U.S. Environmental Protection Agency Industrial Environmental Research EPA-600/7-77-031
Office of Research and Development Laboratory
Research Tridncjle Park. North Carolina 27711 April 1977
HIGH-TEMPERATURE
DESULFURIZATION
OF LOW-BTU GAS
Interagency
Energy-Environment
Research and Development
Program Report
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EPA-600/7-77-031
April 1977
HIGH-TEMPERATURE
DESULFURIZATION OF LOW-BTU GAS
by
G.P. Curran, BJ. Koch, B. Pasek,
M. Pell, and E. Gorin
Consolidation Coal Company
Library, Pennsylvania 15129
Contract No. 68-02-1333
Program Element No. EHB529
EPA Project Officer: S.L. Rakes
Industrial Environmental Research Laboratory
Office of Energy, Minerals, and Industry
Research Triangle Park, N.C. 27711
Prepared for
U.S. ENVIRONMENTAL PROTECTION AGENCY
Office of Research and Development
Washington, D.C. 20460
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ABSTRACT
A process for the desulfurization of low-Btu fuel gas is described.
Processing consists of first desulfurizing the gas at high temperature in a
fluidized bed of half-calcined dolomite. The gas is then cooled to 700°C
and passed through high-pressure drop cyclones to remove particulates and
alkali. The intended end use of the gas is as fuel to gas turbines in
combined cycle power generation. The sulfur acceptor is regenerated with
steam and C02. A liquid-phase Glaus reactor is used to process H2S in the
regenerator offgas into elemental sulfur.
Experimental data are presented in the following areas:
1. Desulfurization and regeneration activity of Canaan,
Tymochtee and Buchanan dolomites as a function of cycles.
The Buchanan and Tymochtee stones were subjected to
various degrees of hardening.
2. Batch studies to elucidate variable effects and rate data.
3. Removal of particulates and alkali at elevated temperature.
A. Chance reaction studies. The Chance reaction is a low
temperature wet process for desulfurizing spent stone and
rendering it suitable for disposal.
5. Process improvement studies.
Two economic studies of the process were done. The first showed an
incentive accruing to the process vis-a-vis a conventional wet desulfuriza-
tion scheme. The second was an update of the process economics which showed
that a plant designed in 1975 for 1980 operation would desulfurize gas from
a high-sulfur coal for 38^/MM Btu (HHV + sensible heat) delivered to a power
station.
This report was submitted by Consolidation Coal Company (also known as
Conoco Coal Development Company) in partial fulfillment of Contract No.
68-02-1333 under sponsorship of the Advanced Process Branch of the Energy
Assessment and Control Division, Industrial Environmental Research Laboratory
of the Environmental Protection Agency. The work was done during the period
July, 1973 to January, 1976.
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TABLE OF CONTENTS
I. SUMMARY
II. CONCLUSIONS AND RECOMMENDATIONS
III. ECONOMIC EVALUATION
A.
B.
D.
E.
F.
H.
Introduction
Summary
Table III-l - Summarized Economic Evaluation - Regenerative
Acceptor Desulfurization Process
Figure III-l - Simplified Schematic of Expander Turbine
plus Conventional Boiler and Steam Cycle
Process Description
Figure III-2 - Block Flow Diagram - Overall System
Figure III-3 - Process Flow Diagram, Low-Sulfur Boiler Fuel
Regenerative Acceptor Case, Dwg. AF-3667
Design Basis
Table III-2 - Mass and Heat Balance - Gas Desulfurizer
Material and Heat Balances
Plant Investment and Operating Cost Estimates
Table III-3
Table III-4
Table III-5
Table III-6
Table III-7
Table III-8
- Mass and Heat Balance - Acceptor Regenerator
Mass and Heat Balance
Mass and Heat Balance
Mass and Heat Balance
Mass and Heat Balance
Acceptor Stripping Column
Liquid-Phase Glaus
Reactor
Sulfur Combustor
S02 Absorption Tower
- Overall Mass and Heat Balance
Evaluation of System Economics
Table III-9
Table III-10
Table III-11
Investment Summary
Direct Operating Cost Summary Ex. Acceptor Cost
Economic Evaluation - Regenerative Acceptor
Desulfurization Process
Figure III-4 - Sensitivity Analysis
Artificially Hardened Sorbents
Figure III-5 - Potential Impact of Combined Cycle Power Station
Efficiencies on Desulfurization Economics
Hardening Process
Mass and Heat Balance - Hardening Desulfurizer
Producer Gas Compositions with and without
Hardening
Figure III-6
Table III-12
Table III-13
Technical Uncertainties
1
7
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TABLE OF CONTENTS - Cont'd.
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IV. EQUIPMENT AND PROCEDURE 44
A. Gas Desulfurizer-Regenerator System 44
B. Hardness Testing 45
C. Fine CaC03 Runs 45
Figure IV-1 - P&I Flow Diagram, Continuous Unit, Acceptor
Reaction Loop, Dwg. AF-3411 46
Figure IV-2 - CS2 Supply System 47
Figure IV-3 - Drop Tester for Acceptor Hardness - Dwg. DM-3461 48
D. Hot Removal of Particulates and Alkali Fume 49
Figure IV-4 - Flow Diagram for Desulfurization by Chalk Slurry 50
Figure IV-5 - Flow Diagram for Desulfurization by Fine
Limestone and/or Chalk 51
Figure IV-6 - Flow Diagram for Ash and Alkali Fume Removal
Study 52
Figure IV-7 - Filter for Alkali Fume Removal System, Dwg. DV-3592 53
Figure IV-8 - Heating Chamber - Alkali Fume Removal System with
Filter Clean-up Assembly, Dwg. AV-3586 55
E. Chance Reaction 56
F. Lab-Scale Batch Acceptor Cycling Reactor 57
Figure IV-9 - Reactor Assembly - Lab-Scale Acceptor
Cycling Reactor, Dwg. AV-3427 58
G. Reduction of CaS04 by H2S 59
Figure IV-10 - Flow Diagram - Batch Acceptor Cycling Reactor,
Dwg. CF-3446 60
Figure IV-11 - Sulfur Condensing System No. 1 61
Figure IV-12 - Sulfur Condensing System No. 2 62
V. MATERIALS: SOURCE, PREPARATION AND ASSAYS 63
A. Sulfur Acceptors 63
B. Acceptor Assays 63
Table V-1 - Results of Raw Stone Assays 64
C. Low-Temperature Ash 65
Table V-2 - Elemental Analyses of Dolomites 66
Table V-3 - Ash from Illinois No. 6 Coal 65
Figure V-1 - Preparation of Low-Temperature Ash 67
Table V-4 - Coal Based Materials for Alkali Fume Removal Study 68
Table V-5 - Analyses of Dolomite Feeds to Alkali Fume
Removal Study 70
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TABLE OF CONTENTS - Cont'd.
VI.
VII.
SAMPLE CALCULATIONS
A. Calculation of Required Make-Up Rate
Table VI-1 - Computer Program for Cycling Acceptor
Table VI-2 - Computer Output for Canaan Dolomite
Figure VI-1 - Equilibrium CaC03 in Canaan Dolomite as a
Function of Conversion and Make-up Rate
Figure VI-2 - Equilibrium CaC03 in Hardened Tymochtee 9
Dolomite as a Function of Conversion and
Make-up Rate
B. Calculation of Normalized Density
Figure VI-3 - Equilibrium CaC03 in Hardened Buchanan Dolomite
as a Function of Conversion and Make-up Rate
TABULAR CHRONOLOGICAL HISTORY OF RUNS
Table VII-1
Table VII-2
Table VII-3
Table VII-4
Table VII-5
Table VII-6
Table VII-7
Table VII-8
Table VII-9
Summary of Initial Runs with Canaan Dolomite
Summary of Early Hardening Runs
Summary of Runs with Sorbent from Various
Hardening Conditions
Summary of Chance Product Runs
Summary of Main Body of Continuous Runs
Batch Studies of Process Variables
CaS04 Reduction by CaS
Low Temperature Ashing of Coal
Alkali Fume Removal
VIII.
CYCLIC CONTINUOUS RUNS WITH CANAAN DOLOMITE
A. Introduction
B. Runs without Makeup
Table VIII-1 -
Table VIII-2 -
Figure VIII-1 -
Figure VIII-2 -
Figure VIII-3 -
Figure VIII-4 -
Figure VIII-5 -
Figure VIII-6 -
Figure VIII-7 -
Runs with Makeup
Conditions and Results for Gas Desulfurizer
with Canaan Dolomite Feed
Conditions and Results for Regenerator with
Canaan Dolomite Feed
Run A20A - Mol 7. CaS in Exit Solids
Deactivation of CaS in Canaan Dolomite
at 704°C (1300°F)
Run A21 - Mol 7» CaS in Exit Solids
Run A22A - Mol % CaS in Exit Solids
Deactivation of CaS with Temperature as a
Parameter
Run A36A - Mol 7» CaS in Exit Solids
Run A48 - Mol % CaS in Exit Solids
Figure VIII-9 - Run A51 - Mol 7. CaS in Exit Solids
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TABLE OF CONTENTS - Cont'd.
Figure VIII-10 - Three Possible Correlations for Deactivation
of CaS in Canaan Dolomite at 704°C (1300°F)
Figure VIII-11 - Run A51 - Mol 7, CaS in Regenerator Exit
Density Measurements on Cycled Canaan Dolomite
Figure VIII-12
Figure VIII-13
Figure VIII-14
Table VIII-3
Run A52A - Mol % Gas in Exit Solids
Run A52A - Predicted CaS Contents for
Several Correlations
Run A52A - Mol 7» CaS in Regenerator Exit
Density of Sulfided Canaan Dolomite
E.
F.
Figure VIII-15 - Normalized Density of Sulfided Canaan Dolomite
Attrition Data
Conclusions
Figure VIII-16 - Normalized Density of Sulfided Canaan
Dolomite from Makeup Runs
Attrition Data for Canaan Dolomite
Attrition Data for Runs with Makeup
Figure VIII-17
Figure VIII-18
IX. CYCLIC CONTINUOUS RUNS WITH HARDENED DOLOMITES
A. Introduction
B. Batch Hardening Tests
Table IX-1 - Conditions for Batchwise Dolomite Sulfiding
Table IX-2 - Conditions for Batchwise Oxidation of
Sulfided Dolomite
Table IX-3 - Properties of Materials from Hardening Study
Table IX-4 - Results of Drop Tests
Table IX-5 - Conditions and Results .for Gas Desulfurizer
with Tymochtee Dolomite
C. Continuously Hardened Tymochtee 10 Dolomite
Table IX-6
Figure IX-1
Figure IX-2
Table IX-7
Table IX-8
Figure IX-3
Figure IX-4
Figure IX-5
Figure IX-6
Figure IX-7
Figure IX-8
Conditions and Results for Regenerator with
Tymochtee Dolomite
Run A23 - Mol 7. CaS in Exit Solids
Run A23 - Deactivation of CaS at 704°C (1300°F)
Conditions for Continuous Sulfidation of
Tymochtee Dolomite
Conditions for Continuous Oxidation of
Sulfided Tymochtee Dolomite
Run A24 - Mol 7o CaS in Exit Solids
Run A24 - Deactivation of CaS at 704°C (1300°F)
Run A25 - Mol 7o CaS in Exit Solids
Deactivation of Hardened Tymochtee 10
Dolomite at 704°C (1300°F)
Run A26 - Mol 70 CaS in Exit Solids
Run A27 - Mol 7= CaS in Exit Solids
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TABLE OF CONTENTS - Cont'd.
D.
H.
Figure IX-9
Figure IX-10
Figure IX-11
Figure IX-12
Tymochtee No. 9 Dolomite
Run A28 - Mol 7. CaS in Exit Solids
Run A29 - Mol °L CaS in Exit Solids
Deactivation of Tymochtee 10 Dolomite at
704°C (1300°F)
Run A30 - Mol 7, CaS in Exit Solids
Figure IX-13
Figure IX-14
Table IX-9
Run A32 - Mol 7. CaS in Exit Solids
Deactivation of CaS in Hardened Tymochtee 9
Dolomite at 704°C (1300°F)
Particle Size Consist of Overhead Fines
Tymochtee No. 11 Dolomite
Figure IX-15
Figure IX-16
Figure IX-17
Figure IX-18
Figure IX-19
Run A35 - Mol 7» CaS in Exit Solids
Run A37 - Mol 7. CaS in Exit Solids
Deactivation of Tymochtee 11 Dolomite at
704°C (1300°F)
Run A38 - Mol 7. CaS in Exit Solids
Run A42 - Mol 70 CaS in Exit Solids
Buchanan Dolomite
Figure IX-20
Figure IX-21
Table IX-10
Table IX-11
Table IX-12
Table IX-13
Figure IX-22
Figure IX-23
Run A45 - Mol 70 CaS in Exit Solids
Run A46 - Mol 7» CaS in Exit Solids
Conditions for Continuous Sulfidation of
Buchanan Dolomite
Conditions for Continuous Oxidation of
Sulfided Buchanan Dolomite
Conditions and Results for Gas Desulfurizer
with Buchanan Dolomite
Conditions and Results for Regenerator with
Buchanan Dolomite
Deactivation of CaS in Hardened Buchanan
Dolomite at 704°C (1300°F)
Run A33 - Mol 7, CaS in Exit Solids
Density of Buchanan Dolomite
Figure IX-24 - Run A43 - Mol 7. CaS in Exit Solids
Figure IX-25 - Run A44A - Mol 7. CaS in Exit Solids
Figure IX-26 - Run A47 - Mol 7, CaS in Exit Solids
Figure IX-27 - Normalized Density of Hardened Buchanan Dolomite
Table IX-14 - Density of Hardened Buchanan Dolomite
Conclusions
X. BATCH SULFIDATION AND REGENERATION EXPERIMENTS
A. Batch Sulfidation Studies
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TABLE OF CONTENTS - Cont'd.
Table X-1 - Conditions for Batch Sulfidation Runs
Table X-2 - Properties of Feeds and Products
Table X-3 - Results of Sulfidation Runs
Table X-4 - Mols H2S Fed/Hr Per Mol CaC03 Remaining
at 0.107o H2S in Exit Gas
Table X-5 - ANOVA for Mols H2S/Hr-Mol CaC03 at Breakthrough
Table X-6 - Mol 7, H2S in Dry Outlet Gas with Excess CaC03
Remaining in the Bed
Table X-7 - ANOVA for Mol 7, H2S in Dry Outlet Gas
B. Batch Regeneration Study
Table X-8 - Typical Conditions for Batch Regeneration Runs
C. Regeneration Kinetics
Table X-9 - Results of Regeneration Runs
Table X-10 - Regeneration Results - Percent Regeneration
Table X-ll - ANOVA for Percent Regeneration
Figure X-l - Approach to Equilibrium vs CaS: Reactants Ratio
Figure X-2 - Regeneration of CaS at 649°C (1200°F)
Figure X-3 - Regeneration of CaS at 704°C (1300°F)
Figure X-4 - Regeneration of CaS at 760°C (1400°F)
Figure X-5 - Regeneration of CaS at 816 and 871°C
D. Cyclic Sulfidation and Regeneration
Table X-12 - Initial Rate Constants for Regeneration Kinetics
Figure X-6 - Effect of Temperature on Regeneration Kinetics
Table X-13 - History and Cycling Results for Reactivated
Dolomite
Table X-14 - Conditions and Results for Gas Desulfurizer
Table X-15 - Conditions and Results for Regenerator
Table X-16 - Run A86 - Batch Cycling of Run A22A Dolomite
E. Cycling with Varying Residence Times
F. Lab-Scale Batch Acceptor Cycling
Figure X-7
Table X-l7
Table X-18
Figure X-8
Figure X-9
Table X-19
Table X-20
Figure X-10
Table X-21
Table X-22
Run A128 Deactivation of CaS at 704°C (1300°F)
Conditions and Results for Gas Desulfurizer
Runs A129 to A132
Runs A129 to A132 - Conditions and Results
for Regenerator
Runs A129, A129A - Deactivation of CaS at
704°C (1300°F)
Runs A130 to A132 - Deactivation of CaS at
704°C (1300°F)
Conditions for Batch Acceptor Cycling Reactor
Results of Lab-Scale Cycle Test CR1
Lab-Scale Batch Acceptor Cycling
Results of Lab-Scale Cycle Test CR2
Screen Analysis of Run CR2 After 22 Cycles
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TABLE OF CONTENTS - Cont'd.
Page
XI. HOT REMOVAL OF PARTICULATES AND ALKALI 219
Table XI-1 - Conditions and Results for Gas Desulfurizer 220
Table XI-2 - Conditions and Results for Regenerator 221
Table XI-3 - Material Balances - Alkali Fume Removal Runs 222
Table XI-4 - Molar Ratios in Effluent Downstream of Hot Box Filter 224
XII. CHANCE REACTION STUDIES 225
A. Introduction 225
B. Discussion 225
Table XII-1 - Conditions and Results for Sulfidation of
Chance Reaction Products 226
Table XII-2 - Chance Reaction: Conditions and Results 228
C. Conclusions 229
Table XII-3 - Chance Reaction; Composition of Slurry
and Exit Gas 230
XIII. PROCESS IMPROVEMENT STUDIES 231
A. Introduction 231
B. Use of Fine Calcined Limestone 231
Table XIII-1 - Conditions and Results for Fine CaC03 Runs 232
C. Generation of an Active Acceptor from Coarse Limestone 233
D. Fully Calcined Dolomite 233
Table XIII-2 - Particle Size Analysis of Product
Limestone Fines 234
Table XIII-3 - Analysis of Product Limestone Fines 234
Figure XIII-1 - Run A41 - Deactivation of CaS at 704°C (1300°F) 235
Table XIII-4 - Conditions and Results for Gas Desulfurizer 236
Table XIII-5 - Conditions and Results for Regenerator 237
Table XIII-6 - Results of Sulfidation Runs with Fully
Calcined Dolomite 238
E. Reaction of CaS04 with H2S to Product Liquid Glaus Feed 239
Figure XIII-2 - Run A123 Deactivation of CaS at 704°C (1300°F) 240
Figure XIII-3 - Modified Sulfur Recovery Section 242
Table XIII-7 - Conditions and Results for Reaction of
CaS04 with H2S 243
Table XIII-8 - Acceptor Compositions for Runs A84C and A84D 244
F. Conclusions 245
Table XIII-9 - Calculated Gas Compositions, Run A84 246
Table XIII-10 - Comparison of Dry Gas Compositions, Run A84 247
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TABLE OF CONTENTS - Cont'd.
Page
XIV. PHYSICAL EXAMINATION OF PRODUCT SOLIDS 249
A. Surface Area and Pore Volume Data 249
B. Sulfur Distribution in Sulfided Dolomites 249
Table XIV-1 - Surface Area and Pore Volume Data 250
C. Scanning Electron Microscope Examinations 251
Figure XIV-1 - Images of Sample A41-1 at 180X - Prerun
Sulfided to 15 mol 7» CaS 253
Figure XIV-2 - Images of Sample A42-1A at 200X - Prerun
Sulfided to 15 mol 70 CaS 254
Figure XIV-3 - Images of Sample A42-2 at 180X - Prerun
Oxidized to 14 mol 7. CaS04 255
Figure XIV-4 - Images of Sample A42-2A at 150X - Prerun
Oxidized to 14 mol 70 CaS04 256
Figure XIV-5 - Images of Sample A42-3 at 150X - Regenerator
Sample, 0.7 cycles, 25 mol 7, CaS 257
Figure XIV-6 - Images of Sample A42-3A at 150X - Regenerator
Sample, 0.7 cycles, 25 mol "L CaS 258
Figure XIV-7 - Images of Sample A42-4 at 150X - Regenerator
Sample, 9.6 cycles, 72 mol 7» CaS 259
Figure XIV-8 - SEM Photos of A76 Product 260
Figure XIV-9 - SEM Photos of A81 Feed and Product 261
Figure XIV-10 - SEM Photos of A46 Desulfurizer Bed 262
Figure XIV-11 - SEM Photos of A46 Desulfurizer Bed 263
Figure XIV-12 - SEM Photos of A46 Desulfurizer Bed 264
\
XV. SUMMARY OF VARIABLE EFFECTS 265
A. Gas Desulfurizer 265
B. Regenerator 266
C. Model for the Deactivation Process 266
XVI. DESIGN BASIS 269
A. Gas Desulfurizer 269
Table XVI-1 - H2S:CaC03 Ratio for Continuous Runs 270
B. Regenerator 271
C. Tertiary Particulate Removal 271
Table XVI-2 - Stone Make-up Summary 272
XVII. ENVIRONMENTAL EVALUATION 273
Table XVII-1 - Trace Elements in 90 Illinois Limestone Samples 273
Figure XVII-1- Process Flow Schematic - Materials Consumed
and Generated (TPY) 274
Table XVII-2 - Phosphorous Pentoxide, Manganese Oxide, and Sulfur
Trioxide in Illinois Limestone and Dolomites
Containing more than 95% Carbonates 275
xii
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TABLE OF CONTENTS - Cont'd.
XVIII. REFERENCES
XIX.
APPENDICES
Appendix A - Detailed Investment Cost
Table A-1 - Cost Estimate;
Table A-2 - Cost Estimate;
Section 300, Sulfur Removal
Section 400, Sulfur Removal
Appendix B - Data Tables
Table B-l
Table B-2
Table B-3
Table B-4
Composition of Exit Solids from Cycling Runs
H2S Content of Exit Gas from Cycling Runs
Attrition Data
H2S Content of Exit Gas - Batch Variable Studies
Appendix C - Thermodynamic Data
Table C-1 -
Table C-2 -
Table C-3 -
Table C-4 -
Table C-5 -
Table C-6 -
Table C-7 -
Table C-8 -
Table C-9 -
Figure C-l -
Figure C-2 -
Figure C-3 -
Effect of Temperature and Sulfur Partial Pressure
on Distribution of Sulfur Species
Equilibrium Constants for Gas Reactions
Numerical Values of Equilibrium Constants for
Table C-2
Heat Capacities at Zero Pressure
Mean Heat Capacities Above 60°F, Gases
Heats of Formation at 25°C, Gases
Equilibrium Constants for Solids Reactions
Mean Heat Capacities Above 60°F, Solids
Heats of Formation at 25°C, Solids
Sulfur Vapor Pressure
Equilibria for CaS Reactions
Equilibria for CaS Reactions
Appendix D - Comparison of CCDC Process to Hot Carbonate Scrubbing
Memorandum RM-13451 - Cost of Pressurized Producer Gas
Desulfurization
Appendix E - Conversion Factors - English to Metric Units
Page
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I. SUMMARY
A. Process Description and Economic Evaluation
The process evaluated here is an evolution of the dolomite-based hot clean-
up process described in the last annual report.C1) The economic study was
limited to a consideration of the hot gas cleanup with exclusion of the gasifi-
cation process. A process flow diagram along with summarized heat and material
balances, operating conditions, etc., are given in Figure III-3. The basis for
the study was supplying hot clean producer gas to a boiler capable of producing
1400 MW of electric power.
The study included in the last annual report!1) was modified to take into
account the effect of inflation for a new presumed start up date of January
1980. Further modifications were made in the process design assumptions to
take into account the results of the last two years of experimental work. A
brief economic comparison of the updated previous design and the present design
is given below. Capital charges are at the rate of 18$ of the investment and
the operating factor is assumed to be 70$.
Updated
Present Design Previous Design
Capital Cost, $/KW 58 53.4
Gas Desulfurization Cost,
5^/MM Btu (HHV + Sensible Heat) 38 32.5
B. Continuous Bench-Scale Studies
1. Cycling Studies - Canaan Dolomite
The cycling of Canaan dolomite in the continuous bench-scale unit was
further studied through the gas desulfurization and regeneration units.
The principal extension of previous work was the use of longer runs
to increase the number of cycles, the study of the effect of acceptor particle
size, the investigation of the effect of higher temperature and higher H2S
concentrations on regeneration rate, and the study of cycling with continuous
addition of fresh stone to supplement previous batch cycling studies.
The major result of this program was to demonstrate that Canaan
dolomite showed acceptable life and activity. Hydrogen sulfide was removed in
the gas desulfurization step to closely approach equilibrium even with very low
inventory levels of CaCO3. A high hydrogen sulfide concentration in the
regenerator does not affect the rate of regeneration and close approach to
equilibrium is obtained. Regeneration rates are improved by an increase in
regeneration temperature.
The runs with continuous acceptor makeup showed results consistent
with the correlation for acceptor activity versus cycles developed from batch
data. Acceptor make-up rates can thus be closely estimated from the "batch"
cycling runs.
1.
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Increase of particle size had no effect on acceptor activity, but
iis effect on attrition rate was difficult to assess.
Attrition rates in the cycling runs were uniformly low, i.e., generally
less than one percent.
2. Hardening and Cycling Studies - Other Dolomites
A number of dolomites do not show sufficient attrition resistance for
dolomites.
direct use in the process. A process was developed for hardening of "soft
The hardening process was first developed batchwise and then applied
in continuous equipment. The hardening process consisted of first partially
sulfiding the stone and then oxidizing the sulfided stone to calcium sulfate.
The oxidized stone is fed directly to the gas desulfurizer where the calcium
sulfate is reduced "in situ" to calcium sulfide. All three reactions can be
conducted in the temperature range of the gas desulfurizer, i.e., at 1600-1700°F.
The hardening process is thought to be due to the formation of transient liquids
in the CaS-CaC03-CaS04 system.
The commercial process would carry out the process directly in the gas
desulfurizer by injection of the "hardening" oxidizing air into a draft tube.
Demonstration of this "simplified" version of the process requires pilot plant
demonstration.
t
The process was successfully applied to two "soft" dolomites, i.e.,
three different batches of Tymochtee dolomites from Ohio and Buchanan dolomite
from Virginia.
The attrition rate of the hardened stones in the 35 x 48 mesh size
range was reduced to below the targeted value of one percent per cycle.
Attrition rate decreased on continued cycling. The effect of particle size on
attrition rate was hard to establish due to experimental difficulties.
The hardening process is associated with some loss of acceptor
regeneration activity in the cycling process. The activity of the hardened stone
was also shown to be independent of particle size over the range tested, i.e.,
20 x 28, 28 x 35 and 35 x 48 mesh.
The loss of activity increases with severity of conditions in the
hardening step.
The activity of the hardened stones was similar to the Canaan standard,
and varied between being slightly poorer to slightly better than the standard
depending on the specific batch of stone and hardening conditions employed.
C. Batch Kinetic and Cycling Studies
The above studies were conducted to extend the range of variables and to
aid in developing a kinetic model of the process. The work was not carried
forward far enough to develop a comprehensive kinetic model. It was shown, how-
ever, that the regeneration process may be characterized as a "pseudo" first
order process.
2.
-------
The major effects of the variables in the regenerator and desulfurizer
were similar to those noted in the continuous unit.
The space velocity in the gas desulfurizer defined as mols H2S fed/hr-mol
CaC03 inventory at the point of incipient H2S breakthrough was found to be
statistically independent of the following variables:
CaC03 age
Type of stone but limestones inactive
Bed depth
Particle size
Mol % H2S in the inert gas
However, space velocity increased with increasing temperature.
The regenerator performance in the batch runs was characterized by the
percent regeneration of the CaS. The percent regeneration decreased with in-
creased stone age, and increased with increasing temperature and bed depth.
It was also adversely affected by prehardening of the stone. An important
finding is the acceleration of the regeneration reaction by H2S in the product
gas.
Cycling studies showed the following important results. The rate of
deactivation under similar conditions was approximately the same as in the
continuous unit. High temperature regeneration is not useful, however, as a
reactivation method since a large portion of the CaC03 produced is inactive
in the gas desulfurizer.
The most significant finding of the batch cycling work is the marked
effect of reduced residence time in the gas desulfurizer in improving the
regeneration performance and cycle life.
D. Chance Reaction Studies
The major variables of particle size, temperature, acceptor type, stirrer
speed and C02 flow rate were studied. Runs also were made to simulate two
stage as well as countercurrent operation as visualized in the commercial
process.
Increasing temperature, C02 flow rate and decreasing particle size in-
creased reaction rate. The other variables had a relatively minor effect.
The use of two stage or countercurrent operation appears advantageous.
Substantially complete sulfur rejection is obtained at atmospheric pressure,
90°C, and one hour residence time.
E. Liquid-Phase Glaus Studies
No further experimental work was performed since the last annual report.(
E. Properties of Product Solids
Studies were made on the properties of the product acceptor solids to
shed light on the mechanism of deactivation. These studies included measure-
ments of particle density, surface area, pore volume, scanning electron
3.
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microscopic studies (SEM) and 'distribution of sulfur, calcium 'and magnesium in
the acceptor particles via X-ray and ultraviolet techniques.
The results of these studies although by no means complete are consistent
with the following hypothesis for the mechanism of deactivation.
Deactivation occurs mainly in the gas desulfurizer. Crystal growth of
CaC03 and CaS, but primarily the latter segregates the calcium crystals from
the MgO. The large CaS crystals in particular are relatively inactive in
regeneration and gradually accumulate on cycling.
G. Hot Removal of Particulates and Alkali
Preliminary experimental work was conducted to evaluate in an exploratory
way the "substantially" complete hot removal of alkali fume and particulate
material.
The preferred device, i.e., high pressure drop cyclones, could not be
evaluated due to the small scale of the equipment and limited time available
for the work. A hot porous stainless steel filter operated at 1300°F (704°C)
was used in this work.
The results are encouraging even though highly preliminary due to the
short nature of the program. Particulates were nearly completely removed and
99+$ of the feed Na and K were removed.
Removal of chlorine fed as NaCl, was much less effective than alkali
removal and it appeared as HC1 in the product gas. This indicates that chemical
trapping of Na by reaction with acidic ash components may play an important
role in the alkali fume removal process.
H. Process Improvement Studies
Preliminary experiments were conducted to evaluate several potential
process improvements as follows:
1. Use of the Chance reaction product as an H2S acceptor in an
entrained-type gas desulfurizer.
2. Use of limestone-based acceptors in the partially or fully
calcined form in the gas desulfurizer.
3. Use of a regeneration cycle in which H2S is reacted with
CaS04 to produce a liquid Claus feed having a 2:1 H2S:S02
ratio.
Experimental difficulties plagued the attempt to evaluate the first
scheme. It therefore cannot be definitely ruled out of contention.
The second system appears to be of little interest due to chronic
problems with acceptor agglomeration in the gas desulfurizer.
4.
-------
The third system was shown to be chemically feasible.
J. Environmental Evaluation
The process is very effective in removing sulfur, i.e., the sulfur dioxide
produced by combustion of the cleaned producer gas amounts to only 25% of the
1.20 Ib/MM Btu standard.
No data are available regarding the NH3 content of the cleaned producer
gas and the NOV level in the flue gas after combustion.
A
Particulate removal effectiveness remains to be established in a pilot
scale. The results given here support the conclusion that alkali fumes would
be efficiently removed with the particulates.
The major waste solid product is the spent dolomite. The Chance reaction
is shown to be effective for removal of potential H2S pollution from the CaS
contained in the spent solids.
One of the favorable characteristics of the process from the point of view
of cooling water requirements and resource utilization is its high thermal
efficiency.
5.
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II. CONCLUSIONS AND RECOMMENDATIONS
A. General Status of Development
The economic evaluation shows that the process is attractive with respect
to alternative wet gas cleanup processes. The major elements of the process
have been thoroughly proven out on a laboratory and continuous bench scale. It
is, therefore, strongly recommended that pilot-scale development of the process
be undertaken. !
The major technical uncertainties that exist can only be resolved on a ''
pilot plant scale. These are;
1. Equipment development required to remove residual particulate
and alkali fumes (at 1300°F or higher) to a level compatible
with subsequent use of the gas to power a gas turbine.
2. The level of NH3 in the cleaned gas, and its impact on genera-
tion of NOX on subsequent composition.
3. The effectiveness of operation on a large scale of the hardening
system proposed for the pilot plant.
Continuing laboratory and bench-scale work is also recommended simultaneous
with pilot scale development to effect optimization and process improvements and
an improved understanding of the process fundamentals.
B. Continuous Cycling Studies with Dolomites
The main conclusion from these studies was that all features of the process
have been confirmed.
Broad extension of the process, however, required development of a pre-
hardening step such that naturally soft dolomites can be used. A hardening step
was successfully developed which has no significant effect on the capital cost
of the process. It was observed, however, that the prehardening step in some
cases had a small deleterious effect on acceptor life. This necessitated an
increase in make up requirements in the process economic study from the one per-
cent previously employed to two percent.
The effect of desulfurizer temperature has been only partially studied in
previous work and residence time was not studied at all. The improved acceptor
life obtained in the batch cycling tests at shorter residence times provides a
great incentive to confirm this beneficial effect in continuous bench-scale work.
All of the prior continuous work was also carried out at 160O°F and 1650°F.
The modified process design presented here calls for an increase in desulfurizer
temperature to 1675°F. Further studies in the continuous unit should therefore
be conducted to tie down the effects of shorter residence times and higher
temperatures in the gas desulfurizer.
7.
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C. Batch Kinetic and Cycling Studies
I';
The outstanding result was that mentioned above, i.e., that reduced resi-
dence times in the desulfurizer improved acceptor life.
Another result on the negative side is the finding that reactivation of
spent acceptors by high temperature regeneration is not a viable process.
The "new" CaC03 thus formed is largely inactive in the regenerator.
Valuable data were obtained in the kinetic study, but they were insufficient
in scope. It is recommended that this effort be continued with the aim of
constructing a mathematical model of the process useful for design and scale-up
purposes.
D. Chance Reaction Studies
Experimental work carried out on the disposal of spent acceptor via the
Chance reaction confirmed the major features of the design assumptions used in
the feasibility study. No further work is required short of pilot plant demon-
stration.
E. Liquid-Phase Glaus
The major design features of this essential unit operation were confirmed
by previous work'given in the last annual report.!1) Further bench-scale work
is desirable, though not essential, to provide additional kinetic data for
mathematical modeling and scaleup.
F. Properties of Product Solids
Data were obtained which shed valuable light on the mechanism of acceptor
deactivation. These data were obtained by a variety of physical and instrumental
techniques.
These techniques should be further developed in conjunction with pilot plant
development of the process.
G Hot Removal of Particulates and Alkali
Encouraging results were obtained in preliminary work conducted on the above
problem. Confirmation of the effectiveness of the process can only be obtained
on a pilot plant scale.
Supporting laboratory and literature researches are, however, recommended
to investigate the kinetics and thermodynamics of chemical trapping of alkali
chloride vapors by reaction with acidic oxides present in coal ash.
H. Process Improvement Studies
Three systems were evaluated as set forth in the summary. The use of lime-
stone in the process, even in the calcined form, appears to be of no interest.
The use of the Chance reaction product and production of a liquid-phase
Claus feed by reaction of H2S with CaS04 merit further laboratory investigation
and such is recommended.
8.
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J. Environmental Evaluation
The process is highly effective for the original environmental purpose
visualized, i.e., removal of sulfur at a high level of thermal efficiency.
No outstanding environmental disadvantages are apparent. An unresolved
question, however, is the level of ammonia in the cleaned gas and its impact
on ultimate NOX emissions.
The above and other environmental questions with regard to particulate
control and trace elements disposition can be answered only by pilot plant
development.
9.
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III. ECONOMIC EVALUATION
A. Introduction
The purpose of the economic evaluation is to re-evaluate the cost of
desulfurization via the proposed commercial gas desulfurization process. The
desulfurization cost was evaluated in the initial feasibility study(2) wherein
the process was fully described. The economic study as in the previous study!1)
is restricted to the costs of desulfurizing the producer gas, including sorbent
regeneration and sulfur recovery. The cost of gasification and the cost associ-
ated with generating electrical power from the desulfurized gas are not included
in this economic review.
The most attractive end use of clean, pressurized, hot producer gas in the
long term is in the integration with a combined cycle power station. However,
since the scope of this evaluation was limited to the desulfurization step, the
commercial evaluation is based on a more immediate practical objective; to
supply a desulfurized fuel gas to a conventional, gas-burning power station with
an approximate rating of 1400 MW. With respect to the total heat content of gas,
this is the same plant size that was considered in the initial feasibility
study. (2)
The economic re-evaluation is based on reflecting changes in the technical
design basis since the first study was made as well as effect of escalation in
costs due to inflation.
The design changes are the result of the experimental work conducted since
the last study was made. The principal items are the following:
1. Acceptor makeup rate has been increased from 1% to 2$ of the
circulation rate.
2. A third stage collector has been added to remove residual alkali
fumes and particulate matter prior to expansion of the gas.
3. The composition and temperature of the gas entering the gas
desulfurizer has been changed to reflect new data on gasifica-
tion kinetics.
B. Summary
The economics for the commercial design are based on an arbitrary start-up
date of January, 1980, and are summarized in Table III-l.
The 198O investment cost of the plant facilities to desulfurize producer gas
under pressure by this process is approximately $58 per installed KW, and the
operating costs correspond to about 2.9 mills/KWH of power generated.
A direct comparison with the economics previously presented in the December
1973 reportV*) is also summarized in Table III-l. For consistency, the costs
and capital charges of the previous design have been updated to reflect a July,
1975. base point with a January, 1980, start-up date. The purpose of this
11.
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TABLE III-l
Summarized Economic Evaluation
Regenerative Acceptor Desulfurization Process
7070 Plant Operating Factor (6132 hours/year)
Coal Required:
MM tons/year (6$ Moisture)
Higher Heating Value, Btu/year
Producer Gas to Station (after expansion):
Mols/hour
Temperature, °F
Pressure, psig
Higher Heating Value, Btu/year
HHV + Sensible Heat Content,(2) Btu/year
Desulfurization Plant Investment/3) Total, MM
Desulfurization Plant Investment, $/KW
Desulfurization Cost:
$/ton Feed Coal
j£/MM Btu HHV of Feed Coal
.._« Btu HHV of Product Gas
'MM Btu (HHV + Sens. Heat) Product Gas
Total Power Generated by Expander, megawatts^4)
Power Required for Compression of Process Air
for Gasification Plant, megawatts(5)
proposed Design
3,414,000
81.51 x 1012
236,679
650
10
62.72 x 1012
66.54 x 1012
81.2
58.0
7.40
31.0
40.3
38.0
340
170
Adjusted^1)
Economics of
Previous Design
3,422,000
81.71 x 1012
213,700
660
10
65.19 x 1012
68.68 x 1012
74.8
53.4
6.51
27.3
34.2
32.5
317
157
(i) Previous design economics of "hot" cyclones only case, Reference 1 updated
to reflect 1975 costs and projected escalation, interest rates and operating
rates to 1980.
(?) Sensible heat content above an assumed air heater outlet temperature of 300°F.
(3) January, 1980 operation, includes escalation and interest during construction.
Does not include gasification plant or coal cost.
(A) Assuming 91% of the insentropic efficiency for the expander (see Reference 3).
(5) Assuming 89% of the polytropic efficiency for air compressors.
12.
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comparison is to illustrate that the higher incremental costs of the proposed
design are mostly due to escalation and inflationary factors realized since the
July, 1973, evaluation rather than to design changes. The design changes
which have the most impact on the present economics are: l) acceptor make-up
requirements have been increased from 1% to 2% of the acceptor recirculation
rate, and 2) a third stage particulate removal system has been added.
A sensitivity analysis on the economics reveals the desulfurization process
is sensitive to total investment costs, utility rates, and sulfur credits while
fairly insensitive to dolomite costs, escalation rates and interest rates within
the realistic range of these variables.
As indicated in Table III-l, the pressurized producer gas is expanded from
system pressure (about 200 psig) to 10 psig through an expanding turbine to
generate 340 megawatts of power. The power output from this unit exceeds the
power requirement to compress the air for the coal gasification step. The
excess power so recovered is included within the net 1400 MW capacity of the
power station. This portion of the total station output would be produced at a
higher efficiency than that produced via the conventional steam cycle. Conse-
quently, the incremental reduction in the power station costs resulting from
this higher efficiency could be a substantial credit to the desulfurization
process. This credit has not been taken in this economic evaluation.
After expansion to 10 psig, the gas is delivered to the power station at a
preheated temperature of 650°F. This sensible heat content serves to improve
the efficiency (or the heat rate) of the station proper. In a conventional gas-
burning station, the fuel gas is delivered at ambient temperature, and sensible
heat from the products of combustion is recovered (via combustion air preheat)
to an air heater exit temperature of about 300°F. Therefore, it seems reason-
able to credit the process with the excess producer gas sensible heat above
3OO°F as an approximate means of measuring the improvement in the power station
efficiency. As shown in Table III-l, the credit to the process for this excess
sensible heat is about 2.3^/MM Btu or about 6$ of the system cost. It must be
emphasized, however, that it is beyond the scope of this contract to define the
actual dollar value of these power credits. (Refer to Figure III-l for a
schematic interpretation of the expander turbine plus conventional boiler and
steam cycle.)
The economics presented in Table III-l also apply for the case where a
hard, active naturally occurring dolomite stone is not locally available and it
is economically advantageous to artificially harden a local stone. It has been
shown experimentally that soft stone can be hardened without losing activity.
It is proposed that the hardening step could be accomplished in one gas
desulfurizer by injecting the entire make-up dolomite stone into a draft tube
with a stream of air. In the draft tube the stone is partially sulfided,
oxidized to CaS04, and then reduced back to the sulfide form to complete the
hardening process. As shown in Section H, since the process conditions and
the total producer gas composition are essentially equivalent to the natural,
hard stone case, the process economics are not significantly affected by the
minor process design modification discussed above. It must be emphasized that
all the experimental hardening studies were done in a two-stage process with
the oxidation step being accomplished in a vessel external to the desulfuriza-
tion process. Consequently, there is some technical uncertainty concerning the
residence time involved with artificially hardening a dolomite stone in a draft
tube of a desulfurizer.
13.
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FIGURE IH-1
Simplified Schematic of Expander Turbine
plus Conventional Boiler and Steam Cycle
Approximate Total Rating of Power Station- 1400 MW
Desulfurized
Producer Gas
195 psig - 1300°F
Air
Expander
Turbine
(91% Isentrop^
Efficiency)
Generator - 340 MW
170 - MW to Air Compressors
for Gasification
170 - MW Excess
Preheated Fuel Gas
10 psig - 650°F
Boiler
Stack
Gas
Generator ?d 1230 MW
Steam
Turbine
Heater
Train
Steam
Condenser
14.
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A few other areas of technical uncertainty still exist. It remains to be
demonstrated that particulate matter and alkali content of the producer gas
may be removable to acceptable levels by cyclonic devices alone. A "wet
scrubbing" system at the sacrifices of higher desulfurization costs and reduced
power recovery may be required to insure a "clean" feed gas to the turbo
expander. Furthermore, without a "wet scrubbing" system the ammonia content of
the gas is not reduced. Thus, the hot clean-up system will result in higher
NOX emissions from the power plant although the magnitude of the increase, due
to failure to remove ammonia, cannot be readily estimated.
C. Process Description
The scope of work under this contract covers the experimental develop-
ment and economic evaluation of a novel process for desulfurizing pressurized
producer gas as a fuel to a power station. The simple block flow diagram in
Figure III-2 defines the various steps in the overall process. Design and
evaluation of coal preparation (crushing and sizing), and of coal gasification
(Sections 100 and 200, respectively), are not covered in this study. Section
300, sulfur removal, and Section 400, sulfur recovery have been completely
evaluated and form the substance of this report.
A relatively small quantity of make-up C02 is required in the process
(Section 500). The C02 could be recovered either from a slip-stream of the
desulfurized producer gas before power generation, or from a portion of the
power station stack gas. Preliminary cost studies indicate approximately equal
costs for the two alternative locations. For the purposes of this study, it is
assumed that the make-up C02 is recovered from the pressurized producer gas
before delivery to the power station. A standard hot potassium carbonate
system is assumed to establish capital cost and utility requirements.
Sulfur Removal
A schematic flow diagram of the proposed commercial embodiment of this
gas desulfurization process is shown in Figure III-3 (CCDC Drawing No. AF-3667) .
Gasifier product gas (stream No. 4) flows to the bottom of a fluidized
gas desulfurizer (D-301). The gas fluidizes a bed of sized dolomite at a
temperature of 1678°F and at 15 atm. absolute pressure. Most (96$) of the H2S
in the gasifier gas reacts with the CaC03 component of the dolomite as follows;
H2S + MgO-CaC03 = MgO-CaS + C02 + H20,
The desulfurized gasifier gas then passes through two stages of
cyclones (G-301 and G-302, respectively) to remove substantially all of the
fines including attrited acceptor and entrained ash. A portion of the sensible
heat content of the desulfurized gas is used to preheat boiler feed water and
to generate and superheat the steam required for gasification of the feed coal
in heat exchangers C-3O3, C-3O2, and C-3O1, respectively. Any potential alkali
compounds formed at the lower gas temperatures are removed via a third stage
particulate collector, which also serves as a final process clean-up step. The
desulfurized producer gas at 130O°F and 210 psia (Stream No. 6) is then
delivered to the power station. Subsequent use of this gas is dependent on the
design of the power station, and is discussed in other sections of this report.
15.
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FIGURE III-2
Block Flow Diagram - Overall System
Desulfurized
Producer Gas
Feed Coal
Coal
Preparation
Section 100
Gasification
Section 200
Producer
Gas
Make-up
C02
Remova1
Section 500
Make-up
C02
Sulfur
Removal
Section 300
Power
Station
KW
Sulfur
Recovery
Section 400
I
Ash
Sulfur
-------
PAGE NOT
AVAILABLE
DIGITALLY
-------
Regenerated acceptor (Stream No. 5) at 130O°F is continuously charged
to the gas desulfurizers by gravity feed. The sulfided acceptor is pneu-
matically conveyed by a portion of the required recycle gas (stream No. 9) to
fluidized acceptor regenerators (D-302). Make-up acceptor (Stream No. 8) is
introduced into the system by a lockhopper system comprised of L-301, F-301,
F-302, and L-302. This small stream is also pneumatically conveyed to D-302
by means of recycle gas. Make-up C02 required in D-302 (Stream No. 10) is
pressurized by JC-301 and charged to D-302.
The fluidized acceptor regenerators (D-302) are maintained at 130O°F and
15 atm. absolute pressure. The sulfided acceptor is recarbonated at these
conditions by the reverse reaction,
MgO-CaS + C02 + H20 = MgO-CaC03 + H2S.
The carbonated magnesium component of the make-up acceptor is also
calcined at these conditions,
MgC03-CaC03 = MgO-CaC03 + C02.
Spent acceptor (2% of the circulating flow) is withdrawn from the regenera-
tors via a lockhopper, F-304, and a rotary feeder, L-306 (Stream No. 11). This
spent acceptor must be treated before disposal. Over 75% of the calcium
component of this stream is in the form of CaS. If this were disposed of
directly to the station ash pit, H2S gas would slowly evolve as the CaS was
hydrolyzed. To avoid this condition, the spent acceptor is directly contacted
with C02 and water in three stages of stirred reactors, D-303, to convert the
CaS to CaC03. At these conditions, the MgO component of the dolomite would also
recarbonate. The overall reaction is,
MgO-CaS + 2 C02 + H2O = MgC03-CaCO3 + H2S.
The spent acceptor is conveyed hydraulically by a dilute recirculating
slurry stream to hydroclones, G-306. The underflow slurry (controlled at 35
wt. $ solids) flows to the three series converters or Chance reactors, D-303A,
B and C, respectively. Each converter is agitated by a turbine-type mixer,
L-3O7, and some reaction heat is removed in each stage by coolers, C-3O4. The
spent acceptor, stripped of sulfur and fully carbonated, is pumped, J-301, to
the station ash pit. The overflow slurry from hydroclone, G-3O5, is continu-
ously recirculated via surge drum, F-305, and circulating pump, J-303. The
sensible heat content from the spent acceptor is removed by mixing with 90°F
make-up black water and by cooling in C-3O5. The make-up water is pumped from
the station pond overflow via J-302 (which also supplies water to cool the
gasifier ash and convey it to the station ash pond).
Acid gas resulting from the acceptor stripping operation (Stream No. 16)
is compressed, JC-302, to the liquid-phase Claus reactor pressure.
Sulfur Recovery
The process gas exiting the regenerators (Stream No. 12) is passed
through two stages of cyclones, G-304 and G-305, to remove entrained acceptor.
The sensible heat content of this stream is then exchanged with recycle gas in
18.
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C-401, and with boiler feed water (required for gasification'steam) in C-402.
Electrostatic precipitators, L-401, are provided to remove all entrained dust.
The gas then flows to the bottom of the liquid-phase Claus reactor, D-401.
The concentration of H2S in the gas from the acceptor regenerators
is 88$ of the equilibrium restriction at 1300°F and amounts to only about 3.6
volume percent. The liquid-phase Claus reaction developed in this work is
uniquely suited to processing this gas. The liquid-phase Claus reactors,
D-401, operate at 310°F and about 210 psia. Product gas from the acceptor
regenerators (Stream No. 12), gas from the reject acceptor stripping section
(Stream No. 16), and dilute aqueous H2S03 (Stream No. 17) are charged to the
bottom of the liquid-phase Claus reactors. Flow is upward through a sparged
reactor containing only aqueous sulfurous acid as was demonstrated during the
experimental work described in the previous report.l1) Liquid sulfur is pro-
duced by the reaction,
2 H2S + H2S03 = 3 S + 3 H20.
Liquid sulfur and liquid water flow from the reactor to decanter-type
separators, F-402. Unreacted gas (Stream No. 9), saturated with water vapor at
310°F, is compressed by JC-401, reheated in exchanger, C-401, to 935°F and
returned to the acceptor regenerator reactors. The sensible heat content of the
liquid water from F-402 (Stream No. 19), is exchanged with the feed acid (Stream
No. 17) in C-4Q3, further cooled to 90°F in C-404, and charged to the S02 , ,
absorption towers, D-402. A slip stream of water is rejected to the black water
pond to maintain water balance.
Approximately one-third of the sulfur from F-402 is burned with
stoichiometric air (Stream No. 20) to produce S02 in the pressurized combustor,
B-401. Excess heat is removed to boiler feed water via cooling tubes in the
walls. The exit gas from the sulfur combustor (Stream No. 21) flows to the
base of the SO2 absorption towers, D-4O2. Water (Stream No. 19) flowing down
through the packed towers absorbs the S02 in the gas by,
S02 + H20 = H2S03 aqueous.
Most of the exothermic heat of reaction is removed by side stream
coolers, C-405. The vent gas from the absorption towers (Stream No. 22) is at
90°F and 205 psia. It probably would contain some residual S02 (assumed in
this case as 0.3 volume percent). The most practical means of disposing of
this gas is to bleed it to an intermediate stage of the expansion turbine
required for the desulfurized producer gas in the power station. This would
decrease the net sulfur removal efficiency of this process from 95.8$ to
95.3$, but would increase the power output of the station slightly and would
eliminate an entire stack gas scrubbing installation following the S02 absorp-
tion step.
D. Design Basis
Gasification
Although the detailed design and evaluation of the gasification
system is not covered in this study, it is of interest to at least relate the
composition of the raw gasifier gas to the input feed streams of the gasifica-
tion section.
19.
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1. Feed Coal
This study is based on the use of a high-sulfur content, highly
caking Pittsburgh Seam coal. The analyses of a typical coal (Ireland Mine) of
this seam is shown below;
Feed Coal Analysis
Moisture, as received, Wt. $ 6.0
Ultimate Analysis. MF basis, Wt. ^
Hydrogen 4.8
Carbon 69.8
Nitrogen 1.2
Oxygen 7.6
Sulfur 4.3
Ash 12.3
Higher Heating Value, MF basis 12,TOO Btu/lb
2. Other Design Constraints
The total sensible heat content of the gasifier steam is supplied
by heat exchange within the sulfur removal and sulfur recovery sections. Addi-
tional design constraints for the fluidized bed gasifier and carbon burn-up cell
are;
System Pressure 15 atm
Gasifier Temperature 1775°F
Fixed Carbon Gasification Rate O.O037 Ib C/min-lb C in bed
Steam Conversion in Gasifier 44.2$
Burn-up Cell Temperature 1800°F
Percent Total Carbon to Burn-up Cell 16.4$
Gas Desulfurizer
The sulfur in the producer gas (as H2S) is removed in the gas
desulfurizer by reaction with CaC03 as follows:
MgO-CaC03 + H2S = MgO-CaS + C02 + H20.
The commercial design basis for this study is adopted from the actual
experimental data which are reviewed in a later section of this report.
Pertinent design values are;
20.
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Desulfurizer Temperature 1678 °F
Desulfurizer Pressure 15 atm
Avg . Acceptor Size, Dp arithmetic mean 0.065 inch
Acceptor Density, ps 130 Ib/CF
Mol $ Ca as CaS - Inlet Acceptor 78.6$
Mol $ Ca as CaS - Outlet Acceptor 90.3$
/Mol^CaSN _ (Mls_C^A
\ Mols Ca/ \ Mols Ca /
Avg. Acceptor Residence Time 2O min.
Outlet Fluidizing Velocity 3.4 ft/sec
Approach to Equilibrium -mrW,
(H20)(C02)/H2S °^
Removal of H2S from Producer Gas 95.8$
C02 Driving Force 0.147 atm(a)
(a) pco - equilibrium pco for the reaction, CaC03 = CaO + C02 .
Acceptor-Regenerator
The highly-sulf ided acceptor from the gas desulfurizer is regenerated
(converted to CaC03) at a lower temperature by the reverse reaction,
MgO-CaS + C02 + H20 = MgO-CaC03 + H2S .
Note that small yields of CO and COS were actually detected in the experimental
program, and S2 was found in the exit gas piping when the gas was deficient in
hydrogen. These trace constituents have been ignored in the economic evaluation.
The acceptor make-up requirement was arrived at by extrapolation of
data from the continuous unit. It was found experimentally that the acceptor
activity declined exponentially with number of cycles. The sulfur absorption
was assumed to involve 11.7$ of the total calcium. The rate of decline was
sufficiently small to select 2$ make-up rate.
The required make-up acceptor (as MgC03-CaCO3) is charged to the
acceptor-regenerator. The MgC03 component of the dolomite is completely
calcined at these conditions as follows:
MgC03-CaC03 = MgO-CaC03 + C02 .
Other pertinent design values for the commercial acceptor-regenerator
are:
Regenerator Temperature 1300°F
Regenerator Pressure 15 atm
Avg. Acceptor Size, Dp arithmetic mean 0.065 inch
Acceptor Density, ps 132 Ib/CF
Mol $ Ca as CaS - Inlet Acceptor 90.3$
Mol $ Ca as CaS - Outlet Acceptor 78.6$
Alois CaS\ . /Stole CaS\ (a)
\ Mols Ca/ \ Mols Ca/
Avg. Acceptor Residence Time 60 min.
Outlet Fluidizing Velocity 1.5O ft/sec
Approach to Equilibrium, (H20)(C02)/H2S 88$
(a) Excluding make-up acceptor.
21.
-------
Acceptor Converters (Chance Reactors)
In this process the spent acceptor to be rejected from the system must
be converted from MgO-CaS to MgC03-CaC03 before disposal. In the unconverted
state, the acceptor would slowly release H2S to the atmosphere.
The conversion is by the chemical reaction,
MgO-CaS + 2 C02 + H20 = MgC03-CaC03 + H2S.
The design of the acceptor converters was based on actual experimental data
described in Section XII, plus on work done by the Bureau of Mines staff at
Salt Lake City.
Pertinent design values are:
Conversion Temperature 2OO°F
Conversion Pressure Atmospheric
Wt. % Acceptor in Slurry 45
Total Residence Time 60 min.
No. Stages Required i 3
Utilization of C02 80$
Conversion of CaS to CaC03 Approaching 100$
Liquid-Phase Glaus
The offgases from the acceptor regenerators and the acceptor converters
are passed cocurrently through a reactor with a stream of aqueous H2S03 where the
following reaction occurs,
2 H2S + H2SO3 = 3 S + 3 H20.
The design basis is taken directly from the experimental results with
approximately 20% excess volumetric space rate to insure a 90$ conversion of
H2S to sulfur. Design values follow;
Reactor Temperature 310°F
Reactor Pressure 14 atm
Volumetric Space Rate 1116 hr"1
Avg. Vapor Residence Time 31 sec.
