.13 =4.91 Ib/hr.
Carbon in exit streams:
Lb./hr Percent
Preoxidized Coal 4.55 92.87
Overhead Fines .027 .55
Tar .122 2.48
CO .023 .48
C02 .171 3.48
COS .002 .04
Hydrocarbons (Cj-Cg) .046 .94
100.84
Oxygen Balance
The oxygen balance showed more oxygen in the products
than in the entering air. The difference was assumed to
come from the coal.
The exit gas was assumed to be saturated at 50°F
and 15 atm. Then, in the exit gas:
1Y8
SCFH H20 = ' x 164.4 = O.133, or O.OO6 Ib/hr.
-------
107,
Lb/hr
To Water
Condensate
Moisture in gas
To C02
To CO
To SO 2
To COS
In Tar
Total
The sources of the oxygen were:
Percent
54.7
0.4
38.9
2.7
2.1
0.3
0.9
100.0
From inlet air .938 Ib/hr
From coal (by diff.) .234 Ib/hr
Coal Balance
Percent
Coal In
Preoxidized Coal Out
Overhead Fines
Tar
Carbon in CO + C02
Hydrogen in H20
Hydrocarbons (Cj-Cg)
Sulfur in S02 and COS
Coal Oxygen to Products
Accumulation in Bed (by diff.)
Total
1.
2.
1,
83.76
.50
.91
.39
.00
.72
.38
2.88
6.46
100.OO
In Run 5P2, the carbon balance was 1OO$, but the coal balance was
only 94$. This was due to the accumulation of dense rocky material in the
bed. In Run 6P, the pressure drop across the fluidized bed increased continu-
ously. Microscopic examination and chemical analysis of the final bed
material drained from the reactor showed that ash particles had accumulated
in the bed during the run. Ash and pyritic sulfur balances on the product
coal and bed material showed that 1O.55$ of the dry feed coal had accumulated
in the bed, compared with 10.29$ as obtained by difference in the overall
weight balance. A check of the chart records for Run 5P2 also showed an
increase in pressure drop across the bed, but to a lesser extent than in Run
6P. The accumulation in Run 5P2 was obtained by difference in the overall
weight balance, in view of the good agreement between the measured and
difference values shown in Run 6P.
-------
108.
2. Gasification Runs
Run 5G3 is used as an example to illustrate the procedure for gasi-
fication runs .
a. Temperature Profile
During the run, the temperature was continuously measured by
three thermocouples placed at 10", 16" and 33" in the thermowell. Traverses
were taken hourly, and typical traverses are shown in Figure 29. The nominal
run temperature was taken at the flat zone. For Run 5G3 this corresponded to
170O°F after correction for thermocouple error.
b. Solids Rates
Preoxidized Coal Feed
The feedstock was Run 5P2 composite product, whose
analysis was given earlier. During the final balance
period, 13 Ib were fed in 3.13 hours, or 4.15 Ib/hr.
Product Char
The last timed product collection yielded 1.51 Ib/hr of
product char. This was analyzed to contain 65.64$ carbon,
O.46$ hydrogen and no oxygen.
c. Gas Rates
Inlet Gas Flows
No recycle flows were used. The inlet flows were as
follows:
SCFH
Fluidizing
C02 38
N2 64
Lift Gas
Air 1O5
N2 34
N2 purges to bed 5
N2 purges above bed 6
The fluidizing gas bubbled through a steam generator
whose temperature was held at 325°F. At equilibrium, the
fluidizing gas would have contained 76 SCFH of steam.
Conditions matched those at the bottom of the bed as speci-
fied in the process design.
-------
-------
110.
Outlet Gas Flows
The outlet gas rate was measured to be 334 SCFH. Con-
centrations of CH4, CO and C02 were measured by a gas
chromatograph . No analysis was available for the hydrogen
concentration. The condensate rate was measured as 2.70
Ib/hr . Extensive experience has shown that shift equili-
brium is always achieved in the exit gas. The hydrogen
concentration was determined from the water gas shift
equilibrium, then the inlet steam rate was calculated by
forced hydrogen balance . Exit nitrogen was determined by
difference. The chromatograph gave the following results
for the exit gas:
Component
CH4
CO
C02
d. Hydrogen Balance and Concentration
In the dry outlet gas, we have:
Component Mol % SCFH
CH4
CO
CO,
Mol TO
.86
13.81
14.97
SCFH C
SCFH
.86
13.81
14.97
2.88
46.10
49.95
2.88
46.10
49.95
5.76
98.93
Outlet H2O = 2.70 Ib/hr = 57.95 SCFH.
The total flow just above the bed = dry exit flow -f- steam
- purges above the bed. Thus, the total flow above the bed
was:
333.7 +57.95-6.0 = 385.6 SCFH.
On a wet basis just above the bed we get:
Component Mol ^
CH4
CO
C02
H20
.75
11.96
12.95
15.03
At 17OO°F, the equilibrium constant for the water gas shift
reaction is .6973. Then,
/„ \
(H2) =
_ (Ha)(COa)
- (H20)(CO)
.6973 x 15.03 x 11.96
- 12~95 - =
or 37.29 SCFH.
-------
Ill,
The total hydrogen leaving was:
SCFH
as H2
in CH4
in H20
in Char
.0046 x 1.51 x 387
2.016
37.29
5.76
57.95
1.33
Total H2 leaving 102.33
The hydrogen entering with the feed coal is:
.0329 x 4.1'5 x 387 = 2(J
2.016
The inlet steam rate is therefore 76.12 SCFH
The exit compositions are summarized below:
Component
H2
CH4
CO
C02
N2 (by diff.)
H20
Mol % (dry)
11.17
.86
13.81
14.97
57.54
Mol % (w<
9.67
.75
11.96
12.95
49.64
15.03
e. Outlet Fluidizing Velocity
The gasifier was the same vessel as the preoxidizer, and as
discussed in Section IV-E-l-e, the outlet fluidizing velocity was;
7.55 x 10~5 x 385.6 x 2160 _„ ^ /
V = 12.20 x 15 = -34 ft/S6C
f. Gasification Rate
Carbon fed = 4.15 x .6703 = 2.78 Ib/hr.
Carbon gasified = (98'93 " 38-°0) 12 = 1.89 Ib/hr.
OoY
The weight of the bed removed from the gasifier at the end
of the run was 4.56 Ib. The gasification rate is defined as 104 x Ib carbon
gasified/min/lb carbon in bed.
1.89 x 104
Gasification rate = 4.56 x .6564 x 60 = 105
Percent carbon burnoff = 100 x 1.89/2.78 = 68.
-------
V. RESULTS &
DISCUSSION
-------
112,
V. RESULTS AND DISCUSSION
A. Tabular Chronological History of Runs
The purpose, general run conditions, and major conclusions
drawn for all the runs are given in the tables listed below. All
twenty runs were made at a pressure of 15 atmospheres.
Table No.
28 Preoxidation Runs
29 Gasification Runs
30 Integrated Gasifier-Regenerator Runs
-------
TABLE 28
Run Number
Unit Revisions
Date of Run
Run Duration, hours
Purpose of Run
Temperature, °F
Feedstock
Shutdowns
Results and Conclusions
Summary of Preoxidizer Operations
IP
None
5/7/71
11
High Pressure Preoxidation
of Ireland Mine coal
7OO
Voluntary
1. Some caking occurs during
preoxidation at 70O°F.
2. Product probably not
operable in gasifier.
3. No oxygen breakthrough.
4. 19.5$ preoxidation.
2P1
None
5/20/71
2.0
Preoxidize at 60O°F to
avoid caking during
preoxidation
600
Ireland Mine Coal
Rapid increase in
temperature to 8OO°F
1. Distortion of temp-
erature profile
indicated agglomeration
probably occurred.
2. 19$ preoxidation.
2P2
None
5/25-5/26/71
19.0
Preoxidize at 75O°F
75O
Voluntary
1. No agglomeration or
caking observed.
2. Startup procedure
changed to get oxida-
tion of start up char
to occur before adding
coal.
3. 18.6$ preoxidation.
-------
TABLE 28
(Continued)
Summary of Preoxidizer Operations
Run Number
Unit Revisions
3P
3P1
4P
4P1
Date of Run
Run Duration, hours
Purpose of Run
Temperature, °F
Feedstock
Shutdowns
Results & Conclusions
Added pitch and heavy
tar receiver for use
in preoxidation.
Replaced L-2 feeder
system with ball valve.
6/3/71
1.0
Preoxidize at 8OO°F to
improve operability of
product in the gasifier.
None
6/3-6/4/71
2.0
Repeat 3P approach-
ing 8OO°F more
slowly.
8OO
Ireland Mine Coal
Distortion of tempera-
ture profile indicated
abnormal mixing pattern
Bed caked
1. Large agglomerates
found in reactor.
2. 20$ preoxidation.
1. Bed filled with
agglomerates.
8OO°F is not an
operable preoxida-
tion temperature.
2. 2O% preoxidation.
None
None
6/4/71
6/4/71
3.75
Find preoxidation Effect of 8OO°F
conditions which preoxidation.
would make Ireland
Mine coal operable
in gasifier.
75O 8OO
•« Run 2P2 Composite Product
Change to 4P1
conditions
1. Product pre-
oxidized to
28$ total.
Very low density
material would not
flow reliably
through overflow weir.
Reactor filled and
plugged gas exit line.
1. No agglomerates
formed, but product
had a very low
density.
-------
Run Number
Unit Revisions
Date of Run
Run Duration, hours
Purpose of Run
Temperature, °F
Feedstock
Shut downs
5P1
TABLE 28
Continued]
Summary of Preoxidizer Operations
5P2
None
8/2/71
9.25
Preoxidize Hillsboro
coal at 750°F.
75O
Hillsboro Coal
Plug in coal feed line.
Results and Conclusions 1. No operability
problems encountered.
2. 11.9$ preoxidation.
None
8/3-8/4/71
6.0
Preoxidize Hillsboro
coal at 8OO°F.
8OO
Hillsboro Coal
Could not switch tar
receivers due to plug
in receiver inlet line.
1. No operability
problems encountered.
2. 11.5$ preoxidation.
6P
Installed bellows on
gasifier tailpieces.
8/11/71
5.5
Preoxidize Hillsboro coal
at process conditions;
8.5$ preoxidation, 81O°F.
810
Hillsboro Coal
Voluntary.
1. No operability problems
encountered.
2. 8.7$ preoxidation.
en
-------
TABLE 29
Run Number
Unit Revisions
Date of Run
Run Duration, hours
Purpose of Run
Feedstock
Shutdowns
Results and Conclusions
OG2
None
5/13/71
9.5
Gasification at O. 43
atm.} O2 inlet
Summary of Gasifier Operations
All runs at 17OO°F
OG3
Switch to OG3
Conditions
None
5/13/71
2.0
Gasification at 1.04
atm., 02 inlet
OG4
Char and air inlet
raised into bed
5/13/71
9.0
Gasification at l.O
atm., O2 inlet
-Partially Gasified Disco Char
Voluntary
No evidence of ash slagging—
Switch to OG5
Conditions
OG5
None
5/13/71
2.0
Gasification at 2.4
atm.} O2 inlet
Voluntary
No evidence of ash slagging —
-------
(Continued)
Summary of Gasifier Operations
All runs at 1700°F
Run Number
Unit Revisions
Date of Run
Run Duration, hours
Purpose of Run
Feedstock
Shutdowns
1G3
2G3
4G3
5G3
6G3
Char and Air Inlet
to Gasifier Bottom
5/19/71
0.22
Test IP Product for
operability at OG3
conditions.
Run IP composite
product
Bed caked
Results and conclusions 1. Run IP product is
not operable at
OG3 conditions.
None
5/27/71
0.75
Test 2P2 product
for operability
at OG3 conditions.
Run 2P2 composite
product
Voluntary
1. No caking, but
significant
agglomerat ion
occurred.
None
6/7/71
2.5
Test 4P product at
OG3 conditions.
4P product
Leak in char exit
line
1. No caking
2. No ash slagging
None
5P2 product
Voluntary
None
8/5/71
8.0
Test operability
of 5P2 product
in gasifier.
8/13/71
7.6
i
Test 6P1
product in
gasifier.
6P product
Voluntary
1. No agglomerates
2. No ash slagging
-------
118.
TABLE 30
Results of Integrated Gasifier-Regenerator Operation
Run Number
Unit Revisions
Date of Run
Purpose of Run
Temperature:
Gasifier
Regenerator
Feedstock:
Gasifier
Regenerator
Shutdowns:
Results and
Conclusions
Dl
Installed acceptor sample line between
regenerator and gasifier
8/23/71
and 8/30/71
Demonstrate integrated operation of
regenerator-gasifier at process conditions
1700° F
186O°F
Disco Char
Tymochtee 9 dolomite
Pregasified Disco Char
Cycle 5 - Replaced corroded sample valve,
inspected vessels. Delayed
resumption to attend OAP contractors
meeting.
Cycle 13 - Regenerator thermowell failed due to
localized sulfur attack.
1. No ash slagging in either vessel.
2. No agglomeration or transient liquid
formation in regenerator.
3. Reaction of C02 with carbon in the regenerator
may be slower than predicted.
-------
119.
B. Preoxidation Runs
1. Introduction and Tabulated Data
A major aspect of the adaptation of the C02 Acceptor Process
for making low-sulfur boiler fuel gas is the use of Eastern steam coals,
These coals become fluid when heated through the temperature range of
roughly 700-850°F, and therefore require pretreatment by preoxidation
to prevent caking at gasification conditions.
Numerical data for run conditions and results of all operable
preoxidation runs are given in Table 31. Analytical data on the
properties of the coal feed and products are given in Table 32.
Material balance data and distribution of oxygen in the products are
given in Tables 33 and 34, respectively.
Other tabulated data specific to the topic being discussed
are presented in the appropriate succeeding sections.
Typical relationships among the intensive variables,
severity of preoxidation and inlet oxygen partial pressure, and the
extensive variables, coal feed rate and nominal residence time are
shown in Figure 30. The data shown were obtained from straightforward
material balance calculations, assuming all the inlet oxygen reacts.
Specific conditions used in the calculations are shown in Figure 30.
The reactor used during the runs has an inside diameter of 4.O inches,
with a bed volume of 0.25 ft3.
-------
TABIJB 31
Contltlfl
Systow. Pressure: IS atmospheres (206 palg)
Size Consist : 24 z 1OO Tyler Mesh
IP
2F2
COBl
Tenperatnre, *F
Inlet O, Partial Pleasure, at»
Fluldlxlng Velocity (top of bed), ft/*i
Bolature In Coal, Wt f, aa fed
Input
Coal Feed Bate, Ib/hr (dry basis)
Lift Gas. SCPH
Air
B,
Plnldlzinc Gas
-Ireland Bine-
Recycle
»t
Purges (Ha), SCFH to bed
Purees (S,), SCFH above bed
Exit Gas Bate, SCFHVm)(dry basis)
Exit Gas Opposition, «ol £ (dry basis)
O,
CO,
CO
CH«
C,H.
TOO
.24
1.58
6.5
73 .
84
310
O
S
15.5
169
<.oa
3. S3
1.20
TSO
.40
.34
1.58
7.46
80
TO
390
92
S
15.5
256
<.02
2.56
.85
2P2 Product
TSO
.32
1.03
9.34
54
96
380
O
5
15.5
1S9
C.H.
COS
so,
H,S
B,
S, (by difference)
Plow Bate, SCFH, at top of bed
Water. Ib/hr (corrected for coal Moisture)
Condense te
Bo is tore la Exit Gas
Tar, Ib/hr
Preoxldized Coal, Ib/hr
Percent Preoxidation
Boailnal Solids Retention Time, min
Duration of Steady-State Period, nr
Total Product, Pounds
9S.3
500
.665
.006
e-S
19.5
(•).
11
TO
96.4
CC9
.729
.O1O
.0149
6.96
18.6
5T
19
132
(i) Standard conditions are 29.92 In Hg; and TCP?.
(*) Uncertain because of agglomerate in reactor.
(a) Sot neasured.
(•) 9.3% based on original raw coal.
(») Sot measured. Assumed to be the sane as in Rim SP2.
SP1
750
0.31
0.29
6.80
7.89
54. 0
97
396
O
5
10
SP2
BOO
0.32
0.29
6.8O
8.13
54.0
99
380
O
5
9
6P
81O
0.25
0.29
6.58
8.71
43.5
110
382
O
S
6
161
<.02
3.33
.89
.21
- —
~
--(")
--
--
—
— .
9S.6
552
.265
.006
0
9.14
<.
3.
.
.
.
.
.
.
.
01
03
4O
25
08
01
Ol
01
O2
16
<.OO3
96.03
572
.590
.006
.0934
7.
2O
lO.o(4) 11.9
49
3.75
34
59
1O.O
72
164
3.36
.46
.38
.14
.02
.06
.Ol
.05
.18
<.OO3
95.34
563
.722
.OO6
.155
6.81
11.5
46
6.0
41
161
2.54
.25
.45
(')
".09J
96.39
569
.824
.006
.292
6.76
8.7
39
5.5
37
CO
O
-------
TABLE 32
Run Number
Hydrogen, Wt % (dry basis)
Carbon
Nitrogen
Oxygen (by difference)
Sulfur
Ash
Volatile Matter
Pyritic Sulfur
Sulfate Sulfur
Size, Tvler Mesh
+24
24 x 28
28 x 35
35 x 48
48 x 65
65 x 10O
-100
Mean Diameter, in(2)
Mean Density, Ib/ft8(3)
Pittsburgh Seam
Ireland Mine Feed Coal
Wt 1t
0
11.2
35.4
34.8
16.9
1.2
0.5
5.00
72.24
1.35
5.53
4.52
11.36
40.1
2.53
.07
DensitvC1)
81.0
.0165
81.0
IP
4.14
7O.51
1.45
8.75
4.O3
11. 12
32.7
1.76
.08
Wt
-------
TABLE 32
(Continued)
Properties of Coal and Products — Preoxidation Runs
Run Number
Hydrogen, Wt % (dry basis)
Carbon
Nitrogen
Oxygen (by difference)
Sulfur
Ash
Volatile Matter
Pyritic Sulfur
Sulfate Sulfur
Hillsboro Mine
Illinois No. 6 Feed Coal
5P1
5P2
6P
Size. Tyler Mesh
+24
24 x 28
28 x 35
35 x 48
48 x 65
65 x 1OO
-100
Mean Diameter, incm8/
Mean Density, lb/ft*(4)
(0
) 4.19
6O. 44
1.17
8.81
4.46
2O.93
37.2
2. 2O
.07
(2)
4.18
60.14
1.22
7.67
4.93
21.86
36.4
2.26
.30
3.18
61.72
1.34
7.81
4.O7
21.88
23.4
1.51
.33
3.29
67.02
1.44
7.98
3.72
16.55
21.6
1.11
.27
3.44
70.61
1.36
5.8O
3.40
15.39
22.0
1.37
.07
Wt % Density Wt % Density Wt # Density Wt # Density Wt % Density
0
9.6
39.3
34.2
15.0
1.3
0.6
O
15.8
41.1
29.0
12.7
1.0
0.4
.0166
80.0
.0176
80.0
O. 4
7.9
33.3
34.7
20. 2
3.0
0.5
52.3
57. 0
6O.5
77.1
81.2
1.4
14. 0
35.5
32. 0
15.0
1.8
0.4
37.9
40. 1
44.1
58.8
61.5
72.9
2.O)
17.0)
34.5
29.4
14.3
2.2
0.5
5Q.O
5O. 5
49.2
64.3
58.9
.O158
61.7
.0171
49.8
.0175
51.8
(i) Batch No. 1, used in Runs 5P1 and 5P2.
(2) Batch No. 2, used in Run 6P.
(3) Arithmetic mean.
(4) Measured in mercury at 1 atmosphere. Reciprocal mean.
to
to
-------
TABLE 33
Material Balances for Preoxidation Runs
Basis; 1OO Ib Dry Coal Fed
Run Number 2P2 4P 5P1 5P2 6P
Preoxidized Coal 93.3O 97.85 91.25 83.76 77.60
Overhead Fines .58 1.21 .70 .50 .91
Tar .20 O 1.18 1.91 3.35
Carbon in CO + CO2 3.65 2.23 2.16 2.39 1.6O
Hydrogen in H2O 1.1O .33 .84 l.OO 1.O6
Hydrocarbons (C!-C3) .34(0 .IsC1) .40 .72 .72
Sulfur in S02 and COS .75(2) .26(2) .30 .38 .22
Coal Oxygen to Products 0 , O .61 2.88 4.25
Accumulation in Bed,
Ash and S as FeS2 0_ O_ 2.56(3) 6.46(3) 1O.29(3J4)
Total 99.92 1O2.O3 lOO.OO 1OO.OO 1OO.OO
(i) Only CH4 was determined.
(2) By sulfur balance on coal feed and preoxidized coal product.
(s) By difference.
(4) Ash and sulfur balances on product and bed materials gave 10.55 Ib accumulation.
to
CO
-------
TABT.K 34
Distribution of Oxygen in Products of Preoxidation ,
Run Number 2P2 4P 5P1 5P2 6J>
Lb/hr Percent Lb/hr Percent Lb/hr Percent Lb/hr Percent Lb/hr Percent
Output
To Water
Condensate
Moisture in Gas
To C02
To CO
To SO 2
To COS
In Tar
On Coal (by diff)
Total
Input
From Coal (by diff.)
Lb/hr O2 from Air
(i) Not determined.
.648
.009
.546
.090
.056
[CO]
.O4O
1.389
0
1.389
46.6
.6
39.3
6.5
4.0
-
3.0
1OO.O
.235
.006
.437
.058
.024
[(0]
.178
.937
O
.937
25.1
.6
46.7
6.2
2.5
-
18.9
1OO.O
.524
.005
.402
.027
.021
.001
.006
O
.986
.048
.938
53.2
0.5
40.8
2.7
2.1
O.I
O. 6
0
1OO.O
.641
.005
.456
.031
.024
.OO3
.011
0
1.172
.234
.938
54.7
0.4
38.9
2.7
2.1
O.3
0.9
0
100.0
.732
.OO5
.338
.017
.012
.003
.018
O
1.125
.370
.755
65.0
0.4
30.1
1.5
1.1
0.3
1.6
O
100.0
to
-------
125.
-------
126.
2. Pittsburgh Seam Coal (Ireland Mine)
For the initial experimental program with Ireland Mine coal, a
severe level of preoxidation* (20$) was chosen deliberately with the intent
of suppressing completely any tendency toward caking at gasification conditions.
A long nominal retention time (~ 60 minutes) was chosen to insure adequate
exposure to 02 for those particles at the low end of the retention time distri-
bution which is characteristic of a continuously fed fluidized bed. With these
constraints, the inlet 02 partial pressure fell in the range of 0.4-0.5 atm.
This level of 02 partial pressure is representative of that in an upper portion
of the envisioned process preoxidizer in which the inlet 02 partial pressure
is 3.15 atm at 15 atm system pressure.