S02 Absorption
The dilute H2S03 required in the liquid-phase Claus reactors is
produced by absorption of S02 in water in standard countercurrent flow packed
towers. The design is based on published data for the system.(4>5) Key design
parameters are;
Absorber Temperature 9O-110°F
Absorber Pressure 14 atm.
Liquid Rate 10,OOO lb/hr-ft2
Packing 3" porcelain rings
No. Stages 4
Outlet Vapor Velocity 0.26 ft/sec
Percent Flooding Velocity
22.
-------
Table III-2
to
co
Basis: 1 Hour
Input
CH4
H2
CO
C0a
NH,
HaS
H20(v)
Sub-Total
(J) Acceptor from
Ace. Renen.
MgO CaS
MgO CaC03
Inerts
Sub Total
Totals
Output
(6) Producer Gas
CH4
H2
CO
C02
N3
NH3
H2S
H20(v)
Sub-Total
(T) Acceptor to
Ace. Reran.
MgO CaS
MgO CaCO,
Inerts
Sub-^Total
Datura: 6O"F, H3O (l)
U> MB
26,350 16.04
76,65O 2.O16
1,106,900 28.01
865,650 44.01
2,921,350 28. 02
9,150 17.03
47,850 34.08
545,700 18.016
5,599,6OO
1,015,950 112.46
345,350 140.41
195,400
1,556,700
7,156,300
26,350 16.04
78,5OO 2.O16
1,O8O,75O 28.01
965,950 44.01
2,921,350 28.02
9,15O 17 .03
2,OOO 34 .OS
553,1OO 18.016
5,637,150
1,167,200 112.46
156,550 14O.41
195,400
1,519,150
Mols
1,644
38,012
39,519
19,669
104,259
538
1,4O4
30,289
235,334
9,034
2,460
-
11,494
1,644
38,947
38,594
21,949
104,259
538
59
30,699
236,679
10,379
1,115
-
11,494
Mol 1
O.7O
16.15
16.79
8.36
44.30
0.23
O.6O
12.87
100.00
78. 6O
21. 4O
-
100.00
0.70
16.46
16. 3O
9.27
44. 05
O.23
0.02
12.97
1OO.OO
90.30
9.7O
-
100.00
n
6,60O
76,650
x
X
X
1,600
2,850
61,050
148,750
x
X
x
X
148,750
6,600
78,5OO
x
x
X
1,600
150
61,90O
148,750
x
x
x
X
c
19,750
x
474, 6OO
236,250
x
x
X
x
730,600
x
29, 550
x
29,550
760,150
19,750
x
463,400
263,600
x
x
X
x
746,750
x
13,400
x
13,400
Gas Desulfurizer
Elemental Balance
N
x
X
X
X
2,921,350
7,550
x
X
2,928,900
x
x
x
X
2,928,9OO
x
X
X
X
2,921,350
7,550
x
X
2,928,9OO
x
X
x
X
0
x
X
632, 3OO
629,400
x
x
X
484,650
1,746,350
x
118,050
x
118,050
1,864,4OO
x
x
617,350
702,350
x
X
X
491,200
1,810,900
x
53,500
x
53,5OO
HHV
t lb
s
x
x
X
X"
X
X
45,000
x
45,OOO
289,650
x
x
289,650
334,650
x
x
X
X
X
X
1,850
x
1,85O
332,800
x
x
332,800
Ash
x
X
X
X
X
X
X
x
X
X
X
195,400
195,400
195,400
x
X
X
X
X
X
X
x
X
X
X
195,400
195,400
JteO Ca
x
X
X
X
X
X
X
x
X
726,300
197,750
x
924,050
924,050
x
X
X
X
X
X
X
x
X
834, 4OO
89,650
x
924,050
Temp
°F
1775
1775
1775
1775
1775
1775
1775
1775
13OO
1300
1300
1678
1678
1678
1678
1678
1678
1678
1678
1678
1678
1678
AH
24,507 Btu/Mol
12,235 Btu/Mol
12,962 Btu/Mol
2O, 293 Btu/Mol
12,822 Btu/Mol
24,507 Btu/Mol
17,1O5 Btu/Mol
34,755 Btu/Mol
28,569 Btu/Mol
45,746 Btu/Mol
296.30 Btu/lb
22,661 Btu/Mol
11,512 Btu/Mol
12,164 Btu/Mol
18,982 Btu/Mol
12,O33 Btu/Mol
22,661 Btu/Mol
15,972 Btu/Mol
33,736 Btu/Mol
38,187 Btu/Mol
62,42O Btu/Mol
398.09 Btu/lb
AH unit
MM Btu Btu/lb
40.3 23,878
465.1 61,100
512.2 4,347
399.1 x
1,336.8 x
13.2 9,68O
24.0 7,100
1,052.7 x
3,843.4
258.1
112.5
57.9
428.5
4,271.9
37.2 23,878
448.4 61,100
469.3 4,347
416.6 X
1,254.6 x
12.2 9,68O
0.9 7,100
1,035.7 x
3,674.9
396.3
69.6
77.8
543.7
Total
MM Btu
629.5
4,682.2
4,811.8
x
X
88.7
339.7
x
10,551.9
-
-
-
629.5
4,797.4
4,698.O
x
x
88.7
14.3
x
1O,227.9
-
-
-
Heat of Reaction
MgO CaCO3 +
CO +
Heat Loss
Totals
HjS = MgO CaS +
H2O = CO2 -t- H2
7,156,300
CO2 + H.
,0 (1345
Mols x 30
83O Btu/Mol)
(935 Mols x 1430 Btu/Mol)
-
-
148,750
76O, ISO
2,928,900
1,864,400
334,650
195,400
924, O5O
41.5
1.3
10.5
4,271.9
Note: Circled numbers refer to stream number on CCDC Dwg. AF-3667. Temperature/Pressure may differ depending on location of Identifying flag.
-------
E. Material and Heat Balances
Basic to the process design and economic evaluation of this process on a
commercial scale was the completion of mass and heat balances for each important
processing step. The capacity of the plant in this study (with respect to the
total heating value of the desulfurized producer gas above 60 °F), is the same
as in the initial feasibility study.(2) This plant would supply a modern con-
ventional gas-fired power station of approximately 1400 MW rating.
The computer program developed by CCDC for the integrated system of coal
preoxidation, gasification, carbon-burn-up cell and producer gas desulfurization
was described in the last report.l1)
With these calculations as a starting point, the following mass and heat
balances were evolved for the commercial process design:
Table III-2 - Gas Desulfurizer
Table II1-3 - Acceptor Regenerator
Table III-4 - Acceptor Stripping Column
Table III-5 - Liquid-Phase Glaus Reactor
Table III-6 - Sulfur Combustor
Table III-7 - S02 Absorption Column
Table III-8 - Overall Mass and Heat Balance
F. Plant Investment and Operating Cost Estimates
From the heat and material balances, a process flow sheet was developed as
shown in Figure III-3 (CCDC Dwg. No. AF-3667). A preliminary capital cost
estimate was developed by the CCDC Engineering Department for Sections 300 and
400, and is summarized in tabular form in Appendix A as Tables A~l and A-2.
Note that Section 500, the make-up C02 system was assumed to be a conventional
"hot pot" system. A detailed estimate was not required since the costs and
utility requirements are adequately defined in the literature.(6) Furthermore,
costs that are not directly associated with desulfurization, i.e., air com-
pressors for the gasifiers and the expander turbine, are not included in this
desulfurization investment.
The capital cost estimates developed in Appendix A are defined as inside
battery limits (iSBL). To these costs must be added off-site facilities (OSBL)
which include utility costs, electric substations, cooling water towers, distri-
bution piping and boiler feed water treating. The accuracy of the capital cost
estimates in this study is estimated as ± 25$.
It was necessary in this study to project escalation in the plant invest-
ment and operating costs to reflect a presumed start-up date of January, 1980.
This work was based upon the following assumptions:
1. July, 1975 costs as a base point.
2. Begin design and construction, January, 1976.
3. Begin operations, January, 1980.
4. Escalation at 9-1/2$ per year average of construction
labor and material.
5. Interest during construction at 8-1/2$ simple interest
on cash flow.
6. Escalation of direct operating labor at 7$ per year.
7. Interest on working capital at 8-1/2$ per year.
24.
-------
TABLE III-3
Mass and Heat Balance
Basi
Inpv
©
©
©
©
Outi
©
(?)
(12)
.s: 1 Hour Datum; 60°F, H2O (l)
LL
Acceptor from Gas Desulfurizer
MgO CaS
MgO CaC03
Inerts
Sub-Total
Make-Up Acceptor
MgC03 CaC03
Inerts
Sub-Total
Gas from Liquid-Phase Claus
COZ
H2S
H20 (v)
Sub-Total
Make-Uo CO, Gas
C02
H20 (v)
Sub-Total
Heat of Reaction
MgO CaS + CO2 + H20 = MgO CaCO3 + H
Totals
jut
Acceptor to Gas Desulfurizer
MgO CaS
MgO CaCO3
Inerts
Sub-Total
MgO CaS
MgO CaCO3
Inerts
Gas to Liquid-Phase Claus
CO2
H2S
H20 (v)
Sub -Total
Heat of Reaction
MgC03 CaCO3 = MgO CaCO3 -(- CO2 (230
Heat Loss
Totals
Ac
Lb
1,167,200
156,550
195. 4OO
1,519,150
42,4OO
3.950
46,350
1,027,850
5,100
231,050
1,264,OOO
36,55O
2,75O
39,3OO
2S (1164 mols
2,868,800
1,015,950
345,350
195,400
1,556,7OO
20,300
6,9OO
3.95Q
31,150
1,023,350
44,75O
212,850
1,280,950
mols x 42,850
2,868,800
ceptor Regenerator
Mols
10, 379
1,115
—
11,494
23O
—
230
23,355.1
149.4
12,828.8
36,333.3
831
152
983
x 3O,830 Btu/mol)
—
9,034
2,46O
—
11,494
181
49
—
23O
23,252.1
1,313.4
11.816.8
36,382.3
Btu/mol )
—
Mol fl
9O.3
9.7
—
1OO.O
—
—
—
64.28
.41
35 .31
10O.OO
84.5
15.5
1OO.O
78.6
21.4
—
1OO.O
78.6
21.4
—
10O.O
63.91
3.61
32.48
100.00
Temp.
°F
1678
1678
1678
60
6O
935
935
935
4OO
4OO
1300
13OO
13OO
13OO
13OO
1300
13OO
1300
13OO
AH
38,187 Btu/Mol
62,42O Btu/Mol
398 .09 Btu/lb
/
—
—
9,4O6 Btu/Mol
7,920 Btu/Mol
26,395 Btu/Mol
3,314 Btu/Mol
21,758 Btu/Mol
28,569 Btu/Mol
45,746 Btu/Mol
3OO.65 Btu/lb
28,569 Btu/Mol
45,746 Btu/Mol
3OO.65 Btu/lb
13,955 Btu/Mol
11,735 Btu/Mol
29,932 Btu/Mol
AH
MM Btu
396.3
69.6
77.8
543.7
x
x
X
219.7
1.2
338.6
559.5
2.8
3.3
6.1
35.9
1,145.2
258.1
112.5
58.7
429.3
5.2
2.2
1.2
8.6
324.5
15.4
353.7
693.6
9.9
3.8
1,145.2
Note: Circled numbers refer to Stream Number on CCDC Dwg. AF-3667.
Temperature/Pressure may differ depending on location of identifying flag.
-------
Table III-4
Mass and Heat Balance
Acceptor Stripping Column
Basis: 1 Hour
Datum: 60°F, H20
Input
(ll) Acceptor from Ace. Regen.
MgO CaS
MgO CaCOs
Inerts
Sub-Total
HL3) CO2 from External Source
CO2
H2O (v)
Sub-Total
(14) Water (t)
Heat of Reaction
MgO CaS + C02 + H20 = MgO CaCOs + H2S
(181 mols x 30,830 Btu/mol)
MgO CaCOs + CO2 = MgCO3 CaCOs
(230 mols x 42,850 Btu/mol)
Lb.
20,300
6,900
3,950
31,150
22,650
1,700
24,350
58,800
Mols
181
49
230
514
94
608
-
Temp.
OF
1300
1300
1300
238
238
90
AH or Cp
28,569 Btu/mol
45,746 Btu/mol
300. 65 Btu/lb
1654 Btu/mol
20,449 Btu/mol
29.93 Btu/lb
AH
MM Btu
5.17
2.24
1.19
8.60
.85
1.92
2.77
1.76
Totals
114,300
Output
(15) Reject Acceptor to Pond
5.58
9.86
28.57
X '
(16)
MgCOs CaCOs
Inerts
Water (J)
Sub-Total
Acid Gas to Liquid Phase Claus
C02
H2S
H20 (v)
Sub-Total
Heat Rejected to Cooling Water
42,
3,
56,
103,
4,
6,
11,
400
950
650
000
550
150
600
300
230
-
-
-
103
181
34
318
120
120
120
200
200
200
40.3
0.21
Btu/mol °F
Btu/lb °F
59.86 Btu/lb
1287
1156
1117
Btu/mol
Btu/mol
Btu/lb
,
0.
3.
4.
0.
0.
0.
1.
23.
56
05
39
00
13
21
67
01
56
Totals
114,300
28.57
Note: Circled numbers refer to Stream Number on CCDC Bwg. AF-3667.
Temperature/Pressure may differ depending on location of
identifying flag.
26.
-------
Table III-5
Mass and Heat Balance
to
Liquid-Phase Claus Reactor
Basis: 1 Hour Datum: 6O°F, H2O (l)
Input
(12) Gas from Acceptor Regenerator
©
(IT)
C0a
H2S
HaO (v)
Sub-Total
Gas from Acceptor— Stripping Column
coa
H2S
HaO (v)
Sub-Total
Liauid from SOo Absorber
HjSOj
HaO (1)
Sub-Total
Heat of Reaction
2 HaS + H2SO3 = 3 S + H2O (1345
Totals
Lb
1,O23,35O
44,75O
212,850
1,28O,95O
4,55O
6, ISO
600
11, SOO
55,2OO
721,750
776,950
mols H2S x 42,98O
2,O89,2OO
Mols
23,252.1
1,313.4
11,816.8
36,382.3
1O3
181
34
318
672.5
4O,O75.0
4O,747.5
Btu/mol )
Mol ^
63.91
3.61
32.48
lOO.OO
32.39
56.92
10.69
100.00
1.65
98.35
100.00
Output
(9)
©
coa
HaS
HaO (v)
Sub-Total
Sulfur (1)
ffater (l)
Heat Loss
1,027,850
5,100
231,050
1,264,000
64,700
740,500
23,355.1
149.4
12 f 828. 8
36,333.3
2,O17.5
41,114.5
64.28
.41
. 35.31
100.00
—
Temp.
335
335
335
305
3O5
305
239
239
31O
310
31O
31O
31O
AH or CD
2,621 Btu/Kol
2,3O8 Btu/fool
21,333
2,34O Btu/teol
2,O49 Btu/»ol
2O,756 Btu/faol
1.0 Btu/lb-°F
2,391 Btu/Mol
2,O92 Btu/Mol
2O,797 Btu/^Iol
O.2 Btu/lb-°F
251.86 Btu/lb
Totals
2,089,200
AH
MM Btu
61.0
3.0
252.1
316.1
0.2
O.4-
0.7
1.3
138.7
57.8
513.9
55.8
O.3
266.8
322.9
3.2
186.5
1.3
513.9
Note: Circled numbsre refer to Stream Number on CCDC Dwg. AF-3667.
Temperature/Pressure may differ depending on location of Identifying flag.
-------
Table II1-6
Mass and Heat Balance
Sulfur Combustor
Basis: 1 Hour Datum; 60°F, H20 (l)
Lb
Input
Sulfur(s)
(20) Air
21,800
Mols
679.5
Temp.
60
AH
02
N2
H20 (v)
Sub-Total
21,750 679.5 495
71,600 2,556.2 495
5OO 28.4 495
93,850 3,264.1
Heat Required to Melt Sulfur and Preheat to 28O°F
Heat of Reaction
S + O2 = S02 (679.5 mols x 127.690 Btu/Mol)
Totals
Output
(2$ Gas to Absorption Tower
S02
N2
H20 (v)
Sub-Total
115,650
43,550
679.5 400
71,600 2,556.2 400
500
28.4 400
115,650 3,264.1
Heat Transfer to B.F.W. Preheater
Heat Loss
Totals 115,650
3,O44 Btu/Mol
22,257 Btu/Mol
3,487 Btu/Mol
2,258 Btu/Mol
21,261 Btu/Mol
AH
MM Btu
9.85
0.63
10.48
1.37
86.77
98.62
2.37
5.77
0.60
8.74
78.80
11.08
98.62
Note: Circled numbers refer to Stream Number on CCDC Dwg. AF-3667.
Temperature/Pressure may differ depending on location of
identifying flag.
28.
-------
Table III-7
Mass and Heat Balance
S02 Absorption Tower
Basis: 1 Hour Datum: 60 °F, H20 (l)
Temp .
Lb Mols °F AH or Co
Input
f2l) Gas from Sulfur Combustor
S02
N2
H20 (v)
Sub-Total
@ Water (l)
Heat of Reaction
S02 + H20 = H2S03 (aq.)
Totals
Output
(22) Vent Gas
S02
N2
H20 (v)
Sub -Total
fl?) Dilute H,SO, to Claus Unit
H2S03
H20 (l)
Sub-Total
43
71
115
733
,550
,600
500
,650
,500
(672.5
849
71
72
55
721
776
Heat Re.iected to Cooling Jate
Totals
849
,150
450
,600
150
,200
,200
.750
,950
r_
,150
679.
2,556.
28.
3,264.
mols x 13
—
7.
2,556.
8.
2,572.
672.
40.075.
40,747.
__
5
2
4
1
y
0
2
8
0
5
0
5
400 3,487
400 2, 258
400 21,261
90 29.93
720 Btu/Mol)
9O 285
90 195
90 19,328
Btu/Mol
Btu/Mol
Btu/Mol
Btu/lb
Btu/Mol
Btu/Mol
Btu/Mol
^ > l.O Btu/lb - °F i
AH
MM Btu
2
5
0.
8
21
9
39
.37
.77
.60
.74
.95
.23
.92
X
0
O.
0
!»
7
39
.50
.17
.67
.85
.40
.92
Note: Circled numbers refer to Stream Number on CCDC Dwg. AF-3667.
Temperature/Pressure may differ depending on location of
identifying flag.
29.
-------
Table III-8
Overall Mass and Heat Balance
Basis: 1 Hour
Input
(*) Gaaifier Gas
CH4
Ha
CO 1
CO,
Na 2
NHj
HaS
H,0 (v)
Sub-Total 5
(^ Make-up Acceptor
MgCO, CaCOj
Inerts
Sub-Total
(S5) Air to Sulfur
Combustor
oa
N,
HaO (v)
Sub-Total
Make-up Gas
C0a
HaO (v)
Sub-Total
Water (l) To Ac-
ceptor Stripping
Beat of Cornel-ess
Heat of Reaction
HaS + 1/2 03 =
HaS + 3/2 Oa =
Totals 5
Datum:
Lb
26,35O
76,650
, 1O6,9OO
865,650
,921,350
9,150
47,850
545,700
,599,6OO
42,4OO
3 . 95O
46,350
21,750
71,600
500
93,85O
59,2OO
4,45O
63,650
58,800
1°P
S + HaO
S0a + H20
,862,250
60°F, H,0 (1)
MW
16 .04
2.O16
28 .01
44.01
28.02
17.03
34. OS
18.016
184.42
—
32.00
28.02
18.O16
44.01
18.016
18.016
(1
(7
Hols
1,644
38,012
39,519
19,669
104,259
538
1,4O4
30,289
235,334
23O
—
679
2,556
2S.
3,264
1,345
246
1,591
-
,338 mo la
mols HaS
Mol %
0.70
16.15
16.79
8.36
44. 3O
0.23
0.60
12.87
100.00
-
—
.5 20.8
.2 78.3
.4 O.9
.1 1OO.O
84.5
15.5
1OO.O
-
HaS x 113,
X 242,080
H
6,600
76,65O
x
X
X
1,6OO
2,850
61.050
148,750
x
x
X
X
X
5O
50
x
5OO
5OO
6,6OO
C
19,75O
x
474, 6OO
236,250
x
x
X
X
73O,60O
5,5OO
—
5, SCO
x
x
X
X
16, ISO
x
16,150
x
Elemental Balance .
N
x
x
X
X
2,921,350
7,550
x
x
2,928,900
x
x
X
X
71,600
x
71,6OO
x
x
X
X
0
x
X
632,300
629,400
x
x
X
484,650
1,746,350
18,400
x
18,400
21,750
x
4 SO
22,2OO
43,O5O
3,950
47,OOO
52,200
Ib
g
x
x
X
X
X
X
45,000
x
45,000
X
X
X
X
X
x
X
X
X
X
X
Ash
X
X
X
X
X
X
X
x
X
X
3,950
3,950
x
X
X
X
X
X
X
X
MgQ Ca
x
X
X
X
X
X
X
X
X
18,500
X
18,500
x
X
X
X
X
X
X
X
Temp
°r
1775
1775
1775
1775
1775
1775
1775
1775
60
60
60
6O
6O
14O
14O
90
AH or CD
24,507 Btu/Mol
12,235 Btu/Mol
12,962 Btu/Mol
2O, 293 Btu/Mol
12,822 Btu/Mol
24,5O7 Btu/Mol
17,1O5 Btu/Mol
34,755 Btu/Mol
-
—
-
-
19,O95 Btu/Mol
736 Btu/Mol
19,7O8 Btu/Mol
29.93 Btu/lb
92O Btu/mol)
Btu/mol;
155,900
752,250
3,000,500
1,886,150
45,000
3,950
18,500
AH
MM Btu Btu/lb
4O.3 23,878
465.1 61,1OO
512.2 4,347
399 . 1 x
1,336.8 x
13.2 9,680
24 .O 7,1OO
1.O52.7 x
3,843.4
x -
X
X
X X
X X
O.S x
0.5
1.0 x
4.8 x
5.8
1.8
42.4
152.4
1.7
4,O48.0
UHV
MM Btu
629.5
4,682.2
4,811 .8
x
X
88.7
339.7
x
1O,551.9
-
—
X
X
X
X
X
X
X
-
-------
Table III-8 (continued)
Elemental Balance.
lb
Temp
HHV
Output
(6J Producer Gaa
CH«
H2
CO
coa
NHa
HaS
H20 (v)
Sub-Total
JB) Reject Acceptor .
MgCOs CaCO,
Inerts
Water (l)
Sub-Total
Sulfur (1)
Waste Water (l)
a) Vent Gas - S0a Absorber
S0a
HaO (v)
Sub-Total
Lb
26,350
78,500
1,080,750
965,950
2,921,350
9,150
2,000
553,100
5,637,150
42,400
3,950
56,650
103,OOO
42,900
7,000
45O
71,6OO
150
72,200
MW
16.04
2.016
28.01
44.O1
28.02
17.O3
34.08
18.016
184.42
-
18.016
32.O6
18.016
64.06
28.02
18.016
Hols
1,644
38,947
38,584
21,949
104, 259
538
59
30,699
236,679
230
-
-
1,338
-
7.0
2,556.2
8.8
2,572.0
Mol *
0.7O
16.46
16. 3O
9.27
44.05
0.23
O.O2
12.97
100. OO
-
-
-
-
-
O.27
99.39
0.34
1OO.OO
H c
6,600 19,75O
78,500 x
x 463,400
x 263,600
x x
1,600 x
ISO x
61,900 x
148,750 746,750
x 5,500
x x
6,350 _x
6,35O 5,5OO
x x
800 x
X X
X X
X X
X X
N
x
x
X
X
2,921,350
7,550
x
X
2,928,900
x
X
X
X
X
X
X
71,600
X
71,600
O
x
X
617,350
702,350
x
x
X
491,200
1,810,900
18,400
x
50,300
68, TOO
x
6,200
200
x
150
35O
s
x
x
X
- X
X
X
1,850
x
1,85O
x
X
X
X
42,9OO
x
250
x
x
250
Ash
x
X
X
X
X
X
X
x
X
X
3,950
X
3,950
x
X
X
X
X
X
MeO Ca
x
X
X
X
X
X
X
x
X
18,500
X '
X
18,500
x
X
X
X
X
X
°F AH or Cp
1300 15,876 Btu/Mol
13OO 8,754 Btu/Mol
1300 9,123 Btu/Mol
1300 14,OOO Btu/Mol
130O 9,O37 Btu/Mol
130O 16,014 Btu/Mol
1300 11,735 Btu/Mol
130O 29,926 Btu/Mol
12O 31.3 Btu/Mol- °F
12O O.21 Btu/lb- °F
120 59.86 Btu/lb
310 0.2 Btu/lb- °F
9O 29.93 Btu/lb
9O 285 Btu/Mol
90 195 Btu/Mol
90 19,328 Btu/Mol
MM Btu Btu/lb
26.1 23,878
34O.9 61,10O
352.0 4,347
3O7.3 x - .•;
942.2 x
8.6 9,68O
0.7 7,100
918.7 x
2,896.5
0.4
0.1
3.4
3.9
2.1
0.2
x x
0.5 x
O.2 x
0.7 x
MH Btu
629.5
4,797.4
4,698.0
. x -_
X
88.7
14.3
x
10,227.9
-
-
-
-
x
X
X
X
Sensible Heat Transferred
to Gaaifier Steam
977.1
Sensible Heat to Cooling
Water
Heat of Reaction
CO + HSO = CO 3 + H2
Heat Loss
Totals
(935 mols
5,862,250
x 143O Btu/mol)
-
-
155,900 752,250
3,OOO,500
1,886,150
45,000
3,950
18,500
1O5.6
1.3
60.6*
4,O48.O
Note: Circled numbers refer to Stream Number on CCDC Dwg. AF-3667. Temperature/Pressure may differ depending on location of identifying flag.
* Includes 33.92 MM Btu heat loss from producer gas between desulfurizers and outlet of third stage particulate collector.
-------
The total investment summary for the proposed design is shown in
Table III-9. The installed cost in July, 1975, dollars is $54,100,000.
Escalation of these costs to January, 1980 (as defined above), increases the
total investment to $81,200,000.
Direct operating costs are shown in Table 111-10. Note, that an
average plant operating factor of 70$ or 6132 hours per year was assumed reflect-
ing power station practice. The unit values used to develop operating costs are
shown directly in Table I11-10.
G. Evaluation of System Economics
Base Case
The first report on this project(2) clearly showed that the most
economic application of this process is in combination with a combined cycle
power station. This current work is more restricted in scope and is limited to
a simple exposition of the costs of desulfurizing the pressurized producer gas.
Based on the cost estimates presented in the previous section, the
process economics are estimated as shown in the first column of Table III-ll.
The net annual operating costs are $25.267 MM/year. Since 58$ of this is
represented by capital charges, the economics are very sensitive to the invest-
ment costs. Furthermore, at a value of $25.00/metric ton the sulfur credit is
equivalent to about 11$ of the annual operating costs.
Based on a dolomite make-up rate of 2$ of the circulation rate and on
a dolomite cost of $16/ton, the annual acceptor cost is about 9$ of the net
operating cost. The 1980 acceptor cost was derived by escalating a 1976 quoted
price of a 4 x 16 mesh dolomite stone delivered by rail from Thornton, Illinois
to Hillsboro, Illinois. The 1980 based dolomite cost includes the investment
charge (at 18$/yr) of installing a rail spur and unloading facilities at the
Hillsboro location. Per recommendation of the dolomite supplier, the cost of
raw stone was escalated at 8$/yr. Transportation cost was escalated at
9-l/2$/yr. In the event that a hard, active naturally occurring dolomite is
not locally available, it may be preferable to artificially harden a local stone
to minimize the make-up acceptor costs (refer to Hardening discussion in
Section H).
The system costs presented in Table III-ll are more meaningful if
related to the coal required for the total system, or to the heating value of
the gas delivered to the power station. The net annual operating costs are
equivalent to $7.40/ton of feed coal and 31.0^/MM Btu HHV of the coal.
Considering only the higher heating value of the product gas to the
power station, the corresponding cost of desulfurization is 40.3^/MM Btu.
Allowing for the sensible heat credit above 300°F as discussed earlier, the
desulfurization cost is reduced to 38.0^/MM Btu^ or a reduction of about 6$.
Another credit to the process is the power generated by the expanding
turbine-generator set that is in excess of that required by the gasification
air compressors. As noted previously, it is beyond the scope of this contract
to determine the actual dollar credit associated with this higher efficiency
generated power. However, to at least qualitatively show the potential impact
32.
-------
CO
U)
Plant Section
Utilities Required
Electricity, KW
Cooling Water, gpm
Low-Pressure Steam, Ib/hr
Boiler Feed Water, gpm
Operating Labor Required
Men/Shift
Investment
Erected Cost (ISBL)
Off-sites Sc Utilities (OSBL)
Off-sites
Electrical
Cooling Water
Boiler Feed Water
Misc. Utilities
Total OSBL
Installed Plant Cost
(July, 1975)
Escalation to Jan. 198O
Sub-Total
Interest during Construction
Total Investment
(Jan., 1980)
Table III-9
Investment Summary
300
Sulfur
Removal
2,478
2,975
X
X
4
24,OOO
1,6OO
30O
2OO
X
5OO
2,6OO
40O
Sulfur
Recovery
10,365
7,744
X
1,584
3
( in $1OOO)
16,7OO
1,1OO
1,200
500
1,3OO
3OO
4,4OO
5OO
Make-up
CO2
System
340
16,OOO
78,3OO
X
1
5,OOO
3OO
-
1,000
X
100
1,4OO
Total
13,183
26,719
78,3OO
1,584
8
45,70O
3,OOO
1,500
1,700
1,300
9OO
8,4OO
26,600
21,1OO
6,4OO
54,100
14,100
68,2OO
13.OOO
81,2OO
-------
Table 111-10
Direct Operating Cost Summary
Excluding Acceptor Coat
Plant Section
Direct Operating Labor
Men/Shift
Direct Operating Cost
Basis: 70% Plant
300
Sulfur
Removal
4
Operating Factor
400
Sulfur
Recovery
3
(in $1000/yr)
500
Make -up
C02
System
1
Total
8
1. Operating Labor at
81,400/man/shift/yr
2. Maintenance Labor at
1.6% Installed Plant Cost
3. Direct Supervision
15% of 1 + 2
326
536
129
4. Indirect Overhead
50% of 1 + 2 + 3 495
5. Payroll Overhead
15% of 1 + 2 + 3 + 4 223
6. Maintenance Material
2.4% Installed Plant Cost 804
7. Miscellaneous Supplies
15% Maint. Material 121
8. Insurance & Taxes
2% Installed Plant Costs 670
9. Utilities
Electricity at 25 mills/KWH 380
Cooling Water at 9£/M Gal. 99
BFW at 85C/M Gal. x
Low-Pressure Steam
$2.10/M Ib x
Chemicals & Catalysts
(ex Acceptor) x
Sub-Total Utilities 479
Total Direct Operating
Cost (Jan. 1980) 3,783
244
426
101
386
174
638
96
532
1,589
256
495
2,340
4,937
81
130
32
122
55
194
29
162
52
530
x
993
27
1,602
2,407
651
1,092
262
1,003
452
1,636
246
1,364
2,021
885
495
993
27
4,421
11,127
34.
-------
Table III-ll
Economic Evaluation
Regenerative Acceptor Desulfurization Process
Basis: 70% Plant Operating Factor (6132 hr/yr)
Case
Coal Required
Ton/yr (6% Moisture)
Higher Heating Value, Btu/yr
Desulfurized Producer Gas To Station
Mols/Hr
Temperature
Pressure
Higher Heating Value, Btu/yr
HHV + Sensible Heat Content,
.per year
Cost Analysis
Installed Plant Cost (1975)
Escalation to 1980
Interest during Construction at
Total Investment
Investment per Installed KW
Working Capital
Annual Operating Costs
Direct Operating Cost (1980)
Acceptor at $16/ton
Interest on Working Capital at 8.5%
Capital Charges at 18% Investment
Sulfur Credit at $25/metric ton
Net Annual Cost
Desulfurization Cost Expressed:
In terms of feed coal
$/ton coal
6/MM Btu HHV
In terms of product gas to station:
£/MM Btu HHV
6/MM Btu (HHV -f Sens. Ht.)(1)
Power Generated by Expander
in Excess of Air Compressor Req'ts
(2)
Present Proposal
3,414,000
81.51 x 1012
236,679
650°F
10 psig
62.72 x 1012
66.54 x 1012
54.1 MM
14.1 MM
13.0 MM
81.2 MM
$58.0
2.75 MM
$11.127 MM/Year
2.274 MM/Year
0.234 MM/Year
14.615 MM/Year
($ 2.983 MM/Year)
$25.267 MM/Year
7.40
31.0
40.3
38.0
170 MW
Previous Design
Hot Cyclones Only
3,422,000
81.71 x 10
12
213,701
660°F
10 psig
12
12
65.19 x 10
68.68 x 10
49.8 MM
13.0 MM
12.0 MM
74.8 MM
$53.4
2.48 MM
$10.597 MM/Year
1.038 MH/Year
0.210 MM/Year
13.464 MM/Year
($ 3.024 MM/Year)
$22.285 MM/Year
6.51
27.3
34.2
32.5
160 MW
(1) Sensible heat content above an assumed air heater outlet temp, of 300°F.
(2) Assuming 91% of the isentropic efficiency for the expander, and 89%
of the polytropic efficiency for the air compressors.
35.
-------
of this credit upon processing costs, it is assumed that the real cost of
electric power generated by an expanding turbine is about 25$ less than the
cost of steam-generated power. At assumed projected 1980 electric rates, this
reduction in power plant cost would amount to 6-1/2 mills/KWH of excess power
produced by the expanding turbine-generator set. Therefore, the net credit to
the process would be about 9^/MM Btu. Obviously, this is a significant credit,
but it could only be evaluated with accuracy through a detailed study.
Comparison with Economics of Previous Report
Even though the design presented in the previous reportl1) is no
longer applicable, it serves as a bench mark to determine the effect of escala-
tion and design changes. The second column of Table III-ll depicts the system
costs for the 1973 proposed design with the economics updated to reflect 1975
costs and projected escalation interest rates and operating costs to 1980.
Furthermore, with the exception of the desulfurizers, 1975 ISBL installed plant
cost was based on a 36.7$ escalation increase from mid-1973 to mid-1975. The
desulfurizers were re-estimated on the basis of 1975 cost factors. The 1973
OSBL costs were also adjusted with 1975 factors.
A comparison of the two columns clearly shows that with the exception
of the increased acceptor make-up rate, the design changes proposed in this study
do not markedly affect the economics of the project. On the other hand, escala-
tion in construction and utility costs for the past two years plus higher
projected interest rates and escalations of construction and utility costs for
the future have increased desulfurization costs by about 62$, and account for
about 70$ of the total increased cost of desulfurization when compared to the
economics presented in the previous report.
Economic Sensitivity Analysis
An economic sensitivity analysis was done to determine the effect of
the major cost parameters on the overall process economics. The results are
shown in Figure III-4 in terms of desulfurization costs of product gas including
HHV and sensible heat.
The cost of desulfurization is very sensitive to investment costs.
Within the accuracy limits of the estimate (± 25$), desulfurization cost of the
product gas could vary from 30.4 to 45.6^/MM Btu.
Desulfurization cost is also fairly sensitive to utility rates. With-
in a predictable range of 1980 utility rates, desulfurization costs could vary
from 36.3 to 41.3^/MM Btu.
The effect of the varying sulfur credit on desulfurization cost is
noteworthy. Although a sulfur credit of $25/metric ton was chosen for the base
case, the future sulfur market is so uncertain that sulfur value could easily
range from $20 to $50/metric ton. At $50/metric ton, the sulfur credit would
recover about 21$ of the gross annual operating costs for desulfurization.
Expressed in another way, for every $10 increase in sulfur price, the cost of
desulfurization is reduced by 1-3/4^/MM Btu.
Conversely, desulfurization cost is fairly insensitive to acceptor
cost, escalation rate and interest rate within the projected ranges of those
variables. For example, an acceptor cost of 150$ the base would result in a
desulfurization cost of only about 1.7^ higher than the base case.
36.
-------
37.
-------
Potential Impact of Using Combined Cycle Power
Station on Desulfurization Economics
It is interesting to demonstrate the potential impact of using a
combined cycle power station in lieu of the conventional base case on the
economics of desulfurization. The effect of net station efficiency (basis
HHV coal) on installed cost of desulfurization per KW is displayed in Figure
III-5. The ordinate represents desulfurization costs based on the coal feed
rate presented in Table III-l. The reduction in desulfurization costs is
directly proportional to the increase in net power station capacity. As an
example, a combined cycle power station with a net station efficiency of 43%
(HHV coal) would yield about 1675 MW at a desulfurization cost of $48.50/in-
stalled KW.
H. Artificially Hardened Sorbents
The foregoing analysis assumed that a hard, active naturally occurring
dolomite was available. In the event such a stone cannot be located close to
a desired plant site, it might be preferable to artificially harden a local
stone to minimize transportation costs associated with the raw stone. It has
been shown experimentally that soft stones can be hardened without necessarily
losing activity. Once in hand, the hardened stone would behave in the process
in a manner indistinguishable from that of natural, hard dolomite. The tech-
nique for hardening consists of first partially sulfiding the stone and then
oxidizing it. Both of these steps are carried out at typical gas desulfurizer
temperatures.
The process is illustrated schematically in Figure III-6. One of the gas
desulfurizers would serve as a hardening vessel for the make-up stone. The
entire stream of make-up stone is fed into the hardening desulfurizer with a
stream of air via a draft tube in the middle of the vessel. Raw fuel gas is
fed to the bottom and is desulfurized as the stone is sulfided. Sulfided stone
is drawn into the draft tube where the CaS is partially converted to CaS04. The
fuel gas associated with the stone being drawn into the draft tube is burned.
At the tube outlet, CaS04 is reduced back to CaS and hardening is completed.
The hardened acceptor is withdrawn and sent to the acceptor regenerator.
The reactor operates adiabatically. The heat released by the oxidation
reactions is balanced by the heat required to heat and sulfide the make-up
stone and to calcine its MgC03 content.
A heat and material balance for the hardening desulfurizer is given in
Table 111-12. The gas fed to the hardening desulfurizer represents about 9%
of the total fuel gas; the amount of air added is about 0.3$ of the total gas.
The changes in gas composition are also barely noticeable in the overall fuel
gas composition as shown in Table 111-13 below. There would be no significant
change in total reactor volume, the hardening desulfurizer still performs the
task of H2S removal, and the existing gas requires no further treatment. How-
ever, as the heat and material balance for hardening requires only 9$ of the
gas to go through the hardening desulfurizer, the other desulfurizers would be
proportionately larger.
38.
-------
39.
-------
FIGURE III-6
Hardening Process
Cyclones
Hardening
Desulfurizer
Air
Entire Stream
of
Make-up Acceptor
Gasifier
Gas
To other
Desulfurizers
Desulfurized Gas
* to
Steam Exchangers
Producer Gas
from other
Desulfurizers
Hardened Acceptor
to
Acceptor Regenerator
Recycle Gas
from
Liquid Phase Claus
40.
-------
Table 111-12
Has a and Heat Balance
Basis: 1 Hour
Incut
Gaailier Gaa
CH4
Ha
CO
coa
Na
NH,
HaS
HaO (v)
Sub-Total
HgCOj CaCO,
Inarts
Sub-Total
Ail
Os
Na
H20 (v)
Sub-Total
Heats of Reaction
Ha + 1/2 Oa =
HH3 + 3/4 Oa =
CH4 + 2 Oa =
CO + 1/2 Oa =
CaS + 2 O,
CaSO4 + 4 Ha =
Totals
Output
Producer- Gas
CH,
Ha
CO
C0a
Na
NH,
HyS
HaO (v)
Sub-Total
Hardened Acceptor
to Accentor Ree .
HgO CaCO,
HgO CaS
Inerts
Sub-Total
CO + HaO
MgCOj CaCO, =
MgCaCOj + HaS =
Fraction MgCO,
Heat Loss
Totals
Datum: 6O°F, HaO (l)
La MB
2,400 16.04
7,OOO 2.O16
100,750 28.01
78,8OO 44.O1
265,950 28 .02
85O 17 .03
4,350 34.08
49,7OO 18.016
509,800
42,4OO 184.42
3,950
46,35O
3,250 32.00
10,8OO 28.02
1OO 18.O16
14,150
HaO
Na + 3/2 H.,0
CO, + 2 H20
CO,
CoSO«
4 H30 + CaS
57O,OOO
2,400 16.04
6,750 2.016
98,O50 28.01
98,600 44.01
276,750 28.02
85O 17 .03
200 34 .OS
53,850 18 .016
537,450
15,25O 14O.41
13,650 112.46
3.95O
32,850
COa + Ha
UgO CaCO, + COa
HgO CaS + C02 + H
CaCO, Oxidized = O
570,300
Mola Hoi
-------
Without Hardening
0.70$
16.46
16.30
9.27
44.05
0.23
0.02
12.97
100.00
With Hardening
0.70$
16.34
16.25
9.35
44.09
0.23
0.02
13.02
100.00
Pending further experimental work on the hardening step which would include
actual demonstration of hardening with a vessel containing a draft tube, there
is little incentive to go to a detailed design of the system. By virtue of the
reasoning given above, it is clear that no significant changes in process
economics will be involved. The changes incurred by going to an artificially
hardened stone are well within the accuracy of the current economic estimate.
TABLE 111-13
Producer Gas Compositions
With and Without Hardening
Total Producer Gas Compositions
Component
CH4
H2
CO
CO 2
N2
NH3
H2S
H20 (v)
Total
J. Technical Uncertainties
The proposed commercial design contains a few areas of technical un-
certainty. The quality of producer gas with respect to size and quantity of
particulate matter and alkali that may be charged to an expander is still one
area of concern. This design includes a third stage particulate collection
system to remove any alkali compounds that may form at the lower gas tempera-
tures resulting from the use of a portion of the gas sensible heat content to
generate and superheat the steam required for gasification. As with the two
stages of high pressure cyclones located upstream of the steam exchangers,
removal efficiency is highly dependent upon the contaminants being in a discrete
particle form. Furthermore, the steam exchangers are subject to fouling caused
by the potential plate out of alkali compounds at the relatively low temperature
tube wall conditions. One alternative is to "wet scrub" the gas as thoroughly
discussed in the previous report. (x) However, water scrubbing would result in
reduced power recovery from the expanding turbine-generator set and increased
plant investment costs. Since preliminary experimental results have shown no
traces of alkali bearing compounds in the producer gas once all the particulate
matter has been removed at 130O°F, the system of "hot" cyclones plus the third
stage final clean-up step was selected as the basis for this commercial design.
A review of the "state of the art" in particulate removal technology at elevated
temperatures is discussed in a recent report to EPRI.V7)
Another contaminant in the producer gas is ammonia. A portion of the
ammonia will be converted to NOX when burned in the power plant. The degree of
conversion is unknown but, conceivably the resultant NOX emissions could
exceed "new source" standards. Results of most flame tests and furnace
studies!8) indicate a low conversion of NH3 to NOX, i.e., less than 10$. To
our knowledge, this has not been confirmed in utility steam generators.
Furthermore, the ammonia concentration assumed to be present in the gas (O.23
mol percent) has not been established. It conceivably could be significantly
42.
-------
lower in concentration. Control methods presently used to reduce NOX
emissions in gas, oil and coal-fired utility steam generators such as over
fire air simulation!9) and low-excess air combustion!1O) should also be
applicable for burning the subject gas. If the ammonia content of the
producer gas has to be reduced prior to burning in a power station, then a
"wet scrub" system as discussed above would be one alternate.
There is also some technical uncertainty concerning the residence time
involved with artificially hardening a dolomite stone in a draft tube of a
gas desulfurizer. All experimental work to date on hardening has been done
in a two-stage process with the oxidation step being accomplished in a vessel
external to the desulfurization process. Even though the hardening reactions
proceed rapidly, there is no experimental work to substantiate that the brief
residence time inherent with the draft tube design would be sufficient to
satisfy the hardening requirements. The incentive of a draft tube design com-
pared to a dual-vessel concept are two fold: the installed costs would be less
and the operability should be better. It must be re-emphasized, however, that
further experimental work on the hardening step should include demonstration
in a draft tube of a desulfurizer vessel.
43.
-------
IV. EQUIPMENT AND PROCEDURE
A. Gas Desulfurizer-Regenerator System
The gas desulfurizer-regenerator system used in this study was the same one
which was described in great detail previously.'1) Figure IV-1 is the process
and instrumentation diagram of the unit; the reader is referred to the previous
report in this series for a detailed account of the equipment and its auxiliaries.
Only changes in the system are detailed below.
In order to make the system more reliable, the hot filter system which
removes attrited acceptor fines from the desulfurizer product gas was modified.
Ball valves which can be closed in the presence of solids replaced the gate
valves which tended to become impacted with fines making it impossible to switch
filters. This had been the cause of several shutdowns. The new valves were for
severe service and had hardened ball stem tips.
After Run A132, the gas desulfurizer reactor was replaced with one made from
aluminized ("Alonized") 310 ss. The process of aluminizing consists of heating
the equipment in a bed of fine aluminum and alumina powder. The aluminum diffuses
into the seel and, after further heat treatment, is converted to an outer layer
of aluminum oxide which is purported to be more resistant to H2S corrosion than
the original steel. • The aluminized reactor was not in place long enough to allow
a final verdict on its performance. It is noted that the aluminizing process
made the vessel brittle and difficult to either weld or machine.
H2S Analyzer
An ultraviolet continuous H2S analyzer was purchased to monitor both
vessels. The unit was a DuPont series 400 analyzer equipped with scales of 0-10
and 0-27, H2S. Performance was generally satisfactory, but there was a continued
tendency for the baseline to shift apparently due to deposition of sulfur on the
cell lenses.
CS2 Supply System
The original liquid H2S feed system was replaced with a CS2 feed system.
Some of the various reasons for the substitution were as follows;
1. The steel cylinders in which H2S was delivered contained
not only H2S but also products of corrosion, oil, water,
noncondensable gases and organic compounds of sulfur. An
attempt to distill H2S into new, clean steel cylinders
failed because of the presence of noncondensable gases.
2. Even a micron porometallic filter in the supply line from
cylinders to the metering pump did not prevent solid
matter from entering the chamber of the pump. Solid
matter deposited in the check valves of the pump resulting
in an erractic, unreliable delivery of liquid H2S.
44.
-------
3. At the H2S flashing point, the sudden drop in tempera-
ture caused water in the liquid H2S to freeze and soluble
matter to deposit and plug the feed line.
Tests for decomposition of the CS2 vapor at the desulfurizer reactor
conditions have shown a satisfactory hydrolysis of CS2 in the presence of H2,
H20 and C02 in the inlet gas. When no hydrogen was fed into the desulfurizer
reactor, a small amount of solid carbon was found in the sulfur acceptor and
in the gas preheater tubing.
A schematic of the CS2 feed system is shown in Figure IV-2. Liquid
CS2 was kept under a thin layer of water in a glass graduated holder. CS2
withdrawal from the holder was metered via a Lapp Model LS-20 diaphragm pump
with a jacketed head maintained at 4°C (40°F) to prevent variations in the
specific volume of liquid CS2 and vapor lock.
Hourly readings of the liquid CS2 level in the reservoir were used
for an observation of the metering pump performance.
The calibration of the pump has shown an accuracy within ± 1$ of a
feed rate, the variance being due to the pump double ball check valve slippage.
Liquid CS2 was fed into the gas mixing manifold most of the time.
However, when the steam generator had to be used, the CS2 feed line was
connected to the saturator exit line to prevent CS2 hydrolysis in the
saturator.
B. Hardness Testing
Two techniques were used to check the attrition resistance of stones on a
laboratory scale. The first was a rolling mill test. About 100 grams of
material to be tested was placed in a Roalox tumbling jar together with ten
size OO Burundum grinding cyclinders (13/16" x 13/16"). The jar was tumbled
in a rolling mill for 30 minutes at 51 rpm. The sample was then screened, and
the external surface area calculated from its size distribution. The hardness
was the increase in surface area which occurred due to the test divided by the
increase observed for a standard hard stone. The standard was A6 product, a
Tymochtee dolomite which had been hardened in the C02 Acceptor Process, which
displayed an increment in surface area of 220 cma/gm.
The second method was to perform a drop test with the apparatus shown in
Figure IV-3. A sample of test stone was screened before and after a metal rod
crushed it in a small holder. This latter test was abandoned in favor of the
rolling mill test as the rolling mill test gave results more in agreement with
the stone's actual attrition resistance in the fluidized bed system.
C. Fine CaC03 Runs
A simplified flow diagram for a trial run using Fisher precipitated calcium
carbonate in a water slurry is shown in Figure IV-4.
The regenerator vessel (D-l) was converted into a solids trap by installing
two star shaped porometallic filter elements. The solids receiving space was
maintained at 343°C (650°F) for "freezing" all downstream reactions yet
preventing water vapor condensation.
45.
-------
PAGE NOT
AVAILABLE
DIGITALLY
-------
FIGORE IV-2
CS., SUPPLY SYSTEM
Graduated
Reservoir
Antl-Freeze
Holder
-(PIC
CS, to
Manifold
Metering Diaphragm
Pump with Jacketed
Head
Ng Purge
Out
-------
Keyway
l" ^ Rod x 6" Long
Guide Sleeve
l" Sch. 80 Pipe x 3-1/2"
achine I.D. for Free
Sliding Fit
arden End of Rod
Case Harden
:up 1-1/2" O.D. x
1.02" I.D.
3/4" High - 3/8" Deep
Ring with Set Screw
1/2" R
FIGURE IV-3
Drop Tester for Acceptor Hardness - Dwg. DM-3.461
Project No. 555.02 - 12/12/73
A8.
-------
The chalk slurry feed system consisted of a continuously stirred tank
filled with water and CaC03 to give a ratio of 35 wt $ solids in the slurry.
A Lapp metering double ball check valve diaphragm pump provided slurry
transportation.
Inert bed material was fed into the desulfurizer (D~2) vessel through the
L-4 feeder. Both coating of the ball checks with chalk and segregation of
slurry in the check valves of the pump prevented satisfactory performance of
the apparatus.
A simplified flow diagram for a trial run using Fisher's precipitated
calcium carbonate in a dry form is shown in Figure IV-5.
Tablets prepared by pressing calcium carbonate with 10 wt % of water were
dried, crushed and screened to 20 x 48 mesh size. However, the feedstock was
so soft it did not survive shearing forces between the rotor and the body of
the L-4 feeder which ground lumps of chalk into powder. The pockets of the
rotor gradually become filled thus preventing a uniform feed rate.
A compromise run to test use of calcium carbonate for desulfurization was
made with Nebraska limestone. Since the L-4 rotary feeder could not reliably
feed powdery material, a series of tests was performed in order to find a size
consist as fine as possible so the rates of the calcination and desulfurization
reactions would be limited as little as possible. At the same time, the finely
divided feedstock had to be fed at a uniform rate. Run L-l was made with
65 x 150 mesh Nebraska limestone and, after polishing the rotor and body of the
feeder, 150 x 200 mesh stone was used for Run L-2.
A bed of dead-burned dolomite was established in the desulfurizer vessel
for Run L-l and, after the bed temperature leveled out, the CS2 and limestone
feed systems were turned on. Magnorite was the inert bed material in Run L-2.
D. Hot Removal of Particulates and Alkali Fume
Figure IV-6 shows the flow diagram for the runs to test the removal of
particulates and alkali fume at 7O4°C (l3OO°F). Fluidizing gas and solids were
fed to the bottom of the gas desulfurizer vessel. The gas containing alkali
and particulates passed through the bed and exited through a heated line to a
filter vessel contained inside a heated enclosure. The filter vessel and heated
enclosure, called a hot box, are detailed in Figures IV-7 and IV-8. The system
was set up so that the gas would be maintained at 704°C or above until is passed
through the filter.
The flange of the desulfurizer was a large source of heat loss. When steps
were taken to apply extra heat to the flange, the brass gasket on the outer
shell leaked. The brass was replaced with a stainless steel and asbestos com-
posite ("Flexitallic" brand) which gave improved performance. The flange was
ultimately operated at 621°C (115O°F) with a tolerably small leakage rate of
balance gas.
The filter vessel itself was constructed from a gas sampling cylinder which
was cut open so that a 20 micron opening stainless steel sintered porous plate
filter could be welded into place. All piping materials were 316 ss. Piping
inside the hot box consisted of schedule 80 or 160 pipe or Autoclave Engineers'
medium to high pressure tubing welded to 3000 psi fittings. The line from the
reactor to the hot box was schedule 160 ss pipe. All connections were welded
except for a Grayloc union at the gas desulfurizer outlet.
49.
-------
CO.
H.
CS,
ui <*a
O Oesulfurizer
Recycle Gas
^
••f
CaCO, <- '^
Slurry
in Water
\^
£ 927° C
(1700PF)
Pump
FIGURE IV-4
DIAffiAM FOR DESULFURIZATION BY CHALK SLURRY
Inert
Bed
Material
Filters
C
CO,
\/
CO,
Filter
Vessel
Cooler
Condensate
Receiver
T
Chiller
Condensate
Receiver
Recycle
Compressor
T
Exit
Vent
-------
FIGURE IV-5
FLOW DIAQtAM FOR DESULFURIZATION BY FIJffi LIMESTONE AND/OH CHALK
CO,
CS,
Gas
Desulfurlzer
Recycle Gas
927"C (1700°F)
N2 Lift L
Gas I I
Filters
CO,
C02
Filter
Vessel
Purge
Cooler
Condensate
Receiver
Chiller
| Recycle
Compressor
_
Venf
-------
871°C
16OO°F
n n
CSa H2 CO2 N2 Rec.
Desulf urizer
7O4°C
13OO°F
28 x 35 Dolomite
8c Ash
Regenerator
L-4
Dolomite
Discharge through
L-6
JNa lAi]
Cooler
Chiller
Separator
7O4°C
13OO°F
ir I N2 Purge
FIGURE IV-6
FLOW DIAGRAM FOR ASH AND
ALKALI FUME REMOVAL STUDY
-------
T
I 5/8"
16-5/8"
IN
A
i i
i '
NOTE: For Installation See
Dwg. AV-3586
.1/2" Sch. 80 Pipe
/ / '
A f
f t
' t
/
/ 4" O.D.
' ' /
/ :
^
/ ^_
/
/,
X
/ 3-17/32" I.D.
y
/^
i
/ :
/ i
:
*—
^
,
r
I
1
1
I .
No. 8 HS2250 Hoke
Sampling Cylinder
1/4" Wall
Weld to Penetrate
into Porous
Cut Cylinder to Install
Porous and Weld
Plain ends
Out
1/4" Thick Porous'tf_ x 3-17/32"
Design Conditions:
250 psig at 1300°F
Hydro Test at 1700°F
Material: All 316 s/s
FIGURE IV-7
Filter for Alkali Fume Removal System
Dwg. DV-3592
Project 555.03 5/5/75
53.
-------
PAGE NOT
AVAILABLE
DIGITALLY
-------
The hot portion of the system was extensively instrumented and fitted
with alarms so that the temperature could be closely monitored.
Due to concern about possible contamination of equipment by alkali gener-
ated during start up of a run, an entirely separate condensate system was
appended to the unit. When gas was switched to the hot filter, it exited
through a second set of 200°C filters and a second condensate system.
The procedure was to startup as if an ordinary run were to be made. The
circulation of acceptor was continued for several cycles to build up the level
of sulfur in the stone and to allow any volatiles to be purged from the dolo-
mite. To start feeding fines and alkali, the withdrawn acceptor was collected
in a surge container and 10OO grams were mixed with the additive material.
This was then placed into the feed hopper. Several minutes after that hopper
was started, the exit gas flow was switched to the filter and alternate con-
densate system.
Upon completion of the run, the lines were taken apart and flushed with
dilute nitric acid to remove any alkali from the walls. Fines were removed
from the filter by backflushing with air.
E. Chance Reaction
The liquid-phase Chance reaction,
CaS(g) + C02(g) + H20(l) - CaC03(s) + H2S(g)
was studied batchwise in a glass laboratory apparatus at atmospheric pressure.
The reactor was a 500 ml standard resin reaction flask. For runs at elevated
temperatures, heat was supplied by heating a mantle. The flask cover was pro-
vided with joints for a sampling device, a thermometer, a gas inlet, a stirrer
and a gas outlet through a reflux condenser.
In a typical run, the assembled apparatus was charged with a batch of 4OO
ml of distilled water and 25 grams of acceptor. The run was started after air
was purged from the apparatus by N2 and the temperature leveled off.
Hydrosulfidation was carried out with 100$ H2S until there were no visible
changes in the slurry color and texture. The reaction was always completed in
less than one-half hour. Pure C02 was used for carbonation, and the run con-
tinued until no H2S could be detected in the offgas. After switching to the
reactive gas mix, the course of the reaction was followed by analyzing the exit
gas composition and the sulfur content of the acceptor-water slurry. The run
was terminated by purging out C02 and cooling down the slurry. This was
followed by filtration and analysis of the dry cake for sulfur.
Exit gas was analyzed for C02 and H2S by a calibrated gas chromatograph.
Samples of gases were taken at 12-minute intervals. Slurry and solids samples
were analyzed for sulfide sulfur content by decomposition in an acidified
iodine solution followed by titration with thiosulfate . Product solids were
filtered at ambient temperature and dried in a vacuum oven overnight prior to
analysis.
56.
-------
F. Lab-Scale Batch Acceptor Cycling Reactor
>i'.
An assembly drawing of the reactor is shown in Figure IV-9. The unit was
arranged for pressure balanced operation. The furnaces and the one-inch
reactor were contained in a 3-1/2" ID autoclave shell. Heat was supplied by
two windings of 15 gauge Kanthal A-l wire on an Alundum core. The full gas
flow was maintained at all times. Synthesized fuel gas entered through the top,
and the regenerator feed entered through the bottom. The gas streams merged
between the reaction zones and exited through the annulus connected to the
bottom.
The acceptor sample was held in a ceramic basket suspended from a centered
piece of 1/4" tubing. Originally a platinum mesh basket was used, but it fell
apart under the severe operating conditions in our system. This was believed
to be due to the formation of platinum carbide and/or solutions of carbon in
platinum which form under reducing conditions and which are extremely brittle.
Cycling was effected by moving the central tubing up and down through a pack-
ing gland above the head. A thermocouple embedded in the acceptor sample
measured the temperature.
A flow diagram of the system is given in Figure IV-10. Gas is metered in
via two sets of rotameters. Flow to the bottom section can be switched to a
saturator to provide the required steam. However, in order to simplify the feed
system, the required H20 feed to the upper reaction zone was made via the water-
gas-shift reaction. Gas chromatograph and compensate measurements confirmed
that shift equilibrium was closely approached. , |
Temperature profiles were taken which showed that the desired temperatures
of 160O and 130O°F (871 and 704°C) could be maintained in the two separate
heating zones.
Since two separate reactions were involved, diffusional mixing (interpene-
tration) of the two feed gases was to be avoided. Interpenetration was studied
by substituting a 1/8" tube for the sample thermocouple and taking a small gas
flow to a thermal conductivity cell. Nitrogen was passed through the reference
cell and the sample gas through the second balanced cell. Nitrogen was fed
over the sample gas part, and 50$ H2-N2 through the opposing reaction zone.
The observed response was less than 1$ of the response due to the H2-N2 mixture
indicating that the sampled gas was nearly pure nitrogen.
The operating procedure was to place a weighed sample of acceptor in the
mesh basket, and then withdraw the "support" tube until the stop hit the-pack-
ing gland. The entire upper assembly was then lowered into the assembled auto-
clave and the upper flange bolted tight. In this position, the sample sat in
the "cold" part of the reactor above the autoclave flange. The unit was then
pressured up and brought to the programmed Powerstat settings. As a rule, the
system was kept hot at all times.
The run was started by immersing the sample into the upper, or gas desulfu-
rizer, zone for half-calcining and sulfidation at the programmed conditions of
time, temperature and gas composition. The sample was then lowered into the
regenerator zone to complete a cycle. The procedure was repeated until the
desired number of cycles was completed. The run was ended by withdrawing the
sample from the regenerator zone into the cool zone above the autoclave flange.
Throughout the entire run, the sample temperature was recorded continuously.
57.
-------
PAGE NOT
AVAILABLE
DIGITALLY
-------
The unit was depressured and the acceptor removed by taking out the entire
support rod assembly. The sample was removed from the basket and weighed to
determine the CaS content.
G. Reduction of CaS04 by H2S
The unit arrangement for reduction of CaS04 by H2S was the same as for
desulfurization-regeneration runs.
The only changes made were on the vapor condensing system of the unit to
deal with elemental sulfur vapor in addition to water vapor. The condensing
system using cooling water only did not perform satisfactorily due to plugging
of the heat exchanger gas side by solid sulfur.
The substitution of steam for cooling water in the first stage of conden-
sing system No. 1 as shown in Figure IV-ll was only partially successful. In
this arrangement high viscosity elemental sulfur condensed on the gas side of
the heat exchanger coil and prevented smooth operation of the condensing
system. i
A satisfactory solution to the problem was reached by arranging the equip-
ment as shown in Figure IV-12. Most of the elemental sulfur was trapped as
liquid in the high temperature condenser-receiver vessel. Only minute amounts
of elemental sulfur ended up in the cold parts of the condensing system.
59.
-------
<5*s FE.JS
QAb FE.E.S
(Tl
-t*J*
Sfc.TURM'OK.
\_E-Cat_tJp ;
T> ~ Tfc*P. iMQ!
TlC -Tt^>Ai? 1»JQ.
Tt-
TR- Te.MP.
Pg.t-pitE.fi. Rc-c. I^a.
Pt> -
REVISIONS
CONSOLIDATION COAL CO.
RESEARCH a DEVELOPMENT DIVISION
LIBRARY. PENNA.
Sil.Ao
fl. t *. LEACH CO.. P6M.. PA.
-------
FIGURE IV-11
SULFUR CONDENSING SYSTEM NO. 1
From
Gas Desulfurizer
Steam
Steam
Trap
Liquid
Sulfur Out
High Temperature
Condenser
Water and
Sulfur Out
Chilled Water
to Exit
and Recycle
Compressor
-------
From
Gas Desulfurizer
High Temperature
Condenser-Rece iver
Water
and Sulfur.
c^
NJ
f~
*fx~
^N
-r~>
SULFUR (
^
Liquid Sulfur
Out
FIGURE IV-12
SULFUR CONDENSING SYSTEM NO. 2
C.W.
Water Out
Chilled Water
Gas to Exit and
Recycle Compressor
Water Out
-------
V. MATERIALS: SOURCE. PREPARATION AND ASSAYS
A. Sulfur Acceptors
During operations of the continuous acceptor unit, the following natural
stones were used;
1. Tymochtee dolomite, C. E. Duff & Son, Inc., Huntsville, Ohio.
2. Canaan dolomite, Charles Pfizer, Inc., Adams, Mass.
3. Buchanan dolomite, James River Limestone Co., Buchanan, Va.
4. Nebraska limestone, Hopper Brothers Quarries, Weeping Water, Nebraska.
5. Rapid City limestone, Hills Materials Co., Rapid City, S. D.
Assay results for the stones, as used in the experimental program, are given in
Table V-l.
The acceptors were ground using a mill which was adjusted to maximize the
yield of the desired size fraction. The oversize fraction was reground and the
undersize material was set aside. Usually, the material was then wet screened
to remove all of the fines. This was necessary to make attrition data obtained
in the continuous unit more meaningful. The wet screened acceptor was then dried
under infrared lamps. Canaan dolomite was not wet screened since it produced
little dust when it was ground.
Table V-2 presents results of elemental analyses on calcined acceptors
which were ignited in air. Duplicate analyses on Tymochtee 9, 1O and 11 dolo-
mites are presented. It was thought that the difference in activity observed
between the hardened Tymochtee stones might be explained by some element of
composition. Reference to Table V-2 shows no obvious reason why this should be
so. The Tymochtee stone has impurities consisting mostly of silica and alumina
which is to be expected from clay minerals. It is concluded that the difference
in activity is not readily explained by the elemental composition of the stone.
The formation of inert "CaO" does seem to parallel the presence of silica
impurity in the stone. Tymochtee dolomite had more of this impurity than either
Canaan or Buchanan stone and also displayed higher concentrations of "CaO".
B. Acceptor Assays
The assay of fresh acceptors was discussed in detail earlier^1) A change
was made in the assay procedure for sulfided stone. Previously, the CaC03
content was determined gravimetrically and the CaS content was determined by
difference. In the current reporting period, the procedure was changed to allow
for the direct determination of the CaS content of the stone. It was found that
the total of the CaS and CaCO3 concentrations did not equal 10O$. A term was
therefore introduced to account for the missing calcium values, and the calcium
was assumed to be in the form of CaO since it had earlier been determined that
under the conditions of the lab assay, CaC03 can react with silica in the stone
to produce compounds such as CaO (Si02)x. Under run conditions of the sulfur
acceptor cycle, additional CaC03 is lost via reaction with impurities. This is
over and above what reacts in the laboratory assay of the fresh stone.
63.
-------
TABLE V-l
Stone
Gravimetric Factors
F,(»)
Hols Useful Ca/lOO grams
Raw Stone
Tvi
Batch 9
1.836
0.3894
0.2121
Qochtoe 00X001116
Batch 1O Batch 11
1.813 1.839
O.36O1 O.3924
O.1986 O.2134
Canaan Dolomite
Batch 3 Batch 4 .
1.894 1.9O2
O.4794 O.4778
O.2531 0.2512
Buchanan
Batch 2
1.881
O.4553
O.242O
Dolomite
Batch 3
1.873
O.4461
0.2382
Nebraska
Limestone
1.745
0.7299
0.4183
Rapid City
Limestone
1.751
0.7338
O.4191
0.482
O.451
O.485
O.576
O.571
0.55O
O.541
O.95O
0.952
Examples :
Let
Then,
Wj = weight on MgCO3 -CaC03 basis (CaC03 basis for limestone).
W2 = weight on MgO'CaC03 basis.
W3 = weight on MgO-CaO basis (CaO basis for limestone).
Ws = weight of sulfided stone.
y = fraction of useful Ca converted to CaS.
Ft = »!/W3
F2 = (FJ(F3)
W, = (1 + F3) W./F!
W3 = W2/(l + Fa)
Hols useful Ca
in fresh, fully calcined stone = (W3)(F2)/44.O1
in fresh half-calcined (MgO-CaCO3) stone = (»2)(F2)/(l + Fa)/44.01
Ws = W2 [l - (Fa)(y)(lOO.O9 - 72.14)/44.Ol/(l + F2)]
(1)
= the weight fraction of CO2 equivalent to useful CaCO3 in the raw, dry stone = (W2-W3)/W1.
-------
The procedure for determination of CaS in sulfided stone is to first
weigh the sample into a 1000 ml flask. About 0.5 to 2.0 grams are used depend-
ing on the sulfur content of the stone. To this are added 2OO ml of water and
50 ml of 0.1 normal iodine solution. The flask is stirred to wet the stone,
and then 150 ml of 1:10 diluted HC1 is added to displace sulfur from the stone
and to dissolve it completely. The flask is stirred for 30 minutes and then
titrated with sodium thiosulfate solution. The sulfur content is calculated
according to standard procedures as mols of sulfur or CaS/100 grams of sample.
The CaC03 and CaS04 contents of the sample are determined on a separate
sample via the gravimetric assay method described earlier,!1) and are put in
terms of mols/100 grams of stone. The final piece of data required is the weight
of impurities plus MgO/mol of calcium and this is available from the original
assay of the raw stone. Then for a 100 gram basis,
100 = mols CaC03 x 100.09 + mols CaS x 72.14 + mols "CaO" x 56.08
+ mols CaS04 x 136.14 + grf^ *nerts x (mols CaS + CaCO, + "CaO" + CaS04 ) .
* mol calcium J * '
In the above expression, the only unknown is the mols of "CaO" so that the
unknown is readily calculated. This was done for each assay.
C . Low-Temperature Ash
The low-temperature ash used for alkali fume removal studies was produced
from Illinois No. 6 coal (Consolidation Coal Company's Hillsboro Mine).
The procedures for the ash preparation from unwashed coal are shown in
Figure V-l .
The typical conditions for low-temperature char ashing at 538°C (lOOO°F)
are shown in Table V~3 below;
TABLE V-3
Ash from Illinois No. 6 Coal
Pressure: 7.8 atm. (100 psig)
Feedstock 14 x 65 mesh carbonized
Illinois No. 6 coal
Temperature, °C (°F) 538 (lOOO)
Input. SCFH
Air 6.5
N2 15
The changes in composition which occurred as the coal was processed are
shown in Table V-4. Data are given for both elemental analyses of the ash
and for the alkali and chlorine content of the unignited material. In addition,
a set of analyses performed. by the Illinois State Geological Survey are given.
Their determinations of sodium and potassium parallel those of Conoco, but the
ISGS number for chlorine in the raw coal was considerably higher. This
discrepancy did not extend to the analysis of the carbonized coal .
65.
-------
TABU V-2
Type of Stone
Component. j
NaaO
KaO
CaO
MgO
Fea03
TiOa
SiOa
AlaO,
SrO
SO,
Total Impurities,
Tymochtee 9
O.4S
O.32
57.19
35.OO
O.54
O.OO
O.O5
3.78
2.3O
O.OO
O.38
7.81
O.51
O.38
56.61
35.44
0.61
O.51
O.OO
4.42
1.43
0.11
7^95
Raw data were normalized to 1OO1
Stone was first ignited in air
Tvmochtee 10 Tvmochtee 11
0.62
O.36
55.39
34.92
O.74
O.16
O.O2
5.O4
2.O8
O.OO
O.67
9.69
O.51
O.35
55.58
35.21
O.S8
O.54
O.O1
5.O4
2. OS
—
O.13
9.21
O.4O
0.31
56.55
34.91
O.56
O.OO
O.O4
S.22
1.38
O.OO
O.63
8.54
O.58
O.48
55. 4O
34.25
O.7O
0.21
0.01
6.61
1.63
—
0.13
1O.35
fpnaan 4
O.O2
O.06
57.37
41.48
O.OO
O.06
O.O6
O.53
O.OO
O.43
1.15
O.O2
O.04
6O.69
36.03
0.49
O.08
O.OO
3.85
O.OO
O.56
5.O4
Nebraska
Limestone
O.25
O.15
96.1O
O.6O
O.3O
O.O5
O.14
1.85
O.25
O.34
3.90(0
Rapid City
Lime atone
0.02
O.O8
97.81
O.42
O.7O
O.O9
0.23
0.51
O.OO
O.O5
O.O9
2.19(0
(1) Including MgO
-------
FIGURE V-l
Preparation of Low-Temperature Ash
UNWASHED HILLSBORO MINE, ILLINOIS NO. 6 COAL
AIR DRYING
-1/2"
SCREENING
+ 1"
1" x 1/2"
HAND SEPARATION OF SLATE
106 Ibs
CARBONIZING 538°C (lOOO°F)
GRINDING
+ 1/8"
SCREENING
-1/8"
GRINDING
+14 Mesh
SCREENING
SAMPLE W-
-65 Mesh
-•-{ SAVED
14 x 65 Mesh
5,OOO grams
ASHING 538°C (lOOO°F)
982 GRAMS OF LOW-TEMPERATURE ASH
67.
-------
TABLE V-4
00
Coal Based Materials for Alkali Fume Removal Study
Whole Sample (dry basis).
Ash '
Cl
Na
K
Ash Analysis. (0 Wt . $
Na2O
K20
CaO
MgO
Fe203
Ti02
^2^5
Si02
A1203
S03
MnO
Dry Raw Coal Basis. Wt %
Cl
Na
K
Raw Hillsboro Coal
ISGS
Assay
Wt.#
13.63 — 12.8
O.16 O.O9 O.4O
O.21O
0.167
1.59 1.37
1.65 1.71
3.96
O.7O
31.86
1.O9
0.12
39.22
16.41
3.39
— —
0.16 0.09 0.40
O.161 0.139 O.21O
O.187 0.193 O.167
Carbonized Hillsboro Coal
19.49
—
1.4O
1.77
4.53
O.84
27.47
O.44
O.75
4O.46
18.11
4.23
— —
—
O.141
O . 2OO . .
20. 63
O.13
1.47
1.84
3.99
O.86
25.25
O.51
0.30
44.97
17.58
3.22
— —
O.O86
O.149
0.208
ISGS
Assay
18.2
O.ll
O.227
O.25O
0.077
0.160
O.176
Low Temperature Ash
Without After
Ignition Ignition
0.03
1.25
1.62
3.35
0.82
29.63
1.17
O.4O
40.72
17.48
3.34
0.05
0.004
0.131
0.190
96.60
0.01
1.O2
1.56
O.OO1
O.1O3
0.176
(i) Ignition carried out in air at 8OO°C.
-------
The analyses show a moderate loss of sodium from the coal upon low-
temperature ashing. It is interesting to note the rather large loss of
chlorine which exceeded 90$. It is clear that the low-temperature ash as we
made it did not duplicate the input of chlorine which would be obtained from
raw coal. To compensate, NaCl was added to the ash as described in Section
XI.
Data are given in Table V-5 for the alkali content of Tymochtee and
Canaan dolomite fines which were fed to the process. The Tymochtee stone
displayed far higher levels of alkali and chlorine than did the Canaan dolo-
mite. The data for Canaan dolomite scatter considerably. This is due to the
fact that the analysis used was developed for coal ash, and is set up to
detect sodium and potassium in ash on the order of 1% ± 0.1$. The weakness
of the solids alkali analyses precluded alkali material balances from being
made when the balance required knowledge of changes in the alkali content of
solids.
69.
-------
TABLE V-5
Analyses of Dolomite Feeds to Alkali Fume Removal Study
Tymochtee 11
Canaan 4
Whole Sample, Wt . %
Cl
Na
K
Analysis After Ignition, ( 1 ) Wt . $
Na30
K20
CaO
MgO
Fe203
Ti02
SiO2
AUO,
Raw
0.07
0.58
0.48
55.40
34.25
0.70
0.21
0.01
6.61
1.63
Ignited
0.40
0.31
56.55
34.91
0.56
0.00
O.O4
5.22
1.38
Raw
O.O7
0.19
O.19
Raw Raw Raw
0.01 O.O4
O.O1 O.O2 O.O4
0.02 0.03 0.04
After Lab
Assay(2)
0.01
0.02
Ignited
57
41
SO,
O.13
0.63
0.02
0.06
37
48
0.00
0.06
O.O6
O.53
O.OO
O.43
(i) Ignition carried out in air at 8OO°C.
(2) Maximum temperature = 760°C (14OO°F).
-------
VI. SAMPLE CALCULATIONS
The major thrust of the work performed under this contract was in the
performance of sulfidation-regeneration runs. A detailed description of the
calculations performed in estimating the number of cycles and the specific
run conditions and results was given earlier.!1) The assay calculation
procedure was changed, and this was described in Section V. Calculations
involved in evaluating the results of the lab-scale Chance runs and alkali fume
removal runs were straightforward material balances. Only the estimation of
the design make-up rate and specification of the conversion of the sulfur
sorbent, and the calculation of normalized density are given here.
A. Calculation of Required Make-Up Rate
First, the run data on regeneration activity as a function of cycles were
reviewed. The data were accumulated over an extended period, and in order to
maintain consistency, all the runs were worked up keying on the CaC03 content
of the stone. It was felt this would be more reliable than the data derived
from gas analyses as the H2S analyzer had undergone occasional calibration
difficulties. In addition, the early runs were made before the H2S analyzer
was delivered.
To obtain a mathematical expression for regeneration activity versus
cycles, a least squares analysis was run. The log-log relationship between
activity and cycles which results was used in the cycling calculations.
A computer program was written to carry out the cycling calculations.
The program is shown in Table VI-1. The starting point of the calculations is
an inventory of 100 mols of CaC03 with no CaS. The inventory is first fed
into the gas desulfurizer where it is sulfided at a predetermined level. The
degree of sulfidation is set at a constant number of mols or percentage of
total calcium. The relative percentage conversion of CaC03 to CaS therefore
increases as the CaC03 is depleted during cycling. It was assumed that all
the CaC03 in the gas desulfurizer is euqally reactive regardless of age. Thus,
if 13% is the predetermined conversion of total calcium and 25 mols of CaC03
enter the gas desulfurizer, the fraction 13/25 of the CaC03 in each particle
is reacted independent of its age and percent CaC03 content. For example,
fresh stone would exit at 52$ CaS and a stone entering at 90$ would exit at
95.2$ CaS.
The inventory then goes into the regenerator where each particle is
regenerated according to its age following the fitted relationship which had
been worked out based on experimental data. It is assumed that this relation
holds regardless of ,the specific CaS content. As an example, a stone particle
one cycle old might have 40$ of its CaS content regenerated to CaC03 while an
older stone might have only 8$ of its CaS regenerated. The regenerated in-
ventory is then reduced by the amount of the make-up rate, e.g., 2$, and
refilled with fresh stone. As the cycle by cycle calculations are continued,
the CaC03 of the inventory falls until an equilibrium level is reached.
71.
-------
CYCLER 20-DEC-74 08s53s51
TABLE VI-1
5 C = 13
10 DIM K<300>* J<300>* F<300) Computer Program for Cycling Acceptor
20 K(l > » 100
30 J( 1) = 0
40 N » 1
70 FOR I » 1 TO 200
80 F » .01*EXP(3.77516 - .431135*LOG
90 NEXT I
110 T = 0
115 FOR I « 1 TO N
120 T «T + (1 - JCI> > * K
130 NEXT I
140 FOR I = 1 TO N
150 JU > « J + C /T*(l - J(I) )
160 NEXT I
170 HEM S'JLFIDATION COMPLETED
180 HEM START REGENERATION
210 FOR I » 1 TO N
250 J = U-F> * JCI)
260 | KCI) <= .98 * K(I)
270 NEXT I '
280 REM END REG'N AND PURGE START MAKE'JP
310 N Q N * i
320 K(M) a 2
330 JCN) » 0
390 IF INT< (N-l)/200>*200 » M-l THEN 410
400 GO TO 110
410 PAUSE
411 PRINT "INPUT STEP NUMBER"!
412 INPUT 3
415 PRINT "AGE* CYCLES"; TAB(l5)i "MOLS"; TA8(S5>; "I CAS"; TARC35)!
"MOLS CAC03"
420 FOR I » 1 TO N STEP Q
430 PRINT DECJ DECCK( I ) * 1 7,2 > J DECC 1 00* J< I ) .28* 1 > J DEC«1-J
(I))*K(I)* 38*3)
440 NEXT I
445 PRINT "T « "JDEC(Tp8*a>
450 PRINT
460 PA'JSE
470 GO TO 110
600 END
READY
Where:
K(l) = Number of mols qf calcium aged I cycles.
j(l) = Fraction CaS in particles aged I cycles.
T = Total mols CaC03 entering the gas desulfurizer.
C = Percent conversion of total calcium to CaS in the gas desulfurizer.
N = Number of cycles processed so far.
F(l) = Fractional regeneration of CaS to CaCO3 in Cycle I.
72.
-------
Table VI-2 is the computer output for Canaan dolomite cycled 200 times
at 2$ makeup and 13$ sulfidation per pass. Although every cycle was calculated,
only every fifth cycle was printed out.
Each computer run gives an equilibrium CaC03 value for a given make-up rate,
sulfidation level, and regeneration cycles relation. Runs are made until a line
for the equilibrium CaC03 content versus percent conversion can be drawn. As
the sulfidation per cycle increases, the equilibrium CaC03 content decreases.
Using the data generated experimentally for regeneration activity versus
cycles, the equilibrium CaC03 content for Canaan dolomite, Tymochtee 9 dolomite
and Buchanan dolomite hardened at 35$ conversion were calculated. The results
are shown in Figures VI-1 to VI-3, respectively. However, another calculation
was necessary in order to set the design for the system.
It was determined that, in order to prevent breakthrough of H2S, some
excess CaC03 must be present in the gas desulfurizer reactor. The design con-
straint was that the ratio of mols H2S per hour fed:CaC03 in the bed should not
exceed 3.6. An additional constraint was that the solids residence time in the
gas desulfurizer was to be 20 minutes. Then the following calculation was made;
Let Q = the molar feed rate of total Ca/hr,
then the number of mols of total Ca in the bed are Q/3.
Let Y = the mols/hr of CaC03 leaving the bed,
then the fraction of CaC03 leaving the bed = the fraction of CaC03
in the bed = Y/Q.
and the mols of CaC03 in the bed = Y/Q x Q/3 = Y/3.
Let X = mols/hr of H2S fed ^ mols/hr H2S reacted = mols/hr CaC03
reacted.
Using the design constraint for the H2S-to-CaC03 ratio gives,
„ „ X 3 X X
3'6 = Y/3=T~ °r Y =T^
Now, CaC03 fed = CaC03 reacted + CaC03 leaving,
or CaC03 fed = X + Y = X + T*— = 1.83 X
1 • ^
Stated in words, the CaCO3 fed must be 1.83 times the CaC03 converted.
This is expressed graphically in Figures VI-1 to VI-3 as a line having a
slope of 1.83. The intersection of this line with the line defining the equili-
brium CaC03 content is a point which satisfies both the make-up rate requirement
and the requirement for excess CaC03 in the bed. Results of these calculations
are discussed in Section XVI.
73.
-------
TABLE VI-2
AGE*1 CrCLES
200.
195.
190.
185.
180.
175.
170.
165.
1 60 .
155.
150.
145.
140.
135.
130.
125.
120.
115.
110.
105.
100.
95.
90.
85.
80.
75.
70.
65.
60.
55.
50.
45.
40.
35.
30.
25.
20.
15.
10.
5.
0.
T = 15.52 MC
MOLS
1 .76
0.04
0.04
0.05
0.05
0.06
0.06
0.07
0.08
0.09
0.10
0.11
0.12
0.13
0.14
0.16
0.18
0.20
0.22
0.24
0.27
0.29
0.32
0.36
0.40
0.44
0.49
0.54
0.60
0.66
0.73
0.81
0.89
0.99
.09
.21
.34
.48
.63
.81
2.00
als CaCO_
% CAS
t1) 94.7
94.7
94.6
94.6
94.5
94.4
94.4
94.3
94.2
94.1
94.1
94.0
93.9
93.8
93.7
93.6
93.5
93.3
93.2
93.1
92.9
92.8
92o6
92.4
92.2
92.0
91 .8
91.5
91 .2
90.9
90.5
90.1
89.5
88.9
88.2
87.2
86.0
84.1
81 .2
74.7
0.0
Total
Computer Output for Canaan Dolomite
2$ Makeup Rate,
MOLS CAC03 13$ Sulf Idation,
0.092 200 Cycles
0.002
0.002
0.003
0.003
0.003
0.004
0.004
0.005
0.005
0.006
0.006
0.007
0.008
0.009
0.010
0.012
0.013
(5.015
0 .0 1 7
0.019
0.021
0.024
0.027
0.031
0.035
0.040
0.046
0.052
0.060
0.069
0.080
0.093
0.109
0.129
0.154
0.187
0.234
0.308
0.458
2.000
'1/ This figure is the residual of the starting inventory. It
represents 0.6$ of the total CaC03, and running the program out
to more cycles would not change the total CaC03 content of the
equilibrated inventory.
74.
-------
12
. r
• f •
~r
-f- - FIGURE - VI-l
liL
;;Equilibrium CaCO3 In Canaan Dolomite As A :
-rFuinct ion Of Conversion—And Makeup-Rate
—t :—
LL__-.-L.L 1.
-1: •:•••!•
-I.. .!.
bX -;-r-j— ~
lioJCaS
! |
• 4 t~ :
10
10 11
Conversion to CaS per cycle
75.
-------
-I—
liU-'i:
Do lorn-: te -As^A -Function- Of- Conversion- And Makeup U--r---
Eqv
FIGURE -V-I-2-i.
^J£-:|: . \.•::].- . \ :::..! - .' ::.:-:! .' :: . ^_
ilibrium CaCO3 In^Hardened Tymocfitee 9"
E\:i
: „.;; J__:_J_.:_J
10
8
10 11
Conversion to CaS per Cycle
-------
B. Calculation of Normalized Density
The normalized density is calculated from the observed density and composi-
tion of a stone sample. As the observed density is a function of the molecular
weight of the components as well as the structure of the stone, it is desirable
to normalize the density to a common basis. The basis chosen was to treat all
components as if they were CaO. The procedure is outlined below. Using a basis
of one mol, the molecular weight of the stone is calculated using its known
composition and content of impurities, MgO plus inerts. This is compared with
the weight of the same stone if it were to consist only of CaO and impurities..
The ratio of the CaO weight to the actual weight times the observed density is
the normalized density.
Basis: 1 gram mol of stone
Component
CaCO3
CaS
CaS04
"CaO"
MgO + Inerts
Run A32 Oxidized Feed
Mols Grams
0.698
0.031
0.193
0.078
Density = 1.99 gm/cc
105.08
151.75
69.86
2.24
26.28
4.37
49.00
151.75
Normalized Stone
Mo
56.08
49.OO
105.08
Normalized Density = 7cii"^g x 1-99 = 1.38 gm/cc.
77.
-------
24
22
20
18 -ETCt:
16
14 -
12
10 -
3rium-CaCQ3_Jn Hardened
Function'Of I Conversion And Makeup Rate
^Teraper&tuire :=; 704° C; (130O° F
'
1..83 .^c:eon.versdonL-to-CaS
10 11 12
Conversion to CaS per Cycle
-------
VII. TABULAR CHRONOLOGICAL HISTORY OF RUNS
The purpose, general run conditions, and conclusions drawn for
all runs made in the course of this contract period are given in the
tables listed below. All the runs were made at a pressure of
15 atmospheres.
TABLE
VII-1 Summary of Initial Runs with Canaan Dolomite
VII-2 Summary of Early Hardening Runs
VII-3 Summary of Runs with Sorbent at Various Hardening
Conditions
VI1-4 Summary of Chance Product Runs
VI1-5 Summary of Main Body of Continuous Runs
VII-6 Batch Studies of Process Variables
VI1-7 Summary of CaS04 Reduction by H2S
VI1-8 Low-Temperature Ashing of Coal
VII-9 Alkali Fume Removal
79.
-------
TABLE VI I-1
Summary of Initial Runs with Canaan Dolomite
Run Number
Unit Configuration
A-20
Sulf idat ion-Regenerat ion
(continuous cycling
setup)
A-20A
Suit Idat ion-Regenerat ion
(continuous cycling
setup)
A-21
Suit Idat Ion-Regenerat ion
(continuous cycling
setup)
A-22
Sulf idat ion-Regenerat ion
(continuous cycling
setup)
A-22A
Sulf idat ion-Regenerat ion
(continuous cycling
setup)
Date of Run
1/3-1/4, 1974
1/7-1/18, 1974
2/5-2/15, 1974
1/22-1/23, 1974
1/23-1/31, 1974
Run Duration with H,S Feed,
hrs.
5.O
117
145
4.6
89
Purpose of Run
To establish activity
decline with a lower
conversion per pass than
in a similar run, No. A-7
of 1972
To establish activity
decline with a lower
conversion per pass than
in a similar run, No. A-7
of 1972
To examine regeneration
behavior with a close
approach to equilibrium
in regenerator
To explore activity
decline at a high
regeneration temperature
To explore activity
decline at a high'
regeneration temperature
00
O
Temperature, °C (°F)
Bed Material
871 flSOO) Sulfidation,
7O4 (13OO) Regeneration
Hall-Calcined, sulfided
Canaan dolomite
871 fl6OOj Sulfidation,
7O4 (130O) Regeneration
Half-Calcined, sulf ided
Canaan dolomite
871 (16OO) Sulfidation,
7O4 (13OO) Regeneration
Half-Calcined, sulf ided
Canaan dolomite
871 (1600) Sulfidation,
76O (140O) Regeneration
Half-Calcined, sulfided
Canaan dolomite
871 (16OO) Sulfidation,
76O (14OO) Regeneration
Half-Calcined, sulfided
Canaan dolomite
Feedstock
35 x 48 mesh Canaan
dolomite
35 x 48 mesh Canaan
dolomite
35 x 48 mesh Canaan
dolomite
35 x 48 mesh Canaan
dolomite
35 x 48 mesh Canaan
dolomite
Shutdowns
Level controls not well
tuned in, plug downstream
of reactor
Termination upon
depletion of Inventory
Termination upon
depletion of inventory
Termination upon loss
of inventory by an
error
Termination forced by
massive plug of solids in
transfer line, gas leak
Results and Conclusions
No useful data,
repeated
run to be Activity is satisfactory
The rate-of ..regeneration
.is not affected by a
-close approach to
equilibrium in regenerator
No useful data, run to be
repeated
Activity and attrition
determined*up to •
termination' point
-------
TABLE VI1-2
Summary of Early Hardening Runs
Run Number
Unit Configuration
HR-1
HR-2
HR-3
HR-4
HR-5 through HR-1O
HR-11
Sulfidat ion-Regeneration Sulf idat ion-Regenerat ion Sulfidat ion-Regeneration Su If idat ion-Regeneration Sulf Idat ion-Regeneration Sull idat ion-Regeneration
(batchwise type of (batchwise type of (batchwise type of (batchwise type of (bate (wise type of (continuous cycling
operation in gas
desulfurizer only)
operation in gas
desulfurizer only)
operation in gas
desulfurizer only)
operation in gas
desulfurizer only)
operation)
setup for presulf idat ion,
batchwise type of
operation for oxidizing)
Date of Run
2/19-2/20, 1974
2/21-2/27, 1974
2/27-3/1, 1974
3/4-3/5, 1974
3/6-3/15, 1974
4/9-4/11, 1974
Run Duration with H2S Feed,
hrs.
4.5
4.5
4.5
4.5
Purpose of Run
To test hardening
procedure
To test hardening
procedure
To test hardening
procedure
To test hardening
procedure
To prepare feedstock
for Run No. A-23
Prepare large amount of
presulfided stone for
consequent oxidizing at
different temperatures
for physical examination
Temperature, "C (°F)
00
871 (16OO) Presulfiding
927 (1700) Oxidizing
871 (16OO) Presulfiding
871 (1600) Oxidizing
871 fl6OO) Presulfiding 871 (16OO) Presulfiding
927 (17OO) Oxidizing 982 ( 18OO) Oxidizing
871 (16OO) Reducing
871 (16OO) Presulfiding
982 (18OO) Oxidizing
871 (1600) Reducing
Presulfiding at 871
(16OO), batchwise oxidiz-
ing at 871 (1600), 899
(1650), 927 (1700), 954
(1750), 982 (1800),
respectively
Bed Material
Ha If-Calcined Tymochtee
dolomite
Half-Calcined Tymochtee
dolomite
Half-Calcined Tymochtee
dolomite
Half-Calcined Tymochtee
dolomite
Half-Calcined Tymochtee
dolomite
Half-Calcined, presulfided
Tymochtee dolomite
Feedstock
35 x 48 mesh Tymochtee
dolomite, shipment
No. 9
35 x 48 mesh Tymochtee
dolomite, shipment
No. 9
35 x 48 mesh Tymochtee
dolomite, shipment
No. 9
35 x 48 mesh Tymochtee
dolomite, shipment
No. 9
35 x 48 mesh Tymochtee
dolomite, shipment
No. 9
35 x 48 mesh Tymochtee
dolomite, shipment
No. 1O
Shutdowns
Voluntarily
Voluntarily
Voluntarily
Voluntarily
Voluntarily
Voluntarily
Results and Conclusions
Temperature ran away,
decrease oxygen partial
pressure
Low attrition
resistance, agglomer-
ation, hot spots
Low attrition
resistance, repeat run
at a higher temperature
in the oxidizing step,
agglomerat ion
Partially agglomerated.
In the next run keep
bed fluidized at all
times
Agglomerates in product,
high temperature in the
oxidizing step seems to
promote agglomeration
Agglomeration at high
temperatures
-------
TABLE VII-3
.ry of Runs with Sorbent from Various Hardening Conditions
Run Number
Unit Configuration
A-23
Su If idat ion-Regene rat ion
(continuous cycling
setup)
A-24
SuIf Ida t ion-Regene rat ion
(continuous cycling
setup)
Sulf idat ion-Regeneration
(continuous cycling
setup, preceded by
hardening step)
A-26
Sulf idat ion-Regenerat ion
(continuous cycling
setup preceded by
hardening step)
A-27
Sulf idat ion-Regenerat ion
(continuous cycling-
setup preceded by
hardening step)
Date of Run
3/22-4/9, 1974
4/15-4/23, 1974
4/24-5/3, 1974
5/6-5/14, 1974
5/14-5/22, 1974
Run Duration with H,S Feed,
hrs.
31
Purpose of Run
To determine properties
of hardened stone
To determine attrition
resistance and activity
of stone hardened at
severe conditions
To determine attrition
resistance of stone
hardened at severe
condit ions
To determine attrition
resistance of stone
hardened at milder
conditions than previous
run
To determine attrition
resistance of hardened
stone presulfided to a
high degree at mild
temperature
Temperature, °C (°F)
CO
871 (1600) Sulfidation,
7O4 (13OO) Regeneration
Presulfiding at 871
(16OO), oxidizing at
982 (18OO), sulfidation
at 871 (1600),
regeneration at 7O4
(1300)
Presulfiding at 871
(1600), oxidizing at
927 (170O), sulfidation
at 871 (16OO),
regeneration at 7O4
(1300)
Presulfiding at 871
(16OO), oxidizing at
899 (1650), sulfidation
at 871 (1600),
regeneration at 7O4
(1300)
Presulfiding at 871
(16OO), oxidizing at
871 (1600), sulfidation
at 871 (1600),
regeneration at 7O4
(1300)
Bed Material
Half-Calcined, sulfided
Tymochtee dolomite,
hardened
Half-Calcined, sulfided
Tymochtee dolomite,
hardened
Ha If-Calcined, sulfided
Tymochtee dolomite,
hardened
Half-Calcined, sulfided
Tymochtee dolomite,
hardened
Half-Calcined, sulfided
Tymochtee dolomite,
hardened
Feedstock
Combined product from
Runs HH-5 through HR-1O
35 x 48 mesh Tymochtee
dolomite, shipment
No. 10
35 x 48 mesh Tymochtee
dolomite, shipment
No. 10
35 x 48 mesh Tymochtee
dolomite, shipment
No. 1O
35 x 48 mesh Tymochtee
dolomite, shipment
No. 1O
Shutdowns
Termination upon H,S
breakthrough
Depletion of inventory,
HaS breakthrough
Voluntarily
Termination upon H2S
breakthrough
Termination upon H2S
brea kt hrough
Results and Conclusions
Hardening step overdone
Attrition resistance,
activity satisfactory
CSa pump malfunctioned
giving low output.
Attrition is satisfactory.
Make CS2 feed system more
reliable.
Attrition is satisfactory
Attrition is satisfactory
-------
TABLE VII-3 (Cont'd. )
Summary of Runs with Sorbent from Various Hardening Conditions
Run Number
Unit Configuration
A-28
Sulf idat ion-Regenerat ion
(continuous cycling
setup, preceded by
hardening step)
A29
Sulf idat ion-Regenerat ion
(continuous cycling
setup preceded by
hardening step)
A-30
Sulf idat ion-Regenerat Ion
(continuous cycling
setup preceded by
hardening step without
use of sulfur)
A-30A
Sulf idat ion-Regenerat ioi
(continuous cycling
setup)
Date of Run
5/22-5 /24, 1974
5/28-5/30, 1974
5/31-6/4, 1974
6/4-6/5, 1974
Run Duration with H2S Feed, hrs.
35
25
38
13.2
Purpose of Run
To determine attrition
resistance of hardened
stone presulfided to a
medium degree at mild
temperature
To determine attrition
rate of stones hardened
at mild conditions
To determine, whether
sulfur involvement in the
hardening step is
necessary
To check the attrition
rate of untreated
Tymochtee dolomite
00
U>
Temperature, °C (°F)
Presulf iding at 871 (16CO),
oxidizing at 871 (16OO),
sulfidation at 871 (16OO),
regeneration at 871 (130O)
871 (16OO) Sulfidation,
704 (1300) Regeneration
871 fl6OO) Sulfidation,
7O4 (13OO) Regeneration
871 (16OO> SuLfidation,
7O4 (13OO) Regeneration
Bed Material
Ha If-Calcined, sulfided
Tymochtee dolomite
hardened
Half-Calcined, sulfided
Tymochtee dolomite,
shipment No. 1O, hardened
at mild conditions
Ha If-Calcined, sulfided
Tymochtee dolomite,
hardened without sulfur
Ha If-Calcined, sulfided
Tymochtee dolomite beds
from Run No. A-3O
Feedstock
35 x 48 mesh Tymochtee
dolomite, shipment No. 1O
35 x 48 mesh Tymochtee
dolomite, shipment No. 1O
35 x 48 mesh Tymochtee
dolomite, shipment No. 1O
35 x 48 mesh Tymochtee
dolomite, shipment No.
Shutdowns
Voluntarily
Voluntarily
Voluntary switch to
Run No. A-3OA
Massive plug in solids
transfer line
Results and Conclusions
Lowest hardening step
temperature, attrition
satisfactory
The extent to which stone
has been presulfided
gives material with low
attrition resistance
Sulfur plays its role in
hardening step
Attrition rate lower tb
expected
-------
TABLE VII- 4
Run Number
Unit.Configuration
Date of Run
Run Duration with H2S Feed,
hrs.
Purpose of Run =
Summary of Chance Product Runs
L-l L-2
00
-p-
Temperature, °C (°F)
Bed Material
Feedstock
Shutdowns
Results and Conclusions
Setup for chance product
feed experiment
6/12-6/14, 1974
6.5
To determine, whether
fine limestone can act
as sulfur acceptor in
calcined form
927 (170O) calcination
and sulfidation
28 x 35 mesh dead
burned dolomite
65 x 15O mesh Nebraska
limestone
Voluntarily
Sulfidation does not
proceed to completion,
decrease limestone
size, dead burned
dolomite is active
toward H2S
Setup for chance product
feed experiment
6/18, 1974
3.5
To determine, whether
fine limestone can act
as sulfur acceptor in
calcined form
927 (1700) calcination
and sulfidation
35 x 48 mesh Magnorite
(electrofused, ground
MgO)
15O x 2OO mesh
Nebraska limestone
Voluntarily
Sulfidation did not-
proceed to completion
A-82
Feed system and collecting
system for use of CaCO3
slurry
2/1O-2/12, 1975
H2S and Ca(NO3)2 solution
fed for 48 minutes
Establish feasibility of
using Ca(NO3)2 solution in
water to disperse most of
calcium salt into hot bed
of Magnorite for formation
of fine CaO which would act
as a sulfur acceptor
927 (17OO)
48 x 65 mesh Magnorite
Saturated solution at 7O°F
in water
Involuntarily, slugging and
plugging in the hot section
of feed line
Ca(NO3)2 melts as water
evaporates, and then
agglomerates
-------
TABLE VII-5
Summary of Bain Body of Continuous Runs
Run Nunber
Unit Configuration
A-31
Sulfidat ion-Regenerat ion
(continuous cycling
setup)
A-32
Sulfidat Ion-Regeneration
(continuous cycling
setup)
A-33
Suit idat ion-Regenera t ion
(continuous cycling
setup, the first cycle
preceded by hardening
step)
A-35
Sulf idat ion-Regeneration
(continuous cycling
setup, the first cycle
preceded by hardening
step)
A-36
Suit Idat ion-Regenermtlon
(continuous cycling
setup)
Date of Run
6/21-6/28, 1974
7/15-7/18, 1974
7/19-7/25, 1974
7/29-8/1, 1975
8/1-8/5, 1974
Run Duration with HaS Feed,
hrs.
28.9
34
37
28.3
Purpose of Run
Establish activity
decline, attrition data
and size effect on
performance of Canaan
dolomite
Establish activity
decline and attrition
data for hardened
Tymochtee dolomite
To determine the extent
of sulf idat Ion and
oxidation on the
attrition resistance
and activity
To determine the effect
of particle size of
hardened stone on the
performance
To determine the effect
of particle size on the
stone performance
Temperature, °C (°F)
871 (16OO) Sulfidation,
7O4 ( 13OO) Regenerat ion
871 (16OO) Sulfidation,
7O4 (13OO) Regeneration
871 (16OO) Sulfidation,
7O4 (13OO) Regeneration
871 (16OO) Sulfidation,
7O4 (13OO) Regeneration
871 (16OO) SuLfidatlon,
7O4 (13OO) Regeneration
Bed Material
00
Ul
Half-Calcined, sulflded
Canaan dolomite
Half-Calcined, sulfided
Tymochtee dolomite,
hardened
Half-Calcined, sulfided
Buchanan dolomite,
hardened
Half-Calcined, sulfided
Tymochtee dolomite,
hardened
Half-Calcined, sulilded
Canaan dolomite
Feedstock
2O x 28 mesh Canaan
dolomite
2O x 28 mesh Tymochtee
dolomite, shipment No. 9
2O x 28 mesh Buchanan
dolomite
28 x 35 mesh Tymochtee
dolomite, shipment No. 9
2O x 28 mesh Canaan
dolomite
Shutdowns
Numerous Interruptions
due to the leaks, solids
transfer lines plugging,
unstable control of
pressure balance.
Improper valves for
solids handling
Voluntarily
Shutdown due to
depletion of Inventory
Voluntarily
Numerous breakdowns of
equipment finally farced
termination (leaks)
Results and Conclusions
No useful data has been
obtained, repeat the run
Desulfur izat ion
performance as expected,
regeneration activity
below normal
Low attrition
resistance
The attrition rate Is
the same as for 35 x 48
mesh Tymochtee dolomite
Data of a low value,
repeat of run
-------
Run Number
Unit Configuration
A-36A
Sulfidation-Regenerat ion
(continuous cycling
setup)
TABLE VII-5 (Cont'd. )
Summary of Main Body of Continuous Runs
A-38
Sulf idation-Regenerat ion Sulf idation-Regenerat ion
(continuous cycling (continuous cycling
setup) setup)
A-39
Sulf idat ion-Regeneration
(continuous cycling
setup)
A-4O
Sulf idat ion-Regenerat ion
(continuous cycling
setup)
Date of Run
8/6-8/13, 1975
8/15-8/19, 1974
8/19-8/23, 1974
8/26-9/3, 1974
9/6, 1974
'Run Duration with H,S Feed,
hrs.
71
33
4O
1.7
Purpose of Run
To determine the affect
of particle size on the
stone performance
Establish whether
different shipments of
Tymochtee dolomite
perform the same way;
to find behavior of
different particle size
on regeneration activity
Establish whether
different batches of
Tymochtee dolomite
perform^the-same way;
extend the last
regeneration cycle
To determine whether
partially recarbonated
limestone can act as a
suitable sulfur acceptor
To determine whether
partially dead burned
limestone can act as a
suitable sulfur acceptor
Temperature, °C (°F)
871 (16OO} Sulfidation,
7O4 (13OO) Regeneration
00
871 flSOCO Sulfidation,
7O4 (13OO) Regeneration
871 f!6OO> Sulfidation,
7O4 (13OO) Regeneration
Calcining at 927 ( 17OO) ,
Recarbonation at 7O4
(13OO), Sulfidation at
871 (16OO), Regeneration
at 7O4 (13OO)
Calcining at 927 (17OO)
in dry atmosphere,
Recarbonation at 76O
(140O), Sulfidation at
871 (16OO), Regeneration
at 7O4 (13OO)
Bed Material
Half-Calcined, sulfided
Canaan dolomite
Half-Calcined, sulfided
Tymochtee dolomite,
hardened
Half-Calcined, sulfided
Tymochtee dolomite,
hardened
Calcined and partially
recarbonated Rapid City
limestone
Calcined and partially
recarbonated Rapid City
limestone
Feedstock
2O x 28 mesh Canaan
dolomite
35 x 48 mesh Tymochtee
dolomite, shipment No. 9
2O x 28 mesh Tymochtee
dolomite, shipment No. 11
20 x 28 mesh Rapid City
1imestone
35 x 48 mesh Rapid City
limestone
Shutdowns
Voluntarily
Voluntarily
Voluntarily
Defluidization in the
first Sulfidation cycle
Numerous interruptions
caused by agglomerates.
After admission of H2
into the calcined-
recarbonated limestone,
defluidization occurred
Results and Conclusions
Deactlvatlon curve has
been established
Regeneration activity is
the same within shipment,
the lowest in all
shipments
Activity of Tymochtee
No. 11 is Intermediate
between batches No. 9
and 10, the particle size
does not Influence
regeneration activity
Stone becomes sticky
At 871°C (16OO°F) with
steam in the gas
desulfurizer stone
becomes sticky
-------
TABLE VII-5 (Conf d. )
Summary of Bain Body of Continuous Rims
Run Number
Unit Configuration
A-41
Su If idat ion-Regeneration
(continuous cycling
presulfldlng in
calcined form)
A-42
Sulf idat ion-Regenerat ion
(continuous cycling, the
first cycle preceded by
hardening step)
A-43
Sulf idatIon-Regenerat ion
(continuous cycling, the
first cycle preceded by
hardening step)
A-44
Suit idat ion-Regenerat ion
(hardening step only)
A-44A
Sulf idat ion-Regenerat Ion
(continuous cycling, the
first cycle preceded by
hardening)
Date of Run
9/9-9/12, 1974
9/16-9/19, 1974
9/19-9/25, 1974
9/25-9/27, 1974
9/3O-10/4, 1074
Run Duration with HtS Feed,
hrs.
15
32
53
9 hrs in hardening step
only
43
Purpose of Run
To establish whether an
attrition resistant,
non-agglomerating sulfur
acceptor can be prepared
from limestone
To determine variations
in activity of different
batches of Tymochtee
dolomite
To determine properties
of hardened Buchanan
dolomite
To determine whether
Buchanan dolomite can
be hardened
To determine whether
Buchanan dolomite can
be hardened at a higher
presulf id ing conversion
Temperature, 'C (°F)
00
Presulfldlng at 927
(170O), Sulfidation at
871 (16OO), Regeneration
at 704 (1300)
871 flGOOj SulfIdation,
7O4 (13OO) Regeneration
871 (16OOJ Sulfidation,
704 (13OO) Regeneration
871 (16OO) Presulf idat ion,
871 (16OO) Oxidation
871 (16OO) Sulfidation,
7O4 (13OO) Regeneration
Bed Material
Sulflded Rapid City
limestone
Sulflded, half-calcined
Tymochtee dolomite
(hardened)
Sulfided, half-calcined
Buchanan dolomite
(hardened)
Half-Calcined Buchanan
dolomite
Sulflded, half-calcined
Buchanan dolomite
(hardened)
Feedstock
35 x 48 mesh Rapid City
limestone
35 x 48 Tymochtee
dolomite, batch No. 11
35 x 48 mesh Buchanan
dolomite
35 x 48 mesh Buchanan
dolomite
35 x 48 mash Buchanan
dolomite
Shutdowns
Voluntarily
Voluntarily
Voluntarily
Voluntarily
Voluntarily
Results and Conclusions
Sulfur pick-up Is
satisfactory, sulfur
rejection is poor
Tymochtee dolomite
displays variations in
activity .
Buchanan dolomite
displays good activity
at lower level of pre-
sulfldlng, however, the
attrition rate is high
By an error, preeulflded
stone reached low sulfur
content
Hardened Buchanan
dolomite performs
similarly to Tywochtee
dolomite.
-------
TABU VII-5 fCont'd. )
Summary of HaIn Body of Continuous Runs
Run Number
Unit Configuration
A-45
SulfIdatIon-RegeneratIon
(continuous cycling, the
first cycle preceded by
bardenIng)
A-46
SulfIdat ion-RegeneratIon
(continuous cycling, the
first cycle preceded by
hardening)
A-47
SulfIdat ion-RegeneratIon
(continuous cycling, the
first cycle preceded by
hardening)
A-48
SulfIdatIon-RegeneratIon
(continuous cycling, the
first sulf Idat ion
batchvlse)
A-49
SulfIdatIon-RegeneratIon
(continuous cycling, the
first sulf Idat ion
batchvise)
tete of Run
1O/4-1O/1O, 1974
10/10-10/25, 1974
1O/28-11/18, 1974
11/19-12/5, 1974
12/11-12/18, 1974
Run Duration with H,S Feed,
hrs.
51
157
233
96
30.1
Purpose of Run
To'determine whether, by
lowering the temperature
to 816* C (15OO*F) In the
first phase of hardening,
sulfur will distribute
more evenly on a
microscopic scale. Thus
activity can be enhanced
To demonstrate behavior
of hardened Tynochtee
dolomite In a large
number of cycles
To demonstrate
possibility of use of
Buchanan dolomite
To simulate the
conversion history of
make-up stone In an
industrial scale process
To establish possibility
of limestone use
Temperature, °C (*F)
00
00
871 (16OO) SuLf idat ion,
7O4 (13OO) Regeneration
871 (16OO) Sulf Idat ion,
704 (1300) Regeneration
871 f!6OO) Sulfidation,
7O4 (13OO) Regeneration
871 fl6OO) Sulf idst ion,
7O4 (13OO) Regeneration
871 (1700) Calcination
and Sulf idat lxm,
7O4 (13OO) Regeneration
Bed (Soterlal
SuLflded, half-calcined
Tyoochtee dolomite
(hardened)
Sulflded, half-calcined
Tymochtee dolomite.
( hardened)
Sulf ided, half-calcined
Buchanan .dolomite
( hardened) '
Highly sulf ided, half-
calcined Canaan dolomite
Highly sulf Ided
limestone in calcined
form
Feedstock
35 x 48 mesh Tymochtee
dolomite
35 x 48 mesh Tymochtee
dolomite
35 x 48 mesh
dolomite
35 z 48 mesh Canaan
dolomite
35 x 48 mesh Rapid City
limestone
Shutdowns
Voluntarily
Voluntarily
Voluntarily
Interruption due to snoa
storm, Involuntary
termination due to failure
of electric heater in
desulfurlzer
Oefluldlzatlon of bed In
gas desulfurlzer,
numerous plugging of
solid transfer lines
Results and Conclusions
No change In activity,
attrition rate higher
than la a run at
standard conditions
Behavior in a large
number of cycles as
predicted has been
demonstrated
Feasibility of use of
Buchanan dolomite has
been established
Sufficient data obtained,
stone is less active
than starting at loo
level of sulfIdatIon
The use of Rapid City
lines tone In calcined
form Is not feasible
In the apparatus
-------
TABLE VII-5 (Cont'd.)
Summary of Bain Body of Continuous Runs
Run Number
Unit Configuration
A-5O
SuIf Ida t ion-Regenera t ion
(continuous cycling)
A-51
Sulf idat ion-Regenerat ion
(continuous cycling)
A-52
Sulf idat ion-Regenerat ion
(continuous cycling)
A-52A
Sulf Idat ion-Regenerat ion
(continuous cycling)
Date of Run
1/2-1/3, 1975
1/3-1/8, 1975
1/8-1/9, 1975
1/10-1/22, 1975
Run Duration with HaS Feed,
hrs.