The first three runs were made at conditions shown below:
Run No. IP 2P1 2P2
Pressure 15 atm (206 psig)
Temperature, °F 700 60O 750
Coal Feed Rate, Ib/hr 6.5 7.1 7.46
Fluidizing Velocity, ft/sec .24 .30 .34
Inlet 02 Partial Pressure, atm 0.5 0.4 0.4
Percent Preoxidation « ~ 20 »
Run IP
Run data and properties of the preoxidized coal product are listed
in Tables 31 and 32. The run proceeded without apparent difficulty, except
for a disturbingly large temperature gradient across the fluidized bed, until
the desired amount of product was obtained. Coal, air and recycle fluidizing
gas were fed to the bottom of the preoxidizer at substantially room tempera-
ture . During the run, the temperature at the bottom typically was 580°F and
the profile showed an increase to 7OO°F at about 25" above the bottom with
constant temperature to the top of the bed at 40".
On disassembly of the reactor, it was found to be partially choked
with large chunks of agglomerated coal. Presence of the agglomerates distorted
the normal mixing pattern of the fluidized bed and thus accounted for the large
temperature gradient. Further, the volume occupied by agglomerated coal
decreased the nominal residence time of the coal particles by an unknown
amount. However, all the inlet 02 had been consumed, i.e., the desired
preoxidation level had been achieved.
A laboratory test (see page 133) showed that the product had retained
much of its caking tendency and that it almost certainly would have caked at
gasification conditions. At the time, this situation was ascribed to the low
nominal retention time which allowed some of the coal to pass through the pre-
oxidizer without reaching the average level of 20$ preoxidation. This material
was set aside temporarily.
Severity of preoxidation is expressed as percent preoxidation:
10O (ib 02 consumed/lb dry coal).
-------
127.
Run 2P1
Since coke had formed at 700°F in Run IP, the next preoxidation run
was made at 600°F at the same 20$ preoxidation level, in a continued effort
to produce an operable feedstock for the gasifier which had had a known
history of retention time in the preoxidizer.
Coals with fluidity characteristics similar to Ireland Mine coal
historically have been fluidized at atmospheric pressure and 600°F in various
Consol pilot plant operations, with and without the presence of 02, without
occurrence of caking or agglomeration.
In Run 2P1, a large temperature gradient developed soon after coal
feeding was-started. After two hours of feeding, the temperature at 16" above
the bottom increased suddenly to 800°F at which time the run was ended.
Inspection of the vessel after disassembly showed no evidence of agglomeration.
Yet, the temperature gradient.and sudden rise, clearly were indicative of dis-
ruption of the fluidized bed mixing pattern as if agglomeration had occurred.
To help explain the behavior the following laboratory tests were made,
Samples of of raw Ireland Mine coal were heated in a thermobalance purged with
helium at atmospheric pressure. The coal reached the desired temperature in
about two minutes and was held for an additional 15 minutes. Results are
shown below:
Observations
Temperature, Particle
^F Rounding Agglomeration Coking
650 No No No
700 Yes No No
750 Yes Yes No
80O — -- Yes
The above run data indicate that no agglomeration should have occurred in
Run 2P1, even at 650°F.
Samples of the raw coal were then heated in a small autoclave purged
with helium at 15 atm (206 psig) with the following results:
Temperature, Particle
°F Rounding Agglomeration
600 No No
63O Yes Yes
Thus, at 15 atm pressure the coal agglomerates at a temperature below that at
which no trace of agglomeration occurs at 1 atm.
-------
128.
In Run 2P1 the particle surface temperature probably reached 630°F
while undergoing preoxidation, with consequent formation of agglomerates.
This behavior at 600°F is attributed to small amounts of waxy and resinous
material which are present in nearly all bituminous coals. This material
melts at relatively low temperatures and evaporates at atmospheric pressure
before agglomerates can form. At elevated pressure, evaporation is hindered
sufficiently so that the particles can stick to each other. The agglomerates
so formed must have been very fragile and were broken up during manipulation
of the reactor on disassembly after Run 2P1.
Run 2P2
The results of Run 2P1 and the laboratory tests showed the probable
cause of agglomeration in Run IP. During startup, the initial bed was an inert
char which had been partially gasified at 1775°F in a previous experimental
program. At the 20$ preoxidation level, heat must be dissipated through the
preoxidizer vessel wall. The start-up procedure was to supply heat electric-
ally to the preoxidizer while feeding the inert char and to bring the wall
temperature to that estimated as sufficient to dissipate the strongly .exo-
thermic heat of reaction when the feed was switched to coal. Generally, a bed-
to-wall AT at 60°F was adequate while preoxidizing coal with no electrical heat
input to the reactor. At the end of the start-up period, the inert char bed
temperature was about 500°F. The char was unreactive toward the inlet 02 at
this temperature. In Run IP the temperature of the upper part of the bed
reached 7OO°F in about one hour after the coal feeding was started, but the
bottom of the bed stayed at about 60O°F. It is possible that agglomerates
formed while the bed was passing through the 60O-630°F range during the heat-
up from 500°F. The resulting poor mixing caused a permanent temperature gradient
to be established in which several inches near the bottom of the bed were always
in the 6OO-630°F range. Therefore, it is also probable that agglomerates
continued to form throughout the run.
In Run 2P2, made at 75O°F, the start-up procedure was revised so that
no coal ever would be exposed to temperatures in the range of 600-630°F. The
inert char bed was brought to about 725°F by means of electrical heat. At
this temperature the char was reactive toward the inlet 02 and relatively little
electrical heat was needed. The electrical heat was turned off and, in quick
succession, the air flow rate was cut to one-half the programmed flow and coal
feeding was started, also at one-half the programmed flow. The flow rates were
increased stepwise to the full programmed rates while keeping the air/coal ratio
constant. During this period, the bed temperature rose to 750°F and subsequently
was held at that level by varying the electrical heat input. The maximum tempera-
ture gradient was about 30°F and the temperature profile was essentially flat
over the range of 16 to 40" above the bottom. The run was ended voluntarily when
the desired amount of product was obtained. Inspection of the vessel after dis-
assembly showed no agglomeration whatever. Run data, some properties of the,
preoxidized coal product, the distribution of reacted 02, and a material balance
are presented in Tables 31 through 34. .
Both Runs IP and 2P2 products were fed to the gasifier at process condi-
tions. The Run IP coal melted and coal feeding had to be stopped after 13
minutes. The Run 2P2 material produced numerous small agglomerates. Examination
of the gasifier product showed that the coal was sticking to itself and to the
Disco char particles in the start-up bed.
-------
129.
It was apparent that at the pressure used in the process, even 20%
preoxidation was ineffective in preventing caking in the gasifier. Pretreat-
ment by preoxidation involves dehydrogenation and devolatilization. Elevated
system pressure inhibits these processes, while increased preoxidation
temperature enhances these same processes. Since a higher preoxidation
temperature can compensate for the suppressing effect of elevated pressure on
dehydrogenation and devolatilization, a preoxidation run at 800°F was scheduled.
Run 3P
This run was made at the same conditions as for Run 2P2, except that
the initial temperature of the start-up bed of Disco char was increased to
775°F by decreasing the heat losses via the electrical heaters. As before,
coal feeding was begun at one-half the programmed flows of coal and air. A
stepwise increase to two-thirds of the programmed flows brought the temperature
to 800°F. Over a period of about one hour the temperature of the middle
portion of the bed increased to 850°F, a positive indication that the normal
mixing pattern of the fluidized bed had been disrupted. •
On inspection of the vessel after disassembly, it was found to be
partially choked with large chunks of agglomerates consisting 6f coal particles
which had adhered to themselves as well as to the Disco char particles.
Run 3P1
In this run the initial temperature of the start-up bed was 725°F,
as in Run 2P2. During stepwise increases of the coal and air flows, the bed
temperature was increased to 800°F by increasing the amount of electrical heat
supplied. On reaching full programmed flows after about two hours of feeding
coal, the pressure drop across the fluidized bed increased suddenly, an in-
dication of a massive blockage.
After disassembly of the vessel, it was found to be nearly full of
agglomerates similar in appearance to that in Run 3P.
Run 4P
To establish a benchmark as to the severity of preoxidation for
24 x 1OO mesh Ireland Mine coal which assuredly would provide an operable feed-
stock for the gasifier, the Run 2P2 product was subjected to a second stage of
preoxidation at 750°F in which an additional 10$ preoxidation occurred.
The run proceeded without incident until the desired amount of product
was obtained. Run data, some properties of the preoxidized coal product, the
distribution of reacted 02, and a material balance are presented in Tables 31
through 34.
The smaller mean particle diameter of the Run 4P product, compared
with the feedstock, was caused by decrepitation of some of the rounded, swollen
particles on passage through the rotary feeder. The high particle density of
the Run 4P product is the result of loss of volatile matter.
-------
130.
As shown in Table 34, during a second stage of preoxidation, much
less water forms, in accordance with previous work, and a considerable amount
of the reacted 02 is retained in the coal.
Run 4P1
In order to obtain a gasifier feedstock which was more highly devola-
tilized, at the end of the 750°F operating period of Run 4P, the bed temperature
was raised to 800°F via the electrical heaters while holding all input flows
constant.
Increased particle swelling was apparent immediately. After about
one hour at 800°F, the very low density material would not flow reliably
through the overflow weir with the result that the reactor filled completely
with coal. The exit gas line then became plugged with solids. Inspection of
the vessel after disassembly showed that no caking or agglomeration had
occurred.
However, such low density material (particle density <30 lb/ft3)
would be impractical as a gasifier feedstock from the standpoint of maintaining
an adequate carbon inventory.
Run 4P product was then fed to the gasifier at the standard process
conditions with no evidence of either caking or ash slagging.
3. Conclusions - Pittsburgh Seam Coal
Work with Ireland Mine coal showed that elevated system pressure has
an adverse effect on operability in the preoxidizer as well as on operability
of the preoxidized coal in the gasifier. Attempts to preoxidize raw coal at
800°F at the severe level of 20% preoxidation were unsuccessful. During a
second-stage preoxidation at 80O°F, an intolerable amount of particle swelling
occurred, although there was no caking or agglomeration.
After two stages of preoxidation at 750°F to a total severity level
of 28$, the Ireland Mine coal was operable at gasifier conditions with no caking
or ash slagging. At this level of preoxidation a very large amount of heat
would have to be removed from the process preoxidizer, with the attendant
economic penalty due to lowered thermal efficiency. For relatively coarse,
24 x 1OO mesh, Ireland Mine coal the work to date has not yielded an economi-
cally satisfactory preoxidation procedure. Further research is required to
find techniques which will decrease the severity of preoxidation to lower and
more economic values.
Work is planned which will investigate the possibilities involved in
using a two temperature staged preoxidation. The magnitude of the economic
penalty involved in operating at high levels of preoxidation will also be
investigated.
Additional areas for study are the use of a small size consist of
the feed coal and consideration of a continuously temperature staged entrained
solids reactor which would adiabatically reach high levels of preoxidation.
Still a final possibility is the use of other pretreatment methods which do
not involve preoxidation.
-------
131,
4. Illinois No. 6 Coal (Hillsboro Mine)
The inability to make Ireland Mine coal noneaking using process
conditions of preoxidation, led to the investigation of the preoxidation
behavior of a less severely caking coal. Three runs were made at the condi-
tions listed below;
Run Number 5P1 5P2 6P
Temperature, °F 75O 800 810
Inlet 02 Partial Pressure, atm 0.31 0.32 0.25
Fluidizing Velocity, ft/sec (top of bed) 0.29 0.29 0.29
Percent Preoxidation 11.9 11.5 8.7
Run data, properties of the preoxidized products, material balances
and distribution of oxygen in the products are given in Tables 31 through 34.
Preoxidation Runs 5P1 and 5P2
Since these runs were the first to be made with Illinois No. 6 coal,
a preoxidation level higher than the adiabatic level was chosen deliberately
(10$ versus 8.5$).
Upon completion of the 75O°F portion of Run 5P1, the temperature was
raised to 800°F via the electrical heaters in order to obtain a more highly
devolatilized product. Both runs were completely operable.
The extent of preoxidation in these runs was somewhat higher than the
desired 10$ because the coal feed rate was lower than the calibrated value, due
to a worn out variable speed drive.
Routinely, the product gas was analyzed by an on-line gas chromato-
graph for carbon oxides, methane and nitrogen. In these runs, separate gas
samples were taken for analyses in the laboratory for C2-C3 hydrocarbons, H2,
COS, and H2S. Also, the continuous infrared SO2 analyzer had been received
and was used for the first time.
In contrast with Run 2P2 in which Ireland Mine coal was preoxidized
at 750°F, Run 5P1 (also at 750°F) gave a 1.2% yield of a heavy tar rather than
a 0.2% yield of light oil. It is not clear whether this difference was caused
by the lower extent of preoxidation in Run 5P1 (11.9$ versus 18.6$) or by an
intrinsic difference in the kinetics of formation and of thermal decomposition
of the liquids which form when the Illinois No. 6 and Pittsburgh Seam coals
are heated to 750°F. A run with Ireland Mine coal at the same conditions used
in Run SPl.may provide a clue as to why preoxidation of Pittsburgh Seam coal
is "ineffective at elevated system pressure.
Run 6P - Preoxidation at Adiabatic Conditions
In this run the air input and the defective feeder drive were adjusted
to give the preoxidation level corresponding to adiabatic conditions in the
process preoxidizer. The temperature was increased slightly to 810°F, compared
with Run 5P2 in which the temperature was 8OO°F. The feeder drive since has
been replaced.
-------
132.
In Run 6P, the pressure drop across the fluidized bed increased
continuously. Microscopic examination and chemical analysis of the final bed
material drained from the reactor showed that ash particles had accumulated
in the bed during the run. Ash and pyritic sulfur balances on the product
coal and bed material showed that 10.55$ of the dry feed coal had accumulated
in the bed compared with 10.29$ as obtained by difference in the overall weight
balance, as shown in Table 33. A check of the chart records for Run 5P2 also
showed an increase in pressure drop across the bed, but to a lesser extent
than in Run 6P. In Table 33 the accumulations in Runs 5P1 and 5P2 were
obtained by differences in the overall weight balances, in view of the good
agreement between the measured and difference values shown in Run 6P.
The nominal retention times shown in Table 31 are corrected for the
presence of accumulated ash.
All of the runs discussed here were completely operable, with no
traces of agglomeration in the preoxidizer. The preoxidized coal products
from Runs 5P2 and 6P were fed to the gasifier operated at full process air
flow as in Runs 5G3 and 6G3, respectively. Both runs were completely operable,
with no traces of agglomeration or of ash slagging.
5. Conclusions - Illinois No. 6 Coal
Current results with Illinois No. 6 coal show that this material is
completely operable at 750 and 800°F in the preoxidizer at about the 8-12$
level of preoxidation.
The coal was run through both preoxidation and gasification at process
conditions with no operability problems whatsoever. It appears that there will
be no oxygen breakthrough in the process preoxidizer.
6. Laboratory Screening for Operability in the Gasifier
It is useful to have a screening test which can identify those coals
which will not be operable in the gasifier. The test should be rapid and
provide a quantitative measure of the relative tendency of the coal to cake.
Several tests were used:
a. Gieseler Plastometer
The Gieseler plastometer (ASTM Designation D-2639-67T) often is
used to measure the caking-coking tendency of coals. In effect the method
measures, at standardized conditions, the coal's viscosity while being heated
to about 90O°F. The coal sample (5 grams) is compacted around a small stirrer
to which a constant torque is applied. The sample is heated from 59O°F at
5.4°F/minute and the number of rotations of the stirrer shaft is read out
continuously on a dial. The coal's fluidity (reciprocal viscosity) is given by
the rpm of the stirrer shaft, expressed as dial divisions per minute (DDPM).
The fluidity increases with increasing temperature, goes through a maximum,
then decreases to zero as thermal decomposition causes the coal to form a rigid
mass of coke.
-------
133.
The plastometer data for raw Ireland Mine coal and several of its
preoxidized products are plotted in Figure 31. The relationship between
fluidity (DDPM) and physical changes which occur when raw coal is heated
usually are stated as shown below. The temperatures at which the phenomena
occur for raw Ireland Mine coal are also shown:
Physical Change Appearance of Coal DDPM Temp.t °F
Incipient melting Rounding of sharp edges. 1 71O
Agglomeration Coherent agglomerate, no 5 742
loss of particle
identity.
Fusion Coke, complete loss of Max. 803
particle identity.
These phenomena, and their associated temperatures, correspond closely with
the observations made when Ireland Mine coal was heated in a thermobalance
purged with helium at atmospheric pressure (page 124).
As shown in Figure 31, the Run IP product, made at 7OO°F, retained
considerable fluidity whereas the Run 2P2 product, made at 750°F, had a lower
fluidity. The plastometer information is in accord with the results of the
gasification runs in which the Run IP product melted with complete loss of
particle identity and the Run 2P2 product formed agglomerates with no loss of
particle identity.
Also shown in Figure 31 are data for 28 x 100 mesh Ireland Mine coal
preoxidized at low pressure (1.5 atm) during a previous experimental program.(22)
Somewhat lower fluidity was obtained at much less severe conditions of pre-
oxidation as noted below;
Pressure, Preoxidation
atm Temperature, °F # Preoxidation
1.5 700 5.5
15 750 18.6
On increasing the severity of preoxidation to 13.8$ at low pressure and at
7OO°F, the product showed no fluidity.
The Gieseler run on raw Illinois No. 6 coal showed a maximum fluidity
of about 2.8. Only Run 5P1 preoxidized coal was tested, and this gave no
fluidity reading at all. Runs 5P2 and 6P were not tested, but the results of
a qualitative test discussed below indicated that they were equally inert.
b. Laboratory Fluidized Shock Heat Test
Another test was to heat the coal sample to 15OO°F in about
two minutes while fluidizing it in a stream of nitrogen. These tests gave the
following results:
-------
nlOO
134.
n
(0
o
z
600 62O 640 66O
720
Temperature, °F
-------
135.
Product
From
Run No.
4P
5P1
5P2
6P
Temp.,
750
800
800
810
Preoxidation
27.9
11.9
11.5
8.7
Observation
Slight agglomeration of particles,
Slight agglomeration of particles,
No agglomeration.
No agglomeration.
The IP and 2P2 products showed moderate to severe caking in the same
test. It was concluded that preoxidized coals which showed no evidence of
agglomeration would certainly pass through the gasifier without caking, while
those which agglomerated slightly would probably make it. A coal which caked
severely definitely would be inoperable.
c. Free Swelling Index
A third test which was tried was the Free Swelling Index (FSl).
This test is primarily used to evaluate coking coals. The test consists of
rapidly heating the coal to 1500°F in a special crucible and then examining
the shape of the resulting pellet against a standard chart. We found that
even a treated coal which had an FSI of only 1 (2P2 product) would agglomerate
in the gasifier. The FSI method was considered inferior to the fluidized test
and was dropped.
7. Probable Cause of Ineffective Preoxidation at Elevated Pressure
Our previous experience with preoxidation at atmospheric pressure
predicts that the products from Runs IP and 2P2 would have been completely
operable in the gasifier. In addition, preoxidation of Ireland Mine coal at
essentially Run 2P2 conditions, but at atmospheric pressure, provided an
operable feedstock for the experimental work on IGT's HYGAS Process.!23)
Microscopic examination of the Runs IP and 2P2 products showed no
unusual features. In fact, the appearance of the particles was entirely
similar to those produced during the laboratory study of preoxidation at
atmospheric pressure, which are described in Section VI.
The Gieseler plastometer data for Ireland Mine coal preoxidized at
low pressure would indicate both the Runs IP and 2P2 products should have no
fluidity.
A clue provided in Run 2P2 leads to a plausible explanation of the
ineffectiveness of preoxidation at 15 atm. Previous data for preoxidation
at atmospheric pressureV24) show that, at Run 2P2 conditions of temperature
and severity of preoxidation, about 3 wt % of tar should have been produced.
Such tars contain a large proportion of high boiling (+750°F) material. In
Run 2P2 no heavy tar was produced. Rather, a light-yellow oil amounting to
0.20 wt i<> of the coal was collected with the condensed water. At 15 atm
pressure, evaporation of the tarry matter from the surface of the coal particles
is hindered. Thus, the material remains within the particles, encased in a
rigid outer shell. When exposed to gasification conditions, this material
breaks through the shell and causes caking or agglomeration.
-------
136.
C. Gasification
1. Introduction
At process conditions, there are four gaseous streams which enter at,
or near, the bottom of the gaslfler. The nature of these streams, as determined
by heat and material balance calculations, made for Case I (Table 5), is shown
below;
Mols/lOO Pounds Dry Coal Feed to Preoxidizer
Stream No.
Preoxidizer Gas Glaus Tail Gas
C02
N2
H20
CO
.188
1.330
.558
.076
2.152
1.946
4.329
.071
6.346
Air
4.353
1.157
5.510
Steam
3.732
3.732
Total
2.134
10.012
4.361
.1.233
17.740
In the experimental work, the flows of C02, steam, coal and total inlet gas
always were maintained in the proportions shown above. The tendency toward
formation of local hot spots and, therefore, of ash slagging increases as the
O2 partial pressure at the entry point increases. The 02 partial pressure can
be varied, within limits, depending upon how the four inlet streams are
combined:
Conditions at Feed Point
.of Preoxidized Coal
Complete Mixing of all Streams
Streams 1+3+4
Streams 1+3
Stream 3 only
p , atm
(system pressure = 15 atm)
1.04
1.62
2.41
3.15
Other combinations are possible, but the above list serves to illustrate the
range of partial pressures involved.
The gasifier was operated in the absence of acceptor since the
phenomena of caking and ash slagging are independent of the presence of
acceptor.
The two possible problems anticipated in operating the gasifier are
caking and ash slagging. The first preoxidized feedstock, Run IP product, was
shown in the laboratory to be almost certain to cake in the gasifier. In order
to determine the magnitude of the problem of ash slagging and to train the
operators in running the gaslfler, Disco char was used in the gasifier feed-
stock. The Disco char, made from Pittsburgh Seam coal, had been partially
gasified at 1775°F in a previous experimental program, privately sponsored.
Four runs were made with this material.
-------
137.
2. Operability with Respect to Ash Slagging
a. Runs OG2 and OG3
These runs were made with all the inlet streams fed to the
bottom of the cone in the gasifier boot, as is shown in Configuration A,
Figure 32. Since the two entry points are within one inch of each other,
this configuration corresponds to perfect mixing of the four inlet gas streams.
Run conditions and results are given in Table 35. In Run OG2 the air flow was
adjusted to give an 02 partial pressure of 0.43 atm at the bottom of the
fluidized bed. Temperature profiles showed no evidence of ash slagging.
After about nine hours of operation, the air flow (for Run OG3) was increased
to give an.inlet 02 partial pressure of 1.04 atm. This air flow corresponds
to the total air required at process conditions. Again, there was ho evidence
of ash slagging. After two hours, the run was ended and the vessel was dis-
assembled for inspection. No trace of slag was found.
b. Runs OG4 and OG5
For these runs, the air and char entry point was raised, by
means of a 1/4" OD tube, into the 4" diameter section of the gasifier bed, as
was shown in Configuration B in the aforementioned figure. In Run OG4 the Oa
partial pressure at the inlet tube was 1.0 atm, corresponding to complete
mixing of all process gas streams. After about nine hours of operation, which
gave no evidence of slagging, the air flow (for Run OQ5) was increased to give
an 02 partial pressure at the inlet tube of 2.4 atm, corresponding to mixing
of Streams 1 and 3. Since the temperature profiles again showed no evidence
of slagging, the run was ended after four hours. Inspection of the vessel
showed that no slag had formed. The exit streams for these runs were not
measured since the inlet flows were identical with Runs OG2 and OG3.