16.6
44
19.0
129
Purpose of Run
To equilibrate CaS
content in circulating
sulfur acceptor with
make-up of fresh stone
and to compare calculated
CaS build-up with actual
performance
To equilibrate CaS
content in circulating
sulfur acceptor with
make-up of fresh stone
and to compare calculated
CaS build-up with actual
performance
To equilibrate CaS
content in circulating
sulfur acceptor with
make-up of fresh stone
and to compare calculated
CaS build-up with actual
performance
To equilibrate CaS
content in circulating
sulfur acceptor with
make-up of fresh stone
and to compare calculated
CaS build-up with actual
performance
Temperature, °C (°F)
871 (16OO) Desulfurization,
704 (13OO) Regeneration
871 (16COJ Desulfurization,
7O4 (13OO) Regeneration
871 (16OO) Desulfurization,
7O4 (13OO) Regeneration
871 (16OO) Desulfurizatio
704 (13OO) Regeneration
Bed Material
CO
VD
Half-Calcined, sulfided
Canaan dolomite
Half-Calcined, sulfided
Canaan dolomite
Ha If-Calcined, sulfided
Canaan dolomite
Ha If-Calcined, sulfided
Canaan dolomite
Feedstock
35 x 48 mesh Canaan
dolomite for starting
charge and makeup
35 x 48 mesh Canaan
dolomite, both the
starting charge and
makeup
35 x 48 mesh Canaan
dolomite, both the
starting charge and
makeup
35 x 48 mesh Canaan
dolomite, both the
original charge and
makeup
Shutdowns
Solids feeder breakdown,
operator's error
At steady conditions the
run showed a
sulfidation per pass
that was lower than
intended. Terminated
after 14 cycles
Due to operator's error
in dosage of make up
stone, the run was
terminated
Voluntarily
Results and Conclusions
Low value data,
to be repeated
run has
Data of a limited value,
run has to be repeated
No useful data,
to be repeated
run has
CaS build-up In Inventory
has been established,
validity of toe
computation of the CaS
level confined
-------
TABLE VII-6
Batch Studies of Process Variables
Run Nimber
Unit Configuration
A-53 - A-61
Sulf Idation
A-62 - A-67
Sulf idat Ion
A-68 - A-73
Regenerat Ion
A-74 - Ar79
Regeneration
A-8O
Regenerat ion
Date of Run
Run Duration with H,S Feed,
hrs.
1/24-1/28, 1975
0.82-7.O
1/29, 1975
0.22-1. 48
1/29-1/31, 1975
Regeneration for 4.7-8.4
hours
2/3-2/5, 1975
3.7-10
2/5, 1975
Regeneration for 8
hours
Purpose of Run(s)
To determine amount of
CaCO, required to
prevent breakthrough at
3 BtS inlet concen-
trations and at 3 bed
levels
To determine amount of
CaCOj required to
prevent H,S breakthrough
at 3 HSS inlet concen-
trations and at 2 bed
heights with cycled
dolomite
Establish course of
regeneration of fully . .
sulfided sulfur acceptor,
using 2 bed heights and
3 temperatures
Establish course of
regeneration of spent
sulfur acceptor at
2 bed heights and 3
temperatures
Establish degree of
regeneration at 816°C
(1500°F)
Temperature, °C (°F)
Bed Material
871 (1600)
Half-Calcined Canaan
doloalte
871 (16OO)
Regenerated cycled
Canaan dolomite
649 (1200)-760 (14OO)
Freshly sulfided Canaan
dolomite (l sulf idat ion
cycle)
649 (1200)-760 (14OO)
Sulfided cycled dolomite
816 (1500)
Sulfided deactivated
dolomite
Feedstock
28 z 35 mesh Canaan
dolomite
Regenerator bed of Runs
A-2OA, A-21, A-36A
28 x 35 mesh (nominal)
gas desulfurizer beds of
Runs A-53 - A-55 and
A-59 - A-61
Gas desulfurizer beds
of Runs A-62 - A-67
28 x 35 mesh (nominal)
spent dolomite from
Runs No. A-75 and A-78
Shutdowns
Results and Conclusions
Voluntarily
Data points have been
established
Voluntarily
Data points have been
established
Voluntarily
Data points have be
established
Voluntarily
Data points have been
established
Voluntarily
Data point has been
established
-------
TABLE VII-6 (Cont'd.)
Batch Studies of Process Variables
Run Number
Unit Configuration
A-81
Regenerat ion
A-85 A-86
Sulf idat ion-Regenerat ion Sulf idat ion-Regenerat ion
A-87 - A-95
Sulf idat ion
A-96
SulfIdation
Date of Run
Run Duration with H,S Feed,
hrs.
2/5-2/6, 1975
Regeneration for 8
hours
3/11, 1975
O. 45
3/11, 1975
2O
3/17-3/2O, 1975
O.62-1O. 3
3/20-3/21, 1975
18.7
Purpose of Run(s)
Establish degree of
regeneration at 871°C
(16OO°F)
Establish reactivity of
CaCOj in acceptor
regenerated at high
temperature
Investigate reactivity
of sorbent at a higher
regeneration temperature
Establish reactivity
of CaCOa in half-
calcined Canaan
dolomite at 3 bed
levels and at 3 H9S
inlet concentrations
with 35 x 48 mesh stone
Take samples of sulfided
dolomite acceptor during
sulf idat Ion and see a
microscopic scale growth
of crystals with tine
Temperature, °C (°F)
871 (160O)
871 (16OO) Sulfidation,
7O4 (13OO) Regeneration
871 (16OO) Sulfidation,
76O (14OO) Regeneration
871 (16OO)
871 (16OO)
I-1 Bed Material
Sulfided deactivated
dolomite
35 x 48 mesh highly
regenerated Canaan
dolomite
35 x 48 mesh half-
calcined, sulf Ided
Canaan dolomite
35 x 48 mesh half-
calcined Canaan
dolomite
35 x 48 mesh folly
sulf ided Tycochtee
dolomite under carboBating
condit ions
Feedstock
28 x 35 mesh (nominal)
spent dolomite from
Runs No. A-75 and A-78
Combined product from
Runs No. A-8O and A-81
Product from Run No.
A-22A
35 x 48 mesh Canaan
dolomite
35 x 48 mesh Tynochtee
dolomite
Shutdowns
Voluntarily
Voluntarily
Voluntarily
Voluntarily
One interruption due to
power failure
Results and Conclusions
Data point has been
established
Very poor reactivity
toward HaS
CaCOj made by
regeneration at higher
temperature may be
less reactive than CaCO,
made at lover
temperatures
Data points have been
established
To be determined
-------
TABLE VII-6 (Cont'd.1
Batch Studies of Process Variables
Run Number
Unit Configuration
Date of Run
Run Duration with HaS Feed,
bra.
A-97
Regeneration
3/21-3/24, 1975
2.57, regeneration for
16 hours
A-98 - A-1O6
Sulfidation
3/24-3/27, 1975
0.72-12.6
A-1O7
Sulfidation
4/1, 1975
18.1
A-108 - A-116
Sulfidation
4/2-4/4, 1975
0.78-13.9
A-117 - A-119
Sulfidation
4/7, 1975
0.93-2.53
Purpose of Run(s)
Take samples of sulfided
dolomite acceptor during
regeneration operation
and see a microscopic
scale growth of crystals
with time
Establish reactivity
of CaCOj in half-
calcined Canaan dolomite
at 3 high bed levels and
at 3 HaS inlet concen-
trations at 899° C
Take samples of fully
sulfided dolomite
acceptor and see a
microscopic scale growth
of crystals with time
Establish reactivity of
CaCOj in half-calcined
Canaan dolomite at 3
bed levels and at 3 HaS
inlet concentrations at
927° C
Establish reactivity of
half-calcined Tymocbtee
dolomite at 3 bed levels
Temperature, °C (°F)
871 (1600) SuLfldation, 899 (165O)
704 (1300) Regeneration
899 (1650)
927 (1700)
927 (17OO)
Bed Material
VO
N>
35 x 48 mesh sulfided
Tymochtee dolomite
35 x 48 mesh half-
calcined Canaan dolomite
35 x 48 mesh fully
sulfided Tymochtee
dolomite under
carbonatlon conditions
35 x 48 mesh half-
calcined Canaan dolomite
35 x 48 mesh half-
calcined Tymochtee
dolomite
Feedstock
35 x 48 mesh Tymochtee
dolomite
35 x 48 Canaan dolomite
35 x 48 mesh Tymochtee
dolomite
35 x 48 mesh Canaan
dolomite
35 x 48 mesh Tymochtee
dolomite
Shutdowns
Results and Conclusions
Voluntarily
To be determined
Voluntarily
Data points have been
established
Voluntarily
To be determined
Voluntarily
Data points have been
established
Voluntarily
Data points have been
established
-------
TABLE VII-6 (Cont'd.)
Batch Studies of Process Variables
Run Number
Unit Configuration
A-12O - A-122
Suit idat ion
A-123 A-12 4
Suit idat ion-Regenerat ion Sulf Idat ion-Regenerat ion
A-123 - A-127
Bull Idat ion
A-138
Sulf Idat ion-Rogenerat ion
Date at Run 4/7, 1975
Run Duration with HjS Feed, 3.1-10.1
hrs.
4/9-4/11, 1975
6.5
4/14-4/15, 1975
5.1
4/15, 1975
1.8O-5.7
4/16-4/Z2, 1975
29
Purpose of Run(s)
Establish reactivity of
CaO in fully calcined
Canaan dolomite at 3
bed levels and at low
HaS inlet content
To find behavior of
sulfur acceptor in
su If idat Ion regeneration
cycling under calcining
conditions
Establish reactivity of
CaO in fully calcined
Canaan dolomite
originally regenerated
at high temperature
Establish reactivity of
CaO in fully calcined
Canaan dolomite at 3
bed levels and a high
H,S inlet content
Establish deactlvation
curve of Canaan dolomite
cycled batchwlae
Temperature, °C (°F)
927 (17OO)
927 (170O) Sulfldation 927 (l7OOj
7O4 (13OO) Regeneration 7O4 (13OO)
927 (1700)
871 (16CO) Sulfldation,
704 (1300) Regeneration
Bed Material
U>
35 x 48 mesh fully
calcined Canaan dolomite
35 x 48 mesh sulfided,
calcined Canaan dolomite
35 z 48 mesh fully
calcined regenerated
Canaan dolomite
33 z 48 mesh fully
calcined Canaan dolomlto
35. z 48
dOlODltO
sh sulf idod
Feedstock
35 z 48 mesh Canaan
dolomite
35 z 48 mesh product from
Runs No. A-12O thru A-122
35 z 48 mesh product from
Run No. A-85
33 z 48 •
dolomite
sh Canaan
Product from Runs Ro.
A-87 thru A-9S
Shutdowns
Voluntarily
Agglomeration observed
Interruption No. 1 - loss
of instrument air
Interruption Ho. 2 - power
supply breakdown
Voluntarily
Voluntarily after th*
objective baa boon
reached
Results and Conclusions
Data points have been
established
Regeneration is less
efficient than with stone
in hall-calcined form
Poor regeneration activity
Data points have bean
established
DeoctiTOtlon with
batchwlco cycling Is
about the Mao as
continuous cycling
-------
TABLE VII-6 (Cont'd. )
Batch Studies of Process Variables
Run Number A-129 A-129A A-13O A-131 A-132 A-132A
Unit Configuration Sulfidation-Regeneration Sulfidation-Regeneratio= Sulfidation-Regeneration Sulf idat Ion-Regenerat ion Sulfidation-Regeneration Suitidation-Regeneration
Date of Run
4/21-4/22, 1975
Run Duration with H,S Feed, 1.3
hrs.
4/22-4/24, 1975
3.7
4/24-4/29, 1975
3.7 hours
4/29-5/6, 1975
11
5/6-5/7, 1975
3 hours
5/7-5/13, 1975
11
Purpose of Run(s)
Establish deactivatlon
curve of Tymochtee
dolomite for 2O minutes
gas desulfurizer, and
2O minutes regenerator
residence times per
cycle
Establish deactivation
curve of Tymochtee
dolomite for 2O minutes
gas desulfurizer, and
2O minutes regenerator
residence times per
cycle
Establish deactivation
curve of Tymochtee
dolomite for 2O minutes
in gas desulfurizer and
6O minutes regenerator
residence times per
cycle
Establish deactivation
curve of Tymochtee
dolomite for 6O minutes
gas desulfurizer and 2O
minutes regenerator
residence times per
cycle
Establish deactivation
curve of Tymochtee
dolomite for 6O minutes
gas desulfurizer and 6O
minutes regenerator
residence times per
cycle
Establish deactivation
curve of Tymochtee
dolomite for 6O minutes
gas desulfurizer and 6O
minutes regenerator
residence times per
cycle
Temperature, °C (*F)
871 (16OO) Sulfldation,
7O4 (13OO) Regeneration
871 (16OO) Sulf idat ion,
704 (13OO) Regeneration
871 (160O) Sulfidation,
7O4 (13OO) Regeneration
871 (1600) Sulfidation,
7O4 (130O) Regeneration
871 (16OO) Sulf idat ion,
7O4 (13OO) Regeneration
871 (16OO) Sulfidation,
7O4 (13OO) Regeneration
Bed Material
35 x 48 mesh sulfided
dolomite
35 x 48 mesh sulfided
dolomite
35 x 48 mesh sulfided
dolomite
35 x 48 mesh sulfided
dolomite
35 x 48 mesh sulfided
dolomite
35 x 48 mesh sulfided
dolomite
Feedstock
35 x 48 mesh Tymocbtee
dolomite
35 x 48 mesh Tymochtee
dolomite
35 x 48 mesh Tymochtee
dolomite
35 x 48 mesh Tymochtee
dolomite
35 x 48 mesh Tymochtee
dolomite
35 x 48 mesh Tymochtee
dolomite
Shutdowns
Voluntarily after
discovery of error in
sampling procedure
Voluntarily after the
objective has been
reached
Voluntarily after the
objective has been
reached
Voluntarily after the
objective has been
reached
Voluntarily, unable to
transfer dolomite
between reactors
Voluntarily after the
objective has been
reached
Results and Conclusions
Establish higher
standards for procedures
Residence time affects
deactivation
Residence time affects
deactivation
Residence time affects
deact ivat ion
Fine pressure balancing
of the apparatus is
essential for successful
operation
Residence time affects
deactIvat ion
-------
TABLE VII-7
CaSO« Reduction by CaS
Run Number
Unit Configuration
A-83
Sulfidation (batchwise);
sulfur collecting system
No. 1, Figure IV-11
A-84A
Sulfidation (batchwise);
sulfur collecting system
No. 1. Figure IV-11
A-84B
Sulfidation (batchwise);
sulfur collecting system
No. 1, Figure IV-11
A-84C
Sulfidation (batchwise);
sulfur collecting system
No. 2, Figure IV-12
A-84D
Sulfidation (batchwise);
sulfur collecting system
No. 2, Figure IV-12
Date of Run
2/14-2/19, 1975
3/3, 1975
3/3, 1975
3/6, 1975
3/7, 1975
Run Duration with H3S Feed,
to-s.
3.6
Purpose of Run
To study possibilities
of CaS04 reduction by
H2S at a low temperature
Establish degree of H,S
decomposition
Establish degree of H2S
decompos it ion
To study possibility of
CaSO, reduction by H2S
at a low temperature
To study possibility of
CaSO4 reduction by H2S
at a high temperature
Temperature, °C (°F)
Bed Material
86O (1580)
Sulfated Canaan dolomite
86O (1580)
Magnorite
893 (164O)
Magnorite
86O (158O)
Sulfated Tymochtee
dolomite
893 (1640)
Sulfated Tymochtee
dolomite
Feedstock
35 x 48 mesh Canaan
dolomite
48 x 65 mesh Magnorite
(electrofused, ground
MgO)
48 x 65 mesh Uagnorite
(electrofused, ground
MgO)
35 x 48 mesh Tymochtee
dolomite, sulfided,
oxidized
35 x 48 mesh Tymochtee
dolomite, sulfided,
ox id ized
Shutdowns
Numerous interruptions
due to plugging of exit
system by sulfur
Interruptions due to
plugging of exit
lines by sulfur
Two interruptions due
to plugging of exit
lines by sulfur
Voluntarily
Voluntarily
Results and Conclusions
The use of large amounts
of once through gas is
not feasible, substitute
recyle gas. Corrosion
was severe.
Keep transfer lines from
reactor to condensing
system as hot as possible.
Improvement of sulfur
condensing system is
mandatory.
Reaction rate is rapid.
Equilibrium appears to
be established
Reaction rate is rapid.
Equilibrium appears to
be established
-------
TABLE VI1-8
Low Temperature Ashing of Coal
Run; Low temperature ashing of Hillsboro mine coal (Illinois No. 6 coal).
Unit Configuration; 4" diameter reactor provided with a dip leg extending
to the bottom of reactor for introduction of air and
diluting inert gas.
Date of Run: 5/14-5/30, 1975
Purpose of Run; To prepare feedstock for alkali fume removal type of runs.
Temperature: £ 538°C (lOOO°F)
Bed Material; Batch of char prepared by low temperature carbonization of
Illinois No. 6 coal.
Shutdowns: 1. Higher than anticipated pressure drop through apparatus;
increased working pressure from 3 atm to 7 atm.
2. A poor heat dissipation in stationary bed; increased total
gas throughput for bed fluidization.
Results and Conclusions; Good heat dissipation in fluidized bed of char,
increased air flowrate possible.
96.
-------
TABLE VII-9
Alkali Fume Removal
Ruo Number
Unit Configuration
FR-1
See flow diagram for
Alkali Fume Removal,
Figure IV-6
FR-1A
For Alkali Fume Removal
PR-IB
For Alkali Fume Removal,
new gasket on vessel
flange, more insulation
on exit line, new beat
winding on exit line
FR-2
For Alkali Fume Removal
FR-3 FR-4
For Alkali Fume Removal For Alkali Fume Removal
Date of Run
7/1-7/22, 1975
7/31, 1975
8/6, 1975
8/7-8/8, 1975
8/11-8/12, 1975
8/13-8/14, 1975
Run Duration with HaS Feed,
hrs.
13
1O
10
13
13.75
Purpose of Run
To see if low
temperature asb and
acceptor fines are
alkali fume acceptors
To see if low
temperature ash and
dolomite fines are
alkali fume acceptors
To see if low
temperature asb and
dolomite fines are
alkali fume acceptors
To study alkali fume
removal
To study alkali fume
removal
Overfeed the process
with Nad
Temperature, °C (°F)
vo
871 (16OO) in bed,
7O4 ?13OO) filter,
627 (125O) flange
871 fl6OO) in bed,
7O4 (13OO) filter,
566 (1050) flange
871 (16OO) in bed,
7O4 (13OO) filter,
621 (115O) flange
871 (16OO) in bed,
7O4 (13OO) filter
621 (115O) flange
871 (16OO) in bed,
7O4 (13OO) filter,
621 (1150) flange
871 (16OO) in bed
704 (13OO) filter,
621 (115O) flange
Bed Material
35 x 48 mesh
presulfided Canaan
dolomite
35 x 48 mesh
presulfided Canaan
dolomite
35 x 48 mesh
presulfided Canaan
dolomite
35 x 48 mesh
presulfided Canaan
dolomite
35 x 48 mesh
presulfided Canaan
dolomite
35 x 48 mesh
presulfided Canaan
dolomite
Feedstock
. O42 gm NaCl + 1.O5
gm low temperature
asb + 13.2O gm -ISO
mesh Tymochtee
dolomite/lOOO gm of
c irculatloc
.042 gm NaCl + 1.O5
gm low temperature
ash + 13.2O gm -ISO
mesh Tymochtee
dolomite/lOOO gm of
circulation
.21 gm NaCl + 13.2 gm
Tymochtee dolomite
(-ISO mesh)/lOOO gm
of circulation
.21 gm NaCl + 5.O gm
-ISO mesh Canaan
dolomite/lOOO gm of
circulation
Was not fed
Shutdowns
Involuntarily, pressure
shell flange developed
a leak at 677'C (125O°F)
Voluntarily
Voluntarily
Voluntarily
Voluntarily
Main desulfurizer
flange opened,
Intolerable leak
Results and Conclusions
The pressure shell
flange has to be
operated at a lower
temperature
Unit operated at low
flange temperature,
low exit line
temperature, leak on
the flange, change
diagram, no Na
breakthrough
Sodium does not
break through
filtering device
Sodium does not
break through
filtering device
Sodium does not
break through beyond
filtering device
For reliable operation
the desulfurizer exit
closure has to be
changed
-------
VIII. CYCLIC CONTINUOUS RUNS WITH CANAAN DOLOMITE
A. Introduction
The reaction,
CaC03 + H2S = CaS + H20 + C02
is used for both gas desulfurization and sorbent regeneration. Conditions in
the gas desulfurizer and regenerator are such that the equilibrium H2S content
of the desulfurized fuel gas is low, and the equilibrium H2S content of the
regenerator offgas is relatively high.
Run conditions in this study generally simulated those in process vessels
as given below:
Gas Desulfurizer Regenerator
Temperature, °C (°F) 871 (1600) 760 (1300)
Inlet Gas Composition. Mol $
H2 16
CO 19
C02 8 66
N2 46
H2S ca. 1
H20 10 34
Mol % of Sorbent Ca
Reacted/cycle 10-15
Canaan dolomite served as the criterion to which all other stones were
compared with respect to the critical characteristics of reactivity toward H2S,
regeneration activity, and attrition resistance. To improve operability, a
little hydrogen was added to the regenerator to inhibit dissociation of H2S.
B. Runs Without Makeup
Reactor conditions and results for the gas desulfurizer and regenerator
are given in Tables VIII-1 and VIII-2, respectively. Detailed tables for stone
composition, H2S content of exit gas, and attrition as a function of cycles are
presented in Appendix B. The data for solids composition include a term for
"CaO". This is to account for about 6$ of the calcium which is assumed to
react with impurities in the stone thereby becoming unavailable for reaction
with H2S.
The "CaO" is determined by deducting the assayed values of CaS and CaC03
from 1OO$. In all' Canaan dolomite runs, the attrition rate was about 1$ of
the circulation rate and the H2S concentration in the desulfurizer offgas was
less than O.03 mol % prior to breakthrough.
Run A20A
Run A2OA was similar to Run A7 of 1972, except that the molar con-
version to CaS was 13$ per pass instead of 20$. The progress of the run as
defined by the CaS content of solids leaving the gas desulfurizer and regener-
ator is shown graphically in Figure VIII-1 as a function of the number of
cycles. A total of 32 cycles was achieved before depletion of the inventory
due to attrition and sampling forced shutdown.
99.
-------
TABLE VI11-1
Conditions ond flcnultn for Gas Doaulfiirltor »lth Canaan Doloratto Pood
System Pressure: 19 ata (206 palg)
System: Temperature: 871°C (16OO°F)
Run Number
A20A
A21
A22A
A36A
A481
AS1
AS2A
Cannon Dolomite Butch Number
Acceptor Stxo Consist, Tyler Mesh
Food Rate, (;n/hr (half-calcined basis)
Nominal Solids Residence Tine, alfk
Fresh Dolomite Food Rate, gms/br
Input, SCF1I*3)
Recycle to Oed
H20
»2
CO
C02
«2
H23
Purges (C02) to Bed
Purges (N2> above Bed
Recycle Acceptor Lift Gas, above Bed
Output In Cycle Ho.
Exit Cos Rate, SCFH (dry basic)
Com poo tt Ion, Hoi %
H2
CO
coa
"a
35 X 48
1810
52
33 x 48
1840
SI
35 x 48
1892
50
20 X 28
3120
43
35 x 48
1260
7V
35 x 48
1990*2*
44
84
39 » 48
1930<»
46
O4
148
31
35
33
8.8
66
1.6
.71
33
143
16.5
17.0
9.8
96.6
0.03
148
31
39
33
8.8
66
1.6
71
18
144
16.6
16.8
0.0
96.3
0.03
148.
21
39
23
8.8
66
1.6
71
8
144
16.6
16.1
10.3
97.0
0.03
310
60
55
62
21
152
1.8
110
4
338
17.8
17.0
0.4
96.0
0.04
148
30
36
23
. 0.5
66
s'.e
110
3
141
17.5
18.2
8.7
85 2
0.39
-------
TABLE VI11-2
Conditions and Results for Regenerator with Canaan Dolomite Feed
Run Kunber
Temperature, "C (»F)
Komloal Solids Residence Time, Bin
Input, SCFH
Recycle to Bed
I72O
H2
CO
C02
Purges (N2) to Cod
Purges (Jij) above Bod
Purges (003) above Bod
Output In Cycle Mo. •
Exit COB tutu, 8CPII (dry bails)
Composition, Hoi %
HZ
CO
C02
"2
11^3
cos
Outlet GJJ, Top of Pod
Cnm|/«ji Hloil, Mill IT"
CO
C62
*2
H2S
cog
Flow Rate, SC7H, Top of Bod
Fluldlrlng Velocity, ft /BOO
% 833 _ In Outlet/Equilibrium % H23
Ecgeoeratlon of Acceptor /Pans , Mol % of Tota.\ Ca
System Pressure: IS
A20A
704 (1300)
60
0.0
S3
8.0
O.Q
91
7
10
4
23
120
S.O
3.95
71.8
18.3
0.82
0.04
32.0
3.8
3.0
32.0
7.6
0.62
0.03
158
0.44
0.18
8.8
A21 .
704 (1300)
59
102
47
3.0
0.0
SO
10
O.O
15
18
33
3.9
5.2
83.1
5.2
2.54
0.11 ;
29.0
3.1
4.1
57.5
4.1
2. '03
0.09
169
0.47
0.61
7,4
atm (206 .pale)
A22A
760 (1400)
57
0.0
48
8.0
0.0
86
10
10
4.S
8
116
3.0
3.7
71.4
10.7
0.04
0.04
32.9
3.0
2.8
51.8
8.5
0.72
0.031
1S1
0.44
0.43
9.3
A36A
32
0.0
108
11
O.O
205
4
236
2.1
2.6
04.1
O.C
O.C2
0.03
33.3
1.5
1.8
02.4
0.4
0.44
0.02
333
0.92
O'.ll
ll.'O
A48
85
35
51
5
0.0
70
2
10S
3.0
1.9
02. C
0.3
1.85
0.1G
28.9
2.4
1.5
C5.2
0.3
1.47
0.13
177
0.49
0.39
25
AS1
54
89
SI
3.S
0.0
21. S
14
54
4.1
4.0
88.9
0.4
2.53
O.1G
39.3
3.3
3.1
C1.9
0.3
2.00
o.i:
181
0.50
0.55
11
A52A
55
89
51
3.5
6.0
21.5
38
55
4.0
2.8
88. 8
0.7
2.93
O.18
29.3
3.2
2.2
61.9
O.J
2.31
0.14
182
O.50
O.G-»
13
-------
-------
-------
The stone remained reactive toward H2S throughout the run; there was
no breakthrough of H2S in the gas desulfurizer even at the end when only 10
mol $ of CaC03 was left in the stone. The attrition rate for the entire run
was 0.7$ of the stone fed.
Regeneration activity was correlated using percent regeneration, or;
mols of CaS converted to CaC03/100 mols of CaS fed to the regenerator. The
data were plotted on log-log coordinates using number of cycles as the absicca.
The percent regeneration can be determined in two ways. The first is
to take the difference between the CaS or CaC03 content curves and divide by
the CaS content leaving the gas desulfurizer. Due to the residence time in
the regenerator, the point equivalent to n cycles on the gas desulfurizer
curve is displaced to about n + .3 cycles on the regenerator curve. The method
is most accurate when the curves have zero slope. The other method is to
determine the moles of H2S in the regenerator offgas, put this in terms of the
change in CaS concentration and divide by the CaS content leaving the gas
desulfurizer. The two methods are generally in close agreement. All the
regeneration data in this chapter were determined by difference between smoothed
solid concentration curves.
Figure VIII-2 shows the results of these calculations for Run A2OA
and several others.
The data correlated in Figure VIII-2 fit a straight line. The
decline in activity in going from 3O to 100 cycles is proportionately the same
as going from 3 to 7 cycles when a linear log-log relationship is obtained.
The stone should therefore continue to be active long enough to achieve a low
make-up requirement.
Run A21
Run A21 was made at the same conditions as Run A20A except that
recycle gas was used in the regenerator. This allowed the H2S concentration in
the exit gas to more closely approach the equilibrium value. Run conditions
are given in Tables VIII-1 and 2, and analyses of exit gases and solids are
given in Appendix B. The progress of the run as defined by the CaS content of
acceptor is given in Figure III-3. The exit H2S concentration, as determined
by the continuous analyzer was 55-75$ of equilibrium during the course of the
run compared to ca. 20$ in Run A20A. The data for Run A21 in Figure VIII-2
were in good agreement with those of Run A20A, and it was concluded that the
rate of regeneration was unaffected by the increased H2S concentration. The
run continued for 36.7 cycles making it the longest run with Canaan stone.
There was no H2S breakthrough and no sign of deactivation of CaC03 with respect
to pickup of H2S. The data indicate that deactivation with respect to regenera-
tion was the same in Runs A21 and A20A.
In examining approach to equilibrium, regenerator data for H2S con-
centration were used. Some results are summarized below. Large differences
in approach to equilibrium and kinetic driving force are apparent, yet there
was no effect on the overall rate of regeneration. This is consistent with
the mechanism of inactive CaS being the limiting factor in regeneration.
103.
-------
FIGURE VII1-2
DEACTIVATION OF CaS IN CANAAN DOLOMITE
AT 704°C (1300°F)
111
5 0 7 8 9 10
Number of Cycles
1.5
2.5
6 1 8 9 10
104.
-------
FIGURE VII1-2
DEACTIVATION OF CaS IN CANAAN DOLOMITE
AT 704°C (130O0F)
^mii
S 6 7 8 9 IU
Number of Cycles
8 9 10
104.
-------
Percent H2S Top of Bed Average
Run Cycle Top of Bottom Equilibrium Equilibrium
No. No. Bed of Bed Avg. ' $ %
A21 13 2.50 1.57 2.04 74 47
A20A 20 O.73 0 O.36 22 11
In a run with hardened Tymochfee dolomite (Section IX-B), the
measured H2S concentration in the exit gas was at 99$ of equilibrium. In that
run, there was also no indication that the regeneration rate was slowed by the
high H2S content of the gas. It is concluded that the regeneration reaction
is intrinsically fast for that fraction of the CaS which is reactive, and that
the conversion of stone is not slowed by the presence of product gas at the
solids residence times employed in this system. There clearly is no difficulty
in attaining the equilibrium concentration of H2S in the exit gas at 704°C
(1300°F).
It must be emphasized that the continuous runs without makeup pro-
duced a regenerator feed of varying active CaS content which made it impractical
to maintain an exit gas concentration close to the equilibrium value. As the
stone activity declined, the H2S content of regenerator offgas varied, and
there was no attempt to keep it close to equilibrium.
Run A22A
Run A22A was made with the regenerator at 760°C (l4CO°F). Regenera-
tion at cycle 10 improved from about 16$ at 704°C (13OO°F) to 25$ at 760°C.
Progress of the run in terms of solids concentration is shown in Figure VIII-4.
Due to the improved regeneration, the CaS content of the acceptor rose more
slowly than in previous runs.
Regeneration activity as a function of cycles is presented in Figure
VIII-5. Results are presented together with data for 704°C taken from Figure
VIII-2 and data for 593°C taken from Runs A15 and Aie.C1) The lines drawn are
roughly parallel.
Runs A36 and A36A
Runs A36 and A36A were duplicate runs. Several operating shutdowns
were experienced in Run A36, and it was decided to repeat the run. Stone in
the 2O x 28 size consist was fed to determine the effect of particle size.
As shown in Table VIII-1, the attrition rate was nominally 1.7$, but this
includes ca. 30$ +1OO mesh material. The larger fluidizing velocity required
by the 20 x 28 mesh particles caused solids which would normally be retained
in the bed to be blown out. The problem was aggravated by the straight side
of the vessel (no disengaging zone) and the restriction above the bed entailed
by use of a preheat coil.
Sulfur absorption and regeneration activities were the same as with
35 x 48 mesh stone. The run progress is detailed graphically in Figure VIII-6.
Run A48
Results for this run are presented graphically in Figures VIII-7
and VIII-8.
105.
-------
-------
") X "
EUFI
'! 1ST
: 10 r
r IN u
> i:
-------
FIGURE VIII-5
DEACTIVATION OF CaS WITH
TEMPERATURE AS A PARAMETER
=100= Me Is -CaS :Fec --^r-:
1.3
o 678 l-j 1.5
Number of Cycles
8 9 '0
108.
-------
) X MN . 10 I
EUF ;SEH : IN u
-------
10 X 10 TO H INCH 7 X 10 INCHES
KEUFFEL Ok CSSER CO. MUDt IN g $»
46 1322
TFIT1T
• 11 • •• > •
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-------
Run A48 was made with Canaan stone which had been initially sulfided
to a high conversion. This more closely simulated the conversion history of
make-up stone fed to the process than did the technique of slowly building up
the CaS content.
The run was started by sulfiding the stone batchwise at about 10$/hr
for 7-1/2 hours. The combined sulfided product contained 78 mol $ CaS. This
was then fed to the unit continuously for regeneration. Some 40% of the CaS
was converted to CaC03 leaving a CaS content of 43$. For the period encompas-
sing the first regeneration through cycle 5.6, residence times were about 70
min in the gas desulfurizer and 80 min in the regenerator. For cycle 1-2, the
stone was sulfided to 85$ CaS and then regenerated to 49$ CaS again yielding
about 40$ relative regeneration of CaS.
On the start of the third cycle (2-3), breakthrough occurred. A
sample was not taken at this point, but the stone should have been about 96$
CaS based on the H2S feed rate. Upon breakthrough, the H2S feed was reduced
from 47 mol $ of the calcium feed rate to 23$. The H2S breakthrough disappeared
and cycling continued. Breakthrough was again observed in cycle 5.3, and in
cycle 5.6 the circulation rate was increased. This had the effectrof reducing
the sulfidation conversion to 13$ of the calcium circulation rate and decreasing
the residence time in each reactor to 4O-50 minutes.
Upon further breakthroughs, the H2S feed was reduced to give 8.5$
and 4.0$ sulfidation in cycles 8.9 and 20.6, respectively.
The H2S content of gas leaving the desulfurizer was normally low
during the course of the run except when insufficient CaC03 in the stone led
to breakthrough. In Run A48 the CaC03 contents were: less than 8$ on the first
breakthrough, about 10$ for the second, and less than 5$ on the third and fourth
breakthroughs.
It is believed that the higher CaC03 content at breakthrough observed
early in the run was due to the high H2S inlet rate. This is consistent with
the assumption that breakthrough occurs at a constant space velocity of mols
H2S fed/hr:mols CaC03 in the bed.
As shown in Figure VIII-8, the regeneration step was normal until the
circulation rate was increased after the second breakthrough in cycle 5.6.
From that point on, regeneration was worse than in previous runs with Canaan
dolomite. For example, in cycle 20, the observed regeneration was 6.5$ versus
12.0$ expected. Based on previous experience, one would have expected a step
drop in percent regeneration coincident with a step reduction in H2S input or
increase in circulation rate. The regeneration performance should then have
remained constant for some time. In Run A48, there was no constant level of
regeneration performance after the second and third breakthroughs. The data
taken about the first breakthrough are inconclusive in this report.
The attrition rate observed was 0.2$ which is lower than usual for
Canaan dolomite.
In summary, the results of Run A48 differed from previous runs in that
breakthrough occurred with more CaCO3 present, regeneration was poorer, and
the attrition rate was lower. In addition, reference to Figure VIII-15
(Section XVIII-D), which.is a summary of density measurements on cycled Canaan
dolomites, shows that the density of stone from Run A48 was higher than in
previous runs.
111.
-------
r
FIGURE VIII-8
Run A48 - Deactivation of CaS at 7O4°C (1300°]?)
1.5
2,3
5 6
Number
7 8 9 10
of Cycles
1.5
2.5
678 10
112.
-------
The conditions for Run A48 differ in two respects from earlier runs:
The stone was maintained at a high level of sulfidation from the very first
cycle and the stone was exposed to gas desulfurizer conditions for some eight
hours prior to its initial regeneration cycle followed by doub.le the usual
residence time in each vessel until cycle 5.6.
It is presumed that the longer desulfurizer residence times caused
the stone to sinter during the cycling performance. At the conclusion of cycle
1, the density was consistent with that of earlier runs. The density increased
markedly from cycle 1 to 10 and then increased slowly. The low attrition rate
may be simply attributed to hardening and densification. The CaS in the
unusually dense stone could have been less available for reaction thereby
explaining 'the low regeneration activity,
C. Runs with Makeup
The objective of these runs was to more closely simulate proposed commer-
cial operating conditions. More specifically it was desired to demonstrate
that equilibrium activity of stone in a continuous make-up run can be calculated
from runs where a single batch of stone was continuously cycled through the
process. .
. Run A51
Run A51 was the first attempt to run with makeup. Due to an opera-
ting error, the sulfidation of the Canaan stone per cycle was only 15.3% on
average. This was lower than intended, and the run was discontinued after 14
.cycles. Detailed run results are presented in Tables VIII-1 and 2 and in
Figure VIII-9. Except for the sulfidation conversion, the gas desulfurizer
conditions were the same as for previous runs. In the regenerator, the recycle
gas flow was increased in order to more closely approach equilibrium. A molar
makeup of A.87» fresh stone was fed to the gas desulfurizer.
The attrition rate and desulfurization activity were satisfactory.
The concentration of H2S in the regenerator offgas reached 55$ of equilibrium.
When using makeup, a regeneration-deactivation correlation cannot be
made as it had been previously. Instead, various correlations are fitted to .
the data. For the makeup runs, three correlations, shown in Figure VIII-10,
were used. These were;
a. The correlation based on previous runs with Canaan
dolomite, excepting Run A48.
b. A correlation based on the deactivation data of Run A48,
c. An arbitrary intermediate correlation.
Application of these correlations to the actual conditions of Run
A51, 4.8$ makeup and 15.3$ sulfidation, gave the results presented in Figure
VIII-11. These results were obtained by inserting the run conditions into
the computer program described in detail in Section VI. This program calcu-
lates the composition of an inventory of stone under given conditions of
conversion, makeup rate, and deactivation behavior as a function of cycles.
As the computer program for cyclding does not take into account the presence
113.
-------
CaS in Exit Solids
~- " Run A51 - Mol
114.
-------
FIGURE VIII-10
Three Possible Correlations for Deactivation of
CaS in Canaan Dolomite at 7O4°C (13O00F)
2.5
5 (i 7 8 9 10
Number of Cycles
1.0
2.5
7 8 9 10
115.
-------
10 1
Cycle Number
-------
of "CaO", the actual run data shown in Figure VIII-12 were put on a "CaO"-free
basis. Figure VIII-11 clearly shows that the best fit is obtained with the
Canaan data correlation.
Run A52A
The conditions for Run A52A were changed from Run A51 only by going
to a sulfidation rate of 19.3$ per cycle. The higher sulfidation level meant
that the equilibrated CaS content leaving the gas desulfurizer would approximate
proposed plant conditions. The make-up sorbent would be sulfided to a high
conversion immediately, and there would be just enough excess CaC03 to prevent
breakthrough. At a 5$ make-up rate, the system would be 99$ equilibrated with
respect to CaS content in 4O cycles. The run lasted for 39 cycles. The only
difficulty encountered was that the H2S feed rate was low for cycles ca. 2O-25.
The raw run data on a "CaO"-free basis are plotted in Figures VIII-13 together
with the three curves obtained from the trial correlations. The Canaan data
correlation gives an excellent fit to the data, while the others clearly do not
fit. In Figure VIII-14, results are shown for the inclusion of the temporarily
lower H2S feed rate in the computer input for the Canaan correlation. The fit
of the Canaan correlation to the actual data was thereby improved.
The attrition rate was about 1$. Hydrogen sulfide concentration in
the gas desulfurizer outlet was 0.01-0.03$ until breakthrough. Breakthrough
occurred with 4$ CaC03 remaining. At 20$ sulfidation, this is equivalent to a
CaC03;H2S molar space velocity of 3.5. For most of the run, the H2S concentra-
tion in the regenerator was 60-75$ of equilibrium. As in other tests, no
deleterious effect of a high H2S concentration on regeneration activity was
observed.
It is concluded that running with make-up acceptor to a high CaS
level in the gas desulfurizer has no unpredicted effects on stone character-
istics. The computer program for predicting activity of the inventory provided
a good fit to the data.
D. Density Measurements on Cycled Canaan Dolomite
Density data for sulfided Canaan dolomites are presented in Table VIII-3
and in Figures VIII-15 and VIII-16. The data are normalized to a CaO-MgO
basis in order to correct for the weight changes due to chemical reaction. In
all cases, the density of the stone increased with cycling. There were differ-
ences between the runs: after 20 cycles, the stone from Run A48 had a
normalized density of 123 lb/ft3 versus 99 lb/ft3 for the Run A21 stone. These
represent the high and low ends of the density range.
Densification of stone probably occurs in the gas desulfurizer due to
thermal sintering. As our experimental desulfurizer operated at 871°C (l60O°F)
instead of at 914°C (1678°F), the process design temperature, additional sinter-
ing would be expected at process conditions.
Runs A15, A20A and A22A were made at regenerator temperatures of 593, 704
and 760°C (1100, 1300 and 1400°F), respectively. Although the regeneration
activity differed widely in these runs, Figure VIII-15 shows that the densities
were the same. It is concluded that the regeneration temperature does not
affect stone density.
117.
-------
Cyrfcle Nuuber
1 h A
30 ^—~
20 -
10 -
O
0
-------
0 TC
. EL a .
H
CO.
4' 2
24 26 28 30 32 34 36
40
-------
8 1O 12 14 16 18 2O 22 24 26 28 30 32 34 36 38 40
-------
TAIIU: vi 1 1-3
Penalty of Snlfld-.-d C»na»n Dolomite
A1S
A16
A20A
A21
A22A
A36A
A48
A51
A52A
Hoi $ of Total
Source of Sample Cycle No. CaCOj CaS
C»
"CaO"
Gas Dosulfurlzer 13.9 + 2 hours 1.6 98.4\
Regenerator 0.8
5.7
13.3
87.8 12.2 1
52.0 48.0 I
15.0 85.0 f
13.9 + 3 noun 34.1 75.9 1
Regenerator 0. 4
8.3
13.1
16.0
17.1
Regenerator 10.5
Regenerator 4.0
110.1
19.1
30.0
31.7
Regenerator 5.9
13.4
32.1
36.7
36.7
Regenerator 6.0
7.6
13.4
19.6
32.0
Oaa Desulf urlzer 4,4
110.9
20.0
64.2 35.8V
11.4 88.61
11.4 88.6 \
10.0 90.0 \
10.1 89.9 1
11.7 M.S./
71.2 21.9
50.7 43.7
37.2 39.5
19.1 80.9
16.7 80.9
64.9 34.6
43.1 55.9
22.9 7S.O
21.7 72.2
21.7 73.3
79.0 18.0
75.4 21.4
64.4 35.6
50.8 45.6
45.9 60.8
62.4 33.3
37.6 36.2
10.8 83.6
20. 0 -f 5 hours 12.9 73.5
Regenerator 4.2
110.8
20.0
73.3 22.1
48.2 49.3
23.5 75.3
20. 0 + 5 hours 39.9 56.1
Gas Desulfurlxer l.O
14.5
10.5
22.7
Regenerator l.O
4.4
10.4
22.6
Regenerator 3. 4
4.6
14.0
Gas Desulf urlzer 14.1
Regenerator 4. 5
12.9
20.1
22.7
25.4
28.9
32.5
35.9
38.6
38.6
38.6
22.4 77.6
1O.3 85.5
6.5 87.0
7.O 86.8
S2.O 43.4
28.9 63.8
13.1 79. 0
11.7 82.8
73.9 19.0
67.2 28.9
48.0 46.8
34.0 60.5
61.7 33.0
39.0 53.6
(4) 39.6 67.3
30.7 63.4
26.7 67.8
24. 5 70. 4
20.1 54.3
20.2 56.0
18.6 75.7
18.6 73.7
18.6 73. T
1)
6.9
5.6
3.2
0.0
2.5
0.3
1.0
2.2
6.1
6.1
3.0
3.2
0.0
3.7
3.2
4.3
6.2
5.8
11.6
4.7
3.5
1.3
4.0
0.0
4.2
6.6
6.3
4.6
5.3
5.9
3.5
7.0
3.9
7.3
3.4
5.3
5.4
3.1
6.0
3.5
5.1
3.3
1.4
5.6
5.6
5.6
Density!
gramH/cc
2.16
2.29
2.23
2.26
2.29
2.16
2.09
3.10
3.09
2,09
3.03
3.30
3.36
3.38
3.30
3.39
3.10
3.04
2.00
2.01
1.98
2.34
2.34
2.37
2.38
2.40
3.13
3.07
3.01
2.O1
2.15
3.10
3.O4
2.14
1.98
2.16
2.25
2.34
2.08
2.23
2.30
2.36
2.12
2.06
2.02
1.97
2.07
2.O3
2.06
3.07
2.07
2.10
2.13
2.16
2.23
2.23
2.24
•)
Ui/ft'
134.5
142.6
139.0
141.3
143.0
134.9
130.5
130.8
130.6
130.4
126.3
143.7
141.3.
142.3
143.3
142.7
131.3
127.5
124.7
123.6
123.7
146.2
145.9
147.9
147.4
149.8
132.0
128.9
135.3
135.6
134.0
130.9
137.3
133.5
133.8
134.7
140.5
145.8
130.1
139.3
143.5
147.1
133.3
128.6
125.7
123.8
129.4
126.8
128.6
129.2
129.2
130.9
133.2
134.6
138.9
140.3
140.1
Normal Izod
gra»m/<:c~
1.68
1.63
1.70
1.78
1.88
1.88
1.53
1.58
1.62
1.63
1.60
1.67
1.68
1.75
1.79
1.83
1.37
1.63
1.69
1.70
1.56
1.61
1.66
1.68
1.61
1.82
1.93
1.99
1.59
1.80
1.92
1.98
1.53
1.50
1.55
1.35
1.53
1.58
1.64
1.64
1.66
1.69
1.74
1.75
1.83
1.84
1.84
104.9
1O3.9
106.2
111.0
117.8
117.1
96.6
88.5
101.0
1O1.7
100.1
104.0
104.6
109.0
111,8
114.0
98.2
101.8
100.6
1O6.3
97.3
100.3
103.3
104.6
10O.7
113.6
119.9
134. a
99.2
113.3
119.8
133.8
93.5
93. T
96.5
96.8
95.6
98.7
1O2.3
103.6
103.6
103.3
108.8
109.3
113.8
114.8
114.7
i) Uoiiviirod in nereury at 1 atnosphero.
>) Calculated by normalizing each calcium conponent to a CaO bails while assunlng no chance In sonple volume.
t) KOI- UIOHO annlvHoH, CnCO, wns do to rained by assay and CaS was estimated as the difference Iron 100 mol £.
4) Mnlo; This is not the sane as cycles in a run without makeup. When u«lng makeup, the average age of the stone Is less
than the number of cjxlea.
121.
-------
IV
N>
140
13O
120 -
110
1OO
90|
8 10
12 14 16 18 20 22 24 26
28
3O 32 34 36 38 40
-------
All the runs with Canaan dolomite gave low attrition rates, but the more
dense stones generally exhibited lower attrition rates. The overall attrition
rate for Run A48 was 0.3$ versus 0.8$ for Run A21.
The cause of the variation in densities is not yet known. The high densi-
ties observed in Run A48 may have been due to the protracted initial residence
time in the gas desulfurizer. The low densities observed in Run A21 are more
difficult to explain. The only variable change was the use of extra recycle gas
in the regenerator which increased the H2S concentration. It is thought that
the H2S concentration may affect the crystal morphology thereby influencing the
degree of densification which later occurs in the desulfurizer. The evidence
for this occurring in our system is rather indirect, and the model is discussed
more fully in Section XV.
For Runs A51 and A52A, normalized density as a function of cycles is shown
in Figure VIII-16. The data are on the high side of the range of densities
reported for other Canaan dolomite runs. It is noted that after 40 cycles the
arithmetic average age weighted for the quantity of stone was about 17 cycles.
The equilibrium average age would be 19.5 cycles. At the end of Run A52A, the
normalized density was 1.84 gm/cc (115 lb/ft3). The corresponding numbers for
Runs A21, A20A, and A48 were: 1.59, 1.76 and 2.00 gm/cc (99, 110 and 125 lb/ft3),
respectively.
E. Attrition Data
Attrition data as a function of cycles are presented in Figure VIII-17.
The attrition rate was consistently under 1$. In four out of five cases, the
observed rate of attrition decreased as cycling continued. It is assumed that
this is due to sintering and densification of the stone as indicated by the
increase in normalized density. Run A21 was the only one which showed a level-
ing trend in attrition as it was cycled, and this run showed the least densifi-
cation.
Attrition results for runs with makeup are given in Figure VIII-18. For
these runs, there is an increase in attrition followed by a decrease. The
normalized density showed a steady increase. It is noted that the recycle runs
and Run A21 shared the condition of a high H2S concentration in the regenerator,
and the result of an attrition rate which initially increased and then leveled
off or decreased. It is thought that the H2S effect may have been responsible
for the observed effects via its influence on crystal morphology. However, a
detailed model of the mechanism has not been formulated.
F. Conclusions
The key conclusion is that, experimentally, Canaan dolomite exhibits the
characteristics required to fit into the process design presented in Section III.
Specifically, the following were shown to be true:
1. The gas desulfurizer can be operated without breakthrough until
most of the CaC03 is consumed. This is true even when the stone
has undergone numerous cycles.
123.
-------
150
1OO
90,
0 2
-------
iu /v lO TO --i llNt-H 7 A iu INCHES
KEURFEL & ESSER CO MADE IN u s *
46 1322
Ul
0.2 -
0
35
Number of Cycles
-------
-------
2. The decline in regeneration activity is sufficiently slow to
enable the stone to be recycled for a large number of cycles.
A low make-up rate can therefore be used.
3. There is no influence of a high H2S concentration on the rate
of regeneration under our conditions. It should therefore be
possible to run at a high H2S concentration in the regenerator
offgas and also to maintain the required regeneration rate.
4. The attrition rate is on the order of 1%. This is less than
the required make-up rate to maintain activity.
5. The activity of an acceptor inventory obtained in a system
using makeup can be calculated from non-makeup run data. Thus,
the results of simple cycling runs can be used for process
design calculations.
127.
-------
IX. CYCLIC CONTINUOUS RUNS WITH HARDENED DOLOMITES
A. Introduction
In the previous study,v1/ three of the four dolomites tested displayed
high rates of attrition. These were Tymochtee, Buchanan and Pennsylvania dolo-
mites. It was decided that a method of hardening local dolomites was needed
for those cases where a local hard stone of good activity was not available.
Work on the C02 Acceptor Process!11) had shown that Tymochtee dolomite was
initially soft in the CaC03-CaO cyclic process. However, the stone hardened
in that cycle when fuels containing sulfur were used. It was determined that
CaS04 was involved in producing a transient liquid phase which helped the stone
become attrition resistant. Attrition rates of less than l.O$ were observed in
both bench-scaleC1J) and the pilot plant units.
With this background information, it was decided to harden the stone in
our desulfurization process via an oxidation step involving formation of CaS04.
B. Batch Hardening Tests
Exploratory batch runs were made with Tymochtee dolomite to harden the
stone via formation of low-melting liquids in the system CaS04-CaS-CaC03. Tests
were run batchwise using the gas desulfurizer reactor. The stone was first
sulfided at 871°C (l6OO°F) according to the conditions given in Table IX-1 and
then oxidized according to the conditions given in Table IX-2. The planned
procedure was to sulfide the stone to about 70 mol $ CaS and then oxidize all
the CaS to CaS04. Using this procedure a wide range of compositions from
O-70-30 to 70-0-30 molar CaS04-CaS-CaC03 would be obtained. As a final step
the CaS04 was reduced back to CaS.
In Run HR1, air was fed at a rate sufficient to oxidize the stone within
two hours. However, a severe exotherm developed and the run was shut down when
the observed temperature exceeded 1040°C. Runs HR2 and HR3 were made at much
lower air rates and the temperature was readily controlled at 871°C (16CO°F)
and 927°C (l700°F), respectively. The HR1 product formed an agglomerated mass
in the vessel but the HR2 and HR3 products were free flowing.
Products were tested for attrition resistance in two ways. The first was
to put a screened sample on a rolling mill under controlled conditions and then
rescreen it. Results of these tests for the exploratory runs and for several
standards are given in Table IX-3. The HR1 product was considerably better
than A2 product, a sulfided Tymochtee dolomite, and not quite as hard as Run
A2O product, a sulfided Canaan dolomite. However, Run HR2 and HR3 products
were still about the same as unhardened stone. In the table, all stones are
compared to A6, a Tymochtee product first hardened in the C02 Acceptor Process
cycles and then tested in the present sulfiding-regeneration cycle.(i/
The other method of checking strength was to screen out a 35 x 48 Tyler
mesh fraction, and perform a drop test on a small sample with the apparatus
shown earlier in Figure IV-3. The resulting material was then screened for
over and under 48 mesh weights. Results given in Table IX-4 were in agreement
with the results of the rolling mill test.
129.
-------
TABLE IX-1 ,.
Conditions for Batchwise Dolomite Sulfiding
System Pressure: 15 atm (206 psig)
System Temperature: 871°C (1600°F)
Run Number
HR1
HE2
HR3
HR4 to •••HR10'
Acceptor
Size Consist, Tyler Mesh
Sulfiding Time, min
Input,
Recycle to Bed
H2S
C02
H2
HoO
*2
CO
Purges (N2) above Bed
Recycle Lift Gas (only for feeding)
Fluidizing Velocity, ft/sec
Total Amount of Acceptor Fed, grams
Total Amount of Sulfur Fed, gram mols
T| I I
(1) Input is given after shift reaction and hydrolysis of CS2 have taken place.
f
jf
2800
1 /IP
1 Sfi
80 _ —
01 f
00 f\
-1 =
T/->
2400 2500 2500
S 3R
TABLE IX-2
Conditions for Batchwise Oxidation of Sulfided Dolomite
Rim Number
Acceptor
Oxidizing Time, hours
Temperature, °C
Input, SCFH
Recycle to Bed
Air
C02
N2
Purges (N2) above Bed
Total Amount of 02 Fed in Air,
gram mols 02 (till breakthrough)
System Pressure: 15 atm (206 psig)
HR1 HR2
HR3
HR4 to HR10
Sulfided Tymochtee Dolomite
ca .75
(2)
55
70
32.75
871
(1600°F)
2
20
105
28
927
(1700°F)
1 3fl
2
20
105
- 1 S
9
982
(1800°F)
5
26
90
N.D.
(1)
16.7
14.3
(1) Air was on and off in efforts to control temperature.
(2) Uncontrollable due to agglomeration, max. .temperature > 1O4O°C.
11
130.