3. Operability with Respect to Caking -
Use of Preoxidized Ireland Mine Coal
a. Run 1G3
This run was made at Run OG3 conditions, with all inlet streams
fed to the bottom of the gasifier boot. Run conditions first were established
by feeding Disco char to the gasifier. Preoxidized Ireland Mine coal from Run
IP was used in this run to provide a severe test of the concept of the use of
an 02-bearing carrier gas to prevent caking.
After 13 minutes of coal feeding, the pressure drop across the
fluidized bed increased drastically, an indication that massive caking had
occurred. Examination of the material recovered after disassembly of the
vessel showed that the coal had melted with complete loss of particle identity
and had cemented together the Disco char particles.
b. Run 2G3
This run was made at the same conditions as in Run OG3.
Preoxidized Ireland Mine coal from Run 2P2 was used as feedstock. Almost
immediately after feeding the coal, the product char contained numerous small
-------
138,
FIGURE 32
CONFIGURATION OF ENTRY POINTS
IN GASIFIER
B
4" Dia.
Dia.
PREOX.
AIR I COAL
H20, C02, N2
PREOX.
AIR + COAL
H20, C02 ,N2
-------
139.
TABLE 35
Conditions and Results for Gasification Runs
System Pressure; 15 atmospheres (206 pslg)
Temperature ; 1700°F
Run Number
OG2
OG3
Input
Steam, SCFH
C02
Na
Lift Gas
Air, SCFH
N2
Purges (N2), SCFH to bed
Purges (N2), SCFH above bed
Feedstock
Feed Rate, Ib/hr
Output
Exit Gas Rate, SCFH (dry basis)
Composition. Mol 1o
H2
CH4
CO
C0a
Na (by difference)
Outlet Gas, top of bed
Composition,. Mol #
HaO
Ha
CH4
CO
C02
N2
Flow Rate, SCFH, at top of bed
Fluidizing Velocity, ft/sec
Product Char, Ib/hr
Duration of Coal Feeding, hr
Bed Weight, Ib
Total Gasification RateC1)
Percent Carbon Burnoff(2)
84
38
64
43
99
13
18
28 x
4.93
334
7.93
.40
9.90
13.51
68.26
14.60
7.04
.36
8.78
11.99
57.2
376
.34
3.56
9.5
7.10
45
35
78
35
64
105
34
13
24
100 Disco Char
5.50
332
7.61
.39
10.33
16.21
65.46
13.68
6.82
.35
9.27
14.54
55.3
370
.33
3.53
2.0
7.02
71
46
5G3
76
38
64
105
34
5
6
Run 5P2
4.15
334
11.17
.86
13.81
14.97
57.54
15.03
9.67
.75
11.96
12.95
49.64
386
.34
1.51
8.0
4.56
105
68
6G3
77
38
64
105
34
5
9
Run 6P
4.83
334
11.81
1.07
13.78
14.64
58.69
15.58
10.23
.93
11.94
12.69
48.62
385
.34
2.07
7.6
4.14
110
55
(i) Pounds Carbon gasified/min/lb Carbon in bed x 104.
(2) 100 (Ib/hr Carbon gasified)
(Ib/hr Carbon in with coal)
-------
140.
agglomerates. The amount of agglomerates continued to increase and the run
was ended after 45 minutes. After shutdown, the bed material drained freely
from the gasifier. Microscopic examination of this material showed that the
coal particles had not lost their identity, but that they had adhered to the
Disco char particles, and to themselves.
c. Run 4G3
This run was made at Run OG3 conditions using the Run 4P product
as gasifier feedstock. The history of this feedstock is summarized below:
Temp., Percent
°F Preoxidation
1st Stage (Run 2P2) 750
2nd Stage (Run 4P) 75O
Total
After 2.5 hours of coal feeding, steady state had been reached
with respect to displacement of the start-up bed inventory. At that time a
severe leak developed in the char outlet line, which forced a shutdown before
any measurements were made on the exit streams.
Examination of the reactor after disassembly showed no trace of
ash slagging or caking/agglomeration.
4. Operability During Gasification - Use of
Preoxidized Illinois No. 6 Coal
The preoxidized coal products from Runs 5P2 and 6P were fed to the
gasifier operated at full process air flow in Runs 5G3 and 6G3, respectively.
Run conditions and results are summarized in Table 35, and some properties of
the char products are shown in Table 36.
Both runs were completely operable, with no trace of agglomeration/
caking or of ash slagging.
5. Conclusions
Results indicate that ash slagging in the gasifier does not pose a
problem with either Pittsburgh Seam or Illinois No. 6 coals.
Preoxidized Illinois No. 6 coal and preoxidized Pittsburgh seam coal
have both been fed to the gasifier without caking. As far as the gasifier
vessel is concerned, these materials represent satisfactory noncaking feeds.
However, the amount of preoxidation required for the Pittsburgh Seam coal is
more than three times that which corresponds to adiabatic conditions in the
process preoxidizer.
D. Integrated Operation with Acceptor Circulation
1. Introduction
A "demonstration" run was scheduled in order to give an appreciation
of possible problem areas. The purposes of the run were;
-------
Properties of Feeds and Products - Gasification Runs
Run Number
Hydrogen, Wt % (dry basis)
Carbon
Nitrogen
Oxygen (by difference)
Sulfur
Ash
Sulfide Sulfur
Pyritic Sulfur
Sulfate Sulfur
Mean Particle Diameter, inch
Mean Particle Density, Ib/ft3
Ash Composition. Sulfur-free. Wt H>
A1203
SiOa
Fe208
CaO
MgO
Na2O
K2O
Ti02
5G3
Preoxidized
Coal Feed
3.29
67.02
1.44
7.98
3.72
16.55
.01
1.11
.27
Char
Product
.46
65.64
.39
-.10
2.47
31.14
1.73
O
0
.0171
49.8
.0155
34.1
6G3
Preoxidized
Coal Feed
3.44
70.61
1.36
5.8O
3.40
15.39
.03
1.37
.07
Char
Product
.50
68.94
.55
-1.09
2.22
28.88
1.31
0
0
.0175
51.8
.O155
43.2
Total
99.1
-------
142.
a. To determine whether or not kinetic limitations exist on;
(l) Rate of desulfurization of the char feed in
the gasifier at process conditions.
(2) Rate of pickup of H2S by the acceptor in the
gasifier.
b. To determine magnitude of acceptor activity loss at process
conditions.
c. To determine ash slagging characteristics of fuel char in
the regenerator.
d. To determine the extent and nature of any deposits formed
in the gasifier or regenerator.
e. To obtain semi-quantitative data on rate of gasification.
No air was used in the run in order to simulate the reducing con-
ditions in the upper part of the process gasifier where most of the char
desulfurization and H2S pickup by the acceptor occurs. Added H2 and recycle
gas were used to control the partial pressures. Process heat was supplied
via the electrical heaters and by the acceptor reaction. Ungasified Disco
char, made from Pittsburgh Seam coal, was used as gasifier feedstock in order
to avoid any difficulty caused by caking. The regenerator fuel char was the
same material which had been partially gasified during break-in operations.
Regeneration of CaS was in the manner of Case I in the feasibility
study, except that by use of recycle gas, the ratio of exit gas/total sulfur
in the regenerator was high enough to avoid any equilibrium limitations on S02
partial pressure, (study of the full-blown Case I regenerator conditions is a
separate experimental program and will be undertaken during Phase II-B of the
contract).
2. Run DIB
During the run, thirteen calcining-recarbonation cycles were achieved.
A major purpose of the demonstration run was to sample the gasifier
product gas to determine the extent of desulfurization. During the fifth
cycle, an attempt to obtain a gas sample was thwarted by a corroded sample
valve. At this point, a voluntary shutdown was made in order to replace the
valve. The opportunity was taken to dismantle the vessels for inspection. No
evidence of any deposits or of ash slagging in the regenerator was found.
Due to inexperience in operating the regenerator, the electrical
power input, used to compensate for heat losses, was too low. As a result the
bed temperature tended to level out at 1840°F rather than at the process
temperature of 1880°F. During the first five cycles, the bed temperature
purposefully was held at 1840°F in order to reduce the chance of early termina-
tion caused by ash slagging.
-------
143.
On resumption of the run the bed temperature was increased to 1860°F
since no slagging had occurred earlier. Again, the bed temperature was kept
lower than the process temperature in order to obtain the desired data on the
acceptor sulfur cycle, uncomplicated by possible ash slagging.
Acceptor circulation was continued until the thirteenth cycle when
the regenerator thermowell ruptured. An immediate shutdown was required. The
thermowell, made of Hastelloy X, must withstand the full system pressure of
206 psig. Failure was caused by severe, localized, corrosion by sulfur. In
future work the thin-walled Hastelloy X thermowell will be replaced by a larger
diameter thick-walled thermowell made of Type 310 stainless steel, which has
better resistance to sulfur corrosion.
Inspection of the vessels after the final shutdown again showed no
evidence of deposits or ash slagging in the regenerator.
Summarized run conditions and results are shown in Table 37 for the
gasifier and regenerator. Results are shown for the tenth calcining-recarbona-
tion cycle, at which time the product gas sample was taken. Analyses of the
char feeds and products are given in Table 38.
3. The Sulfur Cycle
a. Acceptor Sulfur Cycle
A discussion of the sulfur reactions which are involved is given
in Section III-B-2 (feasibility study).
The sulfur contents of the recirculating acceptor, sampled peri-
odically during the run, are shown below:
Cycle No. 4 5_ 7_ 10 13
Gasifier Samples
$ CaS content* 9.0 11.2 10.9 9.5 5.2
Regenerator Samples
$ CaS content ND 9.8 ND 6.4 3.6
* mol % of total calcium in form of CaS.
ND not determined.
During the first part of the run (first five cycles), sulfur
accumulated in the acceptor. This is attributed to the low (l840°F) regenerator
temperature which suppressed the rejection of sulfur by the reaction,
3/4 CaS04 + 1/4 CaS = CaO + S02 (3)
During the second part of the run, the increased bed temperature (1860°F)
allowed the accumulated sulfur inventory to be depleted. During the tenth
cycle the data in Table 37 show that the SO2 partial pressure in the exit gas
was 0.035 atm. The equilibrium partial pressure (Figure C-3) for reaction (3)
at 186O°F is O.O70. Since this cycle occurred during the period of depletion,
the resulting APSO driving force of O.O35 atm thus is more than adequate to
reject sulfur by rSaction (3).
-------
144.
TABLE 37
Conditions and Results for Demonstration Run - Gasifier
System Pressure: 15 atmospheres (206 psig)
Temperature: 1700°F
Run Number
Feedstock
Acceptor
28
16 x
DIB
x 100 mesh fresh Disco Char
28 mesh Tymochtee 9 dolomite
Cycle Number
Acceptor Circulation rate, Ib/hr (MgCOa'CaC03 basis)
Acceptor Activity
Char Feed Rate,. Ib/hr (dry basis)
Input
Steam, SCFH
002 to boot
Recycle Gas, SCFH
Fluidizing, to boot
Char lift line, above boot
Purges (N2), SCFH to bed
Purges (N2), SCFH above bed
Output
Exit Gas Rate, SCFH (dry basis)
Composition, Mol %
H2
CH4
CO
CO 2
N2 (by diff.)
Outlet Gas, Top of Bed
Composition,, Mol $>
H20
H2
CH4
CO
CO 2
N2
Flow Rate, SCFH, at top of bed
Product Char, Ib/hr
C02 to Acceptor, SCFH
Fluidizing Velocity, ft/sec
top of bed
boot (acceptor layer)
Duration of Acceptor Circulation, hr
Bed Weight, Ib
Total Gasification RateV1)
i> Carbon Burnof f ( 2)
(0 Lbs. Carbon gasif ied/min/lb Carbon in bed x 1O4
(a) IQO (Ib/hr Carbon gasified)
(Ib/hr Carbon in with Char)
10
12.8
.573
4.54
73
53
53
15
73
47
10
32
324
17.86
1.59
19.40
14.96
46.19
17.97
16.25
1.45
17.65
13.61
33.06
358
3.12
13.1
.32
.64
9.6
7.58
31
31
= mol ratio,
CaC03
CaCO
CaO
-------
145.
Run Number
TABLE 37 - Cont'd.
Conditions and Results for Demonstration Run - Regenerator
System Pressure; 15 atmospheres (206 psig)
Temperature : I8600 F
DIB
Fuel Char
Acceptor
28 x 100 mesh partially gasified Disco Char
16 x 28 mesh Tymochtee 9 dolomite
Cycle Number 1O
Acceptor Circulation Rate, Ib/hr (MgC03-CaC03 basis) 12.8
Acceptor Circulation Rate, Ib/hr (as fed basis) 8.50
Acceptor Activity .573
Acceptor Sulfur Content, Wt % (as fed basis) 2.10
Fuel Char Feed Rate, Ib/hr (dry basis) 1.40
Input
Recycle. SCFH
Fluidizing 153
Char and Acceptor Lift Line 242
Air 115
N2 85
C02 8
Purges (N2), SCFH to bed 5
Purges (N2), SCFH above bed 19
C02 from Acceptor, SCFH 13.1
Output
Exit Gas Rate, SCFH (dry basis) 244
Composition, Mol H>
CO2 17.01
CO .80
S02 .23
N2 (by diff.) 81.96
Flow Rate, SCFH, at top of bed 635
Fluidizing Velocity, ft/sec. 1.04
Acceptor, Ib/hr (MgO-CaO basis) 7.02
Acceptor, Sulfur Content, Wt % 1.74
Overhead Ash, Ib/hr .652
Partial Pressure Driving Forces, atmospheres
AP C02 Bottom of BedC1) 2.74
AP C02 Top of Bed(x) 2.05
AP S02 Top of Bed(2) .035
Inlet 02 Partial Pressure, atmospheres .58
(i) Below equilibrium partial pressure in the reaction;
CaC03 = CaO + C02
(2) Below equilibrium partial pressure in the reaction;
3/4 CaS04 + 1/4 CaS = CaO + S02
-------
146.
TABLE 38
Analyses of Char Feeds and Product
Run DIB
Hydrogen, Wt $ dry basis
Carbon
Nitrogen
Oxygen (diff.)
Sulfur
Ash
Sulfide Sulfur, Wt #
Ash Composition^ Sulfur-free
A1203, Wt
SiO2
Fe203
CaO
MgO
Na20
K20
Ti02
P20S
Feedstock
2.63
76.24
1.77
4.66
1.27
13.43
0
basis
23.2
51.6
13.1
4.7
0.8
0.4
1.8
1.0
0.4
Product
0.52
75.47
.065
0.28
0.45
23.22
0.35
Gasif ier Regenerator
Fuel Char Overhead Ash
0.69 0.3
76.98 50.2
0.91 ~ 0
1.72 ~ 0
0.84 2.54
18.86 46.9
0.09 0.46
Total 97.0
-------
147.
The H2S content of the gasifier product dry gas was below the
detection limit of 30 ppm of the gas chromatographic method used for analysis.
As shown in the feasibility study, the equilibrium content of the wet gas at
process conditions is 230 ppm. Undoubtedly, most of the H2S content of the
product gas was absorbed by the ammonia-rich condensate formed by the unreacted
steam and the ammonia generated from the char nitrogen in the gasifier. The
equilibrium H2S content of the gasifier wet product gas, calculated for the
data in Table 37 and the equilibrium for the reaction, CaO + H2S = CaS + H20
(shown in Figure C-2), corresponds to 4.5 gm/hr of sulfur. Analysis of the
recovered condensate drained from the catchpots below the exit gas coolers,
Cl and C4 (Figure 12), showed that the absorbed H2S, present as NH4HS in the
condensate, amounted to only 2.2 gm/hr sulfur, or roughly one-half of the
minimum possible amount of H2S in the wet gas. Some of the "missing" H2S
flashed off. the condensate as it was being drained from the catchpots at 15 atm
pressure into receivers at room conditions.
In future work a small slipstream of the hot, wet gas taken
downstream of the gasifier solids filters will be throttled to atmospheric
pressure, cooled and then passed through aqueous dilute mineral acid to insure
that all the H2S is swept into a suitable absorber for subsequent analysis.
Back titration of the dilute acid also will give the total ammonia content of
the wet gas.
b. Char Desulfurization in the Gasifier
Sulfur balances based on data in Tables 37 and 38 show that 76%
of the feed sulfur was removed in the gasifier, based on the total sulfur con-
tent of the product char. However, most of the residual sulfur was in the
form of CaS in the char ash, as is shown by the sulfide sulfur analysis in
Table 38. From a process standpoint, it is immaterial whether the sulfur
content of the gasifier feedstock leaves the gasifier as CaS in the acceptor
or as CaS in the char ash. After correction for the CaS content of the product
char, the removal of organic sulfur was'93$.
A high degree of sulfur removal also occurred in Runs 5G3 and
6G3, made with preoxidized Illinois No. 6 coal feedstock in the absence of
acceptor. From Tables 35 and 36, the extent of sulfur removal based on total
sulfur content and organic sulfur content have been calculated as follows:
Percent Sulfur Removed
from Feedstock
Run No. Total Organic
5G3 76 91
6G3 72 86
Comparison of the outlet gas compositions for these runs with equilibria in
the Fe-0-H-C-S system shows that the FeS and FeO are the stable phases at the
conditions used and that sulfide sulfur is formed by the reaction,
FeO + H2S = FeS + H20.
* Organic sulfur defined as; Total sulfur minus sulfide sulfur.
-------
148.
c. Sulfur Balances
As discussed above, the H2S content of the gasifier product gas
could not be determined because of loss from the condensate. Furthermore, the
accumulation-depletion of CaS "flywheel" in the acceptor inventory somewhat
obscures the true behavior of the acceptor sulfur cycle. The following
approach was used in order to present the sulfur balances: (l) In the gasifier
a hypothetical, sulfur-free acceptor was assumed to enter from the regenerator
and the H2S content of the wet product gas was assumed to be at equilibrium.
The sulfur content of the acceptor leaving the gasifier then was obtained by
difference. (2) In the regenerator, measured values for the output streams
were used, along with the aforementioned sulfur content of the gasifier
acceptor stream. (3) The excess sulfur leaving the regenerator then was
rationalized with the rate of depletion of the "flywheel". The calculations
were based on data in Tables 37 and 38. Three additional measurements on
minor streams were made, but are not included in the tables:
(l) COS content of gasifier dry product gas.
(2) SO2 content of condensate from regenerator gas
cooler, C-6.
(3) Sulfur content of regenerator back-up filter.
Results of the calculations are shown in Table 39.
The by-difference value of the sulfur content of the acceptor
leaving the gasifier corresponds to 1.8 mol $ conversion of the total calcium
to CaS, the same conversion as used in the feasibility study (Table 5). The
regenerator balance shows considerably more sulfur out than in. This excess
sulfur is rationalized with the depletion data shown on page 139 at the bottom
of Table 39. As shown, the agreement is good. As can be seen in Table 39,
the amount of sulfur in the "flywheel" is large compared with the total amount
of sulfur entering the system.
In future studies on acceptor activity with respect to the
sulfur and C02 reactions, the temperature and sulfur/gas ratio in the regen-
erator will be adjusted to avoid the accumulation of sulfur in the recirculating
acceptor.
d. Elemental Sulfur Content of the Regenerator Offgas
If equilibrium is assumed for the reaction,
2 CO + S02 = 1/2 S2 + 2 C02, (4)
the partial pressure of elemental sulfur in the offgas can be calculated from
the data in Tables 37 and C-3. The result is O.O0473 atm at 1860°F, which
corresponds to 5.9 gm/hr sulfur. At the end of Run DIB, only a small amount
of elemental sulfur corresponding to about 0.1 gm/hr was found. Its location
was in the pipiug between the back-up filter and the gas cooler C6 (Figure 12) .
-------
149.
TABLE 39
Sulfur Balance - Run DIB
Cycle Number 1O
Grams/hour
Gasifier Sulfur
IS.
Feed Char 26.2
Acceptor * 0
Out
Product Char
H2S to Gas **
COS
to Acceptor (by dif f.)
Regenerator
In
Acceptor 15.1
Fuel Char 5.3
20.4
Out
S02 in Gas 21.1
S02 in Condensate 0.1
Overhead Ash 7.5
Backup Filter 0.9
29.6
excess sulfur = 29.6 - 20.4 =
9.1
Depletion of Acceptor Sulfur Inventory
Fed to Regenerator
Less Sulfur Added in Gasifier
Out from Regenerator
excess sulfur = 65.9 - 55.5 =
10.4
* By assumption.
** Based on equilibrium content.
-------
150.
The success on the Case I sulfur rejection method described in
Section III-B-2 (feasibility study) depends upon close approach to equilibrium
in reaction (4). If this situation occurred in Run DIB, then about 5.8 gm/hr
of elemental sulfur must be accounted for. As shown in Table 38, the sulfur
content of the overhead ash material is about three times greater than that of
the fuel char. At combustion conditions, the fuel char sulfur can be expected
to form S02 rather than to remain in the overhead ash. As shown in Table 39,
the sulfur content of the overhead ash material corresponds to 7.5 gm/hr sulfur
which, after correction for the sulfide sulfur content,* is sufficient to
account for the necessary elemental sulfur.
In the continuous gasification unit, the regenerator offgas
which contains the overhead ash material has been cooled to about 300°F before
reaching the external cyclone. This temperature is favorable for condensation
of sulfur vapor and subsequent absorption by the high surface area ash material.
A single batch extraction of the overhead ash with boiling
toluene, followed by evaporation of the solvent, produced elemental sulfur
crystals. Qualitatively, the amount of sulfur recovered appeared to be too
small to account for the measured sulfur content of the overhead ash. In
future work the overhead ash will be extracted exhaustively in a modified
Soxhlet apparatus to determine whether absorbed sulfur can account for all of
the observed increase in sulfur content .
An -alternate explanation of the enhanced sulfur content is that
organic C-S bonds are formed from the elemental sulfur and the unburned carbon
at regenerator conditions.
4. Acceptor Activity
A major purpose of Run DIB was to determine the effects on acceptor
activity toward the C02 acceptor reaction of exposure to several cycles of
calcining and recarbonation.
During the run, the circulating acceptor inventory was sampled
periodically, and its CaC03 content was determined. The acceptor activity is
defined as the mol ratio,
CaCO,
CaC03 + CaO
For the acceptor samples withdrawn from the gasifier "undercarbona-
tion" , never was observed, i.e., the carbonation ratio always was consistent
with the intrinsic activity, as was shown by the fact that no weight gain
occurred on prolonged exposure to CO2 at conditions which were kinetically
very favorable for the reaction, CaO + C02 = CaC03 .
Acceptor samples withdrawn from the regenerator always were completely
calcined, i.e., any CaC03 present was undetectable.
-------
151.
No acceptor makeup was used during the run. Therefore, the
acceptor activity determined for each sample could be associated with the
completion of a definite number of calcining-recarbonation cycles. A cycle
is completed when the number of mols of CaO fed from the L-4 acceptor feed
hoppers equals the number of mols of CaO contained inside the regenerator,
gasifier, and the standlegs. Results are shown below:
Cycle Activity
4 0.68
5 0.65
7 0.66
10 0.57
13 0.55
An accurate estimate of the equilibrium activity of the recirculating C02
acceptor at the conditions of Run DIB would require the completion of about
4O cycles, based on our previous work for the Office of Coal Research.(1C)
However, comparison of the above results with the OCR data indicates that the
equilibrium activity will be higher than that estimated in the feasibility
study (page 21 of this report).