-------
TABLE IX-3
Properties of Materials From Hardening Study
Hoi % of Total Ca
Density
Normalized
(1)
Source and Description of Dolomite
Run A6, 20 x 35 tt
35 x 48 H
48 x 65 M
35 x 48 H
Run A2
28 x 35 M
Run A20A,
35 i 48 H Canaan
35 x 48 M
Tymochtee 10
Run HR4
35 x 48 M
Tymochtee 10
HR5-HR10
35 x 48 H
Run A23
35 x 48 H
Tymochtee
Raw Canaan
Raw Canaan
Raw Tymochtee
Tymochtee
Regenerator Sample, 1.4 cycles
Regenerator Sample, 31.7 cycles
HR1 Product<3>
HR2 Product
HR3 Product
Half Calcined Feed
Sulfided Feed
Oxidized Feed
Reduced Feed
Reduced Composite Feed
Tymochtee 1O
Desulfurizer Sample, 22.1 cycles
Tymochtee 1O
Run A24, 35 x 48 H Oxidized Feed
Tymochtee 10 Regenerator Sample, 15:2 cycles
Run A25, . Sulfided Feed
35 x 48 M Oxidized.Feed
Tymochtee 1O Regenerator Sample, 36.0 cycles
Run A26 Sulfided Feed
35 x 48 H Oxidized Feed
Tymochtee 10 Regenerator Sample, 21.5 cycles
Run A27 Sulfided Feed
35 x 48 M Oxidized Feed
Tymochtee 10 Regenerator Sample, 20.6 cycles
Run A28 Sulfided Feed
35 x 48 M Oxidized Feed
Tymochtee 10 Regenerator Sample, 9.7 cycles
Ibs
Density
100
100
100
52
16.7
CaS
"CaO"
48
80.9
1.84
114.9
2.5
1.42
88.8
19.6
39.3
52.5
100
17.3
36.5
40.2
36.8
2.4
64.8
12.7
70.5
65.3
13.9
68.1
68.7
0.2
61.7
63.6
12.9
70.4
73.2
34.7
16.4
1.8
0
—
82.7
0
39.3
44.1
85.8
0
77.2
29.5
1.8
75.5
25.0
0.8
93.1
30.2
3.1
79.1
19.1
0.9
52.6
51.7
52.0
40.7
—
—
54.1
5.1
2.7
—
30.2
—
—
30.4
—
—
25.9
—
—
29.6
—
—
19.8
—
12.4
6.9
6.9
—
0
9.4
15.4
16.4
11.8
5.0
10.1
0
2.6
11.5
6.9
4.6
6.7
8.1
3.7
8.0
10.1
6.0
12.7
2.36
2.06
2.04
1.85/1.90
1.67/1.65
2.36
2.03
2.08
1.84
2.14
1.68
1.81
1.99
1.64
1.79
1.94
1.59
1.77
1.89
1.61
1.79
1.89
1.66
147.3
128.8
127.0
116.8/118.5
104.0/102.9
147.5
130.2
129.7
114.8
133.4
1O4.7
112.7
124.2
102.2
111.8
121.2
99.1
110.4
118.0
100.7
111.9
117.8
103.4
1.65
1.39
1.40
1.39
1.42
1.59
1.70
1.72
1.63
1.46
1.45
1.37
1.35
1.42
1.40
1.37
1.41
1.40
1.33
1.40
1.40
1.35
1.39
102.9
86.8
87.2
86.5
88.4
99.3
105.9
107.4
101.5
91.0
90.3
85.5
84.3
88.6
87.5
85.4
88.2
87.5
82.7
87.6
87.6
84.5
86.6
Hardness
Test
Result
1.0
4.0
3.7
3.9
10.6
2.4
3.6
7.9
11.5
17.8
7.4
4.2
3.0
3.3
2.8
6.4
4.3
10.7
7.0
3.6
11.1
8.3
5.2
10.8
7.3
4.8
11.4
10.8
5.3
(2)
(1) Calculated by normalizing each calcium component to a CaO basis while assuming no change in sample volume.
(2) The results of the increase in surface after tumbling in a rolling mill at a standard condition. 1.0 = 220 cm /gm increment, the change observed in a Tymochtee
dolomite which had been hardened in the CO2 Acceptor Process cycle. Test conducted in Roalox tumbling Jar, Size 00, 10 Burundum grinding cylinders, size 13/16
x 13/16" at 51 rpm for 30 minutes.
(3) Taken from the free flowing part of the run product.
-------
TABU IX-3 (continued)
OJ
Ni
Source and Description of Dolomite
Run A29
35 i 48 M
Tymochtee 10
Run A30.
35 x 48 U
Tymochtee 10
Run A30A
35 x 48 H
Run A3Z
20 x 28 H
Tymochtee 9
Run A33
20 z 28 M
Buchanan 2
Run A35
28 z 35 H
Tymochtee 9
Run A36A
20 z 28 Canaan 3
Run A37
35 z 48 H
Tymochtee 9
Run A38
20 z 48 H
Tymochtee 11
Run A41
35 z 48 H
Rapid City
Linestone
Run A42
35 z 48 II
Tymochtee 11
Sulfided Feed
Ozldized Feed
Desulfurizer Sample, 7.0 cycles
Sulfided Feed
Ozidized Feed
Desulfurizer Sample, 12.7 cycles
Desulfurizer Sample, 3.7 cycles
Tymochtee 10
Sulfided-Feed
Oxidized Feed
Regenerator Sample, 10.7 cycles
Sulfided Feed
Ozidized Feed
Regenerator Sample, 11.5 cycles
Sulfided Feed
Ozidized Feed
Regenerator Sample, 13.3 cycles
Regenerator Sample, 20.0 cycles
Sulfided Feed
Oxidized Feed
Regenerator Sample, 9.2 cycles
Sulfided Feed
Oxidized Feed
Regenerator Sample, 11.9 cycles
Sulfided Feed
Regenerator Sample, 3.3 cycles
Sulfided Feed
Oxidized Feed
Regenerator Sample, 9.6 cycles
Properties
of Materials
From Hardening Study
Hoi % of Total Ca
CaCO3
' 85 . 2
85.5
49.5
87.9
87.3
28.8
56.2
68.8
69.8
10.9
86.0
81.1
38.4
72.6
73.3
9.1
23.5
70.8
73.1
10.9
51.3
51.1
9.6
6.9
7.4
69.2
70.4
10.7
CaS
4.4
0
37.9
0
0
60.3
31.2
22.6
3.1
78.2
12.8
0.9
55.1
11.6
0.2
77.2
75.3
18.0
0.2
81.6
37.9
3.5
73.6
85.7
83.4
15.2
O
70.9
CaSO4
6.2
.-;,»•
—
19.3
15.5
16.0
—
13.7
36.9
—
15.8
"CaO"
10.4
8.3
12.6
• 12.1
10.6
10.9
12.5
8.6
7.8
.. . 11.0
1.2
2.5
6.6
11.3
10.6
13.7
1.2
.. . H.2
13. 0
7.5
10.8
8.5
16.8
7.5
9.3
15.6
13.9
18.4
Density
grams
cc
84
87
72
83
82
63
1.76
86
99
70
2.08
2.19
1.95
1.89
1.99
1.71
1.86
1.98
1.71
1.72
2.005
1.67
1.83
1.89
1.81
1.68
Ibs
114.7
116.9
107.4
114.5
113.8
101.5
109.8
116.0
124.2
106.0
129.7
136.8
121.6
117.8
124.2
106.8
116.1
123.8
106.4
107.
125.
104,
114.3
117.8
112.8
105.1
Normallzed(1)
grams
cc
1.40
1.33
1.40
1.39
1.37
1.38
1.42
1.41
1.38
1.46
1.48
1.48
1.55
1.46
1.39
1.48
~
1.41
1.41
1.46
1.37
1.36
1.46
1.43
1.48
1.40
1.47
Density
Ibs
7f3
87.4
82.9
87.5
87.0
85.6
85.8
88.3
87.7
86.0
91.1
92.5
92.1
96.8
90.8
86.9
92.3
—
87.8
87.7
90.9
85.3
85.1
91.1
89.2
92.1
87.2
91.9
Hardness '2*
Test
Result
10.7
11.8
5.4
11.6
12.5
9.7
~
12.9
9.4
6.6
8.6
9.8
8.1
13.1
12.7
6.5
5.2
12.3
12.0
5.1
10.0
7.7
6.6
1.9
11.5
11.7
5.7
(1) Calculated by normalizing each calcium component to a CaO basis while assuming no change In sample volume.
(2) The results of the Increase in surface after tumbling in a rolling mill at a standard condition. 1.0 = 220 cmz/gm increment, the change observed In a Tymochtee
dolomite which had been hardened in the OOz Acceptor'Process cycle. Test conducted in Roalox tumbling Jar, Size 00, 10 Burundum grinding cylinders, size 13/16"
x 13/16" at 51 rpa for 30 minutes.
(3) The stone is reacted with hydrogen and the water evolved is weighed. For convenience, the material reduced was expressed as sulfate even though none was
believed to be present.
-------
TABLE IX-3 (continued)
Properties of Materials From Hardening Study
Penalty
normalized(1) Penalty
Source and Description of Dolomite
Run A43 Sulflded Feed
35 x 48 M Oxidized Feed
Buchanan 3 Regenerator Sample, 12.9 cycles
Run A44A Sulflded Feed
35 x 48 M Oxidized Feed
Buchanan 3 Regenerator Sample, 11.5 cycles
Run A4S Sulflded Feed
35 x 48 U Oxidized Feed
Tymochtee 11 Regenerator Sample, 15.5 cycles
Run A46 Sulflded Feed
35 x 48 M Oxidized Feed
Tymochtee 11 Regenerator Sample, 42.0 cycles
Run A47 Sulflded Feed
35 x 48 U Oxidized Feed
Buchanan 3 Regenerator Sample, 58 cycles
Mol % of Total Ca
CaCOa
84.3
86.4
23.1
65.2
65.9
13.2
74.5
69.9
11.5
67.2
66.8
8.1
63.1
63.3
10.9
CaS
13.4
0.3
76.9
34.8
1.6
83.4
16.9
0.9
70.2
14.4
1.3
77.5
34.8
0.8
84.3
CaSO4
12.4
32.2
18.4
13.2
33.0
"CaO"
2.4
0.2
0.0
0.0
0.3
3.4
8.6
10.8
18.3
13.4
18.7
14.3
2.1
1.0
4.8
grams
cc
2.01
2.14
1.82
1.98
2.23
1.84
1.87
1.96
1.69
1.84
1.95
1.70
1.96
2.23
1.99
Iba
123.3
133.3
113.3
123.3
140.3
114.6
116.6
122.3
103.7
114.3
121.6
106.3
122.4
139.1
124.1
cc
1.44
1.44
1.48
1.47
1.44
1.34
1.41
1.40
1.48
1.42
1.43
1.47
1.42
1.68
Iba
89.7
90.0
92.2
91.6
90.0
93.9
88.1
87.1
92.3
88.7
89.3
93.0
91.3
88.7
104.6
Hardneaa(2)
Teat
Reault
7.7
8.6
4.7
6.9
S.7
4.9
8.9
9.7
4.5
4.3
3.3
(1) Calculated by normalizing each calcium component to a CaO basis while assuming no 'change in sample volume.
(2) The results of the increase in surface after tumbling in a rolling mill at a standard condition. 1.0 = 220 ca2/g» increnent, the change observed in a Tyaochtee
dolomite which had been hardened in the CO2 Acceptor Process cycle. Test conducted in Roalox tumbling Jar, Size 00, 10 Burundua grinding cylinders, size 13/16"
x 13/16" at 51 rpa for 30 minutes.
-------
TABLE IX-4
Results of Drop Tests
Samples: 0.20 gm
Drop Height: 6.3 cm
Feed was 35 x 48 mesh
Product was screened at 48 mesh
Test Material
A2QA Final Regenerator Bed
A2 Final Regenerator Bed
HR1 Product
HR2 Product
HR3 Product
Average % Average
Undersized Material, Absolute
3 or More Drops Deviation
13.5 2.6
27.7 3.5
20.7 1.3
27.3 3.1
39.8 1.5
HR4 Half-Calcined
HR4 Sulfided
HR4 Oxidized
HR4 Reduced
35.0
32.8
26.2
31.6
1.4
1.2
2.5
1.1
HR5 to HR10 Composite
26.6
3.2
134.
-------
Hardening Run HR4 was made at the conditions shown in Tables IX-1 and IX-2.
The oxidation step was carried out at 982°C (1800°F) and lasted about nine hours
as compared to about 30 hours in Runs HR2 and HR3. Oxidation was followed by a
one hour reduction at the same conditions shown in Table IX-1, except that no
H2S was fed. In Run HR4 samples were taken in the course of the run so that
changes in the stone could be followed as processing continued.
Some agglomerates were formed during the run. Examination of the product
under a microscope showed a surface which appeared to have gone through a liquid
stage, i.e., the overall particle shape was retained but the surface had been
smoothed. Products of Runs HR1, HR2 and HR3 had a similar appearance.
As shown by the rolling mill test data in Table IX-3, HR4 product had
indeed been hardened. The drop test data of Table IX-4 do not show much
improvement in hardness. As the HR4 stone was later shown to actually be
attrition resistant, the drop hardness test was no longer used.
Table IX-3 contains data on stone composition and density as it passed
through the hardening process. The density data are presented on a normalized
basis which takes into account the effect of molecular weight changes on
density as discussed in Section VI. In the series HR1 to HR4, the lowest
density is associated with half-calcined stone. The products oxidized at 982°C
(1800°F) are most dense followed by the 927 and the 871°C products. In addi-
tion, the data for HR4 show a steady increase in density in going through the
transformations from half-calcined to sulfided to oxidized to the reduced state.
It is noted that the level of CaCO3 increased in going from the sulfided
to the oxidized state and again in going from the oxidized to the reduced form.
CaC03 can be formed from CaO and C02 following disproportionation of CaS04 and
CaS as well as from CaS by the action of H20 and C02.
Run A23
While the hardening step at 982°C (18OO°F) suffered from potential
operability problems, a feedstock for a continuous sulfiding-regeneration run
was made up in order to quickly determine whether the stone was actually
attrition resistant and to get feedback on the activity of CaC03 and CaS in
the gas desulfurizer and regenerator, respectively. Five batches, Runs HR5 to
HR10 were made at the same conditions as HR4. In four of these runs, there was
evidence of agglomeration. A total of 68O grams of +28 mesh agglomerates were
formed for about 80OO grams of free flowing -28 mesh stone.
Continuous run conditions were chosen to correspond to Run A21, except
for the use of less recycle gas in the regenerator for the first few cycles.
This was because the feed contained sufficient active CaS to possibly cause the
early regeneration to be equilibrium limited. Conditions and results from the
gas desulfurizer and regenerator are given in Tables IX-5 and IX-6, respectively.
Detailed data on solids composition, H2S content of dry exit gas, and attrition
as functions of cycle number are presented in Appendix B for all runs in this
chapter.
From O to 3.9 cycles, recycle to the regenerator was kept low (35
SCFH) in order to prevent the possible large initial regeneration of the heavily
CaS-laden stone from being limited by equilibrium. From 3.9 to 15.3 cycles, a
recycle flow of 102 SCFH was maintained in order to closely approach the design
value of 87$ approach to equilibrium.
135.
-------
TABLE g-5
Conditions and Besults for Gafl Desulfurlger with T?
chtee Dolomite
Run Number
Feed Stone
Feed Reference
Acceptor Size Consist, Tyler mesh
Feed Rate, gn/hr (ba If -calcined basis)
Nominal Solids Residence Time, min.
input,
Recycle to Bed
B,O
H,
CO
CO,
N,
H,S
Purges (CO,) to Bed
Purges (N,) above Bed
Recycle Acceptor Lift Gas, above Bed
Output in Cycle
Exit Gas Rate, SCFH (dry basis)
Composition. Hoi $
H,
,_. CO
LJ CO,
H,S
Outlet Gas. Top of Bed
Composition. Hoi
H,0
CO
CO,
»«
H,S
Flow Rate, SCFH, Top of Bed
Fluldlzlng Velocity, ft/sec
Attrition, % of Feed Rate
Duration of Circulation with H,S Feed, hr
Removal of Peed Sulfur, %
Removal of Peed Plus Recycle Sulfur, %
$ H,S in Outlet/Equilibrium % H,S
Conversion of Acceptor/Pass, nol £
Systea Pressure:
15 ata
(206 psig)
System Temperature: 871°C (16OO°F)
A23
BR5-BR1O
2280
41
6
142
15.8
18.1
9.O
57.1
O.O3
8.5
15.2
17.4
8.7
50.1
O.O3
304
0.43
O.35
31
97.2
94.6
1.3
13.1
A24
A24
2170
44
11
146
17.5
17.2
8.O
57.2
O.O3
8.1
17. 0
16.6
7.7
50.5
O.O3
3O4
0.43
O.77
45
97.3
94.7
1.4
13.8
A25
A25
225O
42
12
146
19.3
16.9
7.1
56.6
0.03
7.9
18.7
16.4
6.9
5O.O
O.O3
303
O.43
O.74
106
97.2
94.6
1.7
13.3
A26 . A27.
Tymochi.ee 1O
A26 A27
233O 232O
41 41
21 2O
143 147
19.8 18.8
17.2 17.4
8.0 8.0
54.8 55.8
O.O3 O.O3
8.1 8.O
19.2 18.2
16.6 16.9
7.7 7.7
48. 0 49.2
O.O3 O.O3
301 304
O. 43 O. 43
O. 6O O. 74
7O 67
97.6 97.6
95.3 95.3
1. 5 1.5
14.8 14.9
A28
A28
2250
42
6
144
16.9
16.5
10.4
56.2
O.O3
7.9
16.4
16. 0
1O. 1
49.5
O.O3
301
0.43
0.66
35
97.6
95.3
1.2
15.4
A29 A3O
A29 A30
2180 2310
43 41
4 1O
143 145
16.5 16.5
17. 1 17. 1
9.9 1O.O
56.5 56.4
O.O3 O.O3
8.1 8.1
16. O 16. O
16.6 16.5
9.6 9.6
49.7 49.7
O.O3 O.O3
3O1 3O3
0. 43 0. 43
0. 87 1. 44
25 38
97.6 97.6
95.3 95. 3
1.2 1.2
15.9 15.0
A32
A32
2O8O
46
GO
55
> 62
+ 21
10
325
20.3
18.5
8.3
52.9
o. is(*)
8.5
19. 0
17.3
7.8
47.3
0. 17
678
0.97
1.7l«'
34
68
52
8.4
1O. 4
A35
A35
2210
43
- 2OO
37.5
36
39
17 5
9
212
18.4
18.2
9.9
53.4
o.ii(')
9.0
17.4
17.2
9.3
47. 0
0. 10
436
O.62
0.75
49
87
78
4.0
12.6
A37.
A38
A42
A45
A46
A37
2310
39
21.4
25. 0
23 O
8 8
a
145
17.5
15.7
11. 0
55.8
O.O2
7.6
17. O
15.3
10.7
49.4
O.O2
301
O.43
0.78
33
98.4
96.8
0.8
13.6
A38
22 OO
41
60
55
62
21
3
326
17.5
18.3
9.2
55.0
O.O3
8.7
16.4
17.1
8.6
49.2
O.O3
681
0.97
0.99(«)
40
94.6
9O.O
1.2
14.3
A42 A45
237O
38
21.4
25
23
8 8
9
144
18.1
17.1
7.2
57.5
o.ioC
8.3
17.5
16.5
7.0
50.6
0. 1O
302
0.43
1.8
32
92
85
5.5
13.0
2160
41
21
27
23
g
15
149
20. 2
16.7
7.1
55.9
' 0.12 •
7.9
19.5
16.2
6.9
49.3
0. 12
308
O. 44
1.4
51
85.1
74.2
7.0
8.6
A46
2OOO
45
21
27
23
g
24
142
18.3
17.1
7.1
57.5
) 0.03
8.0
17.7
16.6
6.9
50.7
0.03
299
0.43
O.8
157
93; 5
87.5
1.7
5.6
!i) Input Is given after shift reaction and hydrolysis of CSa have taken place.
a) 1.2 SCFH for cycles 1-14.
!») Gas aaaple taken after H,S breakthrough.
4) Includes 2O-3O0 + 1OO mesh particles.
-------
Breakthrough of H2S in the gas desulfurizer occurred at 10.8 cycles.
At this point, it was decided to increase the solids circulation rate by about
80$. This reduced the absolute conversion of Ca per cycle, and should have
enabled the regenerator to lower the CaS level and thereby relieve breakthrough.
Upon increasing the circulation rate, the H2S content of the regenerator offgas
started to rise. The maximum observed was 4.36$ via the U.V. analyzer and 4.25$
via G.C. analysis (dry basis). This was substantially the equilibrium H2S con-
centration. Since the regeneration was now equilibrium limited, the recycle
flow was returned to its original level of 35 SCFH at 15.3 cycles in order to
lower the H2S concentration in the regenerator offgas.
After the circulation rate was increased, the H2S concentration in the
gas desulfurizer exit gas fell from 0.12$ to 0.06$. The regeneration activity
of the stone was then so low that breakthrough increased, and the run was shut-
down .
Figure IX-1 shows the CaS content of solids as the run progressed.
Figure IX-2 plots the regeneration activity of the solids. The raw activity
data do not fall on a single line. However, by displacing each point 3 cycles,
a single line can be drawn through the scattered points. It is reasonable that
some deactivation occurred due to the sulfidation, oxidation, and reduction which
were carried out prior to the cyclic run. The data indicate that the hardened
stone deactivated more rapidly than fresh Canaan dolomite. The rapid deactiva-
tion accounts for the lack of success in totally relieving breakthrough by
increasing the circulation rate. The stone just did not retain enough regenera-
tion activity to raise the CaC03 level enough to eliminate breakthrough.
The results of the run indicate.that increasing the circulation rate
is a good method of reducing breakthrough. It has potential as a means of con-
trolling sulfur absorption in a plant.
A highly significant result of A23 was the substantial attainment of
equilibrium in the regenerator.
It is difficult to pin down the exact percent approach to equilibrium.
The inaccuracies of the factors which go into such a calculation must be taken
into account. These are the following: H2S and C02 concentrations in the exit
gas, steam flow rate, exit gas flow rate, temperature at the top of the bed,
and the accuracy of the equilibrium constant itself. All things considered, the
observed value of 99$ is a close approximation of 1OO$ of equilibrium. It is
noted that the process design assumes only 88$ approach.
The attrition rate was quite low, 0.35$ of the circulation rate over
the entire run. The run demonstrated that the hardening procedure was effective
in producing an attrition resistant stone albeit at the cost of reduced activity.
The very low attrition rate indicated that the hardening procedure was overdone.
Run A24 was made at only 3O$ sulfidation in the hope that the attrition rate
would still be less than 1$ and that activity would be improved.
C. Continuously Hardened Tvmochtee 10 Dolomite
Run A24
The hardening conditions for Run A24 were 30$ sulfidation followed by
oxidation at 982°C (l80O°F). Each step was carried out continuously at the
137.
-------
Run Number
Nominal Solids Residence Tine, rain.
Input. SCFH
Recycle to Bed
H,O
H,
CO
CO,
Purges (CO,) to Bed
Purges (cOa)
above Bed
00
Output in Cycle
Exit Gas Rate, SCFH (dry basis)
Composition. Mol %
H,
CO
CO,
BaS
COS
Outlet Gas. Top of Bed
Compos It ion. Ool %
H,0
CO
CO,
N,
H,S
COS
Flow Rate, SCFH, Top of Bed
Fluidlzing Velocity, ft/sec
$ H,S in Outlet/Equilibrium £ H,S
Regeneration of Acceptor/Pass, mol %
TABU g-6
Conditions and Resulta for Regenerator with Tymochtee Dolomite
System Pressure; 15
System Temperature:
A23
Condition A
SI
35.0
5.O
7O.O
1
102
3.O
4.O
91.1
O.O
1.82
O.OS
30.1
2.4
3.2
62.9
O.O
1.43
O.06
175
0.48
0.38
16.9
A23
Condition B
28
102
3.O
5.O
14
35
4.1
4.8
84.6
2.1
4.36
0.19
29.9
3.2
3.7
58. 0
1.6
3.43
0.15
174
O. 48
O.99
7.7
A24
50
11
97
3.2
2.9
91.3
1.5
1.43
O.OS
31.1
2.5
2.3
62.2
1.2
1.11
O.O4
17O
O.47
a 29
12.4
A25
48
12
98
2.1
3.O
92.9
0.8
1.14
0.05
3O.6
1.7
2.3
63.8
0.64
O.89
O.O4
17O
O.47
a 23
9.7
A26
46
21
102
3.2
2.2
92.8
0.8
1.24
O.O4
30.1
3.5
1.7
64.2
0.6
O.97
O.O3
174
0.48
0.2S
10.5
atm (2O6 psig)
7O4*C (13OO°F)
A27
47
35 O
5 O
2O
1O4
3.0
2.2 .
91.5
1.8
1.34
O.O6
28.4
2.4
1.8
64.8
1.4
1.O8
O.OS
173
0.48
O.29
11.7
A28
48
6
1O1
3.O
4.O
91.3
O.O
1.60
O.O7
30.4
2.4
3.2
62.9
O.O
1.25
O.OS
174
O. 48
O. 33
14. O
A29
SO
4
102
3.0
4.1
91.3
O.O
1.63
O.O7
29.1
2.4
3.3
64. O
0.0
1.30
O.O6
172
0.48
a 35
14.9
A30
41
1O
1OO
3.O
4.1
91.3
O.O
1.6
O. O7
29. 0
2.4
3.3
64.1
O.O
1.25
O.06
169
0.47
a 34
13.2
A32
52
0.0
1O8
11
2O5
10
238
2.1
2.6
93.5
1.3
O. 62
.03
33. 0
1.5
1.8
62.3
as
O. 44
a 02
333
a 92
an
12.4
A3S
49
O.O
79. 0
8.O
14O
9
168
3.3
1.4
93.9
O.7
0.68
O.O4
33.2
2.4
1.0
62.3
0.5
0.5O
O.O3
229
a 63
a 12
9.1
A37
47
35.0
51.0
5.0
7O.O
8
1OS
2.9
4.2
91.7
0.0
1.12
O.OS
28.4
2.4
3.4
64.9
O.O
O.9O
O.O4
175
O. 48
0.24
8.9
A38
49
O.O
1O8
11
205
3
239
2.0
2.6
03.1
1.7
O.58
O.O2
32.7
1.5
1.9
62.4
1.2
0.42
0.02
333
O.92
0. 1O
11.5
A 42
46
35.0
51.0
5.0
7O.O
9
1O4
3.4l('
2.0
90. 5
3.1
1.28
O.OS
29.0
2.7
1.6
63.4
2.5
1.02
O.O6
174
0.48
0.28
1O.3
A45
50
35.0
Sl.O
5.O
7O.O
15
1O8
' 3.4
1.6
93.3
O.6
1.24
O.OS
28.6
2.7
1.3
66. 0
O.S
O.99
O.O4
177
0.49
O.26
11.1
A46
54
35. 0
Sl.O
5.0
70.0
24
108
2.8
1.5
94.6
O.S
O. 7O
O.O4
29.1
2.2
a. 2
66.6
0.4
O.S5
O.O3
181
O.SO
O. 14
6.9
(i) These data were corrected for air contamination of the gas sample.
-------
2O X 20 TO THE INCH 46 1242
7 X 10 INCHES «>UF IN U ', «.
!EL I CO.
-------
FIGURE IX-2
Run A23 - Deactivation of CaS at 7O4°C (l3OO°F)
1.5
2.5 3
5 6 7 8 9 10
Number of Cycles
1.5
2 2.5 3
5 6789
140.
-------
conditions shown in Tables IX-7 and IX-8. Reduction of CaS04 was carried out
simultaneously with the first sulfidation of the run's cyclic phase. There
was no evidence of agglomeration or operability problems in the hardening step.
For the cyclic part of the run, a modest recycle flow of 35 SCFH was
used in the regenerator since the feasibility of approaching equilibrium had
already been demonstrated. Analysis of solids showed that over 9O$ of the
CaS04 was reduced in the gas desulfurizer in the first cycle. The CaS04 content
was zero within a few cycles. This result is consistent with the earlier batch
data obtained when the reduction was run separately.
Regeneration activity was better than in Run A23 and, after 12 cycles,
the CaS content of solids leaving the gas desulfurizer was still under 70 mol $.
In order to examine the behavior of the stone at breakthrough, the CS2 (H2S)
feed rate was doubled at cycle 12.4 until breakthrough at cycle 14.5. The CS2
was then reduced to its original value of ca. 13 mol $ per cycle. When the CS2
was cut back, the H2S concentration in the exit fell from 0.24$ to 0.04$. After
about an hour, it rose to 0.08$. Unfortunately, the inventory of acceptor was
too low to allow an increase in circulation rate and the run was shutdown. The
progress of the run as defined by the CaS content of the solids is presented in
Figure IX-3.
The attrition rate for Run A24 was 0.77$, well within the goal of 1$
or less. The regeneration activity was better than in Run A23. The data plotted
in Figure IX-4 indicate that activity was nearly the same as for Canaan dolomite.
Thus, a hardened acceptor was produced with no penalty in activity.
Run A25
The hardening conditions for Run A25 were 30$ sulfidation followed by
oxidation at 927°C (l700°F). As in Run A24, the hardening was carried out con-
tinuously and reduction took place during the run.
Regeneration activity was better than in Run A24 and, after 12 cycles,
the CaS content of solids leaving the gas desulfurizer was under 6O$. In order
to examine the behavior of the stone at breakthrough, the CS2 feed rate was
doubled from cycle 12.4 until breakthrough at cycle 16.1. The CS2 pump setting
was then cut back to its original value, and the H2S fell from 0.12$ to 0.03$
at the gas desulfurizer exit, A second breakthrough did not occur and, after
making a sulfur balance, it was concluded that the pump malfunctioned so that
the molar CSa feed was on the order of 5-6$ rather than 13$ of the calcium
circulation rate. The run was continued to 36 cycles to gain appreciation of
the long term regeneration behavior.
The CaS content of the solids versus cycles is shown in Figure IX-5.
The regeneration activity shown in Figure IX-6 was slightly better than that of
Canaan dolomite. The attrition rate was an acceptable 0.74$ of the circulation
rate.
After Run A25, a calibrated CS2 reservoir was installed so that the
actual delivered volume could be continually monitored.
141.
-------
TABLE Q-7
Run Number
Acceptor Size Consist, Tyler mesh
Feed Rate, gm/br (raw basis)
Nominal Solids Residence Time, mln.
Input.
Recycle to Bed
H,0
H,
CO
CO,
N,
H,S
Purges (CO,) to Bed
Purges (N,) above Bed
Recycle Acceptor Lift Gas, above Bed
Conditions for Continuous Sulf idation of Tymochtee Dolomite
System Pressure: 15 atm (2O6 psig)
System Temperature; 871°C (16OO°F)
A24
35 x 48
1330
94
148
21.4
25.0
23.0
8.8
66.0
1.6
5.0
15.0
71
A25
A26
A27
A28
A29
A3O
133O
94
1400
81
136O
81
2110
57
379O
43
378O
42
ca O.8
0.0
(i) Input is given after shift reaction and hydrolysis of CS, have taken place.
NJ Run Number
•
Acceptor Size Consist, Tyler mesh
Feed Rate, gm/nr (raw basis)
Nominal Solids Residence Time, min.
Input. SCFH(3)
Recycle to Bed
H,O
H,
CO
CO,
N,
H,S
Purges (CO,) to Bed
Purges (N,) above Bed
Recycle Acceptor Lift Gas, above Bed
A32
1470
108
310
60
55
62
21
152
110
A35
237O
69
200
37.5
36
39
17.5
98
A37
2280
68
148
21.4
25.0
23. 0
8.8
66
A38
2O x 28
182O
80
310
6O
55
62
21
152
A42
2210
65
148
21.4
25. 0
23.0
8.8
66
^
A45
2280
es(O
135
19. 0
23. 0
21.O
8.0
58
1 9
A46
2310
6O
148
21.4
25. 0
23.0
8.8
66
1 8
Sulf ided at 15OO°F.
Input is given after shift reaction and hydrolysis of CS, have taken place.
-------
TABLE IX-8
Conditions for Continuous Oxidation of Sulf ided Tymochtee Dolomite
System Pressure; 15 atrn (206 psig)
Run Number
Feed Rate, gm/hr (adjusted to mainatln
Oa breakthrough)
Nominal Solids Residence Time, min.
Temperature, °C
Input. SCFH
A24
A25
A26
A27
A28
A29
A30
A32
A35
A37
A38
A42
A45
A46
Recycle to Bed
C02
N2
Air
Purges (C02) to Bed
Purges (N2) above Bed
Recycle Acceptor Lift Gas
938
98
982 (1800°F) 927 (1700°F)
1210 =
101 ••'
899 (1650°F)
870 880
132 140
871 (16OO°F)
2910
41
2970
40
12 6O
9O
1O7O
1O5
1370
83
553
176
1000
106
827
132
1OOO
1O4
138
49
90
16
5
15
71
340
14O
165
24O
56
98
138
49
90
340 138
140 49
165 9O
110
-------
2O X 2O TO THE INCH 46 1242
7 x 10 INCHES **or it u -: »
KEUFFEL 0. ESSER CO
C
-------
FIGURE IX-4
Run A24 - Deactivation of
CaS at 704° C (13OO°F)
Mo Is CaCO, Produced
TOCTMoTgnCmS- Fed
^ H 9 10
Number of Cycles
145.
-------
2O X 2O TO THE INCH 46 1242
7 X 1O INCHES *»&[ in u •: A.
KfUFFEL a cssnn co.
A2KI-H MoJM* CaS irl
J
-------
Run A26
The feed for Run A26 was prepared via sulfidation to 25 mol $ CaS at
871°C (1600°F) followed by oxidation at 895°C (1650°F). The attrition rate of
the cyclic run was 0.6$ of the feed rate, and the deactivation behavior of the
stone was the same as for Run A25. The run was terminated when breakthrough
occurred. About 4 mol $ of CaC03 remained in the stone at the time of break-
through. The run data are presented graphically in Figures IX-7 and IX-6.
Run A27
The feed for Run A27 was prepared via sulfidation to 30 mol % CaS at
871°C (1600°F) followed by oxidation at the same temperature. The attrition
rate remained low, 0.7$, and the activity was satisfactory. The run data are
summarized in Figures IX-8 and IX-6. The run was continued until breakthrough
occurred at which time the CaC03 content of the stone was ca. 1-2$.
Run A28
The temperature of 871°C (1600°F) was compatible with other process
flows, and it was felt that there was little to be gained by lowering the
temperature of the hardening process any further. Consequently, reduction of
the initial sulfiding level was explored. For the Run A28 hardening step, both
the sulfidation and oxidation were carried out at 871°C (1600°F) at a molar
conversion of 19$ of the calcium.
The attrition rate of the cyclic phase was 0.7$, and the activity was
normal. The run was terminated after 10 cycles; the concentrations of CaS
during the run are given in Figure IX-9. ;
Figure IX-6 represents all the regeneration data for runs in which
the feed was hardened over the temperature range of 871 to 927°C (1600-17OO°F)
and 2O-30$ sulfidation. The results of the four runs are consistent, and
display an activity slightly better than that of Canaan dolomite. The attri-
tion rate was less than 1$ for all of the above runs.
Run A29
The preparation of Run A29 feed was run at 4.4$ sulfidation. This is
the same level which was used in Run A4(1) when oxidation of 4 mol $ of the
calcium was performed at 704°C (1300°F) in each cycle. No hardening was
observed in Run A4. The attrition rate observed in the cyclic phase of Run A29
was 1.1$ for cycles 0-2.2. By the period 4.5-7.3 cycles, the attrition rate
had fallen to 0.7$. As shown in Figure IX-11, the regeneration behavior was
better than that of hardened Tymochtee 10 dolomite. The run was ended after
7 cycles (Figure IX-10).
The data indicate that 4$ sulfidation followed by oxidation is not
sufficient to produce a feedstock which will immediately have an attrition rate
of less than 1$. However, the behavior of the A29 feed was curious in that the
attrition rate steadily decreased. Therefore, a run was made in which the
hardening operation lacked the addition of H2S in the first step.
147.
-------
FIGURE IX-6
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4
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2,5
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1.5
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7
N 6-
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Deactivation of Hardened Tymochtee 1O Dolomite
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Symbok I
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- -•
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y
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.. - )-:-•
. I ....
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:-.-:-
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i • •
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r "
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-
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'-i •- -
L:--i
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-
E :.-
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-^:
^E=
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.jT _"
_:2L
.~rr
-HE
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ii
rr:
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— rr~
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.
—
—
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=5
_ ~-
^
==
J 1,5 2 2.5 3 i 3 6 7 8 9 10 1.3 2 2.5 3 + 5678
S?
•S
£5
-^d
^
^^
:~
::
X = Number of Cycles
148.
-------
N _ . 3 I B ^ ~ .,
VO
BreaJtthroujh
— (100 -
\— •-•*
lias Deijsulfurizer i
T"
-------
01
o
-------
- -------- -i
FIGURE IX-9
-J-—
. . .........
IJUN A2EJ - MQll # Catf Ity EXIT SOLIDS
(1QO r nidi $ CaO)
FIGURE IX-10 —-1
i .... , , i
j A20 - MOL i Cafit IN EXIT SiJLlDS
Qn ; , _
w~ ff=-~ -)_ _ —,_t _ J-_ — — r-ir
! i
"ncjl % Cfijo]~~~r
• •— t- ; 1- r-
-------
FIGURE IX-11
Deactivation of Tymochtee 1O Dolomite
at 704° C (1300°F)
8 « 10
Cycle Number
152.
-------
Run A30
Run A30 was a blank hardening run, i.e., all the process steps of
"sulfidation" and oxidation of Tymochtee dolomite were carried out but no H2S
was fed to the hardening process. The purpose of the run was to check the
possibility that hardening involved reaction of impurities in the stone with
CaS or other forms of Ca. The original hypothesis was that a low melting
CaC03-CaS04 transient liquid was involved.
The initial attrition rate was 2.9$ but this fell to l.O$ by cycle
13. By comparison, Tymochtee dolomite which had been hardened via a 20$
sulfiding-oxidation procedure showed an attrition rate of 0.7% from the first
cycle. It is concluded that, to be immediately effective as a hardening step,
oxidation must be carried out on a sulfided stone. Oxidation per se is bene-
ficial in that hardening apparently takes place as the acceptor is cycled and
the CaS concentration increases. However, such secondary hardening is slow.
The data indicate that hardening involves reaction of CaS or other forms of Ca
with oxidized impurities in the acceptor. With a very impure stone, such a
phenomenon could possibly lead to agglomerate formation as is suspected to have
occurred with limestone 1691 during Run A13.C1)
As in Run A29, the regeneration activity was superior to that of the
hardened Tymochtee 10 stone. It would appear that hardening does cause a loss
of activity even at the relatively mild conditions employed in Runs A26 to
A28. The run data are shown in Figure IX-12, and the regeneration data are
given in Figure IX-11.
Run A30A
Run A30A was made to check the attrition rate of untreated Tymochtee
dolomite. The.run.was initiated by substituting sufficient untreated stone
into the feed hoppers to displace the A30 inventory. The unit was not shut-
down, and conditions for A3OA were the same as for A30. An attrition rate of
2.0$ of the feed rate was observed, and Run A30A was continued only long
enough to check the attrition rate.
The attrition rate of 2.0$ was lower than had been expected since
previous runs under similar conditions had shown attrition rates 2.5, 5.6 and
4.9$. The fact that the current work was done with a different shipment of
stone, albeit from the same quarry, may have been the reason for the difference,
CaO Formation
The data for Run A30 in Table IX-3 show a 12$ loss of CaC03 just on
calcination at process conditions. It must be concluded that, at run condi-
tions, reaction occurs between CaC03 and impurities in the stone which uses up
ca. 12$ of the calcium. "CaO" formation thus is actually loss of CaC03 by
reaction with impurities. For this to be true, the reaction must take place
to an extent beyond that seen in the assay of the fresh stone. In the fresh
stone assay, the solid is exposed to 1 atm C02 for 15 minutes at 1550°F (844°C)
As the run temperature is only 16OO°F (871°C), the additional loss observed in
the continuous unit was not anticipated.
It is noted that in the previous work with Tymochtee dolomite, which
was performed under Contract EHSD-71-15,(1) the CaS content of sulfided stone
was determined by difference based on the measured CaC03 content. If loss of
153.
-------
Mbl: % :CaS in Exit Solids
l|gfjMpL
!Number of Cycles
154.
-------
CaC03 to a nonsulfided form occurred at run conditions, the calculated initial
sulfide content would be high by ca. 12 mol $. Review of the data show this
to be the case. The high calculated sulfide levels had originally been ascribed
to irregularities in the startup of the liquid H2S feed system.
The use of separate analyses for CaC03 and CaS was begun late in the
previous contract period (EHSD-71-15). At that time, Canaan dolomite, a pure
stone, was being fed, and negligible discrepancies appeared when the CaCO3 and
CaS contents were summed.
As it was not practical to run a pressurized high-temperature assay
in the lab, the present assay was retained. For consistency, data on calcium
species is reported as it was since the beginning of this program with the
understanding that "CaO" formation was really loss of CaC03 by reaction with
impurities.
Clarification of the mechanism involved in CaC03 loss and in harden-
ing would be aided by an intense physical and chemical examination of solids.
Extensive work of this nature is not planned. However, the samples are being
saved, and selected materials can be given over to in-depth study later.
D. Tymochtee No. 9 Dolomite
Various batches of Tymochtee dolomite had given uniform results in the
C02 Acceptor Process.l1) The Tymochtee stone used for the work done under
Contract EHSD 71-ls(2) was the ninth batch received from the C. E. Duff quarry
and was labeled Tymochtee 9. This was the feed for Runs Al, A2, A4, A18, A32
A35 and A37. Tymochtee 10 was received early in the reported contract period
and was the feed for Runs A23-A30. The supply of Tymochtee 10 ran out, and
Tymochtee 9 stone was used as the feed in a series of runs designed to evalu-
ate the effect of particle size.
Run A32
The feed to Run A32 was 20 x 28 Tyler mesh Tymochtee 9 dolomite which
had been hardened via sulfidation and oxidation as detailed in Tables IX-7 and
IX-8. Upon cycling, two unexpected results were observed: regeneration acti-
vity was lower and the attrition rate was higher, i.e., 1.7$ versus 0.7$
attrition observed earlier with 35 x 48 mesh stones. Activity of the acceptor
in the gas desulfurizer was normal. The run data are given in Figure IX-13, and
the regeneration activity data are presented in Figure IX-14. The percent
regeneration was 14$ in cycle 3 and 11$ in cycle 10. The activity of 35 x 48
mesh hardened Tymochtee 10 dolomite was 39$ in cycle 3 and 21$ in cycle 10. The
run was shut down upon H2S breakthrough in the gas desulfurizer during cycle 11.
The attrition rate was constant during the run at 1.7$. Analysis of
collected overhead fines (Table IX-9) shows that 20-40$ was +100 mesh where
previously only about 10$ +1OO mesh stone appeared in the collected fines. Even
after factoring out the contribution of coarser overhead stone, the attrition
rate exceeded 1$. Presumably the +1OO mesh overheads had been elutriated due
to the high fluidizing velocity and the lack of freeboard for proper disengagement
Run A35
The changes in attrition rate and activity which occurred in.going from
35 x 48 mesh to 2O x 28 mesh hardened Tymochtee dolomite seemed anomalously
155.
-------
<^C 10 X 10 TO ' INCH - X V. INCHES
IL. KEUFFEL a ES5ER CO "»DE K u 5 *
46 1320
-------
FIGURE IX-14
DEACTIVATION OF CaS IN HARDENED TYMOCHTEE 9
DOLOMITE AT 7O4°C (13OO°F)
1.5
2.5 3
5 6 7 8 9 10
Number of Cycles
1.5
2.5 3
5 6 7 8 9 10
157.
-------
TABLE IX-9
Particle Size Consist of Overhead Fines
Run Number
Material^
i .
A25
Shutdown
A26_
Shutdown
A27.
18.4
Cycles
A27
18.8
Cycles
A27
19.6
Cycles
A28
8.O
Cycles
8.6
Cycles
A28
9.2
Cycles
A29 A30_
Shutdown Shutdown
1.8
Cycles
A32
4 .2
Cycles
4.6
Cycles
A33
l.O
Cycles
A33
3.2
Cycles
A33
5.6
Cycles
A33
8.6
Cycles
Wt * Tvler Mesh
35
48
65
1OO
150
200
+35
x 48
x 65
x 1OO
x ISO
x 2OO
x 325
-325
.5
.5
.5
16.8
29.7
9.9
42.l(2)
O
.3
O
6.2
23.9
17 .O
52. e(2)
O
0
.4
1O.4
17.2
11.2
11.2
49.6
O
0
O
7.3
15.6
16.1
18.7
42.3
O
O
O
1O.5
19.4
13.4
12.1
44.6
O
O
0.8
7.3
19. 0
13.0
12.1
47.8
O
0
.8
7.2
20. 1
14.1
13.7
44.1
0
O
.4
7.2
21.2
15.6
14.8
4O.8
.4 .1
O O
O .1
8.1 3.6
20.0 9.9
15.2 18.2
56. 3(2) 68.l(2'
O
.8
5.6
15.3
13.3
14.5
20.5
3O.O
0
.8
13.3
22.7
13.3
13.3
14.5
22.1
O
8.9
25.3
17.8
9.3
8.1
11.3
19.4
O
0
4.4
28. 0
23.2
18.0
13.6
12.8
0
O
4 .0
32 .0
25.6
18.4
10.8
9.2
O
O
6.1
32.3
24 .O
14.4
12.9
1O.3
O
0
7.2
3O.O
24.8
17.6
11.6
8.8
00
(i) Cycle numbers given are for the midpoint of the collection period of O.4 to O.6 cycles in duration.
(2) Is -200 mesh.
-------
large. A hardened 28 x 35 mesh stone was run in A35 to better define the effect.
The flow rate of lift gas was kept the same as for the 20 x 28 mesh case.
The attrition rate for the 28 x 35 mesh hardened Tymochtee 9 dolomite
was 0.8%, the same as for the 35 x 48 mesh size Tymochtee 10. This result
eliminated the possibility that higher lift gas rate per se was the cause of
increased attrition of the 20 x 28 mesh size.
As given in Figure IX-14, regeneration activity was still signifi-
cantly less than for 35 x 48 mesh Tymochtee 10 stone. Run data are given in
Figure IX-15. '
Run A37
Run A37 was made in order to obtain more data on the behavior of
hardened Tymochtee 9 dolomite. It was thought that low activity obtained with
28 x 35 and 2O x 28 mesh stone in Runs A35 and A32, respectively, could be due
either to a size effect interacting with the hardening procedure or to intrinsic
differences in the stone.
Run A37 was made with hardened 35 x 48 mesh Tymochtee 9 dolomite. As
shown in Figure IX-14, the regeneration activity was again low. The regenera-
tion data is about the same for all three sizes tested, 20 x 28, 28 x 35 and
35 x 48 Tymochtee 9 stone. Run data are given in Figure IX-16.
E. Tymochtee No. 11 Dolomite
Several runs were also made with Tymochtee 11 dolomite. The regeneration
activity of this stone was very similar to that of the Tymochtee 9 stone as can
be seen from Figure IX-17. Run data are presented in Figures IX-18 to IX-21.
Run A38
Run A38 was made with Tymochtee 11 dolomite using 20 x 28 mesh as the
size consist. The regeneration activity was about the same as that of Tymochtee
9. Due to an operating error, sulfidation in the hardening step was about 40$
instead of the desired 20$.
Run A42
This run was made with 35 x 48 mesh Tymochtee 11 stone. As shown in
Figure IX-17, there was no effect of changing the size consist upon regeneration
activity. This confirmed the conclusion drawn from the Tymochtee 9 data that
there is no effect of particle size upon regeneration activity in going from
20 x 28 to 35 x 48 Tyler mesh stone.
Run A45
Examination of sectioned sulfided stone had shown that, initially, the
sulfur is predominantly distributed around the outside of the particle in a shell,
The hardening procedure could conceivably close off the interior of the particle
and thereby reduce the stone's activity. For Run A45, the sulfidation step in
the hardening was conducted at'816°C (150O°F) instead of the usual 871°C. The
lower temperature would increase the rate of H2S diffusion into the particle in
relation to the rate of reaction. Upon cycling, there was no significant change
in activity, and the attrition rate of 1.4$ was slightly higher than usual.
159.
-------
46 1320
FIGURE IX-15
_ . ^Q.
Rn 5 -T-
un AJ
_
Brte akthriougfe
t
!(IOO -i Mol $ CaO) |
* r ~ - ' re "
: . -j .... . __
-------
I TO
:L a
H 7
:O. M
'2
-------
FIGURE IX-17
Deactivation of Tymochtee 11
Dolomite at 7O4°C ( 130O° F)
i r ———1— ——
: _LJ_L
Mo Is CaCO3 Produced
lOO-Mo-ls-eaS-Fed
1
x!0°
i ' '
1
i
i;
• i
i i
• i
!
xicr
1,3
? 8 'J
8 * 0
X = Cycle Number
162.
-------
-^ TO 17 HE'S
.. K, — ft I _o. MA... ,, „,„
4L
to
MHl'llilimuiml
FIGURE IX-18
MOL;% CaS IN EXIT
" -iiiiLLiLliiiLuJLU
-------
FIGURE IX-19
ro
r—I
10
xg
OU
xl
1
: RUN A42 -tMOL jo CaS IN EXIT SOLIDS ~:'
164.
-------
Run A46
This run was a demonstration of the behavior of hardened Tymochtee 11
dolomite for 42 cycles. A larger initial stone inventory was prepared so that
depletion due to attrition would not force the run to end. An initial sulfida-
tion conversion per pass of 12% helped to defer breakthrough until cycle 14.
At that time the H2S feed rate was cut to ca. 6% and no breakthrough occurred
until the run was voluntarily terminated.
The purpose of this run was to get the experimental data on a large
number of cycles, so the process design will rely less on extrapolation. How-
ever, the correlation in Figure IX-17 uses data from Run A46 only through 14
cycles.
When the H2S feed was reduced, a drop in percent regeneration was
observed in Run A46. The explanation for this follows . Consider a situation
in which the stone has a 10$ regeneration capacity and is being sulfided by
10$ per cycle, e.g., entering the gas desulfurizer the CaS:CaC03 ratio is
90:1O. Leaving the gas desulfurizer and regenerator the ratios are 100;0 and
90:10, respectively. If the H2S input is reduced to 5$ conversion, the ratio
leaving the gas desulfurizer will be 95;5. Two possibilities exist for the
regenerator. The CaS level may be reduced by 10$ relative to 85.5$ or, once
having been formed, the CaS is resistant to regeneration and the CaS content
cannot be reduced much below 90$. In the latter case, the observed percent
regeneration would be 5$. Our experience has been that the second case is what
actually occurs. In Run A46, after the H2S input was reduced, the regeneration
fell and remained nearly constant, and the CaS content of stone leaving the
regenerator was not significantly lowered.
F. Buchanan Dolomite
Four runs were made in which hardened Buchanan dolomite was the feedstock.
The hardening conditions are detailed in Tables IX-10 and IX-11. These were
essentially the same as had been run on the Tymochtee stone. Detailed condi-
tions for the gas desulfurizer and regenerator are given in Tables IX-12 and
IX-13, respectively. Graphical data are presented in Figures IX-22 to IX-26.
Run A33
The feed to Run A33 was 20 x 28 mesh Buchanan dolomite which had been
hardened via sulfidation followed by oxidation of ca. 15$ of the calcium. The
regeneration activity was 30$ in cycle 3 and 17$ in cycle 10 (Figure IX-22).
The attrition rate was initially 1.8$ and rose to 6.4$ in cycle 11. The latter
figure is about the same as for Buchanan stone which had not been hardened. As
in Run A32, some 35$ of the overhead solids were 4-1OO mesh. The high attrition
rate caused the run to shut down due to depletion of the stone inventory.
Run A43
Run A43 was a repeat of A33, except for the use of 35 x 48 mesh feed.
The regeneration behavior was essentially the same again confirming that there
is no effect of particle size on this variable. The attrition rate was 2.1$.
165.
-------
CNJ
CM
10
FIGURE IX-20
RUN A45 -: MOL * CaS IN EXIT SOLIDS :: :!:. :.!::2
EEtEETl -fel.:::.. ~: : t:.!-:/ :.: fc: i E
--Break -Through
(-lOO-t—M6-1—&
Desulfuriz
Regenerator
2
6 ;:;8 10 12 14
16
18
: : . :i . :
• - • : .•:!;•
... ., .
::.:.j--::
: :: j :
. . ;
!
. . :j : .
r - - -
" " : : : I
:-.; :
::>!""'
' :::
. . j ...
......
; ; : : i. ' ' :
r|~TT "
i ;
::;;:::":
_ ..._- „
1
: Cycle Number
. . . l
j j . ...:-.
T" ! . - ." ! ". . " i " - '
...... | . : . -
_J_ i.
1 ......
'"TtTTTl" ": '^ |~i
. ... , .....
1, .
-.'.
".
.). . . .. . , .. ; ........ r :
'.i ' ;!: :;r;;;]::--:; ::i ' ; ; • !:" !
i • .... j . • . • I. - . | '.:...; : :
:•..!::..:: .: ;:.-:: :i -:T: j. :: ': [•••.;
. . . 1 ... """" . i . ' . i . . ; . . .:"•'. . : • ~r\ ~ ~
: i ..;.... i ....;... 1 .... 1
1 " | • '
::.[-...": ":;!;;::i;:::i.;: . \ \ .' •, ' :
1 ... ; - . . ( i ... . . : . :
. . J . . 1 . . -1 . . j . .... . .
...(.... j . . i .
. j . . j ... :
I - 1 - --. -- . . . • . i . . i ......
- ; 1 •-;;;;;;;•;;' j - ; i • ; - - ; -
• ' i • ' . !! :::": j :.:'•
166.
-------
01
20 ;
iCycle Number
-------
TABLE IX-10
Conditions for Continuous Sulfidation of Buchanan Dolomite
System Pressure
System Temperature:
Run Number
15
871
A33
Acceptor Buchanan 2
Acceptor Size Consist, Tyler mesh 20 x 28
Feed Rate, gm/hr (raw basis)
Nominal Solids Residence Time, rain
Input, SCFH
Recycle to Bed
H2b
H2
CO
Cp2
N2
H2S
Purges (C02) to Bed
Purges (N2) above Bed
Recycle Acceptor Lift Gas. above Bed
1770
99
310
60
55
62
21.0
152
1.8
5
15
110
atm (206 psig)
°C (1600°F)
A43
Buchanan 3
35 x 48
2090
72
148
21.4
25.0
23.0
8.8
66
1.8
5
15
110
A44A
Buchanan 3
35 x 48
1028
158
148
21.4
25.0
23.0
8.8
66
1.8
5
15
110
A47
Buchanan 3
35 x 48
1129
165
148
21.4
25.0
23.0
8.8
66
1.8
5
15
110
Mol % CaS in Product Solids
13
13
35
35
TABLE IX-11
Conditions for Continuous Oxidation of Sulfided Buchanan Dolomite
System Pressure:
System Temperature:
15 atm (206 psig)
871°C (1600°F)
Run Number
Feed Rate, gm/hr (adjusted to
maintain 02 breakthrough)
Nominal Solids Residence Time, min
Input, SCFH
Recycle to Bed
C02
N2
Air
Purges (CO2) to Bed
Purges (N2) above Bed
Recycle Acceptor Lift Gas
A33
A43
A44A
A47
1200
103
340
140
165
16
5
15
110
1125
82
138
49
90
16
5
15
110
541
233
138
49
90
16
5
15
110
610
198
138
49
90
16
5
15
110
168.
-------
TABLE IX-12
Conditions and Results for Gas Desulfurizer with Buchanan Dolomite
System Pressure:
System Temperature
Run Number
15 atm
: 871°
A33
Type of Dolomite Buchanan
Feed Reference
Acceptor Size Consist, Tyler Mesh
Nominal Feed Rate, gm/hr (half-
calcined basis)
Nominal Solids Residence Time, min
Input, SCFH
Recycle to Bed
H20
H2
CO
C02
No
H2S
Purges (C02) to Bed
Purges (N2) above Bed
Recycle Acceptor Lift Gas, above Bed
Output in Cycle
Exit Gas Rate, SCFH (dry basis)
Composition, Mol %
H2
CO
C02
N2
Outlet Gas , TOJJ of Bed
Composition, Mol %
H20
H2
CO
C02
N2
H2S
Flow Rate, SCFH, Top of Bed
Fluidizing Velocity, ft/sec
Attrition, % of Feed Rate
Duration of Circulation with H2S Feed, hr
Removal of Feed Plus Recycle Sulfur, $
% H2S in Outlet/Equilibrium % H2S
Conversion of Acceptor/Pass, mol $
A33
20 x 28
1870
50
310
60
55
62
21
152
1.8
5
15
110
11
316
19.4
18.5
8.6
53.6
0.03
9.0
18.1
17.2
8.0
47.6
0.03
672
0.96
3.9(
37
90
1.3
15.5
(206 psig)
C (1600°F)
A43
2 Buchanan 3
A43
35 x 48
2100
42
148
21.4
25
23
8.8
66
1.8
5
15
110
16
143
16.4
17.2
9.8
56.5
0.12(0
8.1
15.9
16.7
9.5
49.7
0.12
301
0.43
2) 2.1
53
83
5.0
13.1
A44A
Buchanan 3
A44A
35 x 48
2030
44
148
21
26
23
9
66
1.8
5
: 15
110 .