5. Char Combustion in the Regenerator
A characteristic of the run was that the CO content of the regenerator
offgas could not be increased beyond 0.8 mol $. Attempts to increase the CO
content by temporarily decreasing the air/fuel char ratio were without effect,
showing that the low reactivity of Disco char toward the reaction, C + CO2 = 2 CO,
which must occur in the upper part of the bed above the combustion zone, was
responsible. It is noteworthy that no CaS~CaS04 transient liquid was observed
at this low level of CO content. Previous work with low-rank Western coals,
which have a high lime content in the ash, showed that maintenance of 2-3$ CO
in the regenerator offgas is necessary to suppress formation of the transient
liquid in the coal ash.
The carbon burnout of the regenerator fuel, calculated by carbon and
ash balances based on data in Tables 37 and 38, was 74$ vs 9O$ which was
assumed for the feasibility study. The observed lower burnout is consistent
with the low reactivity of Disco char which led to the low CO level mentioned
above. In future work the bed temperature will be raised above 186O°F to
determine the limit dictated by ash slagging. Higher temperatures will improve
the burnout situation.
6. Nitrogen Removal from the Gasifier Feedstock
Removal of nitrogen from the feed char, calculated from data in
Tables 37 and 38 was nearly complete at 97$. Analysis of the recovered con-
densate showed that at least 30$ of the incoming char nitrogen had been
converted to ammonia.
-------
152.
7. Kinetics of the Gasification Reactions
The five gasification runs made so far were concerned with potential
operability problems. Gasification rate data obtained during the runs merely
were byproducts. As shown in Tables 35 and 37, the rates range from 110 x 10~4
Ib carbon gasified/lb carbon in bed/min at bottom-of-gasifier conditions to
31 x 1O~4 at top-of-gasifier conditions, compared with the overall rate of
75 x 10~4 which was assumed in the feasibility study.
Although it seems possible that the assumed rate can be achieved, a
systematic study will be necessary in which a preoxidized coal feedstock first
will be processed at bottom-of~gasifier conditions. The char product will then
be used sequentially as feed char to succeeding runs in which the partial
pressures of reactants and products will approach those at top-of-gasifier
conditions.
-------
VI. 1AB STUDY OF
KINETICS
-------
153.
VI. LABORATORY STUDY OF COAL PREOXIDATION KINETICS
A. Introduction
The objective of the study was to determine the rate of reaction
of oxygen with Ireland Mine coal, as influenced by the following variables:
1. temperature
2. O2 partial pressure
3. extent of preoxidation
4. coal particle size
Conditions and results for 17 runs are summarized in Table 40.
-------
Summary of Results
Laboratory Study of Kinetics of Coal Preoxidation
1 atm system pressure
Run No.
1363-
39
35
38
32
30
29
36
39
37
44
41
40
°F
650
I
-------
155.
B. Experimental
A schematic diagram of the apparatus is shown in Figure 33. The
reactor was a l" diameter quartz tube containing 2 to 5 grams of coal and 50
grams of 20O x 325 mesh fused periclase. The periclase acted as a heat sink
and diluent, and thereby allowed good temperature control during the highly
exothermic preoxidation reactions. The system was operated at atmospheric
pressure. All runs were made with Ireland Mine coal.
The inlet gas was an 02-He mixture of known composition, at a flow
rate of 1.50 SCFH. The gas entered the top of the reactor, passed to the
bottom through an axial diptube, reversed direction, and fluidized the coal
and periclase. The reactor was heated by being immersed in an air-fluidized
bed of silica sand which in turn was heated by an external electric resistance
furnace. Reaction temperature was measured by a calibrated thermocouple
immersed in the fluidized bed of coal-^-periclase and was recorded continuously.
The recorder also controlled the net rate of heat input to the fluidized sand
bath furnace to hold the reaction temperature to within 3°F of the desired
value. The product gas passed through a series of absorbers which removed the
reaction products, as follows:
Product Absorbent
H20 magnesium perchlorate
CO2 Ascarite
CO cuprous sulfate
The residual gas was then substantially a binary mixture of 02 and He. Its
composition was monitored continuously by a thermal conductivity cell in which
the reference gas was the original O2-He mixture used as the reactor inlet gas.
The cell was calibrated, for each particular inlet gas, by passing known 02-He
mixtures through it. The method of calibration involves the use of a mathe-
matical model of the thermal conductivity cell which was developed during
previous work for the Office of Coal Research (reference Ic, Book 2). With
this method of analysis for 02, a difference in concentration of 0.01 mol $
between the residual and inlet gas easily can be measured.
After assembly of the apparatus, inlet gas was passed through the
system with the reactor at room temperature in order to establish the recorder
baseline of the thermal conductivity cell output. The fluidized sand bath
furnace, which was mounted on pulleys, then was raised quickly to surround the
reactor. The coal-periclase bed reached the desired reaction temperature in
less than two minutes. After about 20 minutes, the fluidized sand bath furnace
was lowered and the reactor was quenched to room temperature by spraying it
with water.
The rate of 02, consumption at any time was calculated in a straight-
forward manner, based on the known flow of 02 content of the inlet gas and
the measured 02 content of the residual gas. The cumulative extent of pre-
oxidation up to any time was determined by numerical integration of the rate
versus time data.
i
The ratio of coal to inlet 02 was chosen to give a low conversion
of the 02, i.e., differential rate data with respect to 02 partial pressure
were obtained.
-------
FIGURE 33
SCHEMATIC DIAGRAM
APPARATUS FOR STUDY OF COAL PREOXIDATION KINETICS
ROTAMETER
TOTC CELL
F ™*~" —"
02-He MIX
CYLINDER
m
.;.'':V:V
REACTOR
FLUIOIZED
SAND BATH
FURNACE
H90
REMOVAL
VENT
THERMAL
CONDUCTIVITY
CELL
MIX GAS
RECORDER
Oi
-------
157,
C. Rate of Reaction of Oxygen with Coal
The desired rate data were obtained reliably for only six runs.
The data are plotted in Figure 34 as percent preoxidation per minute vs.
cumulative percent preoxidation, where percent preoxidation is defined as
100(grams 02 consumed/gram dry coal). The data also are shown in Figure 35
as cumulative percent preoxidation vs. time.
Use of cuprous sulfate (Cosorbent, Burrell Corp.) as the CO absorber
led to difficulties because breakthrough occurred in an unpredictable manner.
We attempted to alleviate the situation by replacing the cuprous sulfate with
an active cupric oxide catalyst (Hopcalite) which oxidized the CO. We believed
that the effect of the reaction, CO + 1/2 02 = C02, on the response of the
differential thermal conductivity cell, which was used to measure the 02
content of the product gas, could have been neglected. However, when the run
data were worked up, comparison of the total 02 consumed, as determined by the
product gas composition, with that determined by the weight gain of the CO2
and H20 absorbers showed that, in general, the use of Hopcalite gave rates that
were too low. The effect of CO oxidation was small only in Run 39, made at the
lowest conditions of temperature and O2 inlet concentration.
The weight gain of the C02 and H20 absorbers represents most of the
02 consumed since previous work has shown that only about 6-8$ of the 02
consumed appears as CO over a wide range of temperatures and inlet 02 con-
centrations .
Subsequently, the cuprous sulfate method was tried again with smaller
coal charges and with two absorbers in series. No breakthrough occurred and
the method now appears to be reliable. Lack of time prevented repetition of
the Hopcalite runs.
1. Effect of Temperature
Only two pairs of runs, 39-44 and 2O-19 (Figure 34 ) shared common
conditions of inlet 02 concentrations and particle diameter. Characteristically,
the decline in rate with increasing extent of preoxidation is less severe at
the higher temperatures.
2. Effect of Inlet 02 Concentration
Only one pair of runs, 44-41, shared common conditions of temperature
and particle diameter. On increasing the 02 concentration by a factor of 2.2,
the rate increased by a factor of 1.5 to 1.7, depending upon the extent of
preoxidation.
3. Effect of Extent of Preoxidation
Characteristically, all the curves in Figure 34 show a high initial
rate of reaction which is compatible with the heterogeneous nature of bituminous
coal. The most reactive portion of the coal preferentially forms water, rather
than carbon oxides, as determined qualitatively by observation of the rate
of appearance of condensed water in the tube connecting the reactor and the H20
absorber during the first few minutes of each run.
-------
-------
-------
160.
4. Effect of Particle Diameter on Rate of Reaction with 02
At the lower temperature level (~ 650°F), comparison of Runs 39
(28 x 35 mesh) and 24 (65 x 100 mesh) shows that the rates for the smaller
particles were higher by a factor of about 4.0, depending somewhat on the
level of preoxidation.
If the rate of reaction were proportional to the external surface
area of the particles, then the rate should have been higher by a factor of
8.0 for the smaller particles. If the rate of reaction were proportional to
the particle diameter (shrinking core model with reaction rate controlling),
then the rate should have been higher by a factor of 2.8. The observed inter-
mediate value indicates that the mechanism of preoxidation probably involves
the external surface area. The particle density (Figure 37) and microscopic
examination both showed that some swelling of the 28 x 35 mesh particles had
occurred.
At the higher temperature level (~ 750°F), the rates for Runs 19*
(65 x 10O)mesh and 41 (28 x 35 mesh) were substantially identical over the
entire range of extent of preoxidation. As stated below the larger particles
were swollen, with numerous wall punctures. Since the particle interior is
accessible to the reacting 02 at these conditions, the rate of preoxidation
becomes independent of particle size.
D. Particle Density
The particle densities of the 28 x 35 mesh product coals from the
Hopcalite series were measured in mercury at 1 atm pressure. Results are
shown in Figure 36,
s~
High particle density from the preoxidizer is desirable when the
coal is gasified, from the kinetic standpoint of maintaining high carbon in-
ventory and gas throughput in the gasifier vessel. From this point of view,
the data in Figure 36 show that the preoxidation temperature should be as low
as practicable and that the inlet 02 partial pressure should be as high as
practicable. Increasing 02 partial pressure, at a given temperature, increases
the rate of preoxidation and, presumably, that rate at which the particle
surface becomes rigid and resistant to swelling.
Microscopic examination of the 28 x 35 mesh, 7OO-750°F products
showed that all the particles, except those rich in fusinoids, were rounded,
swollen, and punctured by small holes. The particle interiors had a sponge-
like structure. In contrast, none of the 65 x 1OO mesh products showed any
swelling or wall penetration. The particle densities (Table 40) of the
65 x 1OO mesh products were higher than that of the raw coal.
The effects of temperature and particle size upon preoxidized coal
density were confirmed by results of the continuous unit which are given in
Figure 37 for Ireland Mine coal. Similar trends were observed for Illinois
No. 6 coal, but the data scattered badly due to the presence of dense, rocky
material in the coal. p
* After correction from 730 to 750°F by means of activation energies
calculated from Runs 19 and 20.
-------
Figure 36
PARTICLE DENSITY AFTER PREOXIDATION
28 x 35 MESH COAL
0 20 40 60 80 100
INLET 02 CONCENTRATION, MOL %
-------
-------
163.
E. Dehydrogenation
Decaking of coal by preoxidation involves dehydrogenation. The
amount of hydrogen removed from the coal by reaction to form water was
calculated for each run from the H20 absorber weight gain. Results are
plotted in Figure 38. The data points are not delineated with respect to
run conditions because examination of the individual data points showed no
trends. Any separate effects of temperature, 02 inlet concentration, or
particle diameter are confounded, since the higher levels of preoxidation
are associated with higher levels of temperature and/or 02 concentration.
F. Coal Weight Loss
In the temperature range of current interest (600-80O°F) for pre-
oxidation, the coal decomposes thermally in an inert atmosphere, with loss
of volatile products. Preoxidation inhibits the loss of volatile materials
by polymerizing them. At the lower temperatures, enough of the reacting 02
remains in the polymerized materials to cause a weight gain. In general,
the temperatures used in this study were high enough to cause a net weight
loss even though some of the reacted O2 remained in the coal. In Run 20, at
600°F, no net weight change occurred although carbon and hydrogen had been
removed from the coal, as shown by the weight gains of the C02 and H20
absorbers.
Percent weight losses for the other runs are plotted in Figure 39.
As with the dehydrogenation data, confounding occurred because the higher
levels of preoxidation are associated with the higher temperatures. However,
there are apparent effects of inlet 02 concentration and particle diameter
on weight loss. As indicated in Figure 39 smaller weight losses occur at
increasing 02 concentrations. Smaller weight losses also occur with the
smaller particles. It is not clear whether this behavior is caused by in-
creased retention of the reacted 02, or by some other mechanism.
-------
-------
VII. BIBLIOGRAPHY
-------
165.
VII. BIBLIOGRAPHY
1. Consolidation Coal Co., Research and Development Report No. 16 to
the Office of Coal Research, U.S. Dept. of the Interior,
Under Contract No. 14-01-0001-415.
a. Interim Report No. 1, "Pipeline Gas from Lignite Gasification -
A Feasibility Study" (February, 1965). U.S. Department of
Commerce, National Technical Information Service PB-166817
(feasibility study), PB-166818 (appendix).
b. Interim Report No. 2, "Low-Sulfur Boiler Fuel Using the Consol
C02 Acceptor Process" (November, 1967). U.S. Department of
Commerce, National Technical Information Service PB-176910.
c. Interim Report No. 3, "Phase II - Bench-Scale Research on CSG
Process" (January, 1970).
Book 1, "Studies on Mechanics of Fluo-Solids Systems."
Gov't. Printing Office Catalog No. 163.1O:16/INT3/Book 1.
Book 2, "Laboratory Physico-Chemical Studies."
Gov't. Printing Office Catalog No. 163.1O:16/INT3/Book 2.
Book 3, "Operation of the Bench-Scale Continuous
Gasification Unit." Gov't. Printing Office Catalog
No. 163.10;16/INT3/Book 3.
d. Interim Report No. 4, "Pipeline Gas from Lignite Gasification -
Current Commercial Economics." Gov't. Printing Office
Catalog No. 163.10;16/INT4.
2. Rudolph, P.F.H., "New Fossil-Fueled Power Plant Process Based on Lurgi
Pressure Gasification of Coal." Preprints of paper presented at
A.C.S., Div. of Fuel Chem., Toronto, Canada, May 24-29, 1970.
Paper No. 9, V-14, No. 2.
3. Robson, F.L. and Giramonti, A.J., "An Advanced Cycle Power System Burning
Gasified and Desulfurized Coal." Economic Commission for Europe
Working Party on Air Pollution, Seminar on Desulfurization of Fuels
and Combustion Gases, Geneva, Nov. 16-20, 197O.
4. Curran, G.P., Fink, C.E., and Gorin, Everett, "Production of Low-Sulfur
Boiler Fuel - Application of CO2 Acceptor Process."
Paper presented before the Second International Conference on
Fluidized Combustion, Hueston Woods, Ohio, Oct. 4-7, 1970.
5. Schroeter, Louis C., "Sulfur Dioxide," pp. 91-94. Pergamon Press, 1966.
6. Batchelor, J.D., Gorin, Everett and Zielke, C. W.,
Ind. Eng. Chem.. 52, 161 (i960).
-------
166.
VII. BIBLIOGRAPHY (Cont'd.)
7. Zielke, C.W., Curran, G.P., Gorin, Everett and Goring, G.E.,
Ind. Eng. Chem.. 46, 53 (1954).
it
8. Rudolph, P.F.H., private communication, Lurgi Warme G.m.b.H.,
Frankfurt, Germany, Sept. 25, 1970.
9. National Bureau of Standards Circ. 564 (1955).
10. Keairns, D.L., private communication, Westinghouse Research Labs.,
Pittsburgh, Pa., Jan. 26, 1971.
lOa. Ibid., Oct. 18, 1971.
11. Westinghouse personal communication, August, 1970.
12. Refinery Prices - Oil and Gas Journal, p. 106 (April 12, 1971).
14. JANAF Tables, Dow Chemical Co., Clearinghouse for Federal Scientific
and Tech. Information, U.S. Department of Commerce.
15. Preuner, G., and Schupp, W., Z. Physik. Chem., ££, 129 (1909).
16. West, J. R., Ind. Eng. Chem., 42, 713 (1950).
17. Yavorsky, P.M., Mazzocco, N.J., Rutledge, G.P., and Gorin, Everett,
Environmental Science and Technology, 4, No. 9, 757 (1970).
18. Squires, A.M., Advan. Chem. Ser.. 69, 205 (1967).
19. Robson, F.L., Giramonti, A.J., Lewis, G.P., Gruber, G., Technological
and Economic Feasibility of Advanced Power Cycles and Methods of
Producing Nonpolluting Fuels for Utility Power Stations."
United Aircraft Research Laboratories Report No. J-970855-13,
Dec., 1970, Under NAPCA Contract CPA 22-69-114.
2O. Vann, H.E., Whitman, M.J., and Bowers, H.I., "Factors Affecting
Historical and Projected Capital Costs of Nuclear Power Plants
in the United States." Fourth United Nations International
Conference on the Peaceful Uses of Atomic Engery.
Geneva, Switzerland, Sept. 6-16, 1971.
21. Bartok, W., Crawford, A.R., and Skopp, A., Chemical Engineering Progress.
67, No. 2, 64 (1971).
22. Curran, G.P., Fink, C.E., and Gorin, Everett, "Coal-Based Sulfur
Recovery Cycle in Fluidized Lime Bed Combustion." Paper presented
before the Second International Conference on Fluidized Combustion,
Hueston Woods, Ohio, Oct. 4-7, 197O.
23. Kavlick, V.J., and Lee, B.S., Advan. Chem. Ser., £9, 12-19 (1967).
24. Struck, R.T., Consol Internal Report, Project 10, Report No. 7 (1953).
-------
VIII. APPENDICES
-------
APPENDIX A
-------
167.
VIII. APPENDICES
APPENDIX A
Detailed Investment Costs
-------
168.
TABLE A-l
\ •
Detailed Plant Investment
Coal Preparation
Case I
No. Cost
Equipment Reg *d. $1000
Coal Storage Bin, 148' x 22' x 16' - 8 Cones C.S. 1 131
Coal Surge Bin, 156' x 12' x 10' - 12 Cones C.S. 1 49
Coal Surge Bins, 18' x 24' x 16' - Double Cone Bottom C. S. 8 16O
Compressor, 1050 ACFM, AP> 120 psi 1 96
Lock Hoppers, 15' D x 32' C.S. 24 444
Gas Drums, 16' D x 40', 350 psig C.S. 2 80
Conveyors, 36" Belt, 400', 300 tph 2 186
Tripper Conveyors, 1O01 4 200
Vibrating Feeders, 83 tph 8 56
Gundlach Crushers, Model 75-2C4R 8 240
Vibrating Feeders, 48 tph 12 84
Gundlach Crushers, Model 50-2C4R 12 324
Collecting Conveyors, 36" Belt, 75' 2 47
Conveyors, 36" Belt, 300', 300 tph 2 150
Distributing Conveyor 440', Multi-tripper points 1 107
Sub-total Major Equipment 2,354
Piling, Foundations, Structural, Electrical, Insulation,
Buildings, Piping, and Instrumentation 1,295
Engineering,. Supervision, Purchasing, Field Expense,
Contractor's Overhead and Profit
Erected Cost
-------
169.
TABLE A-2
Detailed Plant Investment
Gasification
Case I
No. Cost
Equipment Rea'd. $1OOO
Air Heater, 10 MM Btu/hr (incl. air blower) 1 66
Gasifier Steam Superheater - 4,035 ft2 ea. 6 605
Gasifier Steam W.H.B. - 4,920 ft2 ea. 6 442
Gasifier Steam W.H.B. - 3,125 ft2 ea. 2 94
Gasifier Steam BFW Preheater - 2,140 ft2 ea. 2 64
Gas-Gas Exchanger - 3,880 ft2 ea. 12 1,070
Preoxidizers, 6'3" OD x 25' OSS C.S. 12 396
Gasifiers, 27'2" OD x 73' OSS C.S. Shell-Refractory-lined 12 6,558
Regenerators, 24'8" OD x 44' OSS C.S. Shell-Refractory-lined 4 1,800
Acceptor Storage Bin, 30OO Tons C.S. 1 82
Acceptor Lock Hoppers, 10'D x 16' C.S. 2 22
Spent Acceptor Lock Hoppers, 7'D x 12' C.S.
Shell-Refractory-lined 8 75
Regenerator Ash Lock Hoppers, 7'D x 12' C.S.
Sheli-Refractory-lined 8 75
Gasifier Internal Cyclones, Refractory-llned 72 1,040
Gasifier External Cyclones, Refractory-lined 36 792
Regenerator External Cyclones, Refractory-lined 8 2OO
Black Water Pumps, 4250 gpm 2 32
Gasifier Steam Feed Water Pumps, 1463 gpm 2 32
Air Compressors - 68,548 mph ea. 2 8,4OO
Lift Gas Compressor - 4,528 mph 1 102
Claus Gas Recycle Compressors - 34,790 mph ea. 2 608
Acceptor Air Compressor 1 132
Feed Screws 12 72
Char Rotary Feeders 12 131
Char Rotary Feeders 12 48
Acceptor Rotary Feeders 16 240
Acceptor Rotary Feeders 8 32
Ash-Acceptor Rotary Feeders 4 128
Make-Up Acceptor Rotary Feeders 2 16
Acceptor Feed Belt, 150 tph 1 61
Sub-total Major Equipment 23,415
Piling, Foundations, Structural, Electrical, Insulation,
Buildings, Piping, and Instrumentation 15,785
Engineering, Supervision, Purchasing, Field Expense,
Contractors Overhead and Profit 15,600
Start-up Steam Plant (Turn-key) lf20Q
Erected Cost 56,000
-------
TABLE A-3
170.
Detailed Plant Investment
Sulfur Recovery-Solids Disposal
Case I
Equipment
Gas-Gas Exchangers 2,880 ft.2 ea.
Gas-Gas EKChangers 2,610 ft.2 ea.
Sulfur Condensor 6,570 ft.2 ea.
Sulfur Condensor 5,375 ft.2 ea.
Gasifier Steam BFW Preheater 6,200 ft.2
ea.
Reductors 9'3" O.D. x 11'2" OSS C.S.
First-Stage Claus Converters 13'2" O.D. x 16' OSS
Second-Stage Claus Converters 13'2" O.D. x 16' OSS
Coalescers
Acceptor Slurry Stripping Tower 5"6" O.D. x 24' w/Agit.
Sulfur Storage Tank 10'D x 30'
Sulfur Melt Tank
Lock Hoppers
Sump 5000 gal.
Water Pumps 3260 GPM
Sulfur Pumps 80 GPM
Slurry Pumps '460 GPM
Hydroclone Overflow Pumps 3960 GPM
Stripping Gas Compressor
Electrostatic Precipitators
Hydroclones 24" D
Black Water Pond
Sub-total Major Equipment
Piling, Foundations, Structural, Electrical, Insulation,
Buildings, Piping, and Instrumentation
Engineering, Supervision, Purchasing, Field Expense,
Contractor's Overhead and Profit
Erected Cost
No.
Req'd
8
8
16
16
2
4
C.S. 4
C.S. 4
8
-t. 1
1
1
2
1
3
2
4
2
Cost
$1000
207
188
945
774
74
130
264
264
128
36
10
8
22
6
18
8
30
32
4
4
1
102
11,000
-------
Detailed Plant Investment
Coal Preparation
Case III
171,
Equipment
Coal Storage Bin 148' x 22' x 16' - 8 Cones C.S.