7
147
18.4
17.2
7.4
56.9
0.03
8.0
17.9
16.6
7.2
50.3
0.03
304
0.43
0.8
43
95.3
1.7
14.6
A47
Buchanan 3
A47
35 x 48
1945
46
142
21
27
22
9.3
66 ,
0.64(3)
5
15
110
51
144
20.8
16.7
.7.4
55.1
0.01
8.1
20.2
16.2
7.1
48.4
0.01
301
0.43
0.68
233
96
0.6
5.4
(i) Gas sample taken after H2S breakthrough.
(2) Includes 2O-30# +1OO mesh particles.
(3) For cycles O to 32.2 the H2S feed rate was 1.2 SCFH.
169.
-------
TABLE IX-13
Conditions and Results for
Regenerator
System Pressure: 15 atm
with Buchanan
(206 psig)
Dolomite
System Temperature: 704°C (1300°F)
Run Number
'Nominal Solids Residence Time, min
Input, SCFH
Recycle to Bed
H20
H2
C02
Purges (C02) to Bed
Purges (CO 2) above Bed
Output in Cycle
Exit Gas Rate, SCFH (dry basis)
Composition, Mol %
H2
CO
C02
N2
H2S
COS
Outlet Gas , Top of Bed
Composition, Mol %
H20
H2
CO
co2
N2
H2S
COS
Flow Rate, SCFH, Top of Bed
Fluidizing Velocity, ft/sec
% H2S in Outlet/Equilibrium % H2S
Regeneration of Acceptor/Pass, mol %
A33
58
0.0
108
11
205
10
15
11
222
3.2
1.8
94.0
0.4
0.58
0.03
34.5
2.2
1.2
61.3
O.3
0.41
0.02
3L6
0.87
o.io
11.6
A43
51
35
51
5.0
70
10
15
16
107
2.7
1.9
93.3
0.6
1.70
0.07
28.5
2.2
1.5
66.1
0.4
1.36
0.06
178
0.49
0.36
14.5
A44A
53
35
51
5
70
10
15
7
104
2.5
2.1
93.1
0.8
1.66
0.06
28.3
2.0
1.7
66.2
0.6
1.33
0.05
173
0.48
0.36
14.2
A47
55
35
'51
5
70
10
15
51
108
3.1
1.6
94.4
0.4
0.56
O.03
29.7
2.4
1.2
65.9
0.3
0.44
0.02
182
0.50
0.11
5.2
170.
-------
FIGURE JX-22
DEACTIVATION OF CaS IN HARDENED BUCHANAN
DOLOMITE AT 704°C (13OO°F)
1.5
2 2.5 3
X =
5 6 7 8 9 10
Number of Cycles
1.5
2.5 3
5 6 7 8 9 10
171.
-------
10 X 10 TO '. INCH ? X 10 I'NCHES
KEUFFEL & ESSER CO. MADE IN u s »
46 1320
ro
RUN .033'!-::lte-1% daS :OT EXiT
-------
Run A44A
The feed to A44A was hardened at 35$ sulfidation rather than the 15$
used earlier. The attrition rate was 0.8$ and, as shown in Figure IX-22, the
activity was lowered to about the same level as had been observed with
Tymochtee 9 and 11 dolomites.
Run A47
Run A47 was a demonstration of the behavior of hardened Buchanan 3
dolomite for 58 cycles, the longest that any stone had been run. The sulfida-
tion conversion was ca 10$ until cycle 32 when breakthrough occurred. The H2S
feed was then reduced to give 5$ conversion and no breakthrough occurred until
the run was voluntarily terminated. Good CaC03 activity was observed for the
entire run and regeneration proceeded ixormally. The overall attrition rate was
0.7$. The run confirmed that hardened Buchanan dolomite would be suitable for
use in the process.
Figure 1X-22 graphically presents the regeneration data for four runs
with hardened Buchanan dolomite, two at 15$ sulfidation in the hardening step
and two at 35$. It appears that the activity of the stone is reduced due to
the higher conversion in the hardening sequence. However, the higher conversion
was accompanied by proportionately longer residence times in both the sulfiding
and oxidation steps, and the reduction in activity may have been due to the
extended hold at temperature rather than the specific conversion level.
Specifically, in Runs A33 and A43 residence times average 85 and 95 minutes,
respectively, in the sulfiding and oxidizing steps of the hardening procedure.
The comparable numbers for Runs A44A and A47 were 160 and 215 minutes,
respectively.
Data are presented in Figure IX-22 for activity beyond 30 cycles, but
as for Run A46 these data were not included in the correlation.
It had been noted earlier that the attrition rate tended to decline
as cycling continued. Having lasted 58 cycles, Run A47 provided a good check
on this notion. In the beginning the attrition rate was 1$; this dropped to
0.5$ as cycling continued. Similarly, in Run A46, the attrition rate fell from
1$ to 0.4$ in the course of 43 cycles. All else being equal, the average
attrition rate reported in this program is conservative compared to the rate
which would be observed in a long run with makeup. The magnitude of the effect
would depend on the make-up rate since a low make-up rate would give older bed
on the average than a high make-up rate. It is possible that the process harden-
ing procedure can be made less severe if hardening continues to occur as the
stone is cycled.
G. Density of Buchanan Dolomite
As dolomite is cycled through the process it tends to densify. Data are
presented in Figure IX-27 and Table IX-14 for the normalized density of hardened
Buchanan dolomite as it is cycled. The data show a steady increase in density
through cycle 58. Similar results had been observed with respect to Canaan
dolomite.
173.
-------
FIGURE IX-24
—.1'.'... KUN-A43— rMO
174.
-------
-------
- kbii' *LCaSLlK: EXITJ
Rfcdufce H2S
Feed.Ra'te
I
-------
TO
L & i_
4i 2
- 1.70
~:- 1.60
- 1.50
- 1.40
- 1.30
-------
TABLE IX-14
Normalized Density
VI)
Run No,
A33
A43
Source
Cycle No. CaCO3
CaS
gms/cc
Ibs/f t
00
A44A
A47
gms/cc
Ibs/f t3
Regenerator
1
1
Regenerator
,
r
Regenerator
1
I
Regenerator
>
i
3.9
9.1
11.5
0.9
5.2
12.9
16.2
1.2
6.2
11.5
1.2
6.1
16.7
25.2
41.8
58.1
66.1
48.0
38.4
78.2
48.1
23.1
15.2
58.4
34.9
13.2
61.5
43.4
20.9
10.7
10.6
10 o 9
28.2
42.1
55.1
20.6
48.6
76.9
82.3
39.3
65.1
83.4
36.8
53.1
73.3
84.1
83.7
84.3
1
10
6
0
3
2
2
3
1
3
5
5
5
4
.1 4.5
.0
. 6 —~
.8 0.5
.3
0 ;
.6
.0 0.4
0
.4
.4 0.2
.5
.8
.2
.7
.8
2.07
1.95
1.95
1.99
1.89
1.82
1.81
1.965
1.89
1.84
1.96
1.90
•1.86
1.84
1.85
1.99
129
121
121
123
117
113
113
122
117
114
122
118
116
115
115
124
.2
.8
.9
.9
.9
.3
,0
.6
.7
.6
.2
.5
.1
.1
.5
.1
1.50
1.52
1.55
1.43
1.46
1.48
1.51
1.48
1.50
1.54
1.47
1.48
1.53
1.56
1.56
1.68
93.8
94.9
96.8
89.3
91.1
92.2
94.0
92.3
93.3
95.9
91.5
92.6
95.7
97.2
97.6
104.6
(1) Calculated by normalizing each calcium component to a CaO basis while assuming no change in sample volume.
-------
H. Conclusions
It has been shown that a naturally soft stone can be hardened to the point
where it is sufficiently attrition resistant to be used in the process. Both
of the stones which we attempted to harden were successfully processed; there
were no failures. This success means that plant siting will not be severely
constrained by the need to have a rather particular type of dolomite in the
vicinity. Dolomites are rather widespread and it is likely that a suitable
stone which can be hardened can be found nearby.
Three different batches of Tymochtee dolomite were hardened. It was found
that there was some loss of activity in two of the batches, but the other dis-
played an activity slightly better than that of Canaan dolomite, the base case
stone. Buchanan dolomite was hardened with some loss of activity. However, it
is likely that the hardening conditions were somewhat more severe than was
needed, and that the stone can be hardened and still display an activity com-
parable to that of Canaan stone.
It was found that the severity of the hardening step affects both activity
and attrition resistance but in opposite directions. Tymochtee dolomite run at
982°C (l800°F) was very hard but displayed rather reduced activity. An optimal
condition was found at 871-927°C (160O-1700°F) for the oxidation step with
about 2O$ conversion. Runs made at 4% and 0$ conversion showed excessive
attrition rates. Using Buchanan dolomite, it was found that 35$ conversion
gave a very hard stone with reduced activity compared to 15$ sulfidation which
gave a high activity stone with poor attrition resistance.
For regeneration activity, there was no effect noted oi particle size over
the range of 2O x 28 to 35 x 48 Tyler mesh. The effect on attrition was somewhat
more difficult to evaluate due to the carryover problems caused by insufficient
disengaging height, but generally the rate was not different from the 35 x 48
mesh stone .
It is not known at this time why the different batches of Tymochtee stone
showed varying activity. This was not observed for either Buchanan or Canaan
dolomite.
179.
-------
X. BATCH SULFIDATION AND REGENERATION EXPERIMENTS
A. Batch Sulfidation Studies
A total of 45 batch runs were made to study variables affecting operation
of the gas desulfurizer. Base conditions for the batch sulfidation study are
given in Table X-l. These were essentially the same as for continuous runs.
Properties of feedstocks employed in both the sulfidation and regeneration
studies are shown in Table X-2. For sulfidation, a batch of fresh 28 x 35
Canaan dolomite was one of the feedstocks. Another was prepared by riffling
together the regenerator beds of Runs A20A, A21 and A36A, all made with Canaan
dolomite. These runs were carried out to 32, 37 and 2O cycles, respectively.
Tymochtee 11 dolomite was the third feed..
The procedure was to fill the bed with the programmed charge, equilibrate
with respect to temperature and gas flows, and then start feeding H2S. The
outlet H2S concentration was followed on the ultraviolet analyzer, and upon
breakthrough the H2S flow was shutoff. The cold leg of the reactor (L-5) was
purged of solids in order to get an uncontaminated sample of the bed. As the
H2S fed beyond the breakthrough point could be considerable, the CaC03 content
of the stone was corrected from the assayed value using a sulfur balance based
on the recorded analyzer output. The gas residence time and bed height were
calculated for each run based on the weight and density actually observed.
The amount of excess CaC03 required in the bed to prevent breakthrough of
H2S was correlated as a space rate, mols H,S/hr-mol of CaC03 remaining at 0.10$
H2S in the dry exit gas. The level of 0.1$ was chosen for operational conven-
ience. Once breakthrough occurred, the H2S concentration rose rapidly, and the
time to go from 0.05$ to 0.20$, for example, was short. The data for all the
cases are presented in Table X-3, and in a statistical layout in Table X-4.
The data tended to scatter, but the space rate held constant as the following
variables were changed: fresh versus cycled stone, bed depth from 4 to 17
inches, particle size from 28 x 35 to 35 x 48 mesh, and mol $ inlet H2S from
0.3$ to 1.1$. Only temperature was found to have an effect; the average space
rate increased from 1.9/hr to 3.0/hr in going from 871 to 927°C (l600-17OO°F).
Analysis of the data followed the procedures outlined by Hicks.(12) The
analysis of variance (ANOVA) is given in Table X-5 for the runs in which a
variable effect was found.
By way of comparison, Westinghouse'13) examined the reaction under similar
conditions using a thermogravimetric (TG) experimental system. They reported
the rate of reaction to decrease by about 6$ in going 871 to 927°C (16OO-17OO°F)
and to decrease by 3$ for each sulfidation-regeneration cycle. Rate of sulfi-
dation increased with decreasing particle size over the range of about 35 to 5
Tyler mesh. The Westinghouse results do not concur with those reported here,
but the dependent variables differed, the independent variables were changed
over a narrow range, most effects were small, and different experimental tech-
niques were used. Pending further studies, the Conoco results are assumed to
be correct since experimental conditions more closely simulated the process .
conditions and fluidized bed mechanics of interest to our process.
It was found that the concentration of H2S in the outlet gas varied with
conditions even when there was an excess of CaC03 present in the bed. The data
for percent H2S prior to breakthrough are presented in a statistical layout in
Table X-6. The ANOVA tables for the various data sets are given in Table X-7.
181-
-------
TABLE X-l
oo
Ni
Run Number
Temperature, °C
Type of Dolomite
Acceptor Size Consist, Tyler Mesh
Input, SCFH
Recycle to Bed
H20
H2
CO
C02
H2S
Purges
Purges
(C02) to Bed
(N2) above Bed
Output
Exit Gas Rate, SCFH (dry basis)
Composition, Mol %
H2 . ' ..
CO
C02
N2
H2S
Outlet Gas , Top of Bed
Composition, Mol %
H20
H2
CO
co
H2S
Flow Rate, SCFH, Top of Bed
Fluidizing Velocity, ft/sec
Conditions for Batch Sulfidation Runs
System Pressure: 15 atm (206 psig)
A53-A67
871 (1600°F)
Canaan 3
28 x 65
35
4O
1 Q
JLO
OB
1.3-4.3
-
215
17
19
12
52
9
16
18
11
46
439
0.62
A87-A95
871 (1600°F)
Canaan 3
35 x 48
11
4.6
12
1 Q
JLo
Rfi
0.9-3.0
122
3.4
10.5
21
65
4.8
3.4
10.6
21-:
60"
268
0.38
A98-A106
899 (1650°F)
Canaan 3
35 x 48
11
3.5
12
At
0.9-3.0
5
101
3.5
11.8
29
56
4.6
3.5
12
29
51
258
0.38
A108-A116
« 927
Canaan 3
35
^
0.9-3.0
220
4
8
13
75
6
4
8
13
69
272
0.41
A117-A19
Tymochtee
x 48
in
17
on
oo •
3.0
221
4
8
13
7.5
6
4
8
12
69
272
O.41
11
_
(1) Input is given after shift reaction and hydrolysis of CS2 have taken place.
-------
TABLE X-2
Dolomite Identification
Raw Canaan No. 3
Cycled Canaan
Sulfided Fresh
Canaan No . 3
Source of Feed or Product
Raw Stone
Regenerator beds of Runs A2OA
(32 cycles), A21 (37 cycles)
and A36A (2O cycles)
Gas desulfurizer beds of Runs
A53-A55 and A59-A61
Properties of Feeds and Products
Mol Jt Density
CaCO3 CaS "CaO" gm/cc lb/ft3
ICO 0 O
24.2 69.6 6.2 2.11 131.9
25.5 72.3 2.3 2.O1 125.7
% Size Consist , Tvler Mesh
+28 28 x 35 35 x 48 48 x 65 65 x 1OO -1OO
(0
11.5 13.8 18.9 33.4 16.3 5.9
1.4 59.0 31.8 3.9 2.3 1.7
Sulfided Cycled Canaan
No. 3
A8O-A81 Feedstocks
CD
Raw Tymochtee 11
Gas desulfurizer beds of Runs
A62-A67
Regenerator beds of Runs A7S
and A78
Raw Stone
12.8 82. O 5.2
39.6 58.5 2.0
100
2.O7 129.0
2.29 137.2
1O.O 11.4 19.1 35.6 18.3
100
5.6
(i) 28 x 35 mesh for Runs A53-A61, 35 x 48 mesh thereafter.
-------
TABLE X-3
Results
of Sulfldation
System Pressure: 206 pslg
Run No.
A53
A54
A55
A56
A57
A58
A59
A60
A61
A62
A63
A64
A65
A66
A67
A87
ASS
A89
A 90
-, A91
J A9E
A93
A 94
A95
A98
A99
A 100
A101
A102
A103
A104
A105
A106
A108
A109
A110
Alll
A112
A113
A114
A115
A116
A117
A118
A119
Inlet
% HzS
0.33
0.67
1.06
0.34
0.64
1.07
1.05
0.^53
0.36
0.37
0.67
1.06
0.31
0.59
ca. 1.2
0.32
0.31
0.39
0.59
0.58
0.59
1.00
0.97
1.01
0.31
0.35
0.36
0.63
0.61
0.64
1.00
0.99
1.02
0.34
.33
.40
.64
.62
.67
1.13
1.07
1.11
.99
1.06
1.07
Bed
Type of Temp. , Height ,
Dolomite °c(°F) inches
Fresh 871 16.9
Canaan 3 (1600) 16.5
28 x 35
Mesh
\
Cycled
Canaan
35 x 65
Mesh
Fresh
16.9
10.4
10.5
10.5
7.6
7.3
7.6
16.9
17.3
17.5
7.7
7.8
7.7
16.7
Canaan 3 871 9.9
35 x 48 (1600) 4.6
Mesh
16.9
10.5
4.5
16.4
10.4
4.4
899 17.0
(1650) 10.9
4.3
16.4
10.5
4.6
16.0
9.3
4.3
Fresh 927 17.0
Canaan 3 (1700) 10.9
35 x 48
Mesh
Fresh
Tymochtee 11
35 x 48 B
4.8
16.8
10.8
4.8
17.3
10.9
4.6
4.0
11.5
17.9
Superficial
Gas Residence
Time , sec
2.3
2.2
2.3
1.4
1.4
1.4
1.0
1.0
1.0
2.3
2.3
2.4
1.0
1.0
1.0
3.7
2.2
1.0
3.7
2.3
1.0
3.6
2.3
1.0
3.7
2.4
0.9
3.6
2.3
1.0
3.5
2.0
1.0
3.5
2.2
1.0
3.4
2.2
1.0
3.5
2.2
0.9
0.8
2.3
3.6
Length
of Run,
min
419
227
117
262
143
73
49
90
206
89
39
22
26
13
16
715
461
200
347
215
83
195
134
37
756
453
197
354
230
65
216
129
43
832
438
204
371
228
100
200
135
47
56
102
152
% of Ca
Sulfided/hr
12.3
26.9
43.9
20.1
42.0
72.2
93.8
59.5
30.1
14.3
27.9
46.6
28.3
53.4
116
8.4
12.7
30.8
16.0
21.7
60.9
26.9
41.8
106
7.8
14.1
34.0
16.3
24.6
60.5
26.8
44.3
100
8.7
13.2
36.7
16.3
24.7
58.8
28.3
42.7
104
144
52.1
33.5
% H2& Before
Breakthrough
± .01
.04
.05
.05
.04
.04
.05
.06
.06
.05
.05
.04
.06
.05
.06
N.D.
.04
.04
.07
.04
.05
.05
.04
.05
ca. .09
.04
.05
.07
.05
.05
.08
.05
.05
.09
.02
.02
.02
.02
.02
.04
.02
.02
.04
.02
.02
.02
Runs
(15 atm)
% H2S - % H2S
Equilibrium
± .01
.01
.02
.02
.01
• 01
.02 ..
.03
.03
.02
.02
.01
.03
.02
.03
—
.01
.01
.04
.01
.02
.02
.01
.02
.06
.01
.02
.04
.02
.02
.05
.02
.02
.06
.01
.01
.01
.01
-.01
.03
.01
.01
.03
.01
.01
.01
% CaOO3
Remaining Space Velocity at
at 0.10% Breakthrough
H^S in Dry Mols Inlet HgS/hr
Outlet Gas mol CaCOs
8.1
10.2
27.5
7.3
15.2
24.8
57.5
27.0
12.0
8.6
13.0
16.2
15.0
22.6
N.D.
7.6
7.6
29.5
12.0
19.8
61.6
8.2
8.4
76.7
5.5
6.7
16.4
5.7
7.5
63.3
11.1
19.1
78.8
5.3
5.5
10.9
5.2
6.5
17.2
6.9
12.6
53.9
7.6
6.4
3.4
1.87
2.93
1.63
3.43
3.06
3.12
1.75
2.49
3.06
2.04
2.37
2.93
2.43
2.70
: 5.2
.91
.67
.04
.34
.10
.99
.29
.98
.38
.59
2.10
2.08
2.87
3.28
0.95
2.42
2.33
1.27
1.65
2.43
3.39
3.12
3.82
3.40
4.10
3.39
1.93
18.9
8.1
10.0
Density Mol % in Product
Ib/ft3 Ca(X»3 CaS"CaO"
6.0
6.9
15.8
4.6
5.7
15.8
21.6
17.1
4.9
5.6
7.9
10.3
13.9
15.6
12.3
1.88 117.3 6.6
1.87 116.6 6.5
1.93 120.3 17.6
11.3
15.6
34.1
1.94 121.0 5.9
1.94 121.2 5.3
2.04 127.2 48.5
1.88 117.2 4.9
1.88 117.2 4.0
1.91 119.3 0.5
4.0
4.4
40.0
1.96 122.1 6.6
2.01 125.2 7.3
2.08 129.7 35.5
1.90 118.4 4.3
1.91 119.3 4.6
1.89 118.1 8.0
3.7
4.4
9.5
1.91 119.3 4.0
1.93 120.4 8.4
2.06 128.7 32.5
1.61 100.4 1.0
1.67 104.3 4.3
1.69 105.2 2.1
88.4
88.8
82.7
90.8
89.8
80.0
74.1
78.6
89.3
86.9
84.7
84.9
81.2
79.7
83.2
86.7
85.2
76.2
84.2
81.2
65.9
88.5
89.7
51.5
88.2
91.6
85.6
89.4
89.6
57.6
90.0
89.1
63.9
88.8
90.1
86.1
90.0
88:0
83.8
89.8
89.1
67.5
82.9
80.9
82.5
5.6
3.3
1.5
4.6
4.5
4.3
4.4
4.3
5.7
7.5
7.4
4.8
4. 9
4.8
4.5
6.7
8.3
6.2
4.5
3.2
0
5.6
5.1
0
7.0
7.7
4.9
6.6
6.0
2.4
3.5
3.5
0.6
7.0
5.3
6.0
6.5
7.7
6.7
6.2
2.5
O
16.2
14.8
15.5
N.D. = Not Determined.
-------
TABLE X-4
Mols HaS Fed/hr Per Mol CaCO3
Remaining at 0.IO% HgS in Exit Gas
X-4A Runs A53-A67
Bed Depth,, inches
Age of Stone
Mol jo H2S. Inlet
0.34
0.64
1.06
16.5-17.5
Fresh
1.87
2.93
1.63
Cycled
2.04
2.37
2.93
10. 4-10.5
Fresh
3.43
3.06
3.12
7.3-7.8
Fresh
3.06
2.49
1.75
Cycled
2.43
2.7O
2.33*
* Estimated by a sum of squares minimization technique recommended by Hicks.
X-4B Runs A53-A61. A87-A95
Bed Depth, inches
Size Consist,
Tyler mesh
Mol
HgS. Inlet
0.3
0.6
1.0
Bed Depth, inches
Temperature, ° F
Mol % HaS. Inlet
O. 3-O. 4
0. 6-O.7
1.0-1.1
35
0.
1.
3.
16.0-17.0
x 48 28 x 35
91 1.87
34 2.93
29 1.63
X-4C Runs A87-A95.
1600
O.91
1.34
3.29
16.O-17.O
165O 17OO
1.59 1.65
2.87 3.12
2. 42 4. 10
35 x
1.67
9.3-10.9
48 28 x 35
3.43
1.10 3.06
4.98
A98-A106
9.
1600
1.67
1.10
4.98
3.12
. A108-A116
3-10.9
1650 170O
2.1O 2.43
3.28 3.82
2.33 3.39
4.3-7.6
35 x 48 28 x 35
1.04 3.06
0.99 2.49
1.38 1.75
4.3-4.8
16OO 1650 1700
1.O4 2.O8 3.39
0.99 0.95 3.40
1.38 1.27 1.93
185,
-------
TABLE X-5
ANOVA for Mols HgS/hr-Mol CaCOa at Breakthrough
Runs A87-A95, A98-A1O6, A108-A116
Factor
Mol % H2S, Inlet
Bed Depth
Temperature
Error
Total
Degrees of
Freedom
2
2
2
20
26
Sum of
Squares
3.7638
4.1965
6.86O5
17.3250
42.2259
Mean
Square
1.8819
2.0983
3.4303
.8663
F Level of
Ratio Significance
2.17 < 90%
2.42 < 90$
3.96 > 95%
186.
-------
Bed Depth, inches
Age of Stone
Mol % H.»S. Inlet
0.34
0.64
1.06
TABLE X-6
Mol % H2S in Dry Outlet Gas
With Excess CaCOa Remaining in the Bed
16.
Fresh
0.04
0.05
0.05
X-6A Runs
5-17.5
Cycled
0.05
0.04
O.O6
A53-A67
10.4-10.5
Fresh
0.04
0.05
0.06
7.
Fresh
0.05
0.06
O.O6
3-7.8
Cycled
O.O5
0.06
O.O64*
Estimated by a sum of squares minimization technique recommended by Hicks. (12)
X-6B Runs A53-A61. A87-A95
Bed Depth, inches
Size Consist,
Tyler mesh
Mol % HaS. Inlet
0.3
0.6
1.0
16.0-17.0
35
0.
0.
0.
x 48
04
04
04
28
0.
0.
0.
x 35
04
05
05
35
0.
0.
0.
9.3-10.9
x 48
04
05
05
28
0.
0.
0.
x 35
04
04
O5
35
0.
0.
0.
4.3-7.6
x 48
07
O5
09
28
0.
o.
o.
x 35
05
06
06
X-6C (Mol # H2S-H2S Equilibrium)
in Dry Outlet Gas with Excess CaCOa Remaining in Bed
Bed Depth, inches
Temperature, °F
Mol I HaS. Inlet
O. 3-0. 4
O. 6-0.7
1.0-1.1
Runs A87-A95. A98-A1O6, A1O8-A116
16.0-17.0
9.3-10.9
16OO 165O 17OO 16OO 1650 17OO
0.01
0.01
0.01
0.01
0.02
0.02
0.01
0.01
0.01
0.01
0.02
O.O2
0.02
0.02
0.02
0.01
0.01
0.01
4.3-4.8
16OO
0.04
0.02
O.O6
1650
0.04
0.05
0.06
17OO
0.01
0.03
0.03
187,
-------
TABLE X-7
ANOVA For Mol % HaS in Dry Outlet Gas
Factor
Mol % H2S, Inlet
Bed Depth
Age of Stone
Error
Total
Factor
Mol % H2S, Inlet
Bed Depth
Error
Total
Factor
Mol % H2S, Inlet
Bed Depth
Particle Size
Error
Total
Factor
Inlet H2S
Bed Depth
Temperature
Error
Total
Data were
X-7A
Degrees of
Freedom
2
1
1
6
10
Degrees of
Freedom
2
2
4
8
X-7C
Degrees of
Freedom
2
2
1
12
17
coded by multiplying by 10O
Runs A53-A55, A59-A67
Sum of Mean F
Squares Square Ratio
2.4608 1.2304 3.74
2.4571 2.4571 7.46
0. 1704 0. 1704 < 1
1.9762 0.3244
7.0644
X-7B Runs A53-A61
Sum of Mean F
Squares Square Ratio
2.8889 1.4445 13. OO
1.5556 0.7778 7.00
0.4444 0.1111
4.8889 ;"
Runs A53-A61, A87-A95
Sum of Mean F
Squares Square Ratio
3.444 1.722 2.02
14.777 7.389 8.67
O.5OO O.5OO < 1
10.223 0.852
X-7D Runs A87-A95, A98-A106, A1O8-A116
Degrees of
Freedom
2
2
2
20
26
Sum of Mean F
Squares Square Ratio
3.6297 1.8149 2.54
34.7408 17.3704 24.30
9.4074 4.7037 6.58
14.2962 0.7148
62.0741
Level of
Significance
> 90%
> 95$
Nil
Level of
Significance
> 95$
> 95%
Level of
Significance
< 90%
> 99%
Nil
Level of
Significance
< 90$
> 99%
> 99%
188.
-------
The following independent variables were judged to have no effect on outlet H2S
concentration: age of stone, particle size over the range 28 x 35 to 35 x 48
mesh, and inlet H2S concentration over the range 0.3$ to 1.1$. The result on
inlet H2S concentration is not as certain as the others, since one of the data
subsets (Runs A53-A61) did show an effect at the 95$ significance level. For
all other subsets, the effect was less than the 90$ level.
The effect of decreasing bed depth was to increase the outlet H2S before
breakthrough. Increasing the temperature in the range of 871 to 927°C
(1600-1700°F), reduced the outlet H2S concentration. Deeper beds and higher
temperature gave closer approach to equilibrium. It is noted that in deep beds
(>9 inches), the approach to equilibrium with within 0.02$, about the limit of
accuracy of the H2S analyzer. In addition, the closest approach was observed at
1700°F, followed by 1600°F, followed by 1650°F, and this is an anomalous order-
ing. Pending studies capable of improved accuracy, the outlet H2S concentration
will be assumed to be at the equilibrium level for purposes of the process design.
Runs A117 to A119 were made with Tymochtee 11 dolomite rather than with
Canaan stone. Table X-3 shows that the H2S concentration before breakthrough was
about the same as for Canaan stone, but the space rate at breakthrough was much
higher. The cause of the observed difference is not known.
The results of the study are consistent with a model of very fast reaction
in which the exit gas is close to equilibrium. Although tripling the inlet gas
concentration would be expected to increase the outlet concentration, when com-
bined with the equilibrium base of 0.03$, the observed effect would be small.
Decreasing the bed depth would be expected to have an effect due to gas bypassing.
The fact that age of the CaC03 had no effect is consistent with previous
conclusions, and again confirms a key feature of the process.
Overall, the batch runs did not give results as good as those from continu-
ous runs. The H2S concentration prior to breakthrough was slightly higher, and
breakthrough occurred with more CaC03 remaining in the bed. The reasons for the
differences are not understood at this time. Perhaps the regenerated material
falling into the bed from above gives some countercurrency to the continuous system.
A complicating factor in the analysis was that the H2S feed was made from
CS2. It had been shown that all the CS2 is hydrolyzed to H2S in a full bed of
inert MgO (magnorite). However, the rate of hydrolysis and the rate of the
shift reaction which produced the water may become limiting factors in the very
short beds which were employed in this study. It was assumed that sulfur entered
the bed in the form of H2S with the understanding that the runs with four-inch
beds represent a limiting case of how short a residence time can be .investigated
with the existing equipment.
B. Batch Regeneration Study
Base conditions for the regeneration study are given in Table X-8. The gas
desulfurizer beds of the sulfidation study were combined to give two feedstocks
consisting of sulfided fresh stone and sulfided cycled stone, respectively.
Detailed results and variables pertaining to regeneration runs are presented in
Table X-9. The main study involved temperatures of 649, 704 and 760°C (12OO,
1300 and 14OO°F) with full (36 inch) and one-half (16 inch) beds and fresh and
recycled stone.
189.
-------
TABLE X-8
Typical Conditions for Batch Regeneration Runs
System Pressure;15 atm (206 psig)
Run Numbers A68-A79 ABO. A81
Temperatures, °C 649-760 816, 871
(1200-14OO°F) (150O, 160O°Fj)
Nominal Solids Residence Time,, hours 8 8
Input. ,SCFH
Recycle to Bed 0 0
H20 79 69
H2 8 7
CO 00
C02 135 117
Purges (C02) to Bed 10 1O
Purges (N2) above Bed 0 0
Purges (C02) above Bed 15 15
Output
Exit Gas Rate, SCFH (dry basis) 165 145
Composition. Mol jo
H2 2.3 1.6
CO 2.6 3.2
C02 93 94
N2 0.7 0.2
H2S (i) fi)
COS (2) (2)
Outlet Gas. Top of Bed
Composition. Mol %
H20 35 36
H2 1.6 1.1
CO 1.9 2.3
C02 60 60
N2 0.5 0.2
H2S (i) (0
COS (2) (2)
Flow Rate, SCFH, Top of Bed 230 2O3
Fluid iz ing Velocity, ft /sec 0.6-0.7 0.6
(l) See detailed data for each run.
(a) Usually 3-5$ of the H2S.
190,
-------
The procedure was to fill the bed with the programmed charge and then to
turn on the steam feed. The runs were continued for eight hours or until there
was no longer any change in H2S content, whichever came first. As a rule, the
H2S concentration at the end of the run was less than 0.15$ on a dry basis.
The cold leg of the reactor (L-6) was purged of solids prior to taking a sample
of the bed.
The data for percent regeneration are listed in Table X-10. By inspection,
the effect of temperature seemed to be different for cycled stone and fresh
stone. A temperature times age interaction term, TxA, was therefore included in
the ANOVA, Table X-ll. The effects of temperature, bed depth, age of stone, and
temperature times age were all found to be significant at the 1% level. The TxA
effect showed that increasing temperature improved conversion more for cycled
stone than for fresh stone.
The effect of bed depth was unusual. It was expected that since the deeper
bed ran at an H2S outlet concentration closer to equilibrium, there would be
less driving force for reaction and the regeneration conversion would be lower.
In all six test pairs exactly the opposite was found. The runs with the deeper
beds, i.e., higher H2S outlet concentration, gave higher conversions. This is
consistent with earlier conclusions that increasing the H2S concentration does
not adversely affect conversion in the regenerator.
Figure X-l is an attempt to correlate the factors contributing to approach
to equilibrium in the outlet gas. It was assumed that the more mols of solid
reactants per mol of gas, the closer equilibrium is approached; temperature is
an obvious parameter. The data in Figure X-l follow the expected trend, but no
simple functional correlation is discerned.
Two additional exploratory runs were made at 816 and 871°C (1500 and 1600°F)
with cycled stone from the two 130O°F regenerator runs, A75 and A78. Conversions
based on the original sulfided feed were 67% and 81% at 1500 and 160O°F, respec-
tively. Although the equilibrium H2S concentration is low at these temperatures,
it was thought that a reactivation process might be feasible if a slipstream of
stone could be treated at an elevated regeneration temperature. It turned out
that the stone could not be so reactivated, and details are discussed in Section
X-D.
C. Regeneration Kinetics
Regeneration kinetics were examined by recording the H2S concentration in
the exit gas as a function of time as the runs progressed. The total mols of
H2S produced were normalized to match the total mols of CaS reacted, and con-
version as a function of time was then plotted. These results are presented in
Figures X-2 to X-5 in the form of (l - X) versus time, where X is the fractional
conversion of the CaS feed.
In all cases, the rate continuously decreased. It appeared that the reaction
would eventually cease while there was still considerable CaS left. This is
consistent with the notion that much of the CaS is simply unreactive. The initial
reaction rate was faster for fresh stone than for cycled stone at all temperature
levels. Additionally, the rate seemed to slow sooner for the cycled stone. The
curves also show that the effect of bed depth is present even at the early stages
of all reactions; the deeper beds had a higher reaction rate at all times.
191.
-------
TABLE X-9
Run Ho.
A68
A69
A7O
A71
A72
A73
A74
A75
A76
A77
A78
A79
A8O
AS1
Dolomite
Sulflded
Fresh
Sulfided
Cycled
i
A75 + A78
Composite
Bed
Height,
inches
35.3
36.1
36.8
16.1
16.9
16.2
37.9
36.5
38.9
15.9
15.3
17.1
20. 2
20. 5
%
Regeneration
61.6
75.3
80.8
57.5
71.2
78.3
14.5
31.5
55.4
10.9
23.0
42.6
67.l(')
8l!o(O
Maximum
Approach
to
Equilibrium, *t
13.0
47.4
93.7-1O2.8
3.0
13.5
5O.O
2.8
14.2
31.5
0.7
3.4
13.6
19.5
.50
Run
Time,
hours
7.2
8.0
5.7
8.4
8.0
4.7
9.1
8.1
1O.O
3.7
7.8
8.0
8.0
8.0
Results of Regenerator Runs
Maximum
% H,S, Gas
Dry Temperature,
Basis
1.77
2.78
2.68-2.94(0
.41
.79
1.43
.38
.83
.90
.09
.20
.39
.31
.46
°F CO
1200 (649)
13OO (704)
1400 (760)
12OO (649)
1300 (704)
1400 (760)
12OO (649)
13OO (704)
14OO (76O)
1200 (649)
13OO (7O4)
1400 (760)
1500 (816)
1600 (871)
Residence
Time . sec .
4.9
4.7
4.6
2.2
2.2
2.0
5.3 :
4.8
4.8
2.0
2.O
2.1
2.7
2.6
1O3 x Hols CaS
Available
Hols (COa+HaO) Density. lb/fta
hr
20. 6
24.8
26.5
8.6
11.0
11.4
5.39
1O.8
19.9
1.71
3.38
6.8
Actual
137.4
14O.7
142.1
134.2
137. 2(a
142.8
137. 2(2
146.6
149.1
Normilized
99.8
100.3
1OO.4
1O8.4
' 1O6.2(O
1O7.1
1O6.2(2)
1O7.5
1O5.4
Itol
67.8
79.2
CaO
8.0
10.5
8.9
1.5
9.2
3.4
6.9
3.2
11.0
4.3
4.3
5.0
1O.1
1.0
VO
IsJ
(») The high value is extrapolated where a portion of the chart was unreadable.
(a) A75 + A78 combined.
(a) Based on original feed.
-------
TABLE X-10
Bed Depth
Age of Stone
Term., °F (°C)
1200 (649)
1300 (704)
14OO (760)
1500 (816)
1600 (871)
Regeneration Results
Full
Fresh Cycled
61.6 14.5
75.3 31.5
80.8 55.4
- Percent Regeneration
2/3
Cycled Fresh
57.5
71.2
78.3
67.1
81.0
1/2
Cycled
10.9
23.0
42.6
TABLE X-ll
ANOVA for Percent Regeneration
Factor
Temperature
Bed Depth
Age of Stone
T x A
Error
Data for 15OO°F and
Data were coded by
Degrees of Sum of
Freedom Squares
2 15.848
1 1.056
1 5O. 758
2 1.684
6 0. 389
16OO°F were excluded
multiplying by 0. 10
Mean F
Square Ratio
7.924 122.3
1.056 16.3
50.758 783.3
0.842 13.0
0.0648
Level of
Significance
> 99$
> 99$
> 99$
> 99$
Total
11
69.735
193.
-------
ana
\
E
o
10
15 194. 2°
25
30
35
-------
10
9
8
7
6
5
4
3
in
10 :
Iftl
2
1
ALl
mimrt-es
2QQ
25O
5O
1OO
ISO 20O
195.
25O
3OO
350
-------
to
9
8
7
50
IOC
300
6
5
I ep
*i
iff,
3
itllt
2
^ in
I
50
10O
150 200
196.
25O
3OO
5O
-------
10
9
8
7
50
100
150
200
250
300
6
5
4
3
in
-------
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198 .
-------
While the reaction kinetics are clearly very complicated, they may be
approximated by a simple first order rate constant for the short times (about
an hour) of interest to the process. The reaction model would then be:
Kt = -In (1 - X)
where K = rate constant, hr"1, whose function of gas concentration
is yet to be determined.
t = time, hours.
and X = fractional conversion of the CaS fed.
The value of K was taken from the slope of the data from 5 to 50 minutes
in Figures X-2 to X-4, and is tabulated in Table X-12. The data of Table X-12
are plotted in Figure X-6. The activation energies for cycled and fresh stone
are shown to be the same, about 19 kilocalories.
The kinetic mechanism proposed here allows one to calculate the effect of
changing conditions of temperature or residence time upon fractional conversion
of CaS (percent regeneration).
P. Cyclic Sulfidation and Regeneration
Run A85 - Cycling of Reactivated Dolomite
Run A85 constituted two sulfidation-regeneration cycles on the combined
product of Runs A80 and A81. In these runs cycled Canaan dolomite was regenerated
for 8 hours at 816 or 871°C (1500 or 16OO°F). A highly regenerated product (high
CaC03 content) was obtained. Run A85 was set up to determine the performance of
the regenerated stone, i.e., whether any reactivation took place along with
regeneration. Sulfidation conditions were the same as for Runs A52-A67 described
earlier. Regeneration was at 704°C (l3OO°F) with the usual 65$ CO2, 3O# H2O
composition.
The history of the stone and the result of Run A85 are presented in
Table X-13. In two attempts at sulfidation, breakthrough occurred immediately
and almost no CaC03 reacted. Although a considerable quantity of CaC03 had been
formed in the high temperature regeneration, it displayed very poor reactivity.
The data in Table X-13 show a decline in "CaO" as the stone is cycled
through the long regeneration at 7O4°C (l3OO°F), and a further decline after
regeneration at 817 and 871°C (1500 and 1600°F). The "CaO" level started to
rise upon cycling in the normal sulfidation-regeneration mode. It is possible
that the rigorous regeneration conditions encouraged recarbonation of the inert
"CaO". However, it is noted that the data in Table X-9 do not show any effect
of temperature on "CaO" content over the range 649-76O°C (l2OO-14OO°F).
Run A124 - Reactivated Stone at Calcining Conditions
Run A124 consisted of a single sulfidation and regeneration of Run A85
material, stone which had been regenerated at 817^871°C (15OO-160O°F), and which
then displayed poor reactivity toward H2S. The stone was sulfided at 927°C
(1700°F) under calcining conditions. High reactivity toward H2S was obtained.
Upon regeneration, only 11$ relative conversion of CaS to CaC03 occurred. It is
199.
-------
TABLE X-12
Initial Rate Constants for Regeneration Kinetics
System Pressure: 15 atm
Feed: 35$ H20, 60$ C02, 5% H2
K. hr-1
Temperature, °C (°F) Cycled Fresh
649 (1200) .063 .247
704 (1300) .098 .468
760 (1400) .170 .747
200.
-------
FIGURE X-6
0.96
O 98
)2 1.04
10~3/T, 0K-1
1.10
201.
-------
TABLE X-13
History and Cycling Results for Reactivated Dolomite
Material Description Mol %
CaCOa CaS CaO
A. History
Cycled Canaan Dolomite Feed Composited
from Runs A20A, A21, A36A 24.2 69.6 6.2
Sulfided Stone Composited from
Runs A53-A61 12.8 82.0 5.2
Stone Regenerated at 7O4°C (l3OO°F),
Runs A75 and A78 39.6 58.5 2.O
Stone further Regenerated at 816 and 871°C
(150O and 1600°F), Runs A8O and A81 73.4 26.O 0.6
B. Cycling Results - Run A85
Sulfidation No. 1 871°C, till breakthrough 67.9 32.1 0
Regeneration No. 1 704°C, 60 minutes 71.8 27.9 0.3
Sulfidation No. 2 871°C, till breakthrough 7O.8 27.7 1.5
Regeneration No. 2 7O4°C, 60 minutes 72.8 24.0 3.2
202.
-------
concluded that stone reactivated at elevated temperatures would be active in
the CaO form, but that subsequent regenerations would give the same conversions
as for the original deactivated material. Run conditions are given in Tables
X-14 and X-15.
Run A86 - Cycling of Dolomite Regenerated at 760°C (1400°F)
The results of Run A85 implied that CaC03 produced by regeneration at
high temperatures might be unreactive. Preliminary results obtained at The City
University of New York(14) also indicate that unreactive CaC03 might be formed
at 750°C (1380°F). To explore the matter further, dolomite from Run A22A was
cycled at 871°C (l6OO°F) sulfidation and 760°C (1400°F) regeneration. Run A22A
was a continuous run lasting 22 cycles in which the regenerator was run at 760°C
(1400°F).
Results of Run A86 are given in Table X-16. The first few cycles gave
20$ regeneration, the same as had been observed in the continuous run. However,
the H2S/CaC03 space velocity at breakthrough was lower than for previous batch
runs, and the H2S in the outlet gas tended to rise as cycling continued. Eventu-
ally, the initial H2S level prior to breakthrough reached the breakthrough
concentration so that essentially no sulfidation was carried out. After 13
cycles, the inlet H2S concentration was reduced from O.6 to 0.3$ and a plateau
of 0.08$ H2S was obtained prior to breakthrough. The H2S/CaC03 space velocity
was only 0.17 at breakthrough in cycle 15.
It is concluded that some of the CaC03 in the Run A22A material was
unreactive thereby giving the low H2S/CaC03 space velocity and the high H2S
concentration while excess CaC03 was still present.
An additional relevant item of information is that examination of a
few selected samples of cycled stone using a scanning electron microscope (SEM)
equipped for energy dispersive X-ray analysis has disclosed the presence of very
large CaS crystals in a matrix of small MgO crystals. These results are detailed
in Section XIV. Knowing that large crystals of CaS are formed can aid in explain-
ing the results of Runs A85 and A86.
i
It is hypothesized that deactivation of CaS is caused by growth of large
relatively inert crystals. During normal regeneration, the small CaS crystals
are converted to CaCO3 which are reactive by virtue of their size and high
surface area. High temperature, long time regeneration converts the interior of
the large crystals to CaC03 by "brute force". These large CaC03 crystals are
fundamentally unreactive. Thus, the CaC03 fed into Run A85 barely reacted at all.
The feed to A86 contained some reactive and some unreactive CaC03, and there-
fore performed at a reduced efficiency based on its total CaC03 content.
Run A128 - Cycling of Half-Calcined Dolomite
Run A128 was a base case run made with fresh half-calcined Canaan dolo-
mite. The products of Runs A87-A95 were combined and then cycled batchwise for
four cycles. Figure X-7 shows that results were comparable to those from a
continuous run.
203-
-------
Run Number
Temperature, °C
Type of Dolomite
Acceptor Size Consist, Tyler Mesh
Nominal Feed Rate gm/hr (half-
calcined basis)
Input. SCFHC1)
Recycle to Bed
H20
H2
CO
CO 2
TABLE X-14
Conditions and Results for Gas Desulfurlzer
System Pressure:15 atm (2O6 psig)
A124
H2S
Purges
Purges
(C02) to Bed
(N2) above Bed
Output
Exit Gas Rate
Composition.
A128
927 (17OO°F)
Canaan
from
Run A85
28 x 65
, SCFH (dry basis)
Mol %
H2
CO
C02
Na
HS
190
7
7
6
80
871 (1600°F)
Fresh
Canaan
35 x 48
Batch
80
12.6
11.9
13.6
4.9
138
J «;
* 15
148
21
27
22
9.3
66
142
17
17
9
56
Outlet Gas . Top of Bed
Composition, Mol %
H20
H2
CO
C02
N2
H2S
Flow Rate, SCFH, Top of Bed
Fluidizing Velocity, ft/sec
5
7
7
6
75
267
0.40
8
17
17
9
50
299
0.43
(1) Input is given after shift reaction and hydrolysis of CS2 have taken place.
204.
-------
TABLE X-15
Conditions and Results for Regenerator
System Pressure:15 atm (206 psig)
Run Number
Temperature, °C
Regeneration Time, min.
Input, SCFH
Recycle to Bed
H20
Ha
CO
C02
Purges (C02) to Bed
Purges (N2) above Bed
Purges (C02) above Bed
Output
Exit Gas Rate, SCFH (dry basis)
Composition, Mol ^
H2
CO
CO 2
N2
H2S
COS
A124
A128
704 (1300°F)
60
-35
- 47
•5.0
— 0
70
— 10
— 0
-15
97
{
H
94
P.
Outlet Gas. Top of Bed
Composition. Mol %
H20
H2
CO
CO2
N2
H2S
COS
fifi
n
-, ^
Flow Rate, SCFH, Top of Bed
Fluidizing Velocity, ft/sec
-166
•0.46
205,
-------
TABLE X-16
NJ
8
Run A86 - Batch Cycling ol Run A22A(0 Dolomite
Gas DesuLfurlzer
Hoi HjS, Inlet
Cycle Number
tiol %
CaCO,
CaS
"CaO"
Mol
-------
E. Cycling with Varying Residence Times
Runs A129-A132
Four batch cycling runs were made employing the possible combinations
of 20 or 60 minutes residence time for sulfur absorption and regeneration.
Details are presented in Tables X-17 and X-18 and in Figures X-8 and X-9.
Tymochtee 11 dolomite was the feedstock, and each run lasted for 11 cycles. For
runs employing 60 minutes in the gas desulfurizer, breakthrough occurred after
several cycles. The H2S feed rate was then reduced from 3 SCFH to 1 SCFH for
the remainder of the cycles. In the batch mode of operation, the bed volume
continually decreased due to sampling, attrition and handling losses. Data are
given for both Runs A129 and A129A. Run A129 was terminated after four cycles
due to depletion of the inventory. Thereafter, a larger inventory was used and
11 cycles were achieved in Run A129A.
Most striking was the effect of reducing residence time in the gas
desulfurizer. At 1O cycles, going from 60 to 20 minutes in the gas desulfurizer
increased regeneration activity from 15 to 40$ at 60 minutes regenerator time and
from 15 to 33$ at 2O minutes regenerator time. Conversely, going from 6O to 2O
minutes regenerator time had no effect at the 60 minute desulfurizer level.
Another effect observed was that attrition rates were highest at the short gas
desulfurizer residence time, 2.3$ average for 2O minutes versus 0.8$ for 60
minutes.
It is believed that deactivation of stone takes place primarily in the
gas desulfurizer due to crystal growth and particle sintering. This would
explain the dramatic effect of time on activity. The results of the above runs
are highly significant since they indicate the potential for more than doubling
the stone's capacity for sulfur absorption. It would be desirable, however, to
repeat the runs in a continuous operation.
F. Lab-Scale Batch Acceptor Cycling
Details of the construction of the batch cycler were described in Section
IV-F. The unit was set up to allow a small sample to physically be moved from
the sulfiding to the regenerating zone of a reactor which maintained dual
conditions at all times.
Two runs were made at the conditions shown in Table X-19. Results were
based on the weight changes which occur as the acceptor was cycled between the
carbonate and sulfide forms. The reactor internals were subject to a high
temperature corrosive environment which alternates between mildly oxidizing in
the regenerator to highly reducing in the gas desulfurizer. Metal corrosion
products and "scale" formed. The "scale" can distort the sample weight if it
falls into the holder.
The problem was aggravated by the inability to obtain quick delivery of
type 446 stainless steel wire which was needed for the sample support. Type
310 stainless steel wire was used until the middle of Run CR2, when Kanthal A-l
wire was substituted.
207.
-------
FIGURE X-7
Run A128 Deactivation of
CaS at 7O4°C (130O°F)
2.5 3
5 6 7 8 9 10
Number of Cycles
H 'I 10
208.
-------
TABLE X-17
Runs A129
Conditions and Results
System Pressure:
Run Number
Temperature, 6C
Type of Dolomite
Acceptor Size Consist, Tyler Mesh
Nominal Feed Rate, gm/hr (half-
calcined basis)
Nominal Run Time, min.
Input, SCFH(0
Recycle to Bed
H20
CO
C02
N2
H2S
Purges (C02) to Bed
Purges (N2) above Bed
Output
Exit Gas Rate, SCFH (dry basis)
Composition, Mol ^
CO
CO 2
H2S
Outlet Gas, Top of Bed
Compos it ion t Mol 'p
H20
H2
CO
C02
H2S v
Flow Rate, SCFH, Top of Bed
Fluidir.in? Veloritv ft/s*»r.
to A132 -
for Gas Desulfurizer
15 atm (206 psig)
A129, A130 A131, A132
« 871 (l6008F) v
20 60
t 0-1 . .
« 07 ,
* 00 ,
9T ._, ,,., i
« fifi V
A ^ rt
tf e _
4 - IS i
4. 1 4° V
^ 17 V
4> 17 i
« o >
i 56 1
A P '
4 17 \
4 17 l
4 n AX *-
Attrition.
of Feed
2.7, 1.8
0.97, 0.61
Input is given after shift reaction and hydrolysis of CS2 have taken place.
209.
-------
TABLE X-18
Runs A129 to A132 - Conditions and Results for Regenerator
System Pressure: 15 atm (206 psig)
Run Number
Temperature, °C
Nominal Run Time, min
Input, SCFH
Recycle to Bed
H20
H2
CO
C02
Purges (C02) to Bed
Purges (N2) above Bed
Purges (C02> above Bed
Output
Exit Gas Rate, SCFH (dry basis)
Composition, Mol %
H2
CO
C02
N2
H2S
COS
Outlet Gas, Top of Bed
Composition, Mol %
H20
H2
CO
C02
N2
H2S
COS
Flow Rate, SCFH, Top of Bed
Fluidizing Velocity, ft/sec
A129, A129A, A131 A130, A132
« 704 (1300°F) ^
20
- 35
-47
5.0
— 0
—70
-10
— 0
-15
-97
2
-2
166 --
60
•""•
04
0_._, ... , .. .
0 .'
210.
-------
FIGURE X-&
Runs A129, A129A - Deactivation of
CaS at 704°C (130O°F)
Regeijierat ipn ;
Mol Cals Cap Fed
. L_ - , -. -J
a I) 7 8 9 10 1.3
Number of Cycles
211
-------
FIGURE X-9
Runs A130 to A132 - Deactivation of
CaS at 704°C (130O°F)
-- ••-
fcfc 1
I
q
fs : •
x 101
t
i
; "
% Rege
Mols C
100 M
nerat
aCOj
oli C
^
• - - i .--.-.
I
: J"
i • •
j
, J
?
1
.....
.on
'rodxi
iS to
s^
ced
1
j-
i
i
r™*
.~
i
N
S
i
C 1
--
--
• . . i
30
i
"S
—
_
s
i
^
M'
i
*s
t
i
— -
s
—
V
/
^k^
>
)
- -
^s
.....
i
N,
-
**
^-
i
f
I
b
-
\^
S;
......
T'
-
mbol
. *
*
4
•"in
A
A
- •
i
1.3 i 2..') 3 4 5 6 7 8 ) lu I.D 'I i
...
|
in
> •
L3Q
^L
L32
- -
1
1,
; i • i
Safe iteft.
•
--
JO
30
m
[
I
I
i
j
i
I
— jl+~
j i
j
i
i '
j
i
._.
_-
.,:
—
i
....
:
t
i
L_L
nut
. J
es
gfi_
60
2O
60
i
..i _
1
—
. __
— i
.
H
1
..
-_.
-
Number of Cycles
212 .
-------
Run CR1
In Run CR1 the conversion as judged by the sulfided sample weight loss
never reached the equivalent of 98+$. This was apparently due to the presence
of scale in the sample. As our experience has been that the acceptor does
become completely sulfided, conversions are normalized by the addition of a
constant correction factor of 13 mol $, which brought the maximum observed
conversion to 99.5$ CaS. Data for the 6-1/2 cycles are presented in Table X-20
and Figure X-10. The relative percent regeneration (mols CaC03 formed per 100
mols of CaS fed) decreased from over 50$ in cycle 1 to 18$ in cycle 6.
Run CR2
In Run CR2 greater care was exercised to keep scale out of the sample.
Some scale inevitably got in, and this was removed magnetically in cycle 9 and
after cycle 22. After cycle 22, the sample was removed from the reactor and
about one-third of it was missing. The cause of the loss is unknown; the
sample may have been jarred or some material might have "popped" out as a result
of the chemical cycling.
The sample was cleaned of scale, and then part of it was analyzed for
sulfide sulfur content by decomposition in an acidified iodine solution followed
by titration with thiosulfate. The regenerated sample, Run 22R, was found to
contain 92.9$ CaS. Cycle 23S was run, and (by titration) the CaS content had
risen to 99.6$. The change corresponded to 6.7$, and the change via the weight
loss was about 8.3$. The data are presented in Table X-21 and Figure X-10.
Examination of the data shows that up until scale was removed in cycle
9 calculated conversions to CaS are apparently lower than the true conversion.
The CaS form represents a weight loss compared to CaC03, and it appears that
scale in the sample made the weight loss insufficient. When the scale was
removed, the weight loss showed a conversion over 100$. Either some dolomite
was removed with the scale or had been lost earlier and the loss was offset by
scale falling in. While the listing of conversions over 1OO$ is awkward, there
is no scientific way to normalize the data, and the raw results have been
reported. The last cycle results are believed to be accurate and help to pin
down the shape of the deactivation curve shown in Figure X-10.
Batch lab-scale data are compred to data from the continuous unit in
Figure X-10. Considering the difficulties encountered in the lab-scale work,
the results are in reasonable agreement. Additional development work would be
needed before the lab unit could be considered capable of generating precise,
accurate data.
The screen analysis of the sample after the run is given in Table
X-22. Some slight fines generation (3$ -65 mesh) apparently occurred upon
cycling even without the rough treatment of a pneumatic lift line or fluidized
bed as occurs in the continuous unit.
213
-------
TABLE X-19
Conditions for Batch Acceptor Cycling Reactor
Pressure: 15 atm (206 psig)
Charge: 0.4 grams Canaan Dolomite, 20 x 28 Tyler Mesh
Gas Desulfurizer Zone
Time, minutes
Temperature, °C (°F)
Input, SCFH
C02
5% H2S-H2 Mix
N2
Input After Shift, Mol %
CC-2
H2
CO
H20
H2S
N2
Regenerator Zone
Time, minutes
Temperature, °C (°F)
Input, SCFH
C02
H20
45
871 (1600)
1.73
2.10
6.24
7.5
10.2
9.65
9.65
1.0
62.0
45
704 (1300)
3.25
1.75
214.
-------
TABLE X-20
Results of Lab Scale Cycle Test CR1
.a b c = b + 13 d = 10° a
Change in Mol % Mol % CaS c
Conversion, CaS in in Sample, Relative Percent
Cycle No. Mol % Ca Sample Normalized Regeneration
IS 49.2 49.2 62.2
1R 32.4 16.8 29.8 52
2S 56.0 72.8 85.8
2R 46.5 26.3 39.3 54
3S 49.7 75.9 88.9
3R 34.4 41.5 54.5 39
4S 38.5 80.0 93.0
4R 25.8 54.2 67.2 28
5S 30.5 84.7 97.7
5R 17.8 67.0 80.0 18.2
6S 19.5 86.5 99.5
6R 17.8 68.7 81.7 17.9
7S 16.7 85.4 98.4
215.
-------
FIGURE X-1O
Lab-Scale Batch Acceptor Cycling
Sulfidation: 871 °C (l600°F)
Regeneration: 7O4°C (13OO°F)
'' i
<
2,5
2
to
9
8
c
4
2.5.
2_
1.5
:
—
.... . ....j ,
i
— — ~ -p— |
< - J T - - -
^ •}••••
r
: : .
i ... , .
y& Regeri
Mulu Ca
IOO Uol
*;!<#»
|jj
<
erati<
COg. fl
3 Cai
"N^_
> ***
in,
jrmec
*"«d
1^
j
"-ss
4
t ! " 'i ;i
..... i . 1
.*. .!.[
—
t
•^
9
llii
t . ' :
•I : ,
Iji;
3
"*>
,
: r: :
lih
i Hi
dii
•j
D:
j
1 1
ii
nTl
0*»
*s,
f'i
fit)
i! i
n -i
1
s
i
•%
;
<
iijj
q
>
(
Mjj
ii'-
i ' .
j
Sii
5
i ; : :
1 1 •
^K
i
*>
t
::i
|
' J
i!*'-'
: : <
. : j. j • :
. : . " i : l :
. : 1 :
1 i • : ' ! i
; : Lid
. , i . .
—
|
"^
a\
^
V
: 1 ; :
I : : : :
1
i
w.
j{l:ii
¥
iii
. in
! '
i ' !
li
! i i
*
*^s.
:;
o
' x ]
O1
...
— -,-Ke
< Jjlun
«ri
tun
'
Jftn
C
•«^
^%
;ii!
iar
ant
^
s
i
i
! i
1_
c:
"C
p
in
C
"^^
i
—-
• - ••
i
11
12
D!
no
jrc
*^s
"• —
1
>a
18
le
»«.
_i
it
15
d
•-i-
: 1
•
e
. 1 -
.. J-i-
r ~
--
—
...
—
• -
1.5
2 2.5 3
6 7 8 9 IU 1.5
Number of Cycles
216 .
-------
TABLE X-21
Results of Lab Scale Cycle Test CR2
Change in Conversion, Mol % CaS in
Cycle No. Mol % Ca(1> Sample
IS 76.9 76.9
1R 34.8 42.0
2S 52.2 ' 94.2
2R 28.4 65.8
3S 26.7 92.5
3R 21.4 71.1
4S 21.4 92.5
4R 20.2 72.4
5S 20.3 92.7
5R 19.5 73.1
6S 16.7 89.9
6R 16.7 73.1
7S 15.5 88.6
7R 15.3 73.7
9S 13.1 86.4
Removed Scale 104.1
9R 9.2 94.9
US 10.8 105.6
11R 10.8 94.9
13S 9.5 104.4
13R 11.1 93.3
15S 13.3 106.6
15R 9.2 97.4
17S 11.6 108.9
17R 7.0 101.9
Removed Scale
23S — 99.6(2>
23R 6.7 92.9(2)
In 1R the relative regeneration was 45.3%. For all
other cycles, the relative regeneration was assumed
equal to the change in conversion.
(2) Based on sulfur analysis of the solids. The change
in conversion based on weight difference was 8.3%.
217.
-------
TABLE X-22
Screen Analysis of Run CR2 After 22 Cycles
Starting Size: 20 x 28 Mesh
Screen Size, Average Particle Weight
Tyler Mesh Size, Microns Percent
20 x 28 ' 681 44
28 x 35 503 44
35 x 48 356 7
48 x 65 252 2
-65 208 3
218-
-------
XI. HOT REMOVAL OF PARTICULATES AND ALKALI
Introduction
The low Btu gasification process being developed is intended for application
to combined cycle power generation. The gas turbines involved are subject to
corrosion by alkali metals volatilized from coal and dolomite and to erosion by
particulates. Presently it is proposed to remove particulates and alkali fume
at 704°C (1300°F) by means of high-pressure drop cyclones. The use of hot
filters remains as another viable but less desirable possibility.
It is not known how much alkali will be volatilized from the coal as it
goes through the gasification process, but the fraction will certainly be small.
There was no significant change in the sodium and potassium contents of Loveridge
ash going through devolatilization, a gasification kinetics study and finally
the carbon burnup operation.!1) Conoco experience has been that there are no
noticeable changes in the alkali content of either ash or dolomite in the Rapid
City gasification pilot plant. For Illinois coals, alkali metals are found to
be refractory in high and low temperature ashing and "remain entirely in the
ash."(15) Fluidized bed boiler studies by Westinghouse(16) assumed that 1$ of the
coal and dolomite alkali ended up in the gas phase. However, a detailed discus-
sion of the basis for the assumption was not given.
In our process the ratio of coal to makeup dolomite is 50:1, so alkali from
dolomite is expected to be of minor importance. Overall removal of > 99.95$ of
the coal alkali is thought to be required to make the product fuel gas suitable
as a gas turbine feed.
The experimental setup was detailed in Section IV. Briefly, the standard
gas desulfurizer-regenerator cyclic run was made, and then a programmed amount
of alkali-containing material was added to the acceptor fed into the desulfurizer.
The gas exit stream was kept hot until after it was filtered on a stainless steel
sintered porous plate. The downstream effluent was then analyzed for alkali.
Runs FR1A and FR1B
Ordinarily, the gas desulfurizer is run with a gas feed preheat coil in the
upper part of the bed. The preheat coil reduces the temperature of the outlet
gas as it heats the inlet stream. This difficulty could not be overcome by
adding more heat to the reactor walls as they tended to reach their operating
limit. As a compromise measure, the first run, FR1A, was made with a one-half
length preheat coil. While the problem of the high wall temperature was elimi-
nated, the temperature in the gas space above the bed was too low 649-704°C
(l2CO-13OO°F). The exit line from the reactor to the hot box filters was only
427°C (8OO°F). Detailed run condition and results are presented in Tables XI-1
to XI-3.
The procedure was to first circulate Canaan dolomite for several sulfiding
cycles in order to evaporate any volatile alkali, and to have a bona fide
sulfided stone in the bed. When the alkali feed was started, the system was
switched to a separate filter and condensate system. After shutdown, the
various pipes were rinsed with distilled water and dilute nitric acid to leach
any deposited alkali. The condensates and washes were analyzed for alkali via
atomic absorption (AA).
219.
-------
TABLE XI-1
Conditions and Results for Gas Desulfurizer
System Pressure
Run Number
Temperatures, °C (°F)
Bed
Space above Bed (Minimum)
Flange
Exit Line to Hot Box
Hot Box Filter
Solids Feeds, gm/hr
Sulf ided Canaan Dolomite,
35 x 48 Tyler Mesh
Raw Tymochtee 11 Dolomite,
-150 Tyler Mesh
Hillsboro Low Temperature Coal
Ash, -150 Tyler Mesh
NaCl
Gas Input, SCFH(*)
Recycle to Bed
H20
H2
CO
coa
HaS
Air
Purges (CO2) to Bed
Purges (Na) above Bed
Purges (N2) to Bed
Na Lift Gas to Bed
Outlet Gas Top of Bed.(2) Mol $
HaO
Ha
CO
COa
Na
HaS
Flow Rate, SCFH, Top of Bed
Flu idling Velocity, ft/sec
Duration of Fines Feed, hours
Removal of Feed + Recycle Sulfur, %
% H2S in Outlet/Equilibrium % HaS
Conversion of Acceptor/Pass, mol %
: 15 atm (206
FR1A
<•• • "
660 (1220) «—
427 (800) <—
2000
26.4
2.1O
0.084
23
33
26
7.1
O.O
2.1
0.0
/
94
1O
16
15
7.5
53
O.02
292
0.,42
6.5
97.4
0.9
15
PSig)
FR1B
fl71
— 760
Kfl
• \34JL
— 704
« *7r\A
1930
25.5
2.O3
0.081
««
0.02
321
0.46
6.2
97.3
0.6
16
FR2 FR3
i Ron ^ — ».
14OQ^ fc
1 3OO1 . . - »•
2O50 2030
27 0 (1°'2
'•° (Canaan 4
0.0 0.0
0. 43 0. 43
36 *
22 -*
. O •>
.2 m
8-1 • \
Ti »
5k
lOfl in
11 *
10 *
GO •*
0.02 0.03
321 321
0. 46 O. 46
4.9 5.9
97.3 95.9
0.6 0.9
16 16
'i) Input is given after shift reaction and hydrolysis of CS2 have taken place.
[a) Calculated to fit shift equilibrium.
220 .
-------
TABLE XI-2
Cond it ions and
Results for
System Pressure; 15 atm
Run Number
Temperature, °C (°F)
Nominal Solids Residence Time, min.
Input, SCFH
Recycle to Bed
Hf\
2°
H2
f~tf\
CO
C02
Purges (C02) to Bed
Purges (N2) above Bed
Purges (C02) above Bed
Output
Exit Gas Rate, SCFH (dry basis)
Composition, Mol %
H2
CO
C02
N2
H2S
COS
Outlet Gas, Top of Bed
Composition, Mol %
H20
H2
CO
CO 2
N2
H2S
COS
Flow Rate, SCFH, Top of Bed
Fluidizing Velocity, ft/sec
$ H2S in Outlet/Equilibrium $ H2S
Regeneration of Acceptor /Pass, mol *)
FR1A
Regenerator
(206 psig)'
FR1B
FR2
FR3
^ ir\A / 1 orw-^
50
*
*
*
*
*
*
^
*
50
4.5
6.1
86.8
0
2.54
0.1
28.5
3.6
4.9
61.0
0.0
2.03
0.09
175
0.48
O.59
6 9.4
49
an
ell
1 D
Of\
n f\
1O
Of\
.u
1 (^
45
4.9
3.5
86.7
1.3
2.92
0.18
32.0
3.8
2.7
57.8
1.0
2.23
0.14
176
0.49
O.61
9.5
44
47
4.4
3.2
88.7
1.2
2.09
0. 14
32.5
3.3
2.5
59. O
0.9
1.58
0.11
181
O. 50
O.42
7.3
50
^
_
^-
r
^
44
4.4
3.3
88. O
1.1
2.59
O. 17
31.0
3.5
2.6
59.7
0.9
2.01
0.13
172
O. 48
O.55
8.5
221
-------
TABLE XI-3
Material Balances - Alkali Fume Removal Runs
Note: For accounting purposes more
significant figures are given
than is warranted by their
intrinsic accuracy
A. Run FR1A
Feed . grams
Tymochtee 11 Dolomite, -ISO Hesb
Low Temperature Ash, -ISO Hesb
NaCl
Attrition at O. B%('^
Total
Recovery Downstream of 7O4°C Filter, grams
Condensates
Wash of 2OO° C Filters
Total
% of Feed
Mols/mol wet gas
Recovery Upstream of 7O4°C Filter, grams
Filter
-ISO Mesh fines in Beds
Total
NJ
NJ
Ni
Feed, grams
Tymochtee 11 Dolomite, -ISO Mesh
NaCl
Attrition at O.8#')
Total
Recovery Downstream of 7O4°C Filter, grams
Condensates
Wash of 2OO°C Filters
Total
of Feed
Mol/mol wet gas
Recovery Upstream of 7O4°C Filter, grams
Filter
Line to Filter
-ISO Mesh Fines in Beds and Hoppers
Total
Grand Total Recovered, grams
% of Feed
Na
.3315
.1196
.2145
.O1O4
.676
. OOO9
.OO12
.OO21
O.3
4. 1 x KT8
Na
. 25O8
.8264
. OO8O
1.O85
< .OOO1
.OOll
.OO11
O.I
2.6 x 10~8
K
.3276
.1729
—
. 02O8
.521
.OOO4
. OO16
.OO2O
O. 4
2.3 x 1O-"
C. Run
K
. 25O8
O.O
.O16O
.267
.OOO2
.OO15
.OO17
O. 6
2. 1 x 1O"8
Cl
O. 1201
. OO39
.3315
.O1O4
.466
N.D.
N.D.
—
—
—
FR2
Cl
.O924
1.2737
.OO80
1.374
. 169O
. OO94
.1784
7.9
2.5 x 1O~
F lives
13O.o(O
13.7
O.55
1O4.O
248
0.0
O.O
0.0
0.0
—
85. 4
N.D.
85.4
Fines
1OO.O
2.1
80. 0
182
O
O
O
O
6
84. O
15.5
56.8
156.3
156
86
Feed , grams
Tymochtee 11 Dolomite, -ISO Mesh
Low Temperature Ash, -150 Hesh
NaCl
Attrition at O.S%(2>
Total
Recovery Downstream of 7O4" C Filter, grams
Condensates
Wash of 2OO" C Filters
Total
$ of Feed
Mols/mol wet gas
Recovery Upstream of 7O4°C Filter, grams
Filter
Line to Filter
-ISO Hesb fines in Beds and Hoppers
Total
Grand Total Recovered, grams
% of Feed
Feed, grams
Canaan. 4 Dolomite
NaCl
Attrition at O.B%
Total
Recovery Downstream of 7O4°C Filter, grams
Condensates
Wash of 2OO° C Filters
Wash of Line 7O4«C Filter to 2OO°C
Filters
Total
% of Feed
Mols/mol wet gas
Recovery Upstream of 7O4°C Filter, grams
Filter
Line to Filter
-ISO Hesh Fines in Beds and Hoppers
Total
Grand Total Recovered, grams
% of Feed
Na
.3060
. 11O4
. 198O
.O096
.624
. OO16
.OOO9
.OO25
O. 4
5.2 x 1O-"
Na
.OO6O
.9916
. OO96
1.OO7
< .OOO2
.O012
.0010
.0022
O. 2
4.7 x ur8
K
. 3O24
.1596
-—
.0192
.481
.OOO7
.0013
.OO2O
O. 4
2.4 x 1O~8
D. Run
K
.012O
—
. O192
. O31
.OOO7
.OO73
.OO23
.O1O3
33
13 x 10-"
Cl
.11O9
.OO36
. 306O
. OO96
.430
. 1O2
N.D.
. 1O2
6.8
1.4 x 1O-"
FR3
Cl
.OO6O
1.5284
.0096
1.538
.2556
.OO83
N.D.
.2639
17
3.7 x 10-"
Fines
/ , \
12O.O11)
12.6
.50
96. O- -
229
O
O
O
O
—
83.1
19.5
80. 1
182.7
183
8O
Fines
46.6
2.5
96. 0
145
O
O
O
O
~
41
6
54.8
96.4
96
66
(i) MgO-CaCO, basis.
(a) Estimated.
N.D. - Not Determined.
-------
For the first run, NaCl, Tymochtee 11 dolomite fines and low temperature
ash were added to the recirculating stream of acceptor. The low temperature ash
was made batchwise using the gas desulfurizer vessel at a temperature of 538°C
(lOOO°F) or less. The run was made using sodium chloride doping equivalent to
1$ volatilization of alkali in a coal containing 0.17$ sodium. Dolomite fines
and ash were fed at a rate equivalent to 1$ of the feed of these materials to
the gas desulfurizer and gasifier, respectively.
Recovery of fines in the hot box filter was poor, and at the same time there
was a slight increase in the weight of solids comprising the circulating in-
ventory. It is believed that the proximity of the solids inlet line to the
reactor outlet caused some of the fines to be trapped and carried into the
regenerator. It is not known how this affected the overall alkali removal
efficiency.
Based on the condensate recovery rate observed in Run FR1A, the gas compo-
sition corresponded to shift equilibrium of about 7O4°C (l3OO°F). It is believed
that due to the presence of dolomite and ash fines combined with a modest cool
down of gas leaving the reactor rather than the usual quench to about 260°C
(500°F), the gas actually did shift down to a lower temperature equilibrium level.
The data for gas composition given in Table XI-1 are therefore calculated to give
the composition corresponding to shift equilibrium above the bed.
Results for alkali removal are given in Table XI-3. These data taken at
face value indicate that alkali removal from fuel gas presents no problem. How-
ever, temperatures above the bed and in the exit line were not high enough, and
this may have resulted in the deposit oi sodium upstream of the filter. Not
every connecting line downstream of the filter was rinsed, and something may have
deposited on the walls. Glass jars were used for sample storage, and it was
later found that several tenths of a ppm of sodium could be leached out of a
glass sample container using only distilled water.
The one-half length preheat coil was then removed from the gas desulfurizer,
and air was added to the feed gas to provide the needed extra heat. In addition,
the exit line was rewound so that it could be brought to a higher temperature.
With these changes, Run FR1A was repeated as FR1B. Temperature traverses showed
the minimum temperature in the gas space above the bed to be about 760°C (1400°F).
The exit line to the hot filter was maintained at 704°C (l300°F). Based on
analyses of condensates and washes of process lines, approximately 99.6$ of the
feed sodium and potassium were removed, and the molar ratio of Na or K in the wet
product gas was calculated to be on the order of 5 x 10~8.
Run FR2
It was felt that the ash fed in with the salt in Run FR1B might have reacted
with the alkali feeds which would have accounted for the high removel efficiency.
Run FR2 was made with a five-fold increase in salt input and no fine ash feed in
order to provide a more severe test. The results were about the same. The molar
ratio of Na or K in the gas above the bed was estimated to be 2 x 1O~8, and the
removal efficiency was better than 99$.
Run FR3
Tymochtee dolomite is a relatively impure stone containing clay minerals.
As these could react with alkali compounds, a more severe test for alkali removal
was to make a run in which Canaan dolomite fines instead of Tymochtee dolomite
223.
-------
fines were fed to the unit together with salt. This was done in Run FR3, but
again, there was almost no sodium or potassium detected in the effluents down-
stream of the hot filter. However, whereas the potassium levels in the,gas
had been 2 x 10~8 in Runs FR1A to FR2, the figure for Run FR3 was 13 x 10~8.
This may have been due to the lack of Tymochtee dolomite fines in the system.
On the other hand, some volatilization of potassium from the Canaan inventory
may have occurred which would have added to the totals.
Discussion
It is believed that a lot of the fines were trapped in the desulfurizer
bed and carried down to the regenerator. Table XI-3 shows that the -150 mesh
material in the beds must be included in order to obtain a reasonable material
balance on the fines in the system. The issue is complicated somewhat by the
need to make an estimate of the contribution to the fines total made by
attrition of the Canaan stone inventory.
The data in Table XI-3 indicate an enrichment of the effluent gas in
chlorine relative to sodium. These results are summarized in Table XI-4. The
molar feed ratio was about one in all cases while the ratio for the gaseous
effluent was 27 for FR1B and about 100 for FR2 and FR3. It is assumed that in
the gas desulfurizer chlorine is liberated as HC1 which is then washed out in
the condensers. The condensates generally displayed a pH of about 3, a rather
acid condition. The question remains though, what happens to the sodium half
of the NaCl couple? Given below is a reaction which would yield HC1 and a
potentially nonvolatile sodium compound. It is noted that we have not investi-
gated the associated thermodynamics. Similar reactions can be written for
potassium.
2 NaCl + H20 + 3 Si02 = Na20'3 Si02 + 2 HC1.
The data are not significant quantitatively because of their scatter.
However, as far as they go, the data support the contention that the alkali
did not escape the hot filter.
Obviously more work is needed. As far as we know, the above experiments
were the very first to attempt to come to grips with the question of how much
alkali actually would appear in the product fuel gas of a process which employs
a hot cleanup technique.
TABLE XI-4
Molar Ratios in Effluent Downstream
of Hot Box Filter
Run No. K/Na Cl/Na
FR1B 0.47 26.5
FR2 0.91 105
FR3 2.8 78
224.
-------
XII. CHANCE REACTION STUDIES
A. Introduction
The Chance reaction was commercially practiced as early as 1888(17) to
remove sulfur from alkali wastes via the reaction,
CaS + H20 + C02 = CaC03 + H2S (XII-l)
It is similarly applied to our process in order to remove sulfur from spent CaS
so that it may be disposed of without generation of noxious odors. A Chance
reactor is now in place at Conoco1s Rapid City pilot plant where it is desulfu-
rizing spent dolomite from the C02 Acceptor Process which contains a low level
of sulfur. The lab study first examined the possibility of using Chance
reaction product as a sulfur acceptor to eliminate the requirement for makeup
stone, and then went on to survey the rate and efficiency of the reaction for
disposal of spent acceptor in order to provide an improved design basis for the
commercial use.
B. Discussion
Reactivity of Chance Reaction Products as Sulfur Acceptors
The first few runs were made in order to obtain a bona fide Chance
reaction product which could be evaluated for its activity as an H2S acceptor.
It was hoped that a process could be devised whereby spent stone could be reused
indefinitely thus obviating the need for fresh make-up stone. The product was
prepared by running the Chance reaction in two stages:
Hydrosulfidation, CaS + H2S = Ca(HS)2/ s, (XII-2)
and Carbonation, Ca(HS)2/ •> + C02 + H20 = CaC03 + 2 H2S. (XII-3)
It was then tested in a thermobalance at atmospheric pressure using an
H2S-H2-C02 mix as detailed in Table XTI-1.
The thermobalance study showed that even finely divided precipitated
CaCO3 was not an active H2S acceptor. However, the same material after calcin-
ing to CaO was active toward H2S.
The conclusion is, therefore, that the use of the Chance reaction
product in an entrained-type dilute phase reactor, is only potentially useful
as an H2S acceptor if the operating temperature is high enough to effect
calcination of the CaC03. Whether the kinetics are suitable for this type of
process requires more definitive rate data than were obtained in this study.
Study of Chance Reaction Variables
The next step in the program was a rapid survey of the principal
variables controlling the rate of reaction. A detailed kinetic correlation of
the data was not attempted. CO2 reacts not only with CaS or Ca(HS)2 but also
with MgO according to the reaction,
225,
-------
TABLE XII-1
KJ
Conditions and Results for Sulf idation of
Chance Reaction Products
Run
No. Feed Material
CHI Fisher CaO
CH2 Fisher CaS, as received
Hydrosulf id ing Carbonat ion
Temperature, Temperature,
°C (°F) °C (°F)
Ambient Ambient
60 (14O)
Time for 9O%
Reaction of Product
Product withHaS.O) minutes
Surface CaCOs Form CaO Form
Area, m2/g at 9OO°C at 98O°C Remarks
4.1 (2) 65
.N.D. (2) N.D. 1OO$ Recovery of
CH3 Run A7 Gas Desulfurizer 90 (194)
Sample, 1O cycles,
-2OO mesh
CH4 Fisher CaS, as received 88 (19O)
Fresh Tymochtee Dolomite,
28 x 35 mesh
9O (194)
88 (190)
27.4
4.8
63
N.D.
60
5O
CaCO3 from CaS
Recovery of
CaCO3 from CaS
(i) Composition of treating gas: 0.95 vol. % H2S, 12.75 vol. % H2 and 86.3O vol. # C02.
(2) The initial progress of the reaction indicated that conversion would stop before
90$ was achieved.
-------
MgO(g) + C02(g) + H20(l) = Mg(HC03)2(aq),
at low temperature, and the reaction,
Mg°(s) + C°2(g) + H2°(l) = MgC°3(s)'
above 40-50°C. The mechanism of the Chance reaction itself is uncertain with
respect to whether it is mainly a single-step reaction,
CaS(g) + H20(i) + C02(g) = CaC03(g) + H2S(g),
or a two-step reaction,
CaS(s) + H2S(g) = Ca(HS)2(aq),
Ca(HS)2, . + C02 + H20 = CaC03 + 2 H2S.
\ atl )
The time and equipment available were insufficient to provide a complete study
of the system, but show at least the directional effects of the major variables.
At the beginning, the sulfur content of the final dry product and
the H2S content of the exit gas as a function of time were determined. In later
runs, the sulfur content of the slurry as a function of time was also followed.
The data are summarized in Tables XII-2 and XII-3. It is noted that the gas
compositions in Table XII-3 represent averaged values since it took about ten
minutes to produce a 90$ step change in composition at the flow rates employed.
Also, the slurry sampling procedure introduced a variation in the weight of
solids found in a given slurry sample. Three samples of the same slurry gave
5.34$, 5.88$ and 5.18$ solids at an actual composition of 5.88 wt $ solids.
The slurry analysis gave .both sulfide sulfur in the solid and H2S, HS~ and S=
dissolved in the liquid.
The parameters explored in the study were temperature, feedstock,
particle size, C02 flow rate and stirrer speed. The solids: liquid weight ratio
and stirrer speed had little impact on the interfacial area (judged visually)
over the range studied. The data in Tables XII-2 and XII-3 are inconclusive as
to the precise effect of stirrer speed, but it is clear that any effect is
relatively minor.
Runs CHS through CH7A studied the effect of temperature. In going
from 21 °C to 88°C, the rate increased dramatically, i.e., the residual sulfur in
the product dropped from 24$ to 0.35$. Because of this result, it was decided
to investigate only the higher temperature of 88°C (l90°F).
Reaction rate was also dramatically increased by decrease in particle
size as can be seen by comparing Run CH11 with Run CH13. Going from 20 x 35 mesh
to -2OO mesh reduced the sulfur content of dry product from 1.27$ to O.16$. A
similar trend was observed in going from 28 x 35 mesh to 35 x 65 mesh in Runs
CH16A through CH23.
The choice of feedstock also affected the reaction. The Run A6
product did not desulfurize as easily as Run A15 material. However, the A6
material had been cycled through both the C02 Acceptor Process regime and the
sulfur cycle, so it may not have represented a material which had seen only the
sulfur cycle.
227.
-------
TABLE XI1-2
Run No.
CH-5
CH-6
CH-7A
CH-8A
CH-9
CH-1O
CH-11
CH-12
CB-13
CH-14
CH-15
CH-16A
CB-17
CH-18
M CT-19A
ro
oo CH-2O
CH-21
CH-22
CH-23
- Feedstock
Canaan A-15,
Gas Desulfurizer
Bed
Tymochtee A-6,
Gas Desulfurizer
Bed
Canaan A-15,
Gas Desulfurizer
Bed
Chance Reaction: Conditions and Results ••
System Pressure: 1 atmosphere
Batch Feed: 25 g solids, 4OO ml H
Hydro- H3S . Carbons- C02 Run
Size, Sulfiding, Feed Rate, tion, Feed Rate, Duration,
Trier Mesh min. SCFH min. SCFH min.
28 x 35 — — 61
i
20 x 35 — —
16O .56
120 1
-2OO 60 I
D .1
°8 it 35 <«--.,- *
7 6O
120
180
12O
4 — — 60 2.01 60
35 X 65 — — 72 .87 72
6O 2.O1 6O
1 .87
2.O1
,, ' — T 2.O1
28 X 35 — ' — 84 .87 84
i60 2.01 6O
84 2.O1 84
,o
Wt. % S Wt. % S Soluble Stirrer
in Dry in Dry Material, Speed, Temperature,
Feed Product erams rom °C
25.95 23.86 .29 2:
1 21.83 .58
I .35 .35
6.82 5.38 .06
14.97 .11 i
5.46 .06 42
1.27 .27 3C
.76 .23 3C
8.11 .16 .23 22
5 21
49
88
8
7
5 8
5
5
15
25.95 .26 4.59 . 235
25.95 .00 .17 1
26.73 .03 .55
.03 N.D. V
.06 .30 34O
.13 1.78 340
1 .03 .26 275
25.95 .10 .72 235
1 .16 .44 235
\ .79 .48 275 ,
8
7
8
CO2 + HaS mixes; Total Flow Rate .87 SCFH
Min.
O-12
12-24
24-36
36-48
48-6O
HaS
99.4
74.7
50.0
24.7
.0
-------
Prior treatment by a hydrosulfiding step, i.e., by reaction (XII-2)
was helpful in treating such a refractory feedstock as can be seen by comparing
Runs CH10, CH11 and CH12, respectively. Run CH10 did not involve hydrosulfiding
prior to treatment with C02 while in CH11 and CH12, the slurry was first treated
with H2S for 60 and 120 minutes, respectively. Run CH12 showed the lowest
residual sulfur content of all three products in that series.
Run CH14 was made to simulate a possible gas environment of sulfided
acceptor going through a countercurrent Chance reactor system. It first would
see a very high H2S concentration, and at the exit it would see a very high C02
concentration. The sulfur content of the dry solids was reduced to 0.26$ from
26% even though a reduced level of C02 equivalent to only 1.25 times the esti-
mated stoichiometric amount was used. However, the soluble material left in the
liquid was higher than in any other run. This was attributed to Ca(HS)2 which
may not have seen sufficient C02 to be converted to CaC03. The soluble material
is usually assumed to be Mg(HC03)2 formed from MgC03 and residual CO2 as the
reactor is cooled.
During slurry conversion by C02, the appearance of the solid acceptor
changed from individual grains into fine particles. After filtration, the
solids had a fluffy appearance and occupied several times the bulk volume of
the feed. As the run progressed, the viscosity of the slurry increased (visu-
ally judged). This change in viscosity could influence the C02-slurry interface
and diffusion of C02.in the aqueous part of the slurry. Settling characteristics
of converted slurry were good.
C. Conclusions
It was shown that, on a dry basis, 95+$ sulfided spent acceptor containing
26$ sulfur can be converted to a 0.03$ sulfur form for disposal. The reaction
can be run at atmospheric pressure at about 90°C with a residence time of one
hour.
229.
-------
TABLE XI1-3
to
<_0
O
Run Number
CHS Gas*
CH6 Gas
CH7A Gas
CHSA Gas
CH9 Gas
CH1O Gas
CH16A Gas
Slurry**
CH17 Gas
Slurry
CH18 Gas
Slurry
CH19A Gas
Slurry
CH2O Gas
Slurry
CH21 Gas
Slurry
CH22 Gas
Slurry
CH23 Gas
Slurry
Chance Reaction;
Time, Min. -» O
__
—
—
—
—
—
—
1.57
—
1.57
—
1.57
—
1.57
—
1.57
—
1.53
—
1.53
—
1.53
Composition of Slurry and Exit Gas
12
2.77
1.68
29.0
0.50
O. 39
O.5O
25.2
1.28
18.1
.928
55.5
1.01
38.0
.561
22.2
.969
23.2
1.88
19.0
1.27
19.6
1.13
24
0.56
1.05
39. 0
0.27
O. 25
O.48
47. O
.863
18.1
.525
50.0
.385
4.34
.106
18.4
.222
38.0
.847
18.6
.477
14.2
.339 ..
36
0.25
0.8O
30.0
0.22
0.32
0.38
38.0
.317
19.7
.O943
21.4
.0579
0.10
.0285
1.14
.O494
30.8
.342
4.91
.102
0.41
.238
48
0.16
O. 72
17.0
0.22
O.31
0.35
17.1
.070
0.34
.O127
0.93
.0096
O.O5
.0261
O.OO
.0083
16.2
.O751
O.O9
.O353
0.38
.224
60
0.16
0.63
1.16
O. 23
O. 29
O. 35
0.09
.014
O.OO
.OO66
0.14
.OO6O
0.04
.0200
O.OO
.OO55
2.12
.0214
0.03
.O293
0.32
.165
72
84
0.11
,0037
O.27
.157
0.02
0.23
.147
**
Volume % of H2S in exit gas; balance CO2.
Grams of sulfur/lOO grams of slurry.
-------
XIII. PROCESS IMPROVEMENT STUDIES
A. Introduction
Process improvement studies were concentrated in two areas. The first
involved the use of a sorbent other than half-calcined dolomite. This was to
enlarge the number of suitable sorbents, to lessen the cost of makeup, or to
increase the range of the desulfurizer operating conditions. The other area
of study was to produce a liquid-phase Claus reactor feed without resorting to
burning sulfur and then recovering S02 in an absorption tower.
B. -Use of Fine Calcined Limestone
An alternate process concept for hot desulfurization of producer gas would
use very finely divided CaO. Product CaS would then be regenerated with water
and CO2 at low temperature via the Chance reaction. The advantages of such a
process are: The required make-up rate of fresh acceptor would be drastically
reduced since wet regeneration of CaS can be carried out to completion, and
the gasifier can be run hotter and more efficiently as the gas desulfurizer is
operated at calcining conditions.
Attempts to test the concept by using a slurry of precipitated chalk were
plagued with difficulties in feeding the slurried solids. Two runs were then
made feeding finely ground Nebraska limestone into an expanded bed of coarse
solids.
As shown in Figure IV-5, the gas desulfurizer was equipped with a feed
line to the bottom of the vessel through which the fine limestone was fed. The
reacted stone was elutriated from the coarse bed and collected in the empty
regenerator vessel which was equipped with filters to effect gas-solids separa-
tion. The run temperature was 927°C (1700°F) and the collection vessel was held
at 315°C (600°F).
Run conditions for the two runs are given in Table XIII-1, and assays on
the product fines are given in Tables XIII-2 and XIII-3. For Run LI, the bed
material was 28 x 35 mesh deadburned dolomite. Sufficient limestone was fed to
give 80-90$ conversion of solids. Approximately 93$ of the feed sulfur was
removed, but analysis of the bed material showed that the deadburned dolomite
had participated in the reaction. About 18$ of the CaO in the deadburned dolo-
mite was converted to CaS. The elutriated product had been 58$ converted to
CaS.
In Run L2, the coarse bed material was 35 x 48 mesh fused MgO which is
totally unreactive toward H2S at run conditions. Only 76$ of the feed sulfur
was reacted, and the exit H2S concentration was 0.28$. The product stone was
59$ converted to CaS. Shutdown of the run was conducted so as to immobilize
the bed in order to determine the fines content. Based on the fines content,
the fines residence time was estimated to have been 1.6 minutes.
Table XIII-3 gives the solids compositions as a function of particle size.
It is noted that in both runs there was a period in which the stone was fed
without H2S in the vessel. This was approximately 15$ of the acceptor in Run
LI and 8$ in Run L2. In all cases, there was CaC03 present, and the CaC03
content generally increased with decreasing particle size. The CaS content
231.
-------
TABLE XIII-1
Conditions and Results for Fine CaC03 Runs
Pressure:15 atm (206 psig)
Run Number
LI
L2
Temperature, °C
Bed Material
Size Consist, Tyler Mesh
Feed Material
Size Consist, Tyler Mesh
Feed Rate, gm/hr, Raw Stone
Input. SCFH(0
Recycle to Bed
H20
H2
CO
CO 2
' N2
H2S
Purges (C02) to Bed
Purges (N2) above Bed
Purges (C02) above Bed
N, Lift Gas to Bed
927 (1700°F) 941 (l725°F)
Dead Burned Dolomite Magnorite
28 x 35 35 x 48
« Nebraska Limestone *•
65 x 150 150 x 200
270 227
281
21.4
25.0
23.0
8.8
71.0
1.8
5.0
10.0
10.0
65.0
232
51.5
41.5
54.0
11.0
55.0
3.0
5.0
10.0
10.0
65.0
Output
Exit Gas Rate, SCFH (dry basis)
Composition. Mol ^
H2
CO
C02
N2
H2S
Outlet Gas. Top of Bed
Composition, Mol %
H20
H2
CO
C02
N2
H2S
Flow Rate, SCFH, Top of Bed
Fluidizing Velocity, ft/sec
Removal of Feed Sulfur, $
Removal of. Feed Plus Recycle Sulfur, %
% H2S in Outlet/Equilibrium % H2S
Conversion of Acceptor, Mol %
Estimated Acceptor Residence Time, min.
214
9.0
13.9
8.8
68.2
0.06
5.9
8.9
13.6
6.6
65.0
0.06
505
0.76
92.9
85.1
5.2,
77
256
16.5
21.0
11.4
50.8
0.28
10.2
15.5
19.7
8.7
45.7
0.26
521
0.79
76.0
62.4
12.0
66
1.6
(1) Input is given after shift reaction and hydrolysis of CS2 have taken place.
(2) Not available.
232.
-------
generally increased with decreasing particle size. The data are too meager
to do more than speculate on the mechanism which gave the above results.
It is concluded that although significant reaction of fine CaC03 (CaO)
took place, sufficient reactivity to prevent H2S breakthrough remains to be
demonstrated.
C. Generation of an Active Acceptor from Coarse Limestone
Two attempts were made to obtain active CaC03 in a porous matrix starting
with limestone. In Run A40, 35 x 48 mesh Rapid City limestone was calcined at
927°C (1700°F) in nitrogen and then recarbonated at 76O°C (1400°F) in C02. The
recarbonated stone was fed to the gas desulfurizer at 871°C (l6OO°F) with the
standard gas flows. While the bed was being fed in and sulfided, temperature
aberrations were observed and the run was stopped. Upon disassembly of the
reactor, a large agglomerate was found at the bottom of the vessel.
In Run A41, the stone was simultaneously calcined and sulfided at 927°C
(17OO°F) in nitrogen and CS2. The product was 86$ CaS, with some residual
CaC03 and "CaO". This product was then fed through the regenerator and put
into the unit for cyclic operation. Results are detailed in Tables XIII-4 and
XIII-5 and in Figure XIII-1. Pickup of H2S proceeded without breakthrough
until the CaC03 content of the stone was depleted. This occurred with 3-5$
CaC03 remaining in the stone leaving the gas desulfurizer. Attrition was only
0.3$. Regeneration was poor having dropped to ca. 5$ by cycle 3.
Run A49
For Run A49, a limestone feed was to be cyclically sulfided in the
calcined state at 927°C (1700°F) followed by standard regeneration at 704°C
(1300°F). Rapid City limestone (Minnekhata) was the feed for this run. How-
ever, in the presence of sulfur, agglomeration with attendant defluidization
was a continuing problem.
Sulfidation was first attempted in a cyclic operation using the gas
composition and temperature shown in Table XIII-4. However, fluidization could
not be maintained, and inspection of the bed after shutdown showed it to be
agglomerated. A feedstock of sulfided limestone was then prepared batchwise.
The bed was fed in and calcined using only nitrogen and C02 as the fluidizing
medium. Sulfidation was then carried out as shown in Table XIII-4. The
sulfidation was run at about 9$ per hour with satisfactory desulfurization of
the inlet gas until breakthrough at ca. 48 mol $ CaS in the stone. The beds
contained about 0.5$ of +20 mesh agglomerates.
Five batches were combined for cyclic operation. However, attempts
to run in this mode were unsuccessful because agglomerates continually choked
the transfer line between vessels. The run was abandoned at this point. The
attrition rate was qualitatively judged to be 1$ or less for the brief period
when sulfided stone was being fed to the unit.
D. Fully Calcined Dolomite
Runs A120-A122. A125-A127
Six batch runs were made at 927°C (l70O°F) under fully calcined con-
ditions. Detailed run conditions and results are presented in Tables XIII-4
and XIII-6. Table XIII-4 shows somewhat different conditions for A12O-A122 and
233.
-------
TABLE XII1-2
Particle Size Analysis of
Product Limestone Fines
Material
Percent 65
Size 10O
Consist, 150
Tyler 20O
Mesh
LI Product
-65 Mesh
x 100 16.6
x 150 38.6
x 200 30.3
x 325 5.9
-325 8.6
TABLE
LI Bed
-65 Mesh
84.8
13.7
1.0
0.3
0.1
XIII-3 '
L2 Product
-150 Mesh
— [l_
—
82.5
16.0
1.5
L2 Bed
-150
^
-
75
17
7
Mesh
Hm
-
.6
.1
.3
Analysis of Product Limestone Fines
Material
LI Product,
-65 Mesh
LI Bed,
-65 Mesh
L2 Product,
-150 Mesh
L2 Bed,
-150 Mesh
Size Consist,
Tyler Mesh
65 x 100
100 x 150
150 x 200
200 x 325
-325
whole -65 sample
65 x 100
100 x 150
150 x 200
200 x 325
150 x 200
20O x 325
-325
whole -150 sample
150 x 200
200 x 325
-325
CaCO,
5.9
11.0
20.9
31.2
34.8
13.9
2.1
3.0
11.3
13.1
18.7
19.6
31.1
17.9
35.6
22.3
5.9
Mol 1o
CaS
72.2
66.8
54.7
41.4
33.8
57.8
86.4
82.9
62.9
35.0
59.4
59.5
35.8
59.0
50.3
42.7
15.4
CaO
21.9
22.2
24.4
27.4
31.4
28.4
11.5
14.1
25.8
51.9
21.9
20.9
33.1
23.1
14.2
35.0
78.7
234,
-------
RUNA41 I- iDEAGTIVATION OF CaS AT 7O4°C (130O0F)
235.
-------
TABLE XII1-4
Conditions and Results for Gas Pesulfurizer
Run Number
Temperature, °C (°F)
Type of Stone
Acceptor Size Consist, Tyler Mesh
Nominal Feed Rate, gm/hr (hali1-
calcined basis)
Nominal Solids Residence Time, min.
Input.
Recycle to Bed
H20
Ha
CO
C02
N2
H2S
Purges (C02) to Bed
Purges (N2) above Bed
Purges N2 to Bed
Recycle Acceptor Lift Gas, above Bed
Output in Cycle
Exit Gas Rate, SCFH (dry basis)
Composition/ Mol i
H2
CO
C02
N2
H2S
Outlet Gasr Top oi Bed
Composition. Mol 't
H20
H2
CO
C0a
N2
' H2S
Flow Rate, SCFH, Top of Bed
Fluidizing Velocity, ft/eec
Attrition, % of Feed Rate
Duration of Circulation with E2S Feed, hr
Removal of Feed Sulfur, $
Removal of Feed Plus Recycle Sulfur,
H2S in Outlet/Equilibrium,
Conversion of Acceptor/Pass, mol
H2S
System Pressure; 15
_ A41
871 (1600)
•« Rapid
yp
2490
36
145
27
25
29
12
66
1.8
O
IS
5
70
3
164
18.8
18.5
8.7
53.7
0.16(a)
9.1
17.9
17.7
8.3
46.7
O.15
323
O.46
0.3
15
86
76
6.7
7.6
atm (206 psig)
. A49
927(1700)
City Limestone f-
oe u ^ia
Batch
330
50
27
4O
29
8.2
138
1.7
5.0
15
O
0
1
236
16.6
12.7
5.7
65.0
O.01
10.0
15.7
12.1
5.5
56.7
0.01
301
0.45
N.D.
Batch
99
98
0.5
48
A120-A122
63
11.4
13.1
12.4
3.1
138
1.0
4.4
15
0
0
202
6.4
6.6
4.4
82.6
4.8
6.5
6.7
4.4
77.6
263
O.39
A125-A127
•927 (1700)
• Canaan 3 Dolomite•
35 x 48
Batch
80
12.6
11.9
13.6
4.9
138
1.0
5.0
15
0
0
190
7
7
6
80
5
7
7
6
75
267
0.40
50
27
41
29
8.5
138
2.6
4.4
15
0
0
236
18
11
7
64
9
17
11
7
56
298
0.43
(i) Input is given after shift reaction and hydrolysis of CSa have taken place.
(2) Gas sample taken after breakthrough had occurred.'
-------
TABLE XII1-5
Conditions and Results for Regenerator
System Pressure: 15 atm (206 psig)
Run Number A41
Temperature, °C (°F) <• 704 (1300) >
Nominal Solids Residence Time, min. 43 60
Inputt SCFH
Recycle to Bed 0 35
H20 55 47
H2 6.O 5.O
CO 00
C02 94 70
Purges (CO2) to Bed 10 10
Purges (N2) above Bed 5 0
Purges (C02) above Bed 10 15
Output in Cycle 3
Exit Gas Rate, SCFH (dry basis) 136 97
Composition^ Mol %
H2 3.48 2
CO 1.17 3
C02 83.3 94
N2 11.3 0
H2S 0.74
COS 0.03
Outlet Gas, Top of Bed
Composition, Mol %
H20 31.0 30
H2 2.70 2
CO 0.91 2
C02 58.9 66
N2 5.9 0
H2S O.57
COS 0.02
Flow Rate, SCFH, Top of Bed 176 166
Fluidizing Velocity, ft/sec 0.49 0.46
% H2S in Outlet/Equilibrium % H2S 0.16
Regeneration of Acceptor/Pass, mol $ 4.9
237.
-------
TABLE XIII-6
Results of Sulfidatlon Runs with Fully Calcined Dolomite
Run Inlet, Ty
No. % H,S Do
Pressure: 15 atm (2O6 psig) Temperature; 927° C (17OO°F)
% Active % Active
CaO CaO Space Velocity
Bed Superficial Length % HaS Before % H2S - Remaining at O. lO'jt at Breakthrough
oe of Height. Gas Residence of Run. % of Ca Breakthrough, % H,S Equil. at End H,S in mols inlet H,S/hr Density Uol %
Lomite inches Time, sec min Sulf ided/hr ± .01 ± .01 of Run Dry Outlet mol Active CaO gm/ml lb/ft3 CaCOs CaS "CaO"
A120. .40 Canaan 3 14.2 3.3 6O8 9.4 .02 .01 5.4 6.9 1.37 2.O9 130. 7 5.4 89.4 5.2
A121 .39
A 122 .38
A125 .53
A126 .60
A127 .62
9.4 2.2 4O2 13.4 .01 O 9.2 1O.O 1.35 2.13 132.8 9.2 82.5 8.3
4.1 0.9 184 31.1 .01 O 5.9 8.9 3.47 2.O9 130. 4 5.9 88.1 6.O
4.3 0.8 108 55.9 .03 .01 7.8 13.3 4. 2O 2.O5 127.7 7.8 87. 0 5.2
1O.4 1.9 218 22.7 .02 O 2.O 3.6 7.23 2.O6 128.4 2.O 90.3 7.7
15.9 2.9 341 17.6 .02 0 0. 3 1.7 10. 4 2. 04 127.1 0. 3 9O.9 8.8
S3
00
-------
A125-A127. A stuck rotameter float caused the hydrogen flow in Runs A120-A122
to be low. Due to shift equilibrium, this caused the C02 level to be too high,
for which the operator compensated by reducing the C02 flow as well in order
to maintain a calcining condition.
The outlet H2S concentration was at equilibrium, and the H2S;CaO space
velocity at breakthrough was consistent with data reported for half-calcined
stone. However, considerable scatter in the space velocity results was
observed.
When the stone was assayed, CaC03 was found to be present. It is
assumed that recarbonation took place as the material was drained through the
C02-purged exit line.
Run A123
The products of Runs A12O-A122 were combined and then cycled batchwise
in the fully calcined regime. Run conditions for the gas desulfurizer and
regenerator are given in Tables XIII-4 and XIII-5, respectively. Residence time
in the desulfurizer varied as the stone was cycled and is, therefore, not given
under the conditions cited in Table XIII-4. Results of cycling are presented in
Table XIII-6 and in Figure XIII-2. High utilization of the CaO was obtained but,
as shown in Figure XIII-2, regeneration was poorer than had been observed in
continuous runs with half-calcined stone. It is believed that the long hold time
together with the elevated temperature in the gas desulfurizer in cycle 1 may
have contributed to the poor regeneration results.
A continuing problem with CaO based acceptors has been agglomeration.
At the end of the second sulfidation, weak agglomerates were observed. No
difficulties due to agglomerates were experienced in any other cycle.
It is concluded that it may be possible to run the gas desulfurizer
under calcining conditions. Potential problems are agglomeration and consumption
of C02 in the regenerator upon reaction with unconverted CaO.
E. Reaction of CaS04 with H2S to Produce Liquid Claus Feed
Introduction
In reference (l) a process was proposed in which the regenerator offgas
reacts with CaS04 to produce directly a liquid-phase Claus feed containing the
required 2:1 ratio of H2S/S02. A simplified flow diagram of the process is
shown in Figure XIII-3. Potential advantages are that the sulfur combustor, S02
absorption tower, liquor circulation piping, and liquid-liquid heat exchangers,
all shown in Figure III-3, would be eliminated.
In the process, CaS in the spent acceptor is oxidized with air to form
CaSO4 which reacts with the regenerator offgas at about'16OO°F. After reaction,
the acceptor is processed through the Chance reactor before being discarded.
The principal reaction is,
3/4 CaS04 + H2S = 3/4 CaS + SO2 + H20 (XIII-l)
239.
-------
FIGURE XII1-2
RUN A123 DEACTIVATION OF
CaS AT 704°C (1300° F)
1.5
2 2.6 3
5 7 8 a 10 1.5
Number of Cycles
8 9 10
240-
-------
Routine operation of the regenerator has shown that other reactions also occur
which produce small amounts of elemental sulfur, COS, CO, and H2- A brief
experimental program was done to test the feasibility of the process.
Experimental
A batch of Tymochtee dolomite was sulfated using the same techniques
described in Section IX. In a preliminary experiment (Run A83, not otherwise
discussed here) large amounts of elemental sulfur formed which plugged the
piping downstream of the reactor. Before plugging occurred the dry product gas
showed no detectable S02 content (< O.O1 mol $). This fact strongly indicated
that most of the elemental sulfur was formed in the presence of liquid water
condensate by the liquid-phase Claus reaction;
2 H2S + S02 = 3 S + 2 H20 (XIII-2)
The equipment then was modified to the configuration shown in Figure IV-12 by
adding a heated catch pot to retain some of the condensate and nearly all of
the elemental sulfur.
Run A84
The first two parts of Run A84 were made with a fluidized bed of inert
fused periclase (MgO) in order to help distinguish the separate effects of the
gas phase reactions and of reaction (XIII-l).
The third and fourth parts of the run were made with sulfated
Tymochtee dolomite as bed material. Two temperatures, 860 and 893°C (1580 and
1640°F), were used with each bed material.
The input gas to all parts of the run was a mixture of steam and dry
recycle product gas, along with added C02 and CS2. Recycle gas was used to
simulate conditions in an upper portion of the bed in a tall commercial reactor.
Run conditions and results are given in Table XIII-7. Assay results
for the feed and product acceptor are shown in Table XIII-8. The acceptor was
not assayed between Runs A84C and A84D. However, the acceptor weight loss was
the same in both parts of the run. Therefore, the contributions to or from the
acceptor shown in Table XIII-7 were apportioned equally to each run, using the
data in Table XIII-8. Since CaCO3 was present, the reaction,
CaC03 + H2S = CaS + C02 + H20
also occurred in addition to reaction (XIII-l).
In all parts of the run, small amounts of COS, CO, and H2 were found
in the dry product gas, as expected.
Since the compositions of the product gases were altered by the
reactive environment downstream of the reactor, direct assessment of the con-
version of H2S to S02 in the fluidized beds is not possible. Based on
experience gained during development of the C02 Acceptor Gasification Process,
we have found that gas phase reactions involving sulfur species, as well as
reactions of CaS and CaS04, are very rapid at temperatures above about 649°C
(l200°F), especially when catalyzed by coal ash and impurities in the acceptor.
241.
-------
FIGURE XII1-3
Flue Gas Out
Air
CaS in from
Modified Sulfur Recovery Section
D,
O
o
§
•H
+J
rt
•d
•H
X
O
V
CaS04
in
Oxidation
Reactor
~ 1600°F
871°C
CaS
Spent
Acceptor to
Oxidation
Gas Dcsulfurization
CaC03 ovit to
Gas Desulfurization
Fresh
Acceptor
In
1!2S
CO 2
H20
S2
S02
COS
CO
Regenerator
~ 13OO°F
704 °C
C02 in
Liquid-
Phase
Claus
~ 310°F
154°C
Sulfur
Out
Spent
Acceptor
Discard
242.
-------
TABLE XII1-7
Conditions and Results for Reaction of CaS04 with H2S
System Pressure: 15 Atm
Run No.
Temperature, °C
Bed Solids
Size Consist, Tyler Mesh
Duration of Run, minutes
Input , SCFH
Recycle Gas
H20
C02
CS2
Purges to Bed (C02)
Purges above Bed (C02)
Input by Component, SCFHA1)
H20
CO 2
CO
H2
H2S
COS
Contribution from/to Acceptor,
grams /hr
Oxygen from Acceptor
Carbon from Acceptor
Sulfur to Acceptor
Output(2)
Exit Gas Rate, SCFH (Dry Basis)
Composition, Mol %
C02 (by diff.)(3)
CO
H2
H2S
COS
Flow Rate, Top of Bed, SCFH
Fluidizing Velocity ft/sec
A84A
860
(1580°F)
(206 psig)
A84B
893
(1640°F)
Fused Periclase, Inert
_, Ao ..
^ 1O X
4O7
140
135
.72
.
133.6
155.2
1.76
1.89
5.92
.44
0
O
0
29.7
93.88
1.26
1.35
3.20
.31
299.2
.43
nc
DO >
68
140
135
in
.80
5
133.4
154.2
2.36
1.96
6.36
.56
0
0
0
30. O
93.11
1.69
1.40
3.40
.40
299.4
.45
A84C
860
( 158O° F )
Sulfated
* 35
240
128
124
1.00
122.0
148.6
.90
1.04
3.15
.31
34.8
.97
2.60
33.9
97.35
.70
.81
.90
.24
276.9
.40
A84D
893
(1640°F)
Tymochtee 11
X TO ^
240
128
124
^
1. 15
121.7
149. 1
.93
1.05
3.02
.31
34.8
.97
2.60
33.8
97.65
.73
.82
.56
.24
277.2
. 42
(1) After hydrolysis of CS2 via CS, + 2 H20 = 2 H2S + C02.
(2) The actual recovered gas (after condensation of elemental sulfur and water)
which escaped the reactive environment downstream of the reactor.
(s) Taken by difference for consistency in subsequent calculations. Actual
measured values were, respectively, 94.0, 93.0, 97.5, and 98.0$.
243.
-------
TABLE XII1-8
Acceptor Compositions for Runs A84C and A84D
Feed to Run A84C
Product from Run A84D
CaS
CaS04
CaC03
Inerts,
including MgO
wt %
3.16
58.58
7.24
31.02
100.00
Grams
78.7
1459.3
180.3
772.6
2491
wt %
18.13
41.97
5.19
34.71
Grams
403.6
934.2
115.5
772.6
100.00
Distribution of Calcium, Mol
CaS
CaS04
CaC03
8.0
78.8
13.2
41.1
50.4
8.5
Obtained by forced calcium balance. Some of the product acceptor was spilled.