Coal Surge Bin 156' x 12' x 10' - 12 Cones C.S.
Coal Surge Bins 18' x 24' x 16' Dbl. Cone Bottom C.S.
Compressor 3230 ACFM AP 150 psi
Lock Hoppers 18'D x 20' C.S.
Gas Drums 15'D x 30' 150 psig. C.S.
Conveyors 36" Belt 400' 300 TPH
Tripper Conveyors 100'
Vibrating Feeders 83 TPH
Gundlach Crushers Model 75-2C4R
Vibrating Feeders 48 TPH
Gundlach Crushers Model 50-2C4R
Collecting Conveyors 36" Belt 75'
Conveyors 36" Belt 300' 300 TPH
Distributing Conveyor 440' Multi-tripper points
Sub-total Major Equipment
Piling, Foundations, Structural, Electrical, Insulation,
Buildings, Piping, and Instrumentation
Engineering, Supervision, Purchasing, Field Expense,
Contractor's Overhead and Profit
Erected Cost
1
1
8
Cost
$1000
131
49
160
130
16
1
2
4
8
8
12
12
2
2
1
222
14
186
200
56
240
84
324
47
150
107
2,100
1,180
-------
TABLE A-5 172.
Detailed Plant Investment
Gasification
Case III
No. Cost
Equipment Req'd $1000
Air Heater Duty 73 MM Btu/hr. (Incl. Air Blower) 1 298
Gasifier Steam Superheater 3340 ft.2 ea. 8 615
Gasifier Steam W.H.B. 3540 ft.2 ea. 8 651
Gasifier Steam BFW Preheater 1730 ft. 2 ea. 2 45
Plant Steam Superheater 2890 ft.2 ea. 6 434
Plant Steam W.H.B. 4300 ft.2 ea. 6 593
Plant Steam BFW Preheater 3220 ft.2 ea. 2 90
Preoxidizers 12.' OD x 28' OSS C.S. ' 16 451
Gasifiers 50' OD x 119' OSS C.S. Shell-Refractory-Lined 16 13,664
Regenerators 35' OD x 38' OSS C.S. Shell-Refractory-Lined 8 2,680
Acceptor Storage Bin 3000 tons C.S. 1 82
Gasifier Internal Cyclones Refractory-Lined 96 1,387
Gasifier External Cyclones " " 96 1,351
Regenerator External Cyclones " " 96 1,412
Condensate Pump 1320 GPM 2 66
Blow-down. Pumps 80 GPM 2 12
Gasifier Steam Feed Water Pumps 1230 GPM 2 22
Plant Steam Feed Water Pumps 80 GPM 2 16
Air Compressors 70,498 MPH Steam Turbine Drive 80,000 HP ea. 2 5,120
Lift Gas Compressor 6420 MPH . 1 480
Acceptor Air Compressor 1 42
Feed Screws 16 96
Ejectors 16 448
Char Rotary Feeders 16 128
Char Rotary Feeders 16 128
Acceptor Rotary Feeders 32 480
Acceptor Rotary Feeders 16 48
Char Rotary Feeders 8 128
Acceptor Feed Belt 150 TPH 1 61
Acceptor Rotary Feeders 8 24
Sub-total Major Equipment . 31,052
Piling, Foundations, Structural, Electrical, Insulation,
Buildings, Piping, and Instrumentation 22,748
Engineering, Supervision, Purchasing, Field Expense,
Contractors Overhead and Profit 21,500
Start-up Steam Plant (Turn-key) 1,200
Erected Cost 76,500
-------
173.
Detailed Plant Investment
Sulfur Recovery - Solids Disposal
Case III
Equipment
Gasif ier Steam BFW Preheaters 2850 ft. 2 ea.
Plant Steam BFW Preheaters 4125 ft. 2 ea.
Sulfur Condensers 5000 ft. 2 ea.
Water Coolers 200 ft. 2 ea.
Reductors 13'9" OD x 16'6" OSS C.S.
Claus Converters. 17 '3" OD x 20 '9" OSS C.S.
Coalescer 10' OD x 14' OSS C.S.
Acceptor Slurry Stripping Tower 5*6" OD x 24' w/Agit.
Sulfur Storage Tank 10' D x 30'
Sulfur Melt Tank
Sump 5000 Gal.
Water Pumps 4520 GPM
Sulfur Pumps 80 GPM
Slurry Pumps 470 GPM
Black Water Pumps 4370 GPM
Hydroclone Overflow Pumps 4050 GPM
Stripping Gas Compressors
Electrostatic Precipitators
Hydroclones 24" D
Black Water Pond
Sub-total Major Equipment
Piling, Foundations, Structural, Electrical, Insulation,
Buildings, Piping, and Instrumentation
Engineering, Supervision, Purchasing, Field Expense,
Contractor's Overhead and Profit
No.
Req'd
2
4
16
2
4
4
4
1
1
1
1
3
2
4
2
2
2
8
4
1
Cost
$1000
57
182
720
3
142
224
52
36
10
8
6
21
8
32
32
32
20
1,660
24
540
3,809
3,391
2,800
Erected Cost 10,000
-------
APPENDIX B
-------
174,
APPENDIX B
Discussion of Gasification Design Basis
A. Kinetic Restraints
Design conditions chosen for this feasibility study were noted in
Section III-B-1. These conditions were based on information for similar
processes which have been accumulated by Consol over a period of 22 years.
However, no data are available for the exact operating conditions required
in the fuel gas version of the C02 Acceptor Process. A major objective of
Phases II-A and II-B of the contract is to acquire kinetic data to provide
a firm design basis. An outline of the design conditions which are affected
by kinetics- of the various reactions is shown in Table B-l for Case I.
B. Effect of Variations in Design Conditions
Since no data are yet available for the exact operating conditions
required by the process, a major effort to optimize the gasification condi-
tions would be premature. However, a limited study was done to guide the
choice of design conditions which were used in this feasibility study.
In the guideline study, the criterion was the overall heat rate
for an integrated power station using the supercharged boiler cycle. The
steam cycle conditions were substantially those shown in Table 24. However,
for simplicity, no regenerative feed-water heating was employed. All feed-
water heating duty was supplied by the expander exhaust gas.
The base case was the same as Case I, described elsewhere in this
report, except for the method of feed-water heating. The required heat and
material balance calculations are extremely tedious and a computer program,
described in Appendix D, was developed to perform the calculations. Results
are summarized in Table B~2.
As would be expected, higher inlet gas temperature to the turbine
brings a significant reduction in heat rate. Therefore, the 180O°F tempera-
ture level was chosen for this study. The impact of the other variables
which were considered is relatively small. Detailed design and cost estimates,
which are beyond the scope of this work, would be necessary for true optimi-
zation with respect to power costs which are influenced by vessel volumes of
the gasifier and boiler, by frame sizes for the expanders and compressors,
etc.
Another process restriction which probably would reduce the power
cost is to preheat the process air by exchange with the hot gas streams.
However, this restriction was not considered in the guideline study.
-------
TABLE B-l
Outline of Gasification Design Conditions Affected
by Reaction Kinetics - Case I
Preoxidation
% preoxidation( 1)
Temperature, °F
Coal retention time, minutes
Gasification
Fixed carbon gasification rate, atoms
C gasified /a torn C in bed/min.
Methane yield
Fixed carbon burnoff level
j driving force for acceptor
reaction, atm.
driving force for acceptor
reaction
Regeneration
driving force for calcining acceptor
driving force for sulfur rejection
Assumed Design Conditions
Severity adequate to prevent
caking in gasifier
8.5
SCO
10
75 x 10~4 at 17OO°F
,O42 mols CH4/atom total
carbon gasified(2/
65%
0.7 at mid-point of bed
Equilibrium value for
CaO + H2S = CaS + H2O
(96$ removal of sulfur)
l.O atm in exit gas
3O% of equilibrium SO2
pressure for 3/4 CaSO4
+ 1/4 CaS = CaO + SO2
Process Variable to be Changed
to Increase Design Value
Increase severity by increasing
temperature and/or retention time
Increase bed temperature and/or
inlet steam rate
Not independently controllable.
Revise correlation on basis of
Phase II-B data
Increase air to gasifier
Increase system pressure
Decrease steam content of product
gas and/or bed temperature,
Increase bed temperature
Increase CO partial pressure by
decreasing air/(fuel carbon) ratio
(i) 100 (ib 02 consumed)/(lb MF coal).
(a) Based on correlation of data in work described in reference 4.
M
-J
01
-------
TABLE B-2
Effect of Process Variables on Overall Station Heat Rate
Case I with Supercharged Boiler Cycle
Case
$ Fixed Carbon Burnoff
Fixed Carbon Gasification Rate,
atoms C gasified/atm C in
bed/min x 1O4
Interceding on Total Station
air
% Excess Air to Boiler
System Pressure, Atmospheres
Expander Inlet Temperature, °F
Acceptor Make-up Rate, $
Overall Heat Rate,
% Difference from Base Case
Base
65
75
1O
51
75
102
65
56
75
102
75
^
^
30 ^ ^
1 S
O.5
-^ 1C -1 Ci
•* 1 ftrv^ i ^?oo ^ . .1 o'v*
» l.O
50
15
O.5
+ .71 -.41 + .38 -.28 -.56 +2.42 +3.OO +.31 +.13 -1.54
Oi .
-------
APPENDIX C
-------
177,
APPENDIX C
Data and Procedure Used for Sulfur Removal and Recovery
A. Introduction
Presented here is a body of self-consistent thermodynamic data for
use in equilibrium and heat balance calculations.
When catalyzed by an active alumina, the reductor and Claus
reactions approach equilibrium closely. At the higher temperatures (1600-
1950°F) involved in the sulfur-removal reactions, previous operations of the
continuous. C02 Acceptor Gasification Unit have shown, qualitatively, that
equilibrium also is approached closely. Apparently, catalysis is by the
coal ash and/or the sulfur acceptor. Thus, equilibrium calculations provide
a reliable guide to actual plant performance.
A computer program was developed to perform the extremely tedious
equilibrium calculations over the temperature range of 250 to 2000°F.
B. Method for Calculation of Equilibrium in the
Reductor and Claus Reactors
A system containing reactants and products for a specified tempera-
ture and total pressure can be defined in several ways. The one chosen for
this work is described below.
Equilibrium is assumed for the following reactions:
CO + 1/2 S2 = COS (l)
H2 + 1/2 S2 = H2S (2)
2 H2 + S02 = 1/2 S2 + 2 H2O (3)
2 CO + S02 = 1/2 S2 + 2 C02 (4)
3 S8 = 4 S6 (6)
S8 = 4 S2 (8)
These six reactions, in conjunction with the four elemental balances listed
below, are sufficient to define the system.
An eleventh possible component, CS2, has been neglected in this
work because preliminary calculations showed that, at the partial pressures
of the reactants and products involved in CS2 production in the processes
considered here, a negligible amount of CS2 can be formed in the sulfur-
removal portion of the process (1400 to 2000°F). Furthermore, at reductor
and Claus reactor conditions (lOOO to 25O°F), the kinetics of the reaction
2 COS = C02 + CS2 are very poor.(13)
-------
178.
[S] = 2 S2 + 6 S6 + 8 S8 + S02 + H2S + COS (9)
[C] = C02 + CO + COS (10)
[0] = 2 C02 + 2 S02 + CO + H20 + COS (ll)
= H20 + H2 + H2S (12)
where: The chemical symbol for a substance equals the number of mols of
that substance in the product gas,
[s] = atoms total sulfur in the feed gas.
[c] = atoms total carbon in the feed gas.
[o] = atoms total oxygen in the feed gas.
[H] = mols total H2 in the feed gas.
Equilibria in the six reactions can be expressed as follows, where the various
K's are equilibrium constants obtained from experimental data or calculated
from modern free energy values, and P is the system pressure, atm:
1/2
COS ' *" x
CO (S2)
V* (f) . (13)
H
ic
Ke -
(S2)
(18)
In the above equations, W equals the total mols of product gas, which can be
expressed as:
w = [c] + [H] + i + s8 + s6 + s2 + so2. (19)
where; I = mols inert gas (usually N2).
Solution of equations (9) through (19) is accomplished as follows:
1. An initial value for S2 is assumed, using the data in Table C-I
as a guide. The system is not sensitive to small changes in
the value of W and the initial value is obtained by substituting
the term [s]/3, for the terms, S8 + S6 + S2 + SO2, in equation
(19).
-------
179.
2. Equations (10) through (18) are solved algebraically and the
sulfur balance {equation (9)} is determined.
3. The value for S2 is then changed in proportion to the
magnitude and direction of deviation from 100$ closure of
the sulfur balance. W is calculated from the values of
S8 + S6 + S2 + S02 just obtained in Step 2.
4. Steps 2 and 3 are repeated until the sulfur balance closes
to within 100.00 + .05$.
For multi-stage Glaus reactor calculations, a trial and error
method involving the reactor temperature is required. For the first stage,
a reactor temperature is assumed and equations (9) through (19) are solved
as described above. The partial pressure of elemental sulfur then is
calculated and compared with the vapor pressure of liquid sulfur. The
temperature then is changed and the calculation is repeated until the sulfur
partial pressure is about 20% less than the vapor pressure. This procedure
was adopted in order to allow for errors in estimation of the heat exchange
requirements for the feed gas. The Glaus reaction is moderately exothermic.
Also, condensation of liquid sulfur on the catalyst destroys its activity.
Between stages, the product gas is cooled to condense substantially all the
elemental sulfur. For the second and third (if needed) stage calculations,
the elemental sulfur formed in the preceding stage is subtracted from the
value of [s] in equation (9) and the entire process is repeated, except
that the final temperature corresponds to a sulfur partial pressure of about
10$ less than the vapor pressure. The closer approach to saturation
temperature was used because the heat of reaction is much lower in the
second and third stages. The complete calculation for a three-stage process
requires, typically, about 2OO iterations of Step 2, above.
C. Discussion
As noted in the description of the sulfur recovery section for
Case I (Section III-B-2), the kinetics of reactions (l) through (4) become
very slow at Glaus reactor temperatures (<6CO°F). The actual Glaus reactions,
which kinetically are rapid when catalyzed by alumina, can be expressed for
diatomic sulfur as;
2 H2S + S02 = 3/2 S2 + 2 H20,
and
2 COS + S02 = 3/2 S2 + 2 C02.
Equilibrium constants for the above Claus reactions are given by K3/(K2)2
and K4/(K1)2, respectively. Thus, a system at equilibrium with respect to
reactions (l) through (4) also is at equilibrium with respect to the Claus
reactions.
Case I is unique in application of the computer program to the
sulfur recovery section. In this case CO and H2 are brought into the Claus
reactors with that portion of the gas which bypasses the reductor. This CO
and H2 is treated as an inert gas in the calculations.
-------
180.
D. Sources of Data, Gas Reactions
Nearly all of the free energy values, heat capacities at zero
pressure, and heats of formation were taken from the JANAF tables.\14'
The equilibrium constants were calculated from the free energy values.
Exceptions were;
1. The equilibrium constants for reactions (6) and (8) were
taken from the experimental data of Preuner and Schupp.(15/
2. The heat capacity and heat of formation for S6 were calcu-
lated from the JANAF data for S8 and the equilibrium
constants measured by Preuner and Schupp for reaction (6).
The vapor pressure and heat of vaporization of liquid sulfur were
taken from West(16) and the heat content of steam above liquid water were
values given in the Keenan and Keyes steam tables.
E. Presentation of Data. Gas Reactions
The effects of temperature and sulfur partial pressure are shown
in Table C-l. The first entries for each temperature through 600°F
correspond to the vapor pressure of liquid sulfur. The data were calculated
from equations (17) and (18), taking the sum of the mol fractions of S8, S6,
and S2 as unity.
Equilibrium constants are shown in Table C-2 and numerical values
of the constants as a function of temperature are in Table C-3. Heat
capacities are in Table C-4 and the mean heat capacities above 60°F, derived
from the Table C-4 data, are in Table C-5. The heat content of steam above
liquid water at 60°F also is given in Table C-5.
Heats of formation and heats of reaction at 25°F are given in
Table C-6. Equations for the vapor pressure of liquid sulfur are shown in
Figure C-l, which is a plot of the equations. An equation for the heat of
vaporization of liquid sulfur also is given in Table C-6.
F. Data for Solids Reactions
Equilibria for the pertinent solids reactions involving calcium
compounds are given in Table C-7. Mean heat capacities above 6O°F are
shown in Table C-8. Heats of formation and heats of reaction are shown in
Table C-9.
Sources of these data primarily were experimental measurements
of the thermodynamic properties of CaC03, CaO, CaS and CaS04 which were
supported privately by Consolidation Coal Company during the 196O's.
-------
181,
G. Calculations for Sulfur Rejection in the Regenerator - Case I
Equilibrium for some pertinent sulfur reactions involving lime
are plotted in Figures C-2 and C-3.
To calculate the exit gas composition in the regenerator, equili-
brium is assumed for the following reactions;
2 CO + S02 = 1/2 S2 + 2 C02 '
CaS + H20 = CaO + H2S
CaS + CO2 = CaO + COS
H2O + CO = H2 + C02
These four reactions, in conjunction with the four elemental balances for
C, 0, H, and S are sufficient to define the system.*
In a manner analogous to the formulation of equations (9) through
(19), a system of nine simultaneous equations was set up. Solution of the
equations is accomplished as follows;
1. An initial value for the mol ratio C02/C0, is
assumed.
2. The equations are solved algebraically and the
sulfur balance is determined.
3. The value for the CO2/CO ratio is changed in
proportion to the magnitude and direction of
deviation from 1OO$ closure of the sulfur
balance.
4. Steps 2 and 3 are repeated until the sulfur
balance closes to within 1OO.O ± O.1$
The above procedure is incorporated as part of the computer program
for the process heat and material balances described in Appendix D.
* Note that equilibrium is not assumed for the entire system. Rather, the
assumption is made that the kinetics of the above reactions are rapid
compared with those of back-reactions such as;
CaO + 3/4 S2 = CaS + 1/2 S02.
CaS04 + S2 = CaS + 2 SO2.
-------
182.
TABLE C-I
Effect of Temperature and Sulfur Partial Pressure
on Distribution ot Sulfur Species
Mol Fraction
S,
.8579
.8472
.B08O
.7567
.7022
.8249
.8OOO
.7473
.6916
.6230
.7566
.7158
.6501
.5841
.5O72
. 69O3
.5903
.S2O4
.4410
.3682
.6O79
.5651
.4937
.4133
.34O1
.2623
.1958
.4614
.4146
.34O4
.2612
.1934
.1279
.08O1
.3287
.2822
.2117
.1423
.0910
.0499
.0264
.217O
.1751
.1164
.0670
.O369
.0176
.0084
.0016
.0977
.O572
.0288
.01 43
.0063
.OO2 9
.OOO5
.OO5S
.0023
.0010
.OOO4
.0001
.OOOO
s.
.1424
.1527
.1917
.2426
.2963
.1749
.1997
.2520
.3068
.3735
.2428
.2831
.3474
.4108
.4819
. 3O8O
.4045
. 4693
.5374
.5906
.3873
.4274
.4913
.5558
.6019
.6273
.6163
.5159
.5512
.5955
.6175
.6018
.5380
.4431
.5949
.6075
.5997
.5448
.4582
.3398
.2368
.597O
.5726
.SOI 2
.3899
.2822
.1817
.1149
.0398
.4625
.3561
.2423
.1588
.O946
.O571
.O189
.087O
.0523
.0294
.O172
.0055
.OOO7
S,
.OOO1
.OO01
.OOO3
.0007
.0015
.O002
.O003
.OO07
.0016
.0035
.OOO6
.0011
.O025
.0051
.0109
.O017
.0052
.0103
.0216
.O412
.0048
.0075
.0150
.0309
.O58O
.1104
.1879
.0227
.0342
.O641
.1213
. 2O48
.3341
.4768
.0764
.1103
.1886
.3129
.4508
.61O3
.7368
. I860
.2523
.3824
.5431
. 68O9
.8007
.8767
.9586
.4398
.5867
.7289
.8269
.8991
. 94OO
.98O6
.9O75
.9452
.9696
.9824
.9944
.9993
Partial Pressure of
Total Sulfur, atm A
250" F
4.85 x 10-"
3.54 x 10-s
1.24 x 10-*
.396 x 10-*
.142 x 10-"
30O°F
2.33 x 1O"*
1.25 x 10-4
.401 x 10-4
.144 x 10-4
.O482 x 10- 4
400° F
3. 11 x KT*
1.4O x 1O-*
.461 x KT'
.171 x 10"'
.O592 x 1O-1
500* F
2.42 x 10-"
.508 x 10" *
.192 x 10- •
.0680 x 10- *
.O272 x 10- '
6OO°F
9.81 x 1O-"
5.31 x UP2
2. 02 x 1O-2
.726 x 1O-2
.294 x 10-*
.114 x 10-2
.O511 x 10-"
700° F
.1300
.O724
.0294
.O115
.OO517
.OO234
.00124
8OO°F
.1825
.1063
.0472
.0211
.0110
.O06O1
.00378
90O"F
.2764
.1714
. O859
.0448
.0271
.O171
.O118
.OO6O9
1000° F
.3O71
.1749
.1042
.0698
.O475
.0345
.O186
1200° F
.546
.398
.1.87
.216
.120
.O328
Atom Sulfur/Mol Total Sulfur
7.714
7.694
7.615
7.511
7.398
7.649
7.599
7.492
7.377
7.232
7.S11
7.427
7.290
7.148
6.971
7.374
7.160
6.999
6.795
6.572
7.196
7.10O
6.927
6.7O3
6.448
6.O83
5.64O
6.832
6.692
6.424
6.037
5.568
4.919
4.253
6.352
6.123
5.669
5.O33
4.378
3.658
3.106
S.69O
5.341
4.703
3.962
3.35O
2.832
2.510
3.169
4.436
3.768
3.142
2.721
2.416
2.246
'2.O7B
2.381
2.224
2.124
2.072
2.022
2.O03
-------
TABLE C-2
Equilibrium Constants
for Gas Reactions
Temperature in "Kelvin
InK = A + BT + CT2 + D/t
Kl CO + 1/2 S2 = COS
K2 Ha + 1/2 S2 = H3S
K3 2H2 + SO2 = 1/2 S2 + 2H2O
K4 2CO + S02 = 1/2 Sa + 2CO2
K6 3S8 = 4S6
K8 S8 = 4S2
K5 2H2S + S02 = 3/2 S2 + 2H2O
A
-15.5992
-1O.3503
-3.3749
-4.9853
-4.8008
-1.4871
-15.7562
-14.3343
21.9713
-33.2782
1.99482
B
.0122924
„ 0006332
-.0027184
-.OOO7315
.O03597O
-.O02161O
.0052624
.0015907
-.O03582
. 132324
.C0895O2
C x 1O8
-8.O1O42
-. 1O012
1.O1362
.18766
-2.26137
. 43958
-2.65427
-.35856
3.5055
-63.97O7
-4.6424
D
12,O17.2
11, 327. O
1O,O39.6
1O,48O.5
15,353.9
14,786. 0
25, 277. O
25,314.O
-15,135.4
-29, 596. O
-4370.84
Temperature Range,
°K
3OO-9OO
90O-14OO
300-9OO
9OO-14OO
3OO-9OO
9OO-14OO
30O-9OO
90O-14OO
300-8OO
3OO-8OO
300-9OO
00
CO
-------
TABLE C-3
Numerical Values of Equilibrium Constants for Table C-2
*g
20O
30O
4OO
5OO
6OO
700
BOO
9OO
1OOO
1200
1400
16OO
180O
2OOO
Temperature in °F
1.
1.
8.
Kl
05 x 1O8
66 x 1O7
O4 x 1O*
73,9OO
10,6OO
2,09O
517
152
5O.3
7.O6
3.18
1.15
.501
.251
K2
1.55 x 109
2.76 x 10s
1.57 x 107
1.60 x 106
2.49 x 106
5.27 x 1O4
1.42 x 104
4,62O
1,750
353
99.6
35.9
15.5
7.64
K3
1.8O x 1O"
1.45 x 1O14
2.26 x 1O12
8.36 x 1010
5.72 x 1O*
6.15 x 10*
9.29 x 1O7
1.82 x 1O7
4.39 x 10°
4.01 x 1O5
6.35 x 1O4
1.38 x 104
3,900
1,340
K4
5.27 x 1021
8.45 x 1018
9.35 x 1O*6
4.35 x 1014
5.62 x 1O12
1.55 x 1O1*
7.55 x 10e
5.73 x 10s
6.16 x 1O7
1.55 x 10e
9.15 x 1O4
9,28O
1,420
296
K6
3.17 x 10~8
3.88 x 1O~7
2.44 x 1O'8
6.62 x KT4
9.81 x 1O~3
.0937
.644
3.43
15.1
188
—
—
'
—
K8
2.03 x lO'28
2.61 x ID" 2e
5.75 x 10"al
1.56 x 1O"16
8.48 x 1O"13
1.26 x ID"*
6.3O x 1O"7
1.25 x IO~4
.O1O8
9.54
K5
.OO0749
.OO19O
.O0914
.O325
.O925
.221
.461
.856
1.44
3.22
6.39
10.7
16.3
22.9
00
-------
TABLE C-4
Heat Capacities at Zero Pressure
Temperature in "Kelvin
Cp = A + BT + CTa + OT3 + E/T8
185.
B
C x 106
D x 10e
E
Temp. Range,
°K
S8
se
S2
COS
H2S
S0a
Na
CO
coa
Ha
H,0(g)
28.51
23.13
8.633
9.678
7.344
5.855
7.098
5.551
5.101
7.219
7.757
.03771
.01852
.000272
.005978
.O01851
.01534
-.001431
.003185
.015568
-.OO0674
.0000003
-25.25
-3.137
-.1903
3.364
-11.002
3. 490
-.8600
-10.238
.6840
3.219
300-800
300-1000
-87,826 300-140O
-123,658
-1.623
2.842
-1.348
48,814
2.552
-16,003
-1.125
\
-------
186.
TABLE £-5
!eat Capqc;it-igQ Above 60°;
Temperatures in °F
Cp - A + BT + CT2 + DT3
Btu/Lb
B
C x 10e
D x 10B
Temp. Range,
°F
S8
S0
sa
COS
H2S
S03
N2
CO
coa
Ha
H»°^
36.903
27.940
7.680
9.718
8.065
9.274
6.941
6.925
8.632
6.910
7.974
.006734
.004678
.O01251
.003033
.OOO916
.O02798
.OOO037
.000105
.002958
.000135
.O00410
-2.598
-.321
-.640
-1.096
.214
-.898
.252
.266
-.844
-.044
.239
60-1000
60-1300
. 12(5 60-2000
.177
-.070
.121
-.058
-.070
.109
.022
-.040
Heat content of steam (l a :m) above liquid water at 60°F.
Btu/Lb = 1021.08 + .48540T - 35.83 x KT6 T2 + 34.46 x 10'° T:>
Btu/Lb = 1039.25 + .41916T + 46.6O x 10"6 T2
220-1000
1000-1600 *
* Can be extrapolated accurately to 20OO°F.
-------
187,
TABLE C-6
Reaction
1
2
3
4
5
6
8
6a
8a
5a
5b
Heats of Formation atl25°C Gases
cal/g mol
sB
Sfi(g)
S2(g)
COS
H2S
SO 2
CO
CO 2
AHf
+24,200
+25,580
+30,840
-33,080
-4,880
-7O,960
-26,416
-94,052
H20, x -57,798
HjjO^v -68,430
Heats of Reaction at' 25° r. r,aa n^^+^ — „
cal/g mol
CO + 1/2 S2 a COS
H2 + 1/2 S2 = H2S
2H2 + S02 = 1/2 S2
2CO + SO j = 1/2 Sa
2H2S + SO2 - 3/2 S2
3SB = 4Se
S8 = 4S2
1/6 S6 = 1/2 S2
1/8 SB = 1/2 S2
2HaS + S0a = 1/2 Se
2H2S + S02 B 3/8 S8
Heat of Vaporization
+ 2H20, x
+ 2002
* 2Ha°(g)
+ 2H20(g)
+ 2H20(g)
of Sulfur
-22,080
-20,300
-29,220
-48,890
+11,380
+29,720
+99,160
+3.1,160
+12,400
-22,090
-25,800
Temperature in °F
Btu/Lb = 211.1 - .33O2T + 434.6 x 10~s Ta - 214.4 x K>-° T3
-------
TABLE C-7
Equilibrium Constants for Solids Reactions
3/4 CaSO4 + 1/4 CaS = CaO + SO2
1/4 CaSO4 + CO = 1/4 CaS + CO2
CaO + H2S = CaS + H20
CaO + 3/4 S2 = CaS + 1/2 S02
CaS04 + S2 = CaS + 2 SO2
CaC03 = CaO + CO2
CaCO3 + H2S = CaS + C02 + H2O
Temperature in "Farenheit
In K =
A
71.5O76
3.O04O
17.5646
-7.22273
-6.55015
-.840165
16.6700
A + BT + CT2
B
-.O3125O
-.000751
-.O092926
. OO25330
.0118713
.OO74O15
-.0018535
+ D/T
C x 1O6
5.79655
. 01932
1.7716
-.36880
-1.86265
-1.28243
.48O912
D
-67,104
4,232
-1,070
15,200
-12,576
-12,932
-13,977.2
Temperature Range,
^F
15OO-2OOO
15OO-2OOO
13OO-2OOO
15OO-20OO
1500-20OO
all
1100-2000
oo
oo
-------
189.
E C-8
CaS04
CaS
CaO
MgO
MgO-CaO
MgO'CaC03
Char Carbon
Ash*
Coal
**
Mean Heat Capacities Above 60° FT Solids
Temperature in "Farenheit
Cp = A + BT + CT2
Btu/lb Mol/°F
A
23.57
10.14
10.98
9.77
20.75
30.64
1.447
.207
B
.00645
.00230
.O0071
.O0081
.00152
.O05O6
. 00291
Btu/lb/0 F
.000029
Heat Content Above
Btu/lb
C x 10s
-.0383
-.502
-.0772
-.0387
-.116
-.193
-.643
-.0034
60° F
Temperature Range,
°F
60-20OO
60-2OOO
-12.2
.1879
.00025
60-800
* Also used for impurity content of acceptor.
** Including products of pyrolysis.
-------
190.
E C-9
Heats of Formation at 25° C, Solids
cal/g mol
AHf
CaS04 -344,090
CaS -113,550
CaC03 -228,280
CaO -151,900
Heats of Reaction at 25°C, Solids Reactions
cal/g mol
3/4 CaS04 + 1/4 CaS = CaO + S02 +63,600
1/4 CaS04 + CO = 1/4 CaS + CO2 -10,002
CaO + H2S = CaS + H2O, , -14,570
CaO + 3/4 S2 = CaS + 1/2 S02 -20,260
CaS04 + S2 = CaS + 2 S02 +57,780
CaC03 = CaO + C02 +42,330
CaC03 + H2S = CaS + C02 + H20, v +27 760
* * (g)
CaS + 3/2 02 = CaO + S02 -109,310
CaS + 2 02 = CaS04 -230,540
-------
-------
10
1000
1100
1200
13OO
14OO
1500
16OO
1700
-------
CO «
0
in u
o2
-J"
- o
1000
100
.001
-------
APPENDIX D
-------
194.
APPENDIX D
Description of Computer Program for Process
_ Heat and Material Balances _
The computer program described herein was developed to be con-
sistent with the design constraints as set forth on pages 21 and 22.
The elemental balance relationships for the preoxidizer-gasif ier
system lead to a set of simultaneous equations which cannot be solved
analytically. For a given set of process conditions,* the equations are
solved by assuming initial values for the following ratios:
CaC03/EC = (mols CaC03 f ormed)/(atoms total carbon gasified).
C/CaC03- = (lb fuel carbon fed to the regenerator)/(mol CaC03 formed).
CH4/£lC = (mols CH4 formed) /(atoms total carbon gasified).
Subsequent steps are shown in Figure D-l, which is a flow diagram for the
computer program. After convergence of the preoxidizer-gasif ier and
regenerator heat balances, and after establishment of the desired partial
pressure driving forces for calcining of CaC03 .and for sulfur rejection in
the regenerator, the following kinetic restraints are calculated and printed:
1. % burnoff of fixed carbon in the gasifier.
2. ^>C02 driving force for recarbonation at the midpoint
of the gasifier. " '
3. Mean gasification rate of fixed carbon.
If any of the kinetic restraints are not at the value chosen for the design
basis, new values for specific input variables, as shown below, are selected
and the entire calculation is repeated.
Reetraint Variable Changed
1. % burnoff input air to gasifier
2 . A^cO system pressure
3. gasification rate input steam flow and/or gasifier
temperature .
In principle, additional loops could have been incorporated in the program to
force convergence on any desired values of the kinetic restraints. However,
the additional commands involved would exceed the allowable core capacity of
the particular time-sharing computer used by Consol.
After the kinetic restraints are satisfied, all data needed for
presentation of the mass and heat balances, such as those in Tables V and VI,
are printed.
A listing of the program is attached.
Necessary input data involving the process conditions are shown on
the first page of the listing of the program. Input data are based
on 100 pounds dry coal fed to the preoxidizer.
-------
FIGURE D-l
FLOW DIAGRAM FOR COMPUTER PROGRAM
195,
GASIFIER
REGENERATOR
-------
196.
Program to calculate heat and elemental balances
for production of low-sulfur fuel gas by the C02 Acceptor
Process. BASIC programming language is used.
00040 DIM C(10»I2>«N(8»2>»Q<10»3>«W<6»2>»FC6»3>»S<6»3>
00041 DIM X(10*2),T(20)»R(8»1)»0(6»1)*I(6»3)
00042 DIM JC6«3>»V(6»4>»Y<4»4>»H<25>»L(25>
00044 PRINT "NEED K'/K?? 1»YES*0«NO"!
00046 INPUT Y9
00047 IF Y9«0 THEN 55
00048 PRINT "K'/K GASIF"!
00050 INPUT XI~
00055 PRINT "NEED INT TYPEOUT??*1«YES 0=NO"!
00060 INPUT ?7
00061 PRINT "NEED FULL KIN?? 1-YES 0=NO"!
00062 INPUT Y8
00065 PRINT TAB<5>!"REGENERATOR"
00070 PRINT "XCO OUT"! '
00075 INPUT C7
00080 PRINT "XBURNOUT"!
00085 INPUT C8
00100 PRINT "MOLS LIFT GAS"!
00105 INPUT R5
00110 PRINT "LIFT GAS T"!
00115 INPUT T(7)
00136 PRINT "NEED DP CALC?? 1«YES*0=NO"!
00137 INPUT 01
00138 IF 01*0 THEN 141
00139 PRINT "DP C02"l
00140 INPUT P2
00141 PRINT
00145 PRINT "PRESS"!
00150 INPUT P
00155 PRINT "GASIF T"J
00160 INPUT T<1>
00165 PRINT "REGEN T"!
00170 INPUT T(2)
00250 PRINT "BTU/LB"!
00255 INPUT H4
00295 PRINT "INPUT SOLIDS"
00300 PRINT nH "J
00305 INPUT HI
00310 PRINT "C "I
00315 INPUT C0
00320 PRINT "N "I
00325 INPUT N0
00330 PRINT "0 "I
00335 INPUT 00
00340 PRINT "S "J
00345 INPUT S0
00350 PRINT "ASH"!
00355 INPUT A9
00360 PRINT "ZACCCPT MAKEUP"!
00365 INPUT Ml
00366 IF P<5 THEN 380
00370 PRINT "ACCEPT ACTIV"!
00375 INPUT X9
00380 PRINT "X IMP IN MGO-CAO"!
-------
00385
00390
00395
00425
00430
00435
00436
00437
00438
00439
00440
00443
00444
00447
00448
00450
00455
00460
00465
00466
00467
00470
00475
00480
00485
00490
00525
00530
00535
00536
00538
00540
00542
00544
00546
00548
00555
00560
00580
00585
00620
00625
00630
00635
00654
00655
00660
00665
00670
00695
00700
00705
00710
00736
00740
00765
00770
M77S
00780
00785
00790
00795
00800
00805
00810
00815
INPUT It
PRINT -X H IN GASIF CHAR-J
INPUT S3
IF P<5 THEN 436
PRINT -INPUT GAS T-J
INPUT T<3)
PRINT TAB<5>J-PREOXIDIZER"
PRINT -X PREOX-J
INPUT «6
PRINT -X 02 UTIL-J
INPUT 87
PRINT -AIR T-J
INPUT T<9)
TC8>BT<9>
PRINT
PRINT -INPUT H20 T-J
INPUT T(6)
IF P<5 THEN 470
G(4)«.3
G(5)».7
GOTO 525
G(2>B*24
G(3)=.23
6<4>s«09
6<5>B.'34
G(6)=.2
PRINT
PRINT -CH4/C-J
INPUT Al
IF P>5 THEN 555
BCl>B4E-5
C«1E6
I0».3
T<3*«TU>
X9-1E-5
GOTO 630
PRINT "CAC03/C-J
INPUT BCD
PRINT -LB FUEL C/CAC03-J
INPUT C
PRINT -MOLS INLET GAS-J
INPUT 10
PRINT -MOLS H20 TO GASIF"J
INPUT 01
PRINT
PRINT -HEAT LOSS GASIF-J
INPUT LCD
PRINT -HEAT LOSS RE SEN "I
INPUT L2
PRINT -X INPUT C REJECTED TO PRODUCT"*
INPUT Z3
PRINT -NEED FULL TYPE?? 1»YES 0*NOMJ
INPUT S9
K5»EXP<71«5076-*03125*TC2>+5*7965SE-6*T<2>«2-67104/T<2»
K4«EXP< 17. 5446-9.2926E-3*T(l»1.7716E-6*T(l)t2-1070/T(l»
Kl-EXP(-1.12894-9.33l7E-4*T(l)*i.55385E-7*T»1615*EXP(44*096-9*084E4/T)
X(2»I)=18.52*EXP<29.179-6.011E4/T)
»0S*EXP< 31 • 548-6*499EVT)
f.4*EXP( 10. 738-2. 212EVT)
41.8*EXPC39.587-8.I5SE4/T)
•491*EXP(4.6053-.9487E4/T)
•0S*EXP<9.99S-2.059E4/T$
19V.
X(6*I)
X(7*I>
-------
00820
00825
00885
00890
00895
00896
00898
00900
00905
00910
00915
00920
00921
00922
00925
00930
00935
00940
00945
00950
00951
00952
00975
00980
00982
00990
01035
01055
01056
01057
01058
01059
01060
01095
01096
01100
01105
01110
01115
01120
01125
01130
01135
01140
01145
01150
01155
01160
01165
01170
01175
01180
01185
01190
01195
01200
01205
01206
01207
01208
01209
01210
01215
01880
01225
01230
X (9, I >=f.4*£XP(9. 7766-2.0 14E4/T)
X(10,I>=.7*EXP(11.17-2.301E4/T>
FOR I-1T02
198..
P»EXP«-4i.5964+.0041409*T-.3279E-6*T»2+43498/T>/-1.9872>
NEXT 1
FOR I-0T03
C(l»n*7.976+4.374E-3*TCI>-*4424E-6*Tt2
C(2*I>«6.868+.d045E-3*T(I)+.0045E-6*T(I)t2
C(3«I> «6.843+.4645E-3*T-•3579E-6*T"6.410+*5811£-3*T(I>-.007E-6*TU>«2
NEXTI
FOR I«1T03
C(6*I)»1026* 53*.438232*T(I)*35.7628E-6*T(I> 12+2«15368E-9*T(I> 13
C(7*I)».207*.03f7E-3*T(I)-.0034E-6*TCI)t2
C(8»I)»1.447+2.91lE-3*T(I)-.643E-6*T(I)t2
C(9»I)-21. J6-M.083E-3*T(I)
C( 10» I) -3L31 + 4. 333E-3*T( I >
NEXT I
M(7)»9.274*.002798*1(2)-.898E-6*T(2)t2+.l21E-9*T(2)»3
M(8)«7;68+.00125l*T<2>-.64E-6*T<2)t2+.l26E-9*T<2)t3
Cl«l026.53+.438232*T(6)+35.7628E-6*T(6)t2+2.15368E-9*T(6)t3
I»l
R0«Ml/(2/X9-2.34+l.l7*Ml)
Z(1)=C0*23/1201
G2»Hl/2.016-.0705+(G(6)+G(2)+2*G(l))*I0-T3
W8=W6/32
IF W7i0 THEN 1060
W9»W6/32*(100/W7-1>
B4»11.76*W6/W7
G3=00/16+(G(3)+2*G(4»G(6))*I0+2*(W8+W9) + T4
A4-AI/((1-A1)*(1-B(I))-A1*BCI))
F3»F3*1
B1»1+B(I)*(1+C/12.01)/(1-B(D)
G4"C*<1+A4>*BCI>*H3/<201*6*<1-B(I>»
G9-C12.01*Z(1)+A9)*H3/201.6
G5-G3+D1
66»G2+D1-G5
G7=G4-2*B(I)*(l+A4)/( 1-8(1 »*8*A4
G8=«(Gi-Z(l>>/(Bi*(l+A4»
S1«G6-67*G8+G8-G9
S2=G5-G8-2*B(I)*(l+A4)*G8/(!-B(I))
S3=»S1+K1*(S2+G8)
C9°K1*S2*G8
IF ABSO-K1X.001 THEN 1185
IF (S3t2*4*C9*(l-KI»>0 THEN 1175
PRINT II-NEG"
PAUSE
N<4«I)»<-S3*SQR
GOTO 1190
N(4,I)-C9/S3
N(3,I)«G8-N(4,I)
N(e»I)«G6-6?*68+G8-G9+N(4,I)
N<6«I>«65-G8-N<4»I)-2*B*<1+A4>*G8/<1-B*6C5>*I0*B4
T3*N(6»1>/K4
T4-S0/32.06-T3
TS-T4/CT3>T4>
M(n»(N<4*I)*N(3»I»*5 THEN 1326
I0»D1*M(1)/(50*(N(3»I)+N(2*1»-N(6*1 ))
FOR K-1T06
6»N(K*1)/M(1)
-------
01235
01326
01327
01328
01332
01345
01405
01410
01415
01420
01425
01430
01435
01436
01437
01540
01545
01550
01555
01560
01565
01570
01575
01585
01590
01595
01615
01630
01640
01670
01675
01705
01720
01722
01724
01725
01745
01750
01755
01760
01765
01770
01775
01780
01785
01790
01795
01800
01815
01818
01819
01820
01825
01835
01840
01850
01855
01870
01880
01885
01955
01960
01961
01965
01975
01980
NEXT K
05»1I.76*W6/W7+W8+U9
06-M3*05/(!-M3>
07»H<8>*06*f8.02
e<3»l>=CI8.02*CDl-06)-6.383>*Cl
T«TU>-60
R<8,1>=0
FOR K-1T05
R(K*l>-I0*6*C*-60>
R<8»1>»R<8»1>+RCK»1>
NEXT K
R=6<6>*C<6»3>* 18-02*10
R(8»1>*R<8»1>+R<6«1>
B7»B4/.79*-6fl)*<6»41+.5811E-3*TC8>-.007E-6*T<8>»2>
R(8»l J«R(8*D*B7*07
0<7»I)»A9*C<7»I>*T
E(I)=-0
FOR K-IT05
S(K»I)«N*C(K»I)*T
E*E(I>+S
NEXT K
S<6»I>»N<6»I>*C<6»I>*18*02
199.
I4»H3*C12«01*ZC1>+C*UC1>+A9>/(2*016*C100-H3»
Q(8»l>aI4*CC2»I>*T
0(5»I)«(ZC1)*C*UU)/12.01)*C(8»I)*T
W(1»I)«N(3»I>*121760-»>NC1*I>*383020>N(2»I>*123180
U(2» I >-(C*U( I )/l 2*01 +Z(1»*<178540-4*T< I»+ 125190*1 4
W°U(l»I>+7750+T3*242000
W(0»I)»W(1,I)*W(2»I)-W(5,I)
B(0>»T4* 196540
B5"G(3)*I0* 121 760*6(2)* 123180* I0*I0*G<1)*383020
W(0»I)=W(0,I)-B5*BC0)
W<6#I>=E»U(6»I>+769
F(3»I)«CC9*2)*(T<2)-60)-C(9*I)*T
15-11*96. 4/U00-ID
F(5/I)=«I5*(C(7,2)**«F(3*I>+FC5»I»/R0+F(l»I»-U
Z1«jN(3»I)iN<4»I>lN(6»I)|M(I)IN(5»I>
PRINT
IF V<50 THEN 1818
PAUSE
IF Y>0 THEN 1820
IF ABS(Z1X.05 THEN 1955
IF ABSCZ1X.0002 THEN 1955
IF V>50 THEN 1785
IF ABSCZ1X.5 THEN 1870
B
V«V*i
60TO 980
B(I)-(1-Z1/2)*B
-------
01985
01990
01991
02000
02005
02010
02011
02015
02016
02020
02021
02022
02023
02024
02026
02135
02140
02145
02146
02150
02155
02156
02160
02165
02166
02167
02170
02171
02172
02173
02174
02175
02176
02177
02178
02180
02305
02315
02316
02317
02320
02330
02340
02350
02355
02365
02370
02371
02372
02375
02380
02381
02382
02383
02386
02387
02388
02390
02395
02400
02405
02406
08408
02409
02410
02411
H2»P*N(6*1)/M<1>
C3«P*N<4*1)/M(1)
IF S9»0 THEN 2024 ZOO.
K7«EXPC-5»41543-9. 53865E-3*TNCl*l)4>N<6*l>/N(3*l)/N<2*l)t3*/P>t2
PRINT MKVK"JDEC«N7/K7*i0*2)
PRINT "Ar"jDEC+N<3*1)+N<4*1)+UC1)-I0*<6<3>+G<4))
PRINT "CAC03/C-JDECCIK 1)/Z7»12*4)1 TAB(20)I"CO"!DECCN(3*1>*26*4)
PRINT "C QASlF"jDEC»TABC20>JnC02MJD£CCN(4*l)*26*4>
PRINT "C TO REdEN**lOEC((C*U(l>/12*01)» 13*4)1 TAB(20)INH20**I
PRINT DEC*26*4)
PRINT MSBNJDEC
CS*P*N<3*1)'M<1>
C6*P*NU*1)/M<1>
IF S9-8 THEN 7000
R2»X<1*I)*>/<1+XC3*1>*H2+.2*C4+X<4*1)*C5)t2
H(12>*X(5*1)*
R3«H(l2)/(l*X(7,l)*H2*X(8,l)*C4+X<9»l)*C5*X<10*l)*C6)t2
PRINT -RC MIDEC(R2*14*1)ITAB(20)I-OELTA HMIOEC(W(0*1>*37*0)
PRINT **RCH4 "I DEC CR3»M, 1)1 TAB (20) I "GAS SENSM«DEC(E( 1>*769*37*0)
D0=W9*.79*P/<11.76*W6/U7)
Yl«W9*f*/(11.76*U6/W7/.79*Dl)
PRINT MAOLS 6ASMlDEC(M(l)*14*3)lTAB(20)lMACC OUTYMI
OSC
-H60* CAC03-IOEC
nCAS/M60.CAOMfOEC(T4*R0/X*14*4)l
TAB(20)fAIRN2**lOEC(ll»76*W6/W7*28*4)
PRINT
PRINT
PRINT
PRINT
PRINT
PRINT
PRINT
PRINT
PAUSE
T»T<2)-60
T2»T(l>-60
R4»0
U(0)-0
I6»<0<0>+U<1»/R0*M1/100
01-(U(|)-I6*R0J*(C(1(J*2)*T-C(10*1)*T2)
03»(U(l)/R0-I6)*I5*(C(7,2)*T-C(7,l)*T2)
Q5«(U(I)*C1/R0-1)-I6*(1-R0))*CCC9»2)*T-C<9»1>*T2)
06-*<75215.8*l.l6669*T(2)-.0085774*T<2)t2)
Q8*I6*(46000+C<10*2y*T)«I6*CC4«2)*T
09»I6*I5*C(7*2)*T
U2*(U(1))*C(4*2)*T
U3»2*I6*C(4«2)«T
04»T4*R0/X*I6
T9»T4-04
S4«T9
IF P>5 THEN 2390
M3-«256/CP*14*7)
80 TO 2395
M3».949/(P*14.7)
E|-C8*UO)*C/1201+U<
E2-e*I6*(2-R0)-»-2*U( 1 )-3*T9-I4
A2«C7/(|fl0-C7)
IF A2>0 THEN 2412
L/< 1
N(4*2)«E1-*>LC15>
N(3*2)M
GOTO 2490
8809 5/El*<2*El-E2»
-------
02412
02415
02422
02485
02430
02435
02440
02455
02460
02465
02470
02475
02480
02485
02486
02490
02495
02500
02505
02510
02515
02520
02541
02542
02545
02550
02555
02556
02560
02565
02570
02575
02580
02585
02590
02595
02596
02600
02602
02604
02605
02606
02607
02608
02609
02610
02611
02612
02613
02614
02615
02617
02618
02620
02622
02624
02626
02628
02635
02640
02645
02646
02647
08648
02650
02651
IF B>0 THEN 2420
A«10
E3-E1/<1*A) 201
L<15)-R5*(1-M3)/(1/A*1*1.88095*(2*1/A-E2/A/E3»
N(3»2)-E3*LO5)/A
B2»((A*E3)*L(15))*(1*1.88095*(2*1/A-E2/(A*E3)))
Z9»CB2-N(3»2>/A2)/B2 " ' "
IF ABSIZ9X*00001 THEN 2485
A»A*
B-B*I
IF B<50 THEN 2420
PRINT "B"JZ9
PAUSE
N(4»2>»A*N(3»2>
B»0
L(12)»N(3»2)*L<15)/N(4,2)
A3«(2*N(4,2)*N(3.2)-L<12)-2*L(15)-E2)/.42
L(I6)"(*79*A3*L(15»/(N(4»2>*L(15»
N(6»2>«r4*R5*M3*A3*M3/(l'M3>
N(5»2)-.79*A3*L<16)
V3»N(3»2)*N(4*2)*N(5,2)*N<6,2)*T9
L(13)-R5*M3
IF P«5 THEN 2545
GOTO 6000
FOR I-7T08
H*U(1)*C/I2*01*(178540-4*T(2)>
V4«N(3»2>*121760
U1-L<12>*121760
IF P<5 THEN 2620
LO5)- ^ '
N<6»2)"H3*A3/(|-M3>*M3*R5-H0-S7*I4
R7»(N<3.2>*2**S4>*N(6,2>*S5*T9>
B3»A3*<»42*M3/<1-M3»*2*I6*(2-R0)*2*U*(T<9>-60>
C(3»7)»6.843*.4645E-3*T(7)-.0027E-6*T<7)»2
C(4,7)-8.891*2.283E-3*T(7)-.3579E-6*T<7)te
C<5.7)«6.41*.58llE-3*T(7)-.007E-6*T(7)t2
IF P<5 THEN 2650
U8-*C(4,7))
GOTO 2652
H(18)»R5*(l-M3>*(T(7)-60)/(V3*(l-M3»
U8-H(18)*(N<3*2)*C(3»7)*N<4.2)*C(4»7)*N(5»2)*C(5,7»
-------
02652 IF P>5 THEN 2680
02654 U9-T9*196508
02680 V8-A9*C(7,2)*T
02685 U7«<|-C8/100>*U(1>*C/12.01»C<8»2>*T
02690 M9»0
02695 FOR K-2T05
02700 M(K>>N*C(K«2>*T
02705 M9*N9+M(K>
02710 NEXT K
02715 M<6>»N(6»2>*H<6>*18»02
02716 IF P>5 THEN 2720
02717 M9«M9+MC6>
02718 GOTO 2725
02720 M9»M9+MC6)+T**S4+MC8>*S6>
02725 U0-M9+U7+Q7+L2-U8-U6-M7-V7+V8-U2-U3
02730 H6«V5+U|-V6-V4*U9
02735 Z4»/U0
02736 F4-F4+1
02737 G»0
02740 IF Y7«0 THEN 2755
02745 PRINT Z4JM9JU0IH6JQ7IC
02750 PRINT N(3»2>JN<4»2>IN(5»2>JN<6«2>JV3
02751 PAUSE
02755 IF Y< 50 THEN 2765
02760 PAUSE
02765 IF ABS(Z4X.0001 THEN 2810
02770 IF Y>50 THEN 2745
02775 IF ABS<*2 THEN 2795
02780 C«<1*Z4/2>*C
02785 Y«Y*1 "
02786 GOTO 4000
02795 C*<1+Z4>*C
02800 Y-Y+1
02801 GOTO 4000
02810 Y«l
02812 IF Y7=l THEN 2985
02822 IF 01«0 THEN 2826
02824 GOTO 8140
02826 PRINT "VI"! V1*"FI"JFU "HMJHJ"S"JS
02827 PRINT "ITR"lF3+F4i-GASIFMJF3JMREGEN"f F4
02828 V1»Y»F3*F4*H«S«F1=0
02840 IF S9-0 THEN 1985
02841 PRINT TAB(S>JMRE6EN"
02845 U5«N(4*2>«P/V3
02850 PRINT MCMID£C(C»5»3)ITAB(20>IMDUTYMIOEC(H6»32»0>
02851 02»T8/CV3-T6>*100
02855 PRINT" "DP C02"IDEC( (P(2)-W5)* 12*3) J TAB(80)I °'XCO"I DECC02* 28*2)
02880 PRINT "MGO.CAO*** DECCX/R0, H, 3>
02890 PRINT "MOLS GASMIDECJTAB(20)IMDP S02M«
02891 PRINT DEC(K5-S4*P/V3*30»4)
02912 Y6«N(1»1)*383020
02914 N2-N(3»1>*121760
02916 N3-N<2,1>*123180
02918 H(2J)"Y6*N2*N3*7750*T3*242000
02934 PRINT "T"IDEC/<100*H4>>»14»4>
02951 IF 01<1 THEN 2970
02955 PRINT -STEAM CONV***
02940 N4»CD1-N<6»1»/D1
02965 PRINT OECCN4* l4»4>lTAB(20>iMR0**lOEC
02970 N8*X*M|*/<100*R0>
02975 PRINT "LB STONEMIDEC(N8»14*2>l
02976 B4»162S00*N8/W(1«1)
02977 PRINT TAB(20)lMC/MMMfOEC
-------
02980 GOTO 1975
02985 PAUSE
04000 IF P«S THEN 980
04002 X5»2*N<4*2>+N<3«2>+N<6*2>+2*S4-2*LC15>-L<12>-L<13>
04004 X6«N<6»2>*H0+S7-L<13>
04006 X7«NC4»2>*N(3«2»S5-LC15>-L<12>
04008 N<2»2>».67*H0
04010 NC6»2>»X6-NC2»2>
04012 N<4»2>«X5-NC6*2>-X7
04014 N(3»2>«X7-N<4»2>
04016 I0»N<2»2>+N<3»2>+N(4»2>+NC5»2>+N<6»2>-LC16>
04017 M(5»2)»N(5»2)-L<16)
04018 FOR K-2T06
04020 6(K)-N(K»2)/ie
04022 NEXT K
04024 GOTO 980
04070 N<2»2)«H0
04075 NO,2)»T8
04080 N<4*2>*A*T8
04085 NC6»2>»T6
04105 GOTO 2545
06000 I-e
06005 T7«N(3»2>+N<4»2>
06010 0-N(3«2>+2*N(4»2>+N(6»2>+2*T9
06015 D«N<6»2>
06020 IF S»0 THEN 6030
06025 W*T7+D+N<5»2>*«6*T9
06030 K"0
06035 Tl»(T(2)-t>460>/1.8
06040 K8«EXP<.278777-2.73705E-4*Tl*4.37722E-8*Tlt2-l1554.5/T1>
06045 K2=EXP(-6.24562*1. 71 757E-3*Tl-3.52622E-7*Tlt!>*5198/Tl)
06050 K3«EXPC17.5646-9.2926E-3*T(2>M.7716E-6*T<2>1?2-1070/TC2»
06055 K9«EXP(-1.4871-2.|6103E-3*Tl+4.39584E-7*Tlt2*l4786/Tl)
06056 K5-EXP(7l.5076-.03125*Tfe)*5.79655E-6*T(2)t2-67104/T(2))
06060 IF S>0 THEN 6070
06065 A"10
06070 Bi»l+2*K8
06075 G«G*1
06080 B2«2*2*K8
06085 B3»1*2/K3
06090 T6»A*0/0 THEN 6120
06110 A-A/1.1
06115 GOTO 6090
06120 S6»(K9*S4)t2*(CD-T6*B3)/T6)t4*P/W
06125 Sl«2*S6+2*K8*A*T7/(l*A*Bl»2*T6/K3+S4
06130 Z-0 THEN 6160
06150 IF
-------
062 iS
06220
06225
06230
06235
06240
06245
06250
06255
06260
06261
06262
06265
06270
06275
06280
06285
06290
06295
06300
06305
06310
06312
06325
06330
06335
06340
06342
06344
06346
06348
06350
06355
07000
07005
07010
07015
07020
07025
07030
07035
07036
07040
07041
07045
07046
07047
07050
07055
07060
07065
07070
07075
07080
07085
07090
07095
07100
07105
07110
07115
07120
07125
07130
07135
07140
00 TO 6090
1-1*1
U«T7+D+N(5«2>+S4+S6
IF Z<0 THEN 6250
A«*995*A
IF K100 THEN 6090
GOTO 6195
A»1«006*A
IF f«100 THEN 6090
GOTO 6195
A>(1-Z/20>*A
GOTO 6090
T8«T7/U+A*B1>
H0»D-T6*B3
S5*2*K8«A*T8
S7«2*T6/K3
E«K5-S4/W*P
E6-.3*KS
Z6"
GOTO 6342 *
PRINT MOttlOlMZ6t*lZ6
Q-0
PAUSE
Y3*0*(l-Z6/50>
NC5»2>»1.8B1*+NC5»2>
0»Y3
H»T7+D+N<5»2>+S4+S6
S»S*I
IF G<20 THEN 6075
GOTO 6325
J0*N+N(3»l>+N(4»l>+U(l>-I0*+6<4»
Jl»(G2-G(6)*I0-I4)/4
J2»+2*G<4>+G<6»-2*(W8+W9»/4
J3«Al/4
J4»J0/4
FOR K-3T01 STEP-1
I(l«K>«J3*J4*Kt2
L3«K*J4*(1-B(1»-K1»K)
L3»=L3*I0*
L4-L4*I0*6(6)
L5«D1 *K* J2-2*K* J4*B< 1 >
H(23>«I0*CG(3>*2*6<4> + G<6»+2*
L5«L5*H(23)
L6»L5-L4-L3
L7-L3-L6
L8«L7*L4
L9-L7+L4
J6-K1*L8
J5-L6*K1*L9
J7-K1-1
IF ABS(J7)«.001 THEN 7100
H2»K)-*J6/J5
I<4»K)-I(2»K)*L6
I(3»K)«L3-K4,K)
I(6»K)«L4-I(2»K)
2O4,
FOR L-1TO 6
J(L*K).I(L»K)/HCK)*P
NEXT L
NEXT K
-------
07141
07142
07143
07146
07147
07148
07149
07150
07152
07154
07155
07165
07170
07175
07180
07181
07185
07190
07195
07200
07205
07206
07210
07215
07220
07225
07230
07231
07235
07240
07245
07250
07251
07255
07260
07265
07270
07275
07285
07290
07295
07296
07300
07305
07310
07315
07320
07325
07330
07335
07340
07341
07345
07350
07355
07356
07360
07361
07362
07365
07425
07430
07440
07475
07480
07483
PRINT
IF S9*l THEN 7M6
Y8»0
Z8«D1+I0+11.76*W6/W7
205.
FOR L-IT04
J(L»0>«I0*GCL>*P/Z8
NEXT L
PRINT "H20"IDECJ
IF Y8»0 THEN 7175
FOR K-3T00 STEP-1
PRINT DEC(J(6»K>»<40-8*K>»3>J
NEXT K
PRINT
PRINT MH2"JDEC»3>J
NEXT K
PRINT
PRINT "COMJDECJ
IF Y8=0 THEN 7225
FOR K-3T00 STEP-1
PRINT DEC»3>J
NEXT K
PRINT
PRINT MC02MJDEC(C3»8»3>J
IF Y8»0 THEN 7250
FOR K=3T00 STEP-1
PRINT DEC(J(4»K)#<40-8*K>»3>f
NEXT K
PRINT
IF P«5 THEN 7285
PRINT "DP C02MJD£C«C3-PC 1 »» 10»3>J
FOR K»3T00 STEP-1
PRINT D£C«J(4»fO-P< 1 »«<40-8*K>«3>J
NEXT K
PRINT
V(6*4)«.9*H2*«1*JC6»3)
V<2»4>*«9*C4+*1*J<2«3>
V<4*4>»*°9*C3+«!*JC4»3>
VC1*4)«.9*C6*.|*J<1»3)
FOR K»3f 02 STEP-1
FOR L-1T06
V>H<19>/l>*VC6»L>»*2*V<2»L>+X<4»l>*V<3»L»t2
H(20)«X(5,1)*(V(6,L)*V(2*L)-V(3»L)*V(1»L)/X(6,1))
H<22>«1*X<7»1>*VC6»L>+XC8»1>*V<2»L>
Y(2»L)«H<2i)/(H(22)*X(9»l)*V(3»L)+X(10»l)*VCl,L))»2
NEXT L
FOR L.-4T01 STEP-1
Y(3»L)»Y<1»L)*2*Y(8/L)
NEXT L
J8M/Y(3»4>+1/Y(3»3>*1/YC3»2>+1/Y<3»1>
PRINT "R BAR"I DECC 4/J8» 1 2, 1 >
Y8-1
-------
07485
08000
08005
08010
08015
08080
08025
08026
08027
08030
08035
08040
08045
08050
08055
08130
08140
08145
08150
08155
08160
08165
08170
08175
08180
08185
08190
08195
08200
08205
09000
09004
09005
09010
09015
09020
09025
09030
09035
09036
09040
09045
09046
09050
09060
09065
09070
09075
09076
09080
09096
09097
09098
09100
09105
09111
09115
09120
09125
09130
09135
09140
09145
091S0
09155
09160
GOTO 2985
K7«EXP<"5.41543-9.53865E-3*T(l)*1.39652E-6*Tt2*16629/T*N<6«i)/l«3*l>/N<2»l>f 3*/P>t2
X2-N7/K7
Pli(Xl-X2)/Xl
If ABS(P1><«01 THEN 8050
Al-AI*(l*Pl/2)
IF Al>0 THEM 8030
A1-.05
F-F+1
IF F<20 THEN 980
PRINT "F"J"A1"JA1
PAUSE
F1«F1+F
p.0
GOTO 2305
W5»N(4»2>/V3*P
P4»PC2>-W5
P5-/P2
IF ABSCP5X.01 THEN 2826
IF PS«0 THEN 8190
T(2»T(2>+20*PS
H-H+1
IF H<20 THEN 885
PRINT "H"JDEC(T<2>»8»0>
PAUSE - • •
T<2)»T<2)+28*P5
H-H+1
IF H<20 THEN 885
GOTO 8180
PRINT TAB<30>I"RE8£N"
TAB(37)I"HT. CONTENT"
TAB(22)I "MOLSMlTAB(37)|-OR DELTA H"
TAB<5>I"CIN3"
-CHAR-
M CMlOEC/12«01»22»4)lDEC(OC5»l )»39,0)
* HNJDEC»39»0>
M ASH-|DEC
206.
PRINT
PRINT
PRINT
PRINT
PRINT
PRINT*
PRINT
PRINT
PRINT
PRINT
PRINT
PRINT
PRINT
PRINT
PRINT
PRINT
PRINT
-DRY AIRMIDEC»13»0>I-FMDECJDEC
" H20MJDEC»22»4>JDECCM4*39*0>
" TOTALMlOEC«U6-*-A4>»46«0>
"LIFT GAS"
•* C02MIOEC»22»4>
H COMiDEC(L<12>»22«4>
" N2"|D£C(L(16)*22»4)IDECCU8»39»0)
M HSOMIDEC+LU3>+LC15>+L<16>
PRINT
PRINT
PRINT
PRINT
PRINT
PRINT
PRINT
* TOTALM>DEC(N5»22»4>lDECfDEC(U8+M6»46»0>
•TOTAL IN**JDEC«U8+U6+M7*V7>»52»0>
TAB(5)I"COUT1"
"ACCEPTOR DUTY"
" M60»CAO"!DEC(X/R0*22»4>
" INERTCLB>MIDEC
F7«U(l)*CC(10>2)*T-C<19»l)*Tfi)
F8«*<1/R0-1»*(C<9»2>*T-CC9»1>*T2>
PRINT " SENSIBLE MG*CA"IDECC(F7«F8«Q0«04>»39»0>
F9«I5/R0*CU(1>*CCC7«2>*T-C<7»1>*T2>+U<0>**T-C<7*0>*T1»
PRINT
PRINT
PRINT
PRINT
PRINT
SENSIBLE INERT"! DEC
HT OF RCACT"fDEC(X*Q6/CX+I6>»39»0>
MAKEUP**iDEC
INERT(LB>"IDECCIS*I6»24»2>
-------
09165
09170
09175
09180
09 185
09186
09190
09200
09210
09215
09220
09225
09230
09235
09240
09245
09250
09251
09252
09253
09254
09255
09256
09257
09261
09262
09275
09280
09285
09290
09295
09300
09305
09310
09311
09315
09320
09335
09345
09350
09355
09360
09365
09425
09426
09427
09428
09430
09435
09440
09445
09450
09455
09456
09460
09465
09470
09475
09476
09480
09S10
09515
09520
09525
09530
09535
PRINT " SENSIBLE+HT REACT*4! DEC« I8+Q8«-Q9> ,39,0)
PRINT
PRINT " REJECT-
PRINT " MGO.CAO"JDEC-04*22.4>
PRINT " MGO.CAC03"!DECU6*R0,22,4)
PRINT " MGO.CASMJDEC<04,22,4)
PRINT " SEN5IBL£"JDEC(-(F7+F8+F9>*M1/100»39»0>
PRINT " TOTAL ACCEPTOR DUTY"! DEC(Q7*46»0>
PRINT "CHAR-
PRINT " C"!DEC<(C1-C8/I00)*C*U<1)/12.01),22,4)!DEC*46*0>
PRINT "GAS-
PRINT " CO"JDEC*22»4>
PRINT " C02"JDEC»22»4>
PRINT " N2M!DEC(N(5,2),22»4)
PRINT " H20"JDEC»22»4>
PRINT " H2"JDEC(H0,22,4)
PRINT " S02"JDEC
IF P<5 THEN 9257
PRINT - S2"JDEC(S6,22,4)
PRINT " H2S"JDEC(S7»22*4>
PRINT " COS"!DECCS5,22,4)
PRINT " TOTAL"JDEC»DEC(M9*46»0>
PRINT "HT CONTENT OF C02"
PRINT "«46»0>
PRINT "LOSS**IDEC
PRINT "TOTAL OOTM|DEC«M9*U7*V8*Q7*L2-U2-U3>*52»0>
PRINT "NET DUTY"! DECCU0* 52*0)
PRINT "HT. OF COMB."
PRINT " CHAR IN-JDEC
PRINT " CO IN"JDEC
PRINT " CHAR OUT"JDEC<-V6»46»0>
PRINT " GAS OUTMIDECC-V4*46»0>
PRINT "S REACTIONS"JDEC
PRINT "NET HEAT SUPPLIED"! DECCH6* 52»0)
PRINT TAB (30) I "GAS IF"
PRINT TAB(5)J"CIN3W
PRINT "COAL"
PRINT " C"JDEC(C0/12«01»22»4>
PRINT - H"lDEC(Hl/2. 016*22*4)
PRINT " 0"IDEC(00/16«22«4>
PRINT " ASH(LB)"lDEC(A9»24»2>
PRINT "STEAM"! DECCD1 -06-6. 383/ 18*02* 22* 4>JDEC(Q(3»
PRINT M HOIST"! OECC. 3542.22* 4>
IF P>5 THEN 9435
PRINT "REDUC TAN T"J DEC (10,22,4)
GOTO 9465
PRINT "CLAUS GAS-
PRINT " C0-|DEC(G(3>*I0»22»4>
PRINT " C02-JDEC(G<4>*I0»22»4>
PRINT " N2"JOECCG(5)*I0»22*4>
PRINT " H20"IDEC(G(6)*I0,22,4)
PRINT " H2"IDEC(6<2)*I0,22,4)
PRINT " TOTAL SENS"! DECC (R(8, 1 ) -87-07)* 46*0)
PRINT "AIR-
PRINT M N2-JDEC<11»76*U6/W7»22»4)
PRINT - 02"IDEC(W8*W9,22,4)
PRINT " H20-JOEC<06»22*4)JOEC<07,39«0>
PRINT " TOTAL"IOECCB7«07»46«0>
PRINT
PRINT "TOTAL IN"IDEC,S2,0)
PRINT
PRINT TAB(5)J"COUT3-
PRINT -CHAR"
PRINT " C-JDEC((C*U(l)/12.il*Z(D)»22,4)lDEC(0<5,
1),46,0)
1),39,0)
207.
-------
09540
09 545
09550
09 555
09560
09565
09570
09575
09580
09590
09600
09601
09605
09625
09630
09640
09645
09646
09648
09650
09655
09660
09665
09670
09680
09695
09705
09710
09715
09720
09725
09730
09735
09745
09746
09755
09760
*
PRINT " H*VDECCI4,22,4>JDEC(Q<8,1),39,0)
PRINT ** ASHCLB)"!DECCA9,24*2)!DECCQC 7,1), 39*0)
PRINT " TOTAL"! DECCQC5,1)+QC8,1)+QC 7,1), 46,0»
PRINT "GAS** -
PRINT ** CH4"!DECCNC1,1),22*4)
PRINT " H2"!DECCNC2,1>,22,4)
PRINT " CO"IDECCNC3,1),22,4)
PRINT " C02"!DECCNC4, 1>,22*4>
PRINT " H20"!DECCN(6, 1>,22,4)
PRINT" N2"JDECCNC5, 1>»22,4)
PRINT " NH3"! DECC .047, 22, 4)
PRINT " H2S"!DECCT3,22,4)
PRINT " TOTAL"! DECCMU), 28*4)! DECCE(l), 46,0)
PRINT "HT. OF REACT**
PRINT ** CHAR OUT"! DECC WC 2,1), 39,0)
PRINT " GAS OUT"JDEC(WC1,1),39,0)
PRINT ** COAL IN"!DECC-W(5,1),39,0)
IF P>5 THEN 9655
PRINT ** REDUCTANT"! DECC -85, 39,0)
GOTO 9660 '
PRINT " CLAUS GAS"! DEC (-85,39,0)
PRINT " S REACTIONS"! DECCBC0), 39,0)
PRINT " NET HT OF REACT"! DECC WC0, 1 ), 46,0)
PRINT "LOSS"! DECCLC1), 46*0)
LC20>«EC1)+WC0*1) + RC7*1)*QC5*1)+OC7*1)*QC8,1)-»>LC1)
PRINT "TOTAL OUT"! DECCLC20), 52,0)
PRINT "NET DUTY"! DEC(W( 6,1), 52,0)
PRINT "ACCEPTOR DUTY-
PRINT ** SENSIBLE MGO.CAO"!DEC(U(1)/R0*FC3, 1),39,0)
PRINT " SENSIBLE INERT"! DECCUC 1 )/R0*FC 5, 1 ), 39,0)
PRINT " HT OF REACT"! DEC CUC1)*FC 1,1), 39,0)
PRINT " CO 2 SENS HT"! DECC -UC 1 )*C TC 1 )-60)*CC4» 1 ), 39,0)
PRINT " NET ACCEPTOR DUTY"! DECCF(0,1)*52,0)
PRINT " HQO.CAC03 OUT"! DECCUC 1 ),22, 4)
PRINT ** M GO. CAS OUT"! DECCT4,22, 4)
80 TO 2985
END
208,
-------
APPENDIX E
-------
209.
APPENDIX E
Literature Review on Preoxidation of Caking Coals
A. Introduction and Summary of Prior Consol Work
Because of its specific purpose, the Bibliography for this Appendix
has been kept separate (as Section C) from the main Bibliography for this
Report.
The gasification process under development for GAP calls for the
use of Eastern bituminous coals which are highly caking. When fed into the
gasifier without pretreatment, these coals agglomerate, making operation of
the fluidized bed very difficult.
A fluidized bed coal gasification pilot plant was operated in 1949
at the Research Division of Consolidation Coal Company in a joint program
with Standard Oil Development Company.V1) It was demonstrated in the course
of this work that Pittsburgh seam coal could be successfully fed to an air-
blown fluidized bed operated at 175O°F without pretreatment.
These results, however, do not show that operation without pre-
treatment is feasible for the presently envisaged adaptation of the C02
Acceptor Gasification Process. The demonstrated coal feed rate, for example,
was 20 lb/hr-ft2 of gasifier cross section which is much too low for the
present application. The above gasification work was also carried out at
atmospheric pressure. Subsequent work indicates that the need for pretreat-
ment increases with the operating pressure and particularly as the hydrogen
partial pressure increases.
Work carried out by Consol for the Office of Coal Research in a
continuous bench-scale unit used for development of the C02 Acceptor
Process(2) showed, for example, that mild preoxidation (4$) was required to
prevent coking in fluid bed hydro devolatilization of subbituminous coals
at 1500°F and 20 atm pressure (H2 partial pressure ca. 5.5 atm). No such
pretreatment is required in the corresponding operation at atmospheric
pressure.
The nature of the atmosphere at the point of injection of the coal
also affects the coking tendency and need for pretreatment. The same
subbituminous coal could be fed successfully without coking to the identical
gasifier vessel operated at 1775°F and 2O atm pressure as a. "CO maker."(3)
The principal difference between this operation and the previous hydro-
devolatilization is the relatively low partial pressure of hydrogen existing
at the coal inlet. Thus, it is concluded that a highly reducing atmosphere
increases the need for pretreatment.
Caking bituminous coals in contrast to the subbituminous coals were
found, however, to require a high degree of preoxidation when fed under the
same conditions to the same gasifier operated as a "CO maker." In fact even
when 13$ preoxidized coal was used* some caking was observed in the gasifier.
The high degree of pretreatment required in this case is associated both with
the high operating pressure and the reducing environment existing in the
gasifier at the point of coal injection.
* % preoxidation = lCO(lb 02 reacted)/lb dry coal.
-------
210.
Experimental evidence indicates, however, that the degree of pre-
treatment required to sustain operability is reduced when the coal is
delivered into the gasifier in an oxidizing atmosphere. This is the method
of operation proposed in the present adaptation of the C02 Acceptor Process.
This is thought to be the reason why the air blown fluidized bed gas producer
operation mentioned previously(*) could be carried out without pretreatment,
even at the low rates demonstrated.
Similarly, highly caking Pittsburgh seam coals were used success-
fully as fuels with only mild preoxidation (5.3 wt #) for fluidized bed
reduction of sulfated dolomite at temperatures up to 1875°F at a pressure of
8 psig.v4) The success in this instance is likely due to the highly oxidi-
zing atmosphere existing at the point of injection of the coal into the bed.
In this instance, straight air was used as transport gas to carry the coal
into the bed.
The preoxidation process itself was studied extensively by the
Research Division of Consolidation Coal Company in the 1950's. The
objective of the work was to use preoxidation as a pretreatment method to
render the coal operable when fed to a fluidized bed carbonization process
operated in the temperature range of 90O-950°F.
A comprehensive survey of the preoxidation literature prior to
1952 was made\5' at that time. A survey of the preoxidation literature
since 1952 was made under the GAP contract and this is given here. It
concentrates on work carried out in the temperature range of 6OO-1OOO°F. The
reason for this will be evident in what follows. Some of the highlights of
the previous literature(5) search are also given.
The earlier Consol laboratory program on preoxidation was conducted in
three continuous bench-scale units with preoxidation vessels having internal
diameters of 1-1/2",(6) 4",(7) and 6ft,(8' respectively.
The laboratory study was made in the temperature range of 65O-825°F.
The study included a determination of reaction kinetics, heat of reaction,
effect of process variables such as particle size and determination of the
effect of preoxidation conditions on properties of the product. The latter
included determination of caking properties, bulk densities and other
properties of the preoxidized coals. The general findings are discussed
later in this survey in conjunction with a comparison to published literature
data.
The main thrust of the program was, however, directed' towards maxi- .
mizing tar yield in the overall process. The approach used was to minimize
the amount of preoxidation required for operability in subsequent fluid bed
carbonization. This objective is achieved by combining preoxidation with
thermal pretreatment, i.e., by operating at temperatures above 650°F wherein
pyrolysis of the coal occurs simultaneously with preoxidation. The maximum
temperature that can be employed is dictated by the operability limits imposed
by the coal itself, since the preoxidation step usually must be operated below
the point at which the "untreated" coal attains its maximum fluidity. In most
-------
211.
cases, depending on the caking properties of the coal employed, the preoxida-
tion temperature thus must be below 825°F. With this method an "optimum"
preoxidation temperature is defined which corresponds to the minimum oxygen
input to the preoxidizer required to maintain operability in both the pre-
oxidation and carbonization steps. The optimum temperature is a function of
the particular coal employed but is generally within the range of 70O-80O°F.
The above principle of an optimum preoxidation temperature in the
plastic range of the coal was successfully demonstrated in a 30 TPD pilot
plant. It is outside the scope of this survey to discuss the operation of
the pilot plant. A good description of the process as finally developed
is given, however, in a U. S. Patent granted to Consolidation Coal Company.(9)
The above background data on preoxidation forms a useful guide to
our impending experimental program for OAF. It minimizes but does not
eliminate further experimental work required to satisfy our present objective.
It is likely that in this case also it will be desirable to operate the
preoxidizer within the plastic range of the coal, i.e., 70O-800°F. It is,
however, no longer necessary to minimize the amount of preoxidation as in the
case of the low-temperature carbonization development. It is merely suffi-
cient not to exceed the oxygen input as required by heat balance to sustain
temperature in an adiabatic preoxidizer. The operability limits in the
preoxidizer itself are affected by operation under pressure and will thus
have to be determined separately. Likewise, the extent of pretreatment
required for the gasification operation and how it is affected by the method
of introducing the coal into the gasifier still remains to be determined
experimentally.
Finally, an important new consideration here is the desirability
of producing a relatively high density char for the gasifier operation.
This is a requirement which likely calls for maximization of the amount of
preoxidation. Practical considerations, as discussed above, limit the amount
of preoxidation to that which is consistent with the heat balance in the
preoxidizer.
B. Discussion of Literature
1. Oxidation Kinetics
a. Mechanism of Oxidation
The oxidation of coal is extremely complicated. Depending
upon the severity of reaction conditions, every organic
oxidation product from peroxides to water and carbon dioxide
may be observed. An exact model would involve taking into
account a number of simultaneous and consecutive reactions.
The reaction progress is generally evaluated using the amount
of oxygen reacted per pound of coal fed or the fraction of
unreacted material remaining in the coal.
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212.
b. Dependence of Rate Upon Degree of Coal Conversion
Struckv6) investigated kinetics of the oxidation of
Disco char, Montour No. 10 coal, and Arkwright coal from
about 550 to 765 °F. Conversion was expressed as oxygen
consumed, wt % of MAF coal. At 65O°F, a log-log plot of
rate vs. oxygen consumed was linear up to 20$. The slope
was about -1, thus indicating the rate to be inversely
proportional to the extent of reaction.
Weisman'10) studied the rate of dehydrogenation of
bituminous coal during oxidation in the range of 3OO-710°F.
He found the rate to be first order in unreacted hydrogen
remaining in the coal up to about 20% loss of the hydrogen
in the coal. The rate expression then became more complex
as secondary reactions attained importance.
One may expect the reaction rate to be proportional
to the material remaining available for reaction. The
problem is what basis to use for the reactive material.
Weisman uses the total hydrogen but divides it into more
reactive and less reactive fractions. Struck used the
amount of 'oxygen consumed to follow the reaction progress.
Still another method is to measure the SCF oxygen reacted
per pound of coal. One percent oxidation equals about
O.12 SCF/lb coal.
c. Effect of Oxygen Concentration
In his literature survey,(5) Struck found the rate of
oxidation below 4OO°F to be correlated by (PQ ) where h
was anywhere from O.5 to O.7. Struck assumed, however, a
first order relationship with respect to oxygen partial
pressure in kinetic interpretation of his own data but his
data were inadequate to confirm this assumption. Szuba,
et al.t11) report the oxidation to be second order with
respect to oxygen. They investigated the oxidation of
Silesian hard coals with air over the range of 34O to 52O°F.
Gaberman and KhalitovC12) oxidized coal with air in a fluid
bed up to 15 atm at about 35O°F. They found the reaction to
be markedly faster under pressure.
It is noted that when the reaction is diffusion control-
led, the rate will be first order with respect to oxygen
concentration. Many experimenters varied inlet oxygen
partial pressure, but operated with zero exit concentration,
and this made evaluation of kinetic effects impossible.
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213.
d. Effect of Temperature
Literature values for the activation energy of the
coal-oxygen reaction vary from 9 to 35 Kcal/mole. StruckC5)
reports the data of Sebastian and Mayers I13/ as given
below:
Lignites and low-rank bituminous 12,000 cal/mole 02
High-rank bituminous and anthracite 18,000 cal/mole 02
High temperature coke 30,000 cal/mole 02
Struck's own values(6) were about 12 Kcal/mole for Pittsburgh
seam coal and 33 Kcal/mole for Disco char. Szuba, et al.(14)
report an activation energy of 9.3 Kcal/mole for Silesian
hard coals.
A problem associated with evaluation of temperature
effects is whether significant diffusion and heat transfer
effects are present. Diffusion effects tend to lower the
activation energy. Heat transfer limitations raise the
internal solid temperature above the nominal, system
temperature.
e. Effect of Particle Size - Diffusion
Struck(e) found that at 650°F, reactivity tended to
increase as particle size was reduced. For Montour No. 10
coal, the 1OO to 200 mesh material reacted five times as
fast as the 35 to 65 and 65 to 150 mesh sizes. For
Arkwright coal, the 65 to 150 mesh material was noticeably
more reactive than the 35 to 65 mesh coal. Szuba, et al.*11'
attributed reduced activation energy in the range of
340-520°F to diffusion effects. Adrzejak and IgnasiakC15)
reported that for O.06-1.5 mm particles at 270°F, oxygen
diffusion is controlling for bituminous and anthracite
coals but not for brown coals. Peters and Juentgenv16)
claimed that activated diffusion of oxygen into the coal is
the rate determining step in the oxidation process.
Jenkins^17' found the rate of oxygen consumption at about
7OO°F to be roughly proportional to the external surface
area of the coal particle.
f. Heat of Reaction
Jenkins!17) reported the heat of reaction to be 87,OOO
Btu/lb mole of oxygen. Struck(6) reported values of
181,000 and 190,OOO Btu/lb mole of oxygen evaluated at
65Q°F. Curran, et al.(18) estimated the heat of reaction
to be 205,000 to 256,000 Btu/lb mole of oxygen reacted at
5OO°F. These latter two values are not necessarily in-
consistent since Struck dealt with a bituminous coal and
Curran studied a subbituminous coal.
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214.
Lee, et al.(19) reported on a calorimetric study of
the heat of reaction of a Pittsburgh seam coal at 700°F
and 800°F under pressure. Their data are presented as a
function of the weight loss of the coal; the amount of
oxygen consumed is not given. They noted that at 800°F,
the rate of devolatilization may compete with the rate of
oxidation.
g. Effect of Coal Source
The reactivity of coal generally varies greatly even
among coals of the same general type. Struck!6) found
marked differences in reactivity between Montour No. 10
and Arkwright coals even though both are Pittsburgh seam
coals. The large variation of activation energy among
coals has already been discussed. As a broad generaliza-
tion, reactivity decreases with increasing rank of the
coal.
h. Physical Changes Occurring During Preoxidation
When Pittsburgh seam coal is heated over about 700°F,
swelling takes place. Masom20) and Mason and Schora(21)
describe the effect of preoxidation at about 750°F as
"inflation of the particles to cenospheres ... development
of a high-reflectance skin differing in thickness from
particle to particle." Forney, et al.(22) reported a
volume increase of 1OO$ for 10 to 14 mesh particles and
40$ for 48 x 1OO mesh particles, both preoxidized at 750°F.
The smaller particles swelled only 1O$ when the pressure
was increased to 3OO psig. StruckC6) reports results
showing that the bulk density of preoxidized coal decreased
with preoxidation temperature, i.e., from about 0.55 gm/cc
after 7OO°F treatment to only about 0.25 gm/cc after 8OO°F
preoxidation.
2. Use of Preoxidation to Inhibit Agglomeration
a. Mechanism of Agglomeration
Hiteshue, et al.(23) felt that high molecular weight
oils melted and wet the coal. Subsequent cracking forms a
sticky residue which can glue particles together.. Mason(2°)
claimed that when preoxidized coal agglomerated, it was due
to discharge of material from the interior of the particle
and from some particles that may have escaped pretreatment.
It is sufficient for only a small fraction of a charge
to be agglomerating in order to make a reactor inoperable.
Struck'6/ used as little as 25% pretreated material mixed
with coke in order to maintain operability in a carboniza-
tion step. Experience gained during development of the
CO2 Acceptor ProcessC2) shows that 18$ bituminous coal mixed
with non-caking lignite could Cause agglomeration in the
gasifier.
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215.
Forney, et al.\22/ showed that agglomeration would occur
if helium, steam, or prepurified nitrogen were used to replace
a gas stream containing only 0.3% oxygen which had success-
fully destroyed the agglomerating tendency of the coal.
The Bureau of Mines(24) reported that oxidation produces
a skin on the surface of a coal particle which is not wetted
by sticky material in the coal.
b. Effect of Preoxidation on Agglomeration
Struck(6) used operability of the Curran carbonization
assay as a criterion of successful preoxidation. The assay
procedure is to raise the temperature of a fluidized sample at
3°F/min. up to 932°F. He found that at 785°F, 0.6% oxidation
was required for Montour No. 1O and 8.5% for Arkwright coal.
The operability criterion is whether dilution of the feed coal
with char is required to prevent agglomeration and defluidi-
zation.
Gasior, et al.(25) treated Pittsburgh seam coal in a
fixed bed from 80 to 200 minutes at 660 to 950°F. The char
did not agglomerate at 750°F and 3,OOO psig in the presence
of hydrogen. Forney, et al.(22/ used criteria of a free
swelling index (FSl) of less than two and no caking in the
presence of hydrogen at 11OO°F or 1650°F. His experiments
were conducted on Pittsburgh seam coals. Successful pre-
oxidations were obtained using a gas containing Q.2% oxygen
at about 775°F with a residence time of 5 minutes in a batch
fluidized bed reactor. However, the same coal required 2.3$
oxygen, 20 minutes and 820°F in a continuous feed reactor.
Using the same criterion for successful preoxidation as
Forney, Kaviick and Lee(2e) found they required 725-750°F,
an oxidation level of 1.O-1.5 SCF 02/lb of coal, and a
residence time of 1-2 hours. Forney, et al. and Kaviick
and Lee remarked that their fluid bed products contained
particles which had undergone relatively little treatment.
Another technique for effecting solid-gas contact is via
entrained flow or free fall. Steinmuellerv27/ reports that
pulverized coal containing 26% volatile matter can be made
non-agglomerating by preoxidation while entrained in oxygen
containing gas at 7OO-75O°F for 1-3 seconds.
National Fuels Corporation(28) used an entrained coal
reactor to preoxidize caking coals prior to briquetting.
Coal of -28 mesh was entrained by 5-15 pounds of air per
pound of coal for 2 to 4 seconds at 57O to 885°F. Various
pilot plants (more than 5 TPD) were operated from 1938
through 1951.
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216.
Gasior, et al.(29) preoxidized Pittsburgh seam coal in
a free fall system using 5.5-12.7$ oxygen, two seconds
contact time, 300 psig pressure, and 1040-1250°F. The coal
expanded and became sticky during the treatment. Decaking
was considered successful if coking did not occur upon
exposure to hydrogen at about 1100°F for five minutes *
Forney, et al.'3°/ describe a staged system operating
at about 600 psig. The coal is first pretreated with
0.3-0.8 SCF 02/lb coal in a free fall preheat section at
752°F for about one second. It is then treated in a carboni-
zing section where an additional 4 SCF 02/lb of coal further
oxidizes the coal. The final gasification is carried out at
1742°F.
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217.
C. Preoxidation Bibliography - Consol Work and Literature Since 1952
1. Consol Internal Reports, Project No. 29.
2. Office of Coal Research, U. S. Department of the Interior
Research and Development Report No. 16
Interim Report No. 3 - Bench-Scale Research on C02
Acceptor Process - Books 1, 2 and 3.
3. Curran, G.P., Fink, C.E., and Gorin, Everett, "Production of Low-Sulfur
Boiler Fuel - Application of C02 Acceptor Process."
Paper presented before the Second International Congress
on Fluidized Combustion, Hueston Woods, Ohio, Oct. 4-7, 1970.
4. Curran, G.P., Fink, C.E., and Gorin, Everett, "Coal-Based Sulfur
Recovery Cycle in Fluidized Lime Bed Combustion."
Paper presented before the Second International Congress
on Fluidized Combustion, Hueston Woods, Ohio, Oct. 4-7, 197O.
5. Struck, R.T. - Consol Internal Report, Project 10, Report 6 (1953).
6. Struck, R.T. - Consol Internal Reports, Project 1O.
7. Struck, R.T. - Consol Internal Reports, Project 86.
8. Struck, R.T. - Consol Internal Reports, Project 14.
9. Sylvander, N.E. - U.S. Patent 3,O7O,515 - Dec. 25, 1962.
1O. Wiesmann, Udo, Brennstoff Chem., 49. 3, 87-92 (1958).
11. Szuba, J., Gubrynowicz, L. and Stromich, T., Koks, Smola, Gaz, 11,
5, 161-168 (1966) C.A. 66 -67721s.
12. Gaberman, E.G. and Khalitov, I.Z., Tr. Khim-Met. Inst. Akad. Nauk
Kaz. SSR .8, 81-87 (1969) C.A. 72-45847j.
13. Sebastian, J.J.S. and Mayers, M.A., Ind. Eng. Chem., 2.9, 1118-24 (1937).
14. Szuba, J., Gubrynowiz, L., and Stromich, T., Koks, Smola, Gaz,
12, 3, 45-49 (1967) C.A. 67-92589h.
15. Andrzejak, A. and Ignasiak, B., Freiberg. Furschungsh.
A429, 7-26 (1968) C.A. 69-79116y.
16. Peters, W. and Juentgen, H., Eng. Gas-Solid React. Proc. Symp. 11-17 (1968)
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218.
17. Jenkins, G.I., Int. Conf. on Chem. Eng. in the Coal Industry,
25-39 (1956).
18. Curran, G.P., Fink, C.E., and Gorin, E., Report to OCR by Consol,
R & D Report No. 16, Interim Report No. 3,
Book 3, p. 178 (1969).
19. Lee, A.L., Feldkirchner, H.L., Schora, F.C., and Henry, J.J.,
I&EC Proc. Des. and Dev., 7., 2, 244-249 (1968).
20. Mason, D.M., I&EC Proc. Des. and Dev., JJ, 2, 298-303 (1970).
21. Mason,'D.M., and Schora, F.C., Jr., Advan. Chem. Ser. No. 69s
18-30 (1967).
22. Forney, A.J., Kenny, R.F., Gasior, S.J., and Field, J.H.
USBM R.I. 6797 (1966).
23. Hiteshue, R.W., Friedman, S., and Madden, R., USBM R.I. 6376 (1964).
24. Ibid. USBM I.C. 7794, 78 (1957).
25. Gasior, S.J., Forney, A.J., and Field, J.H., USBM R.I. 6605 (1965).
26. Kavlick, V.J., and Lee, B.S., Advan. Chem. Ser. No. 69f 12-19 (1967).
27. Steinmueller, L. and C., G.m.b.H., German Patent 1,170,365 (1964).
28. Barrit, D.T. and Kennaway, T., J. Inst. Fuel, 27f 232-240 (1954).
29. Gasior, S.J., Forney, A.J., and Field, J.H., ACS Div. Fuel Chem.
Preprints .10, 4, 123-130 (1966).
30. Forney, A.J., Gasior, S.J., Haynes, W.P., and Katell, S.J.,
USBM Coal Gasification Program, Tech. Prog. Report No. 24 (1970).
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