Actual recovered weight was 2096 grams. Operator estimated 100 grams spillage.
244.
-------
Equilibrium calculations therefore provide accurate estimates of the gas compo-
sitions at the top of the bed in Run A84. The exit gas compositions corre-
sponding to equilibria in the following reactions were calculated.
CO + 1/2 S2 = COS
H2 + 1/2 S2 = H2S
2 H2 + S02 = 1/2 S2 + 2 H2O
CO + H20 = C02 + H2
Results are shown in Table XIII-9. Although some conversion of H2S
to S02 occurred in the inert beds, transfer of oxygen via reaction (XIII-l)
provides a closer approach to the desired goal of a 2;1 ratio of H2S/S02.
The nature of the downstream reactions can be inferred by comparison
of the dry gas compositions taken from Table XIII-9 with the measured compositions
of the recovered dry gas. Inspection of the gas compositions, termed Conditions
I and III, in Table XIII-10 show that the dry, top-of-bed gases all contain more
CO and H2S, less COS, and about the same amount of H2 than do the actual product
gases. If the assumption is made that all of the incremental COS in the product
gas is formed.by. an overall reaction which occurs at nonequilibrium conditions
in the downstream environment and which can be expressed as,
3 CO + S02 = COS + 2 C02 (XIII-3)
and if all the remaining S02 reacts in the heated catch pot via reaction (XIII-2),
then the dry gas compositions shown as Condition II in Table XIII-10 are obtained.
Very good agreement between Conditions II and III was obtained in all four cases.
Other combinations of possible reaction paths such as formation of thionic acids
(e.g., H2S + 3 SO2 = H2S406) or the reactions;
3 H2 + S02 = H2S + 2 H20
2 COS + S02 = 3 S + 2 C02
gave poor concordance in all cases.
F. Conclusions
Little success was obtained in attempts to use raw limestone as an H2S
sorbent. Calcined limestone was a good sorbent, but regeneration was poor. It
may be possible to use calcined dolomite as a sorbent as rates of desulfurization
and regeneration appear to be satisfactory. However, there may be some diffi-
culties due to agglomeration.
The process for reacting the regenerator offgas with CaS04 to produce a
liquid-phase Claus feed having the required 2; 1 ratio of H2S/SO2 was shown to
be chemically feasible.
245
-------
TABLE XII1-9
Calculated Gas Compositions, Run A84
Run No. A84A A84B A84C A84D
Temperature, *C (°F) 860 (1580) 893 (1640) 860 (1580) 893 (1640)
Gas Composition, Top of Bed,
Mol %
H20
C02
CO
H2
H2S
COS
S02
S2
100.003 99.994 100.004 100.006
H2S/S02 15.7 12.4 3,14 2.26
44.70
51.57
.944
.694
1.864
.081
.119
.031
44.56
51.16
1.204
.804
1.968
.095
.159
.044
44.30
53.52
.575
.403
,869
.040
.277
.020
44.16
53.52
.717
.453
.760
.039
.336
.021
246-
-------
TABLE XIII-1O
-C-
-J
Run Number
Temperature, °C (°F)
Condition
Gas Composition. Mol j>
CO2
CO
H2
H2S
COS
S02
Comparison of Dry Gas Compositions.* Run A84
A84A
86O (158O)
I
93. 3O
1 .71
1.26
3.37
.15
.21
II
93.93
1.23
1.26
3.27
.31
O
III
93.88
1.26
1.35
3.20
.31
O
A84B
893 (164O)
I
92.36
2.17
1.45
3.55
.17
.29
II
93.2O
1.5O
1.46
3.45
.40
0
III
93.11
1 .69
1.40
3.4O
.40
O
A84C
86O (1580)
I
96.12
1.03
.72
1.56
.07
.50
II
97.58
.54
.73
.91
.24
O
III
97.35
.70
.81
.90
.24
O
A84D
893 (1640)
I II III
95.87 97.64 97.65
1.28
.81
1.36
.07
.60
.80
.82
.50
.24
O
.73
.82
.56
.24
O
* After condensation of elemental sulfur and water.
I - Equilibrium dry gas composition at top of bed (data from Table XIII-9).
II - Calculated dry gas composition after completion of downstream reactions (see text)
III - Actual measured composition of final product gas.
-------
XIV. PHYSICAL EXAMINATION OF PRODUCT SOLIDS
A. Surface Area and Pore Volume Data
Pore volume and surface area data are presented in Table XIV-1. Most of
the surface areas reported were in the range of 2-3 M2/g. It appears that
.surface area decreases slightly as the stone is cycled, but the changes are
insignificant compared to the changes observed in regeneration activity.
Similarly, pore volume up to 5000 angstroms was in the range of 0.1O to 0.15
cc/g, but porosity changes occurring as a result of cycling are small compared
to the change observed in activity. The data in Table XIV-1 go up to 10,000
angstroms for pore volume. The data in this range show a slight increase in
pore volume with cycling, however, it is considered that pores of this size
may include microcracks as well as actual pores.
B. Sulfur Distribution in Sulfided Dolomites
Figures XIV-1 to XIV-7 show the unspecific electron microprobe images and
sulfur, magnesium, and calcium X-ray images of particles present in the
sulfided dolomite. The particles were mounted in Bakelite and polished so
that cross sections of the particles are shown. When more than one particle
was imaged, an A was added to the sample number.
At low conversions, the samples varied in sulfur penetration. Figure
XIV-1 shows a prerun sulfided particle for which the sulfur is present in a
shell on the outside. Figure XIV-2 shows a particle from the same sample which
displays deep sulfur penetration which left three small areas nearly sulfur
free. Figures XIV-3 and XIV-4 present oxidized particles from the hardening
procedure. In both stones, the sulfur is concentrated on the outside, but is
present throughout. Figures XIV-5 and XIV-6 are from a sample which had
experienced about one cycle in the process. The sulfur is present throughout
the stone, but in Figure XIV-6 an area of high magnesium content can be seen
which is relatively free of sulfur. The cycled stone presented in Figure
XIV-7 shows a uniform distribution of sulfur along the entire cross section.
In another study a number of acceptor samples were mounted in resin,
sectioned, and examined under low power magnification illuminated by short
wave ultra violet (UV) light. Calcium sulfide fluoresces under short wave UV
light and the distribution of sulfur in the solid can be visually observed.
At low conversions and few cycles, an outer shell of CaS was observed in
sulfided samples from Runs A27, A32, A33 and A36A. The list includes a 20 x 28
mesh Canaan dolomite sample, and 20 x 28 and 35 x 48 mesh Tymochtee dolomite
samples. A series of cycled Canaan dolomite samples regenerated at 593°C,
704°C and 76O°C (llOO, 1300, 1400°F) showed the 593°C sample to fluoresce at a
qualitatively different shade than the others.
The results indicate that, at low conversions, the sulfur may be distri-
buted in a shell on the outside of the particle. This would be consistent
with a mechanism of either shrinking core diffusion or kinetic control of the
reaction. However, it is also clear that cracks in the particles can allow
sulfur to penetrate the interior of particles even at low conversions. At
high conversion, or after several cycles, sulfur is uniformly distributed in
249.
-------
TABLE XIV-1
Surface Area and Pore Volume Data
Run No.
Cycle No.
Source of Dolomite
Origin of Sample
A27
A28
A32
A36A
A38
A42
— 2.5 20.6
-Tyoochtee 1O
3.1 1O.8
Sulflded
Prerun
Surface Area, Ha/g 2.9
Pore Volume, cc/g
0-80 J .01
0-100 .01
O-12O .01
O-15O .01
0-200 .01
O-25O .01
0-350 .01
O-5OO .02
0-8OO .02
0-1000 .03
O-2OOO .03
0-5000 .16
O-1O,OOO . 19
Oxidized
Prerun
2.9
Gas Reg.
Das.
Sulfided
Prerun
O.O O.O
1
.01
.01
.01
.02
.02 .01
.03 .01
.03 .01
.04 .01
.09 .11
.09 .15
O
.01
.01
.01
.01
.01
.01
.01
.01
.01
.02
.13
.24
Sulflded
Prerun
2.2 2.6 5.2 2.6/2.1
-Tymochtee 9
Oxidized
Prerun
2.O
N.D.
Gas Gaa
Des. Des.
N.D. 2.3
.O
31
01
01
01
Ol
03
12
15
O.O
.01
.01
.ox
.01
.01
.01
.01
.01
.03
.03
.13
.15
O.O O.O
i
.01
.01
.01
.01
.01
.03
.03
^
.03 .01
.03 .02
.11 .13
.11 .19
Reg.
2.5
N.D.
N.D.
0.0
.01
.03
.08
.10
.10
•Tyoochteo 11-
Reg.
1.2
0.0
Sulflded
Prerun
3.7
N.D.
Sulflded
Prerun
3.1
0.0
.01
.10
.11
.01
.01
.03
.13
.18
Oxidized
Prerun
1.5
0.7
Reg.
9.6
Reg.
2.4
0.0
.01
.01
.01
.01
.03
.03
.03
.03
.03
.04
.14
.13
O.O 0.0
.01
.01
.01
.01
.01
.03
.03
.03 .01
.03 .01
. IB . 13
.30 .20
10
Ul
O
N.D. - Not determined.
-------
the particles. In addition, examination of cycled samples from the regenerator
do not show any preferential distribution of the residual inert CaS in the
stone.
The above considerations lead one to the conclusion that it might be
possible to ascribe a shrinking core reaction model to the stone on its first
or second cycle. However, this does not follow for subsequent cycles. The
presence of natural imperfections in the stone may further complicate the
reaction mechanism. It is suggested that in a bed containing particles with a
distribution of ages, the kinetics may well be too complicated to handle with a
rigorous kinetic calculation. However, a simplified model which assumes the
reaction to occur uniformly in the particles could conceivably lead to a usable
kinetic correlation for the desulfurization reaction.
C. Scanning Electron Microscope Examinations
Two series of scanning electron microscope (SEM) pictures of dolomite
samples were obtained. The pictures were combined with energy dispersive X-ray
analysis to determine the elemental nature of the materials being observed.
Figure XIV-8 shows A76 product, a sulfided cycled stone. The picture shows the
the presence of small MgO crystals together with large crystals of CaS. Figure
XIV-9 shows A81 feed and A81 product. The feed was cycled dolomite which had
been regenerated for 8 hours at 704°C (1300°F). The product was regenerated for
an additional 8 hours at 871°C (1600°F) and contained very little residual CaS.
The feed material shows the presence of large CaS crystals and small crystals of
Ca and Mg without sulfur, i.e., MgO and CaC03. The product shows no evidence of
sulfur in the crystals which were examined. The large crystals which formerly
contained CaS had been desulfurized to CaC03.
The next series of photos was obtained by the Illinois State Geological
Survey. The sample analyzed was the gas desulfurizer bed of Run A46. This was
a highly sulfided stone which had undergone 42 cycles. Individual grains could
not be identified, and only the relative elemental mol ratios over an area of
the photo are given. A further difficulty in evaluating the photo was induced
by the uncertainty of the depth to which the analytical X-ray beam penetrated.
Figure XIV-10 gives a view of the exterior surface. Large and small grains can
be seen. It appears that there is somewhat more sulfur associated with the
larger grains. Figure XIV-11A is a view of the fine grained area on the outer
surface, and a lower sulfur content is indicated. In Figure XIV-11B, a view
of large grains from a fracture surface on the particle interior shows a higher
sulfur level. Figure XIV-12 represents an entirely different phenomenon.
Rather large crystals of a sulfur-free material high in Mg, Ca and Si are
observed. It is believed that these are magnesium-calcium silicates which are
formed when the dolomitic structure reacts with silica impurities in the stone.
There is insufficient data to confirm whether these are immediately formed upon
calcining, or whether they continue to grow as the stone is cycled.
The key piece of information relevant to a deactivation model is that the
CaS is present as very large crystallites in the cycled stone. It is assumed
that as the stone is cycled, the calcium species grow and the CaS form
does not regenerate readily. The photo of the regenerated material which was
the A81 feed shows large grains of CaS which apparently could not be regener-
ated along with small grains of CaC03 which were regenerated. Under the brute
force conditions of an 871°C (1600°F) regeneration even the refractory CaS was
251
-------
converted to CaC03 as indicated by Figure XIV-9B. The model which relates
deactivation to crystal growth will be elaborated upon in Section XV which
discusses variable effects in the system.
Runs A97r A98r and A107 - Preparation of Samples
If it could be determined in which reactor crystal growth actually takes
place, it might be possible to change conditions to retard such growth and
thereby prolong sorbent life. Toward this end beds of dolomite were held at
reactor conditions for 16 hours while samples were taken for future analysis.
Runs A97 and A1O7 were at gas desulfurizer conditions of 871 and 899°C (1600
and 1650°F), respectively. Run A98 was at regenerator conditions at 704°C
(1300°F). These samples remain in storage as reference materials, and no
measurements have been made on them to date.
252,
-------
Unspecific image,
reversed sample current
S X-ray
Mg X-ray
Ca X-ray
FIGURE XIV-1
Images of Sample A41-1 at 18OX. Prerun sulfided to 15 mol $ CaS.
253
-------
Unspecific image}
reversed sample current
S X-ray
Mg X-ray
Ca X-ray
FIGURE XIV-2
Images of Sample A42-1A at 200X. Prerun sulfided to 15 mol % CaS.
254
-------
Unspec if ic image,
reversed sample current
S X-ray
Mg X-ray
Ca X-ray
FIGURE XIV-3
Images of Sample A42-2 at 180X. Prerun oxidized to 14 mol % CaS04.
255
-------
Unspec if ic image,
reversed sample current
S X-ray
Mg X-ray
Ca X-ray
FIGURE XIV-4
Images of Sample A42-2A at 150X. Prerun oxidized to 14 mol % CaSO,
256
-------
Uns pec if ic image,
reversed sample current
S X-ray
Mg X-ray
Ca X-ray
FIGURE XIV-5
Images of Sample A42-3 at 15OX. Regenerator sample, 0.7 cycles, 25 mol % CaS.
257
-------
Uns pec if ic imag e}
reversed sample current
S X-ray
Mg X-ray
Ca X-ray
FIGURE XIV-6
Images of Sample A42-3A at 150X. Regenerator sample, 0.7 cycles,, 25 mol % CaS.
258
-------
Uns pec if ic image _,
reversed sample current
S X-ray
Mg X-ray
Ca X-ray
FIGURE XIV-7
Images of Sample A42-4 at 150X. Regenerator sample, 9.6 cycles, 72 mol % CaS.
259
-------
FIGURE XIV-8
SEM Photos of A76 Product
CaS
A. A76 Product x 25,OOO
0. 4 u.
CaS
B. A76 Product x 71,OOO
MgO
0. 14 |JL
260
-------
FIGURE XIV-9
SEM Photos of A81 Feed and Product
CaS
Mg, Ca (no sulfur)
A. A81 Feed x 5,OOO
B. A81 Product x 5,OOO
261
-------
FIGURE XIV-vLO
SEM Photos of A46 Desulfurizer Bed
Elemental ratios given are relative rather than absolute
\3/Ca = .79
A. Exterior Surface x 77 1O
Whole Area; S/Ca = .72, Mg/Ca
= .37
B. Interior Surface, closeup of fine grained in 1OA
Whole Area; S/Ca = .67, Mg/Ca = .2
-------
FIGURE XIV-11
SEM Photos of A46 Desulfurizer Bed
Elemental ratios given are relative rather than absolute
A. Fine Grained Area on Outer Surface x 8095
Whole Area: S/Ca = .51, Mg/Ca = .37
B. Fracture Surface (interior of Particle) x 3O86
Entire Area: S/Ca = .80, Mg/Ca = .32
-------
FIGURE XIV-12
SEM Photos of A46 Desulfurizer Bed
Elemental ratios given are relative rather than absolute
Mg/Ca = .62 No sulfur
Hi Mg, Ca, Si
Mg/Ca = .41
External Surface x 3310
Whole Area; S/Ca = .3, Mg/Ca
= .34
264
-------
XV. SUMMARY OF VARIABLE EFFECTS
. This section is a summary of the variable effects in the gas desulfurizer-
regenerator system. The data relating to the system are scattered through a
prior report and several sections of this report; this section is intended to
tie them together and summarize them.
A. Gas Desulfurizer
Temperature - Increasing the temperature improves the desulfurization effi-
ciency due to the improved equilibrium constant. Operation of the unit would
also be affected if increasing temperature caused the equilibrium partial pressure
of C02 needed to maintain the stone in the half-calcined state to exceed that in
the gas stream thereby calcining the stone. As the temperature is increased, the
allowable space rate, mols H2/hr-mol CaC03 in the bed, needed to prevent break-
through increases. On the bench-scale equipment, the approach of the outlet gas
to equilibrium improved as the temperature increased. However, it is noted that
at 927°C (l700°F), the exit gas was at equilibrium, and the commercial unit is
assumed to operate at equilibrium.
Conoco data on the effect of desulfurizer temperature on regeneration
activity are sparse. In a previous study,v1) there was no effect in going from
871 to 899°C (1600-1650°F) in a short run. Data from the City College Clean
Fuels Institute! u) indicate that increasing the desulfurization temperature from
8OO to 900°C (1470-1650°F) reduces the regeneration activity, 100 x mols CaC03
produced/mol CaS fed, after 15 cycles from 9% to 4$.
Residence Time - There was no effect of solids residence time on the
desulfurization step, however, increasing residence time produces a more severe
deactivation of the stone with respect to the regeneration reaction. Experi-
ments on the effect of gas residence time were confounded with bed height, but
increasing gas residence time improved the approach to equilibrium.
Inlet H2S Concentration - there is no effect of the inlet H2S concentration
since the outlet concentration is equilibrium controlled. However, the inlet
H2S concentration does affect the allowable H2S;CaC03 space rate if all else is
held the same.
Bed Height - Increasing the bed height improved the approach to equilibrium.
However, the greatest effects were observed when the bed was reduced to only
four inches. At that low level, bypassing may have been the controlling factor.
Particle Size - Over the range of 35 x 48 to 20 x 28 Tyler mesh there was
no effect of particle size.
Age of CaCO3 - There was no effect on desulfurization as the stone was
cycled. As noted earlier, Westinghouse( 13) had observed a 3$ per cycle decline
in rate of desulfurization in their system.
Type of Stone - All dolomites behaved about the same in the desulfurization
step including the hardened stones. Limestones were not active in the desulfu-
rization step.
265.
-------
B. Regenerator
Temperature - Temperature has an immense effect on the'conversion of CaS
to CaC03, the conversion increasing with temperature. The effect on equili-
brium H2S concentration is to decrease the H2S as the temperature increases.
Residence Time - Conversion of solids increases with residence time, but
the conversion rate drops markedly as a function of conversion at constant
temperature.
Gas Concentration - There was no noticeable effect of variation of the
inlet C02:H20 ratio on the rate of reaction. Of course, the reactant concen-
tration influences the ultimate concentration of product H2S which is attainable
within the equilibrium limitation. The role of H2S is strange; data on hand
suggest that increasing the inlet H2S concentration allows the stone to reach a
higher level of conversion.
Bed Height - Experiments with variation in bed height were confounded
with effects of H2S concentration and gas residence time. However, the effect
was that increased bed height led to increased conversion and closer approach
to equilibrium.
Particle Size - As in desulfurization, there was no effect of particle
size.
Age of Stone - Without exception, the regeneration conversion fell as the
stone was cycled. A log-log relation was found to exist between percent
regeneration and number of cycles.
Type of Stone - Fresh dolomites displayed similar deactivation behavior.
Hardened stones were sometimes less active than fresh ones depending upon the
severity of the hardening process. Limestones were generally inactive.
Kinetic Model - The kinetic model, Kt = - In (l - X), is discussed fully
in Section X and can be used to predict the effects of changes in residence
time or temperature over a narrow range.
C. Model for the Deactivation Process
The regeneration reaction suffers from deactivation of the CaS even from
cycle one. The model postulated below fits the behavior of the system so far.
The part dealing with regenerator variables was developed jointly with City
College Clean Fuels Institute.
Deactivation of the dolomite occurs mainly in the gas desulfurizer due to
the elevated temperature. Two mechanisms are operating. Crystal growth of
CaCO3 and CaS segregates the calcium crystals from the MgO and produces a
material tending toward a limestone rather than a dolomite. Simple sintering
and densification take place as well, and this both hardens the stone and tends
to reduce porosity. However, the observed changes in surface area and pore size
are small compared to changes in activity.
266-
-------
In the regenerator, densification does not occur. However, the reaction
starts rapidly and then tails off to an insignificant rate. Using the same
feed, it can be shown that the conversion at which reaction essentially ceases
is a strong function of temperature, the lower temperature producing the lower
conversion. It has further been observed that introducing product H2S does not
retard the rate of reaction but conversely increases the ultimate level of
conversion attainable .
The phenomenon of the products of reaction increasing the conversion of
calcium in the sulfur-calcium system is not new. Pell( la ) found that H20 in-
creased the conversion of CaO to CaS in the reaction of H2S with calcined
dolomite, and Ruth(19) showed that either C02 or H20 could increase the con-
version of CaC03 to CaS in the reaction of H2S with half-calcined dolomite.
Furthermore, both studies demonstrated the effect of the reaction flagging at
lower and lower conversion levels as the reaction temperature decreased.
Ruth^ ' proposed the following mechanism after examination of electron micro-
scope results on various samples. "High levels of CO2 may promote the
formation of large numbe'r of fine crystallites of CaS that do not protect the
underlying CaCO3.... On the other hand, at low levels of C02 a smaller number
of large CaS crystals form and grow together to close off the surface, thereby
protecting the remaining CaC03 from further reaction."
Ruth's careful analysis of his data showed that in the presence of reaction
products, the reaction was indeed slower initially. However, within a short
time the conversion associated with the product-rich environment soon surpassed
that in the product-poor environment.
It is postulated that a similar model fits the regeneration reaction. At
low levels of H2S, the CaC03 formed grows in large crystals shutting off the
interior of the grain from further reaction. At high levels of H2S, the
reaction nucleates at many sites to produce many small crystals which leave the
interior of the grain open to further reaction. Via such a mechanism, the
presence of H2S would enhance conversion in the regenerator. This is in full
concordance with our experimental results.
267,
-------
XVI. DESIGN BASIS
This section contains a summary of the considerations entertained before
selecting the specific design constraints used for the economic evaluation.
A. Gas Desulfurizer
The gas desulfurizer should be operated with the shortest possible resi-
dence time in order to maintain good regeneration activity of the dolomite
sorbent . The previous residence time had been 30 minutes. Based on batch
experiments, 20 minutes appeared completely sufficient to effect the required
conversion. Since deactivation of the stone is thought to occur at the
elevated temperature of the desulfurizer, the shorter time was chosen for the
commercial design case. If it can be shown that the stone can be converted in
a shorter time, the residence time might well be reduced further. The
ultimate objective would be to desulfurize in an entrained-phase pipeline type
of reactor. However, it has not been possible to set up an experiment to test
this concept on the bench scale. Actual data on continuous runs have used a
residence time on the order of 40-60 minutes. It is anticipated that going to
the shorter residence time would ameliorate the decline in regeneration activity.
Four key factors affect the temperature of the gas desulfurizer: heat
balance, deactivation, and equilibrium with respect to both H2S removal and
calcination of CaC03. In practice, the desulfurizer temperature is fixed by
heat balance considerations between the hot fuel gas and the cooler stone from
the regenerator. This worked out to be 1678°F (914°C) and is consistent with
the other constraints.
The temperature of the bench-scale experiments was limited to 1600°F (871°C)
by corrosion considerations, and the need to avoid excessive wall temperatures in
the unit which might calcine the stone. Both from data and the model of deacti-
vation, going to the design temperature of 914°C (l678°F) should result in more
severe deactivation than would have been observed at 871°C.
For the design case, it was assumed that the deleterious effect of the
higher temperature would be balanced by the ameliorating effect of the reduced
residence time. The deactivation relationship used in the design was therefore
the one experimentally determined for Canaan dolomite.
The makeup and circulation rates of dolomite sorbent were specified to
maintain activity toward H2S in the desulfurizer. The outlet H2S concentration
was assumed to be that determined from the equilibrium relationship. In order
to achieve that equilibrium without breakthrough, a certain excess of CaC03 is
required. It has been determined that the amount of excess was such that the
H2S;CaCO3 molar hourly space velocity was a constant. Table XVI-1 is a summary
of the data on this ratio for continuous runs in the bench-scale unit. The
average ratio of 3.6 was chosen as the design constraint. This should be con-
servative since the gas residence time in the commercial vessel would exceed
that in the bench-scale unit by several hundred percent. Using the ratio of
3.6, the H2S feed rate and the 20 minute residence time as it affects, the inventory
of CaC03 in the unit, the makeup and conversion matrix shown in Table XVI-2 was
calculated. This lists the acceptable design conditions for several stones of
varying regeneration activity.
269,
-------
TABLE XVI-1
Run No.
A7
A16
A20A
A23
A25
A26
A27
A32
to A35
O A37
A38
A42
A43
A44A
A45
A46
A46
A47
A52A.
Mol CaC03
at Break-
through
ca.
14
4
< 10
10
< 5
3
2
ca. 2.5
2
4
5
2
5
3
4
3.5
< 2
2
3
H2S:CaCO3 Ratio
Mols CaC03 in
Reactor at
Breakthrough
2.5
O.72
< 1.9
1.6
< .74
.40
.26
.40
.30
.55
.64
.24
.73
.48
.50
.45
< .26
.32
.58
Mols/hr
H2S
3.
3.
1.
1.
1.
1.
1.
1.
0.
1.
2.
Feed
0
O
4
4
4
5
r
o
0
55
0
O
for Continuous
Mol Ratio,
H2S:CaCO3,
Hrs"1
(ca. 1 )(0
4.1
(> -7)
(0.9)
(> 1.8)
3.8
5.9
3.8
5.O
2.8
2.4
6.3
2.1
3.2
2.0
2.3
(> 2.1)
3.1
3.5
Runs
Origin of
Canaan 1
Canaan 1
Canaan 2
Tyraochtee
v
Tymochtee
1
Tymochtee
1
Buchanan
i
Tymochtee
1
J
Buchanan
Canaan 3
Stone
1O
9
11
3
11
3
3.6 average
Remarks
649°C (12OO°F) regeneration
No breakthrough
982°C (18OO°F) hardened
No breakthrough
No breakthrough
(i) Numbers in parentheses were not included in the evaluation of the average.
-------
The choice was made on the basis of conservatism with respect to the
possibility of the attrition rate exceeding 1$. A 1$ makeup rate would be
desirable from considerations of makeup cost and the cost of the disposal
train. However, since the total process cost increases by only about 5$ for
each 1% increase in makeup acceptor, the 2% level was chosen. This permits
a greater variety of natural stones to be employed without having to resort to
a hardening step. A finding to the effect that natural stones were prevalent
which could indeed be used with an attrition rate below 1$ would lead to the
reinstatement of the 1% makeup rate as the design basis.
B. Regenerator
Residence time in the regenerator was chosen as 60 minutes, essentially
the same time which had been employed in the experimental continuous runs. The
temperature of 704°C (l300°F) was a compromise between two factors. If the
temperature were higher, increased conversion to CaCO3 could be obtained. How-
ever, the equilibrium H2S concentration would be reduced and we would require a
larger liquid-phase Claus system. Reducing the temperature would yield a
higher equilibrium H2S concentration at the cost of reduced conversion of stone.
An optimization was not carried out, and the design temperature is an estimate
of what would give a reasonable result.
In bench-scale runs, the equilibrium H2S concentration was obtained in the
outlet gas on at least one occasion, and all the data pertaining to approach to
equilibrium in the regenerator indicate that approach to equilibrium should
present no problem. In fact, higher H2S concentrations may enhance the attain-
able conversion beyond what had been observed in the continuous unit. As there
would be little change in the design by raising the previous estimated approach
to equilibrium of 88$, the same number was retained for the present design.
C. Tertiary Particulate Removal
The tertiary particulate removal system was sized and costed as a cyclone.
While it is recognized that a more elaborate system may be required, work in
this area is not sufficiently advanced to warrant design of a particular alter-
native. It is expected that the efforts underway in the area of pressurized
fluidized bed combustion will eventually result in workable high temperature
tertiary particulate removal system. Its substitution for the hot cyclone
assumed here should not markedly affect the overall economics.
271
-------
TABLE XVI-2
Stone Makeup Summary
Regenerator Temperature = 704°C (l300°F)
Material
Canaan Dolomite
Makeup Rate,
$ of
Circulation
1.0
1.5
2.0
Conversion,
Mo Is CaS Formed
100 Mols Ca Fed
to Desulfurizer
8.9
10.4
11.7
Equilibrium,
Mols CaC03
100 Mols Ca Fed
to Desulfurizer
16.2
19.1
21.4
Mol Ratio,
CaS Makeup:
H2S Absorbed
.11
.13
.17
Hardened Tymochtee
9 Dolomite
Hardened (35%)
Buchanan Dolomite
1.0
1.5
2.0
1.0
1.5
2.0
8.0
8.9
9.7
7.5
8.6
9.5
14,
16,
17.9
13.7
15.7
17.4
.12
.17
.21
.13
.18
.22
272-
-------
XV11. ENVIRONMENTAL EVALUATION
Introduction
The purpose of this evaluation is to describe the impact of the desulfuri-
zation process on the environment. The integration of the desulfurization
process into the coal gasification and power plant scheme is depicted in Figure
XVII-1. The quantity of materials consumed and generated is reported in tons/year
(6132 operating hours). It must be emphasized that even though this is an inte-
grated process, the scope of this environmental evaluation is limited to the
desulfurization processing step. Refer to CCDC Drawing AF-3667 in Section III
for more detail of the proposed process.
Thermal Efficiency
Since thermal efficiency measures efficiency of resource utilization, it is
a key environment factor. As used herein, thermal efficiency is the percentage
of the total input heating value (feed gas plus external fuel) that is recovered
as net product (producer gas to power plant) and by-product heating value (gasi-
fier steam and sulfur) . As calculated from the overall mass and heat balance
(Table III-8), the process has a thermal efficiency of 95.8$. The high thermal
efficiency is the single most important attribute of this desulfurization process .
Waste Generation
About 0.14 million TPY of spent dolomite will be generated from the desul-
furization process for deposit in a slurry pond. To date, a process to recover
and recycle dolomite from its attrited sludge form has not been developed. On
the other hand, there are several outlets for the dolomite sludge. The material
can be used for land fill since it has the same basic chemical constituents as
in the original quarry stone. Other potential markets include road fill material,
cement production, and agricultural lime.
Make-up dolomite is calcium magnesium carbonate with some inert noncarbonate
compounds such as clay, silica and silt. Trace elements in the dolomite will
vary with the stone type and source. Typical trace elements and oxides for
Illinois limestone and dolomite are listed in Tables XVII-1 and XVII-2(21) below.
The morphology of the trace elements through the process and the potential impact
of their final state on the environment have not been studied.
TABLE XVII-1
Trace Elements in 90 Illinois Limestone Samples
Trace
Element
Barium
Boron
Chromium
Copper
Iron
Lead
Manganese
Molybdenum
Amount
Cla bv Wt.
0.0260
0.0018
O.O011
0.0018
1.13
0.0026
0.14
O.OOO1
Trace
Element
Nickel
.Potassium
Sodium
Strontium
Titanium
Vanadium
Zinc
Amount
1a bv Wt.1
0.0015
0.16
0.07
0.049
0.04
tt
0.004
Present as trace in only three samples.
273.
-------
FIGURE XVII-1
Process Flow Schematic
Materials Consumed and Generated (Tons/Year)
Coal
3.41 x 106
Coal
Gasification
Gasifier
Gas
1.72 x lO7^
Dolomite
1.42 x 10£
Sulfur Removal
and
Sulfur Recovery
T
Air
2.88 x
10
Water
1.80 x 105
Sulfur to Atmosphere 6.44: x 1O3
+ NOX + Particulates
1400 MW
Power
Plant
Elemental Sulfur
1.32 x 105
Spent Dolomite 1.42 x 105
Waste Water
1.95 x 10"
Sludge
Pond
-------
TABLE XVII-2
Phosphorus Pentoxide, Manganese Oxide, and Sulfur
Trioxide in Illinois Limestones and Dolomites Containing
More than 95 Percent Carbonates
No. of Samples Average (%} Range (jo]
Phosphorus Pentoxide (P2°5)
Limestone 26 .039 .OO4-.097
Dolomite 16 .006 .000-.032
Manganese Oxide (MnO)
Limestone 12 .015 .002-.044
Dolomite 14 .040 .010-.050
Sulfur Trioxide (S03)
Limestone 29 .13 .OO-.88
Dolomite 39 .17 .O2-.59
Elemental Sulfur
The elemental sulfur product of 1.32 x 105 short tons per year will be
marketed at a substantial dollar credit to the process. Consequently, there is
no disposal problem associated with the sulfur.
Water Use
About 1.80 x 105 tons per year of recycle pond water are required for the
cooling and conversion of reject acceptor. In addition, annual cooling water
make-up requirements total approximately two million tons or 1500 acre-feet.
Noise and Dust
Without a detailed design, noise and dust levels are difficult to predict.
If crushing and sizing of the dolomite is done on the plant site, some noise
and dust will be generated with the feeder and milling operations. Also, the
process includes several large compressors which must be designed to meet OSHA
noise levels. These areas must be considered with other pertinent design
criteria.
Emissions from the Power Station
A rigorous review of the environmental evaluation of the power station
using the desulfurized gas is beyond the scope of this study. However, since
the atmospheric emissions of the power plant are a direct function of the
incoming gaseous fuel quality, the impact of the desulfurization process on the
producer gas quality will be discussed here.
The Clean Air Amendments of 1970 have resulted in the promulgation of air
quality control regulations covering emissions from new stationary sources which
contribute significantly to air pollution, or cause or contribute to the
endangerment of public health and welfare. For a coal-derived, gas-fuel-
275.
-------
burning installation rated at 250 x 106 Btu/hr input or greater, the "new
source" emission standards expressed in pounds per million Btu of input fuel
are as follows: Sulfur dioxide - 1.20; particulate matter'"1 O.10; and ':
nitrogen oxides - O.^Q* Although a number of states and local communities
have adopted more stringent specifications, this discussion is limited to the
general standards 4!--''
;V>
Sulfur Oxides Pollution
The basic purpose of this study is to desulfurize a low-Btu coal-derived
gas for fuel to power stations. The total sulfur in the producer gas (including
the vent gas from the S02 absorption column) is equivalent to approximately 25$
of the 1.20 Ib/MM Btu input sulfur dioxide emission standard.
Particulates
The state of the art in particulate removal technology at elevated tempera-
tures has advanced markedly in the past several years.(7>22' The incentive to
develop this area has not been for environmental reasons alone. To permit
greater utilization efficiency of our coal resources, considerable development
work is being done on the low Btu gas combined cycle. Hot fuel gas filtration
is necessary to protect the gas turbine components of the combined cycle power
system. In fact, the turbine specifications for particulate matter are generally
more stringent than the environment "new source" standards...
The proposed design scheme includes three stages of particulate collection.
First, the desulfurized gasifier gas passes through two stages of cyclones to
remove substantially all of the fines including attrited acceptor and entrained
ash. After heat exchange to generate steam for coal gasification, the gas is
sent through a third stage particulate collector to remove any potential alkali
compounds formed at the lower gas temperatures and to serve as a final process
clean-up step. To maximize efficiency, the second and third stages can be
operated at high pressure drops due to the inherent high pressure of the producer
gas.
Preliminary experimental results have shown no traces of alkali-bearing
compounds in the gas once all the particulate matter has been removed at 7O4°C
(1300°F). It is theorized that the particulate matter serves as nuclei for
alkali deposition. Consequently, removal of particulates eliminates any potential
alkali problem. Further work is needed to confirm these preliminary results.
Nitrogen Oxides Pollution
All high temperature combustion processes of fossil fuels with air result
in the formation of nitrogen oxides (NOX). Nitrogen oxides will be formed in
the high temperature combustion zone by the direct reaction of atmospheric oxygen
with atmospheric nitrogen and with nitrogen containing compounds in the fuel.
In the producer gas from coal gasification the nitrogen-bearing compounds include
about 44% nitrogen and about 0.23$ ammonia. One disadvantage of the hot gas
cleanup process is that the ammonia is not removed. This will effect an in-
crease in the NO emissions from the power plant.
A
The level of ammonia conversion to NO in the power plant cannot be pre-
dicted with any degree of certainty. The results of laboratory experiments and
furnace studies indicate a low conversion, i.e., less than 30%.(8>36)
276.
-------
However, the results have not been confirmed under power plant conditions.
Furthermore, experimental work at low ammonia concentrations (less than 170) is
very limited. The potential impact of various ammonia conversion levels on
NOX is shown below;
70 Ammonia NOX Level
Conversion to NOX (Ib/MM Btu of Coal Input)
2 0.04
10 0.19
50 0.95
At any conversion level above 37% the "new source" standard of 0.7 Ib/MM Btu
of fuel input is exceeded by the ammonia contribution alone.
Based on the results of laboratory experiments and furnace plant tests,(8>36)
it appears the conversion of ammonia to NOX can be limited by the same
principles as presently applied to control NOX emissions from gas, oil, and coal-
fired utility steam generators.'9>23' Typical control techniques include staged
combustion (also known as "over-fire air simulation" or "off-stoichiometric
combustion"), flue gas recirculation and water injection. These methods result
in varying (1) reductions of NOX emissions, (2) degradation in boiler efficiency,
and (3) yields of other emissions such as carbon monoxide.
In the case of gas turbine combustion systems, NOX levels have been reduced
by 85% by either direct water injection in the combustor primary zone or by
adding steam to the combustion air.'24) In the first method, further benefits
included a marked reduction in smoke and a net increase in turbine power.
In conclusion, assuming the ammonia conversion to NOX can be controlled at
a relatively low level, i.e., less than 10%, then it appears that the principles
and techniques required to achieve NOX specifications are relatively well known.
This is true whether the conventional gas burning power station or the gas
turbine of the advanced combined cycle burns the fuel gas. As indicated pre-
viously, a thorough investigation of the level and control of NOX emissions from
the power station is beyond the scope of this study.
277.
-------
XVIII - REFERENCES
1. Curran, G.P., Clancey, J.T., Pasek, B., Pell, M., Rutledge, G.D., and
Gorin, E., "Production of Clean Fuel Gas from Bituminous Coals."
Report to Office of Research & Development, U. S. Environmental
Protection Agency, Under Contract EHSD 71-15, Period; March, 1972
to June, 1973, EPA Report No. 650/2-73-049, NTIS No. PB 232-695/AS.
2. Curran, G.P., Clancey, J.T., Fink, C.E., Pasek, B., Pell, M., and
Gorin, E., "Development of the C02 Acceptor Process Directed Towards
Low-Sulfur Boiler Fuel." Annual Report to Control Systems Division,
Environmental Protection Agency, Under OAP Contract EHSD 71-15, Period;
Sept. 1, 1970 to Nov. 1, 1971, NTIS Accession No. PB 210-840.
3. Robson, F.L., Giramonti, A.J., Lewis, G.P., Gruber, G., "Technology 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.
4. Sherwood, T.K., Ind. Eng. Chem.. _17, 745 (1925).
5. Adams, F.W., Trans. Am. Inst. Chem. Engrs., ^5, 424 (1933).
6. Katell, S., and Faber, J.H., Petr. Ref., 39, No. 3, 187-190 (1960).
7. Stone and Webster Engineering Corp., "Purification of Hot Fuel Gas From
Coal or Heavy Oil." Interim Report to Electric Power Research
Institute, EPRI 243-1, Nov., 1974.
8. Wendt, J.O.L., and Sternling, C.V., (Shell Development Co.), "Effect of
Ammonia in Gaseous Fuels on Nitrogen Oxide Emissions." Journal of
Air Pollution Control Association, Vol. 24, No. 11, Nov., 1974.
9. Blakeslee, C.E., and Burbach, H.E., (Combustion Engineering), "Controlling
NOX Emissions from Steam Generators." 65th Annual Meeting of the Air
Pollution Control Association, Miami Beach, June, 1972.
10. Beychok, M.R., (Fluor Engineers & Contractors, Inc.), "NOX Emission From
Fuel Combustion Controlled." The Oil and Gas Journal, 71, No. 9,
52-56, Feb. 26, 1973.
11. Curran, G.P., Fink, C.E., and Gorin, E., "Phase II, Bench-Scale Research
on CSG Process, Operation of the Bench-Scale Continuous Gasification
Unit." Office of Coal Research R&D Report No. 16, Interim Report
No. 3, Book 3, GPO Catalog No. 163.10:16/INT3/Book 3 (1969).
12. Hicks, C.R., "Fundamental Concepts in the Design of Experiments," Holt,
Rinehart and Winston, New York (1965).
13. O'Neill, and Keairns, D.L., "Selection of Calcium-Based Sorbents for High-
Temperature Fossil Fuel Desulfurization." Paper presented at AIChE
Meeting, Boston, Sept. 7-10, 1975.
279.
-------
14. Squires, A.M., City University of New York, Personal Communication,
Oct., 1975.
15. Rao, C.P., and Gluskoter, H.J., "Occurrence and Distribution of Minerals
? in Illinois Coals." Circular 476, Illinois State Geological Survey,
Urbana, 111. (1973).
16. Keairns, D.L., Archer, D.H., et al.} "Fluidized Bed Combustion Process
Evaluation, Phase II - Pressurized Fluidized Bed Coal Combustion
Development." EPA Report No. EPA-650/2-75-027C (1975).
17. Chance, A.M., "The Recovery of Sulphur and Alkali Waste by Means of Lime
Kiln Gases." J. Soc. Chem. Ind., 163-179, March 31, 1888.
18. Pell, M., "Reaction of Hydrogen Sulfide with Fully-Calcined Dolomite,"
Ph.D. Thesis, The City University of New York (1971).
19. Ruth, L.A., Squires, A.M., and Graff, R.A., "Desulfurization of Fuels with
Half-Calcined Dolomite: First Kinetic Data." Env. Science & Tech.,
6, 1009-1014 (1972).
20. Ruth, L.A., "Reaction of Hydrogen Sulfide with Half-Calcined Dolomite,"
Ph.D. Thesis, The City University of New York (1972).
21. Lamar, J.E., Handbook on Limestone and Dolomite for Illinois Quarry
Operators, Bulletin 91, Illinois State Geological Survey, Department
of Registration and Education, 1967.
22. Westinghouse Research Laboratories, Report to the Office of Air Programs,
Environmental Protection Agency, under Contract No. CPA 70-9 (1971),
Volume III, Appendix M.
23. Bartok, et al.} Esso Research and Engineering Company, "Systematic Field
Study of NOX Emissions Control Methods for Utility Boilers."
Environmental Protection Agency, under Contract No. CPA 70-90,
National Technical Information Service PB 210-739, Dec. 31, 1971.
24. Singh, et al., Westinghouse, "Formation and Control of Oxides of Nitrogen
Emissions from Gas Turbine Combustion System." ASME Paper
No. 72-6T-22, March, 1972.
25. JANAF Tables, Dow Chemical Company, Clearinghouse for Federal Scientific
and Technical Information, U.S. Dept. of Commerce, No. PB 168-370
(1965) and No. PB 168-370-1 (1966).
26. Preuner, G., and Schupp, W., Z. Physik. Chem., 6£, 129 (1909).
27. West, J.R., Ind. Eng. Chem.. 42, 713 (1950).
28. Curran, G.P., Fink, C.E., and Gorin, E., "C02 Acceptor Gasification
Process - Studies of Acceptor Properties," in "Fuel Gasification,"
Advances in Chemistry Series No. 69, ACS, Washington, D.C. (1967).
29. Zawadski, J., Z. Anorg. Chem.. 205. 180 (1932).
280.
-------
30. Hill, K.J., and Winter, E.R.S., "Thermal Dissociation Pressure of Calcium
Carbonate." J. Phys. Chem.. 60, 1361-1362 (1956).
31. Uno, T., Tetsu to Hagana, J37, 14-17 (1951).
32. Rosenqvist, T., "A Thermodynamic Study of the Reaction CaS + H20 = CaO + H2S
and the Desulfurization of Liquid Metals with Lime." J. Metals
Trans., AIME, 3, 535-540 (1951).
33. Batchelor, J.D., Yavorsky, P.M., and Gorin, E., J. Chem. & Eng. Data, 4,
(3), 241 (1959).
34. Terres, E., and Schaller, A., Gas and Wasserfach, 65, 761 (1922).
35. Batchelor, J.D., CCDC unpublished data.
36. Lisauskas, R.A., and Johnson, S.A., Chem. Eng. Prog., 72, (8), 76-77,
(1976).
281,
-------
Appendix A
Detailed Investment Costs
283.
-------
Cost Estimate:
Table A-l
Section 300, Sulfur Removal
fO
00
Equipment Number
Major: Steam Superheater C-301
Waste Heat Boilers C-302
BFW Preheater C-303
Intercoolers C-304
Quench Water Cooler C-305
Water Slurry Cooler C-306
Gas Desulfurlzers D-301
Regenerator Reactors D-302
Acceptor Converters 0-303
Acceptor Rcvg. Hopper F-300
Acceptor Surge Bin F-301
Acceptor Lock Hoppers F-302
Particulate Lock Hoppers F-303
--"•*'
Acceptor Lock Hoppers F-304
Quench Water Surge Drum' F-305
Prod. Gas 1st Stg. Cyclones G-301
Prod. Gas 2nd Stg. Cyclones G-302
3rd Stg. Particulate Collectors C-303
1st Stg. Regenerator Cyclones G-304
2nd Stg. Regenerator Cyclones G-305
Slurry Hydroclones G-306
Slurry Pumps J-301
Black Water Pumps J-302
Quench Water Circ. Pumps J-303
CCj Gas Compressor JC-301
Acid Gaa Compressor JC-302
CCj Blower JC-303
Acceptor Conveyor L-300
Description
2 Ea. 1455 S.F., Q = 23.9 MM,
316 S.S. Tube, 28CW
10 Ea. 2110 S.F., Q - 65.4 MM,
2-1/4 CH-1 Mo Tube, 280#
1 Ea. 1315 S.F., Q = 34.7 MM,
2-1/4 CR-1 Mo Tube, 280#
3 Ea. 362 S.F., Q = 5.7 MM,
304 S.S., Tube, 75#
1 Ea. 125 S.F., Q = 1.1 MM,
304 S.S., Tube, 15W
1 Ea. 990 S.F., Q = 5.4 MM,
304 S.S., Tube, 75#
6 Ea. 22'-6" Dia. x 41' O.S.S.,
Elllp. Top, Cone Btm. 250ft
2 Ea. 20'-9" Dia. x 67' O.S.S.,
Ellip. Top, Cone Btm. 250#
3 Ea. 8'-0" Dia. x 13' O.S.S.,
F.&D. Hds., 304 S.S., 15#
12' x 12' Cone w/grating, 1/4"
C.S. Reinf.
2 Ea. 9' Dia. x 12' O.S.S.,
Cone Hds., 1/4" C.S. PI., Atm.
4 Ea. T Dia. x 12' O.S.S., Ellip.
Hds., C.S. PL., 250# w/valve
2 Ea. 6' Dia. x 9' O.S.S.. Elllp.
Hds., C.S. PL., 250* w/valve
2 Ea. 8' Dia. x 13' O.S.S., Elllp.
Hds., C.S. PL., 25O* w/valve
1 Ea. 10' Dia. x 18' O.S.S., F.&D.
Top, Cone Btm., 304 S.S.,10#
24 Ea. 92" Dia., 17,200 CFM,
Refr. Lin., 250#, 1700°F
24 Ea. 87" Dia., 17,200 CFM,
Refr. Lin., 25O*. 1700°F
8 Ea. 100" Dia., 44,200 CFM,
Refr., Lin., 2SO#, 1400°F
4 Ea. 82" Dia., 11,850 CFM,
Refr. Lin., 2500, 1400°F
4 Ea. 77" Dia., 11,850 CFM,
Refr. Lin., 25O#, 1400°F
8" Dia., 590 Gpm w/rubber lln.,
304 S.S. piping
2 Ea. 160 Gpm, 45 psl, Aust. I.,
Centrlf., 200°F
3 Ea. 1438 Gpm, 60 psl, CR-Stl,
Centrif., 200°F
2 Ea. 555 Gpm, 120 psi, Aust. I,
Centrlf., 200°F
5630 ACFM, 14O"F, 4—>23O psig,
2nd stg. w/intercool.
2100 ACFM, 20O°F, O—»215 psig,
2nd stg. v/coolers
3485 ACFM, 140°F, 4 >16 psig,
1st stg. Centrif.
2 Ea. 23.5 TPH, 12" Dia. Screw,
45 rpn. 43' Ig., 30° slope
Weight
Material
$
Labor
$
Insulation Electrical Structural Foundation Piling
S.F. H.P. It C.Y. 50 T Ea.
51,600
250,000
25,000
9,000
2,000
5,000
4,368,000
1,888,400
18,060
8,850
13,760
60,800
25,840
49,080
8,200
1,086,000
985, 000
300,640
126,800
111,240
8.0OO
2,600
10,200
6,000
65,000
50,000
8,500
7,300
88,600
403 , 000
28, 9OO
38, 400
5, 100
21,800
1,822,000
747,600
31,350
3,370
8, 520
54,000
22,500
41, 100
13,500
607,000
534,000
257,200
71,800
63, 120
14,300
8,420
19,260
13,560
187,000
104,000
15, 100
8,400
3,000
20,000
1,500
750
150
400
1,442,400
568,800
1,740
3,020
1, 1OO
5.0OO
2,260
4.0OO
800
75, 100
68,500
25,120
9,280
8,040
1,300
300
1,050
600
10,000
6,000
1,000
1,100
1,260
7,800
600
29,400
12, 140
15,400
13,900
6,240
1,000
920
20
300
150
2,500
700
25O
10
27,000
85,000
9,000
800
1,000
3,300
6,400
6,000
7,500
84,000
79,200
32.0OO
6,000
5,600
14,280
38
20
236
135
36
108
60
3
15
8
30
24
8
8OO
-------
Table A-l (Continued)
Cost Estimate: Section 300, Sulfur Removal
to
00
Ul
Equipment
Ma.jor: Acceptor Elevators
Rotary Feeders
HI Temp. Rotary Feeders
Hi Temp. Rotary Feeders
Hi Temp. Rotary Feeders
Hi Temp. Rotary Feeders
Converter Agitators
Minor: Piling
Foundations
Structural
Electrical
Insulation
Buildings
Piping
Instrumentation
Markup, Incl. Eng'r., Purch.
Number
Description
L-301 2 Ea. 23.5 TPH, 90' C-C Webster
C-16824 C.S.
L-302 4 Ea. 290 CFH, 8" x 8", 5* AP,
250# at 100°F
L-303 6 Ea. 4050 CFH, 20" x 20" Ref . Lin.
304 SS, 250#, 1700°F
L-304 24 Ea. 64 CFH, 16" x 16" Ref. Lin.
3O4 SS, 250», 1700°F
L-305 2 Ea. 243 CFH, 8" x 8" Water cool.,
250#, 650°F
L-306 2 Ea. 500 CFH, 12" x 12" Water
cool., Ref. Un., 250», 650°F
L-307 3 Ea. 36" Dia. Imp. at 120 rpm
304 SS, 15*, 200°F
Total Major Equipment
M + L = 400
M = 80, L = 80
M = .36, L = .30
M = 80, L = 68
M = 1.10, L = 2.30
M = .80, L = .80 x 50,000 C.F.
Mp = .30 Uaj. M., L = .70 Mp
Mj = .06 Maj. M., L = .40 Mj
Total Minor Equipment
Total Material + Labor
Field, Home Office, Profit & Contingency
Total Investment
Weight
t
32 , 900
12,800
48, 000
108,000
9,000
19, 000
13, 500
Material
$
35,600
40,700
124,200
236,000
17,160
36,200
30.450
5,753,210
200,000
92,000
580,000
353,000
286,000
40,000
1,720,000
344,000
3,615,000
14,528,820
9,471, 180
Labor Insulation Electrical Structural Foundation Piling
$ S.F. H.P. # C.Y. SOTEa.
5,800 30 16,030 6
4,800 12 3,200
8,400 360 30 9,000
22,800 720 48 19,200
1 , 800 6 1 , 600
2,200 6 1,600
1,500 45 3,000
2,309,610 89,740 4,107 421,510 547 198
500
92,000 1,150
483,000 1,610,000
300,000 4,420
598,000 260,000
40,000
1,200,000
138,000
2,851,000
$24,000,000
-------
N3
CD
Equipment
Major: Sulfur Combust or
Claus Gas Reheaters
BFW Preheaters
Acid Heaters
Water Coolers
Acid Coolers
Liquid Phase Claus Rxs
SO2 Absorber Columns
Dust Hoppers
Separators
Liquid Sulfur Tank
Claus Gas Compressors
Comb. Air Compressor
Sulfur Pumps
Acid Circ. Pumps
Acid S. Stream Pumps
Boiler Feedwater Pumps
Absorber Recycle Pumps
Electrostatic Precip.
Minor Piling
Foundations
Structural
Electrical
Insulation
Buildings
Piping
Instrumentation
Cost Estimate:
Number Description
B-401 Q = 88 MM, 14' Dia. x 34' O.S.S.,
Refr. Lin., 250*, W.C.
C-401 6 Ea. 5050 S.F., 2500, 1300°F,
Max. 316 SS Tube
C-402 4 Ea. 5575 S.F., 2500, 400CF,
1-1/4 CR-1/2 Mo. Tube
C-403 4 Ea. 4325 S.F., 250#, 350°F,
316 SS Tube
C-404 4 Ea. 8260 S.F., 2500, 200°F,
C.S. Tube b Shell
C-405 2 Ea. 3915 S.F. 250*, 250°F,
316 SS Tube
D-401 2 Ea. 20' Dia. It 30' O.S.S.,
316 SS Clad, 2500, 450°F
D-402 2 Ea. 7'-3" Dia. x 57' O.S.S.,
316 SS Clad, 250#, 450°F
F-401 2 Ea. 5' Dia. x 10' O.S.S.,
C.S. PL., 250#, Ellip.HD.
F-402 2 Ea. 14'-6" Dia. x 16'-6" O.S.S.
316 SS Clad, 2500, 450°F
F-403 1 Ea. 5' Dia. x 15'-0" O.S.S.
C.S. PL., ATM, 400°F
JC-401 2 Ea. 1,250 ACFM, 184-» 234
psig, Centrif.
JC-402 20,600 CTM, 0 —> 205 psig,
2nd Stg. Centrif.
J-401 2 Ea. 25 gpm, 215 psig,
Rotary gear, 304 SS
J-402 3 Ea. 750 gpm, 190—»213 psig,
Centrif, Aust. I.
J-403 3 Ea. 750 gpm, 194—>200 psig,
Centrif., Aust. I.
J-404 3 Ea. 793 gpm, 0 >250 psig,
Centrif., C.S.
J-405 3 Ea. 750 gpm, 0 >110 psig,
Centrif., Aust. I
L-401 2 Ea. 12,250 CFM, 250*>, 400°F,
C.S., 8" 0 Tubes
Total Major Equipment
M + L = 400
M = 80, L = 80
M = .36, L = .30
M =80, L = 68
M = 1.10, L = 2.30
M = .80, L = .80 x 91,500 C.F.
1L = .25 Maj. M, L = .70 M^
M^ = . 06 Maj. M, L = . 40 M^
Table A-2
Section 400, Sulfur Recovery
Weight - —Material Labor Insulation Electrical Structural Foundations..-^ - Piling
S.F. H.P. # C.Y. 50
Total Minor Equipment
Total Material + Labor
Markup. Incl. Eng'r., Purch., Field, Home Office, Profit & Contingency
Total Investment
3 53,000
235,000
168,000
160,000
268,000
80,000
542,400
308, 780
13,600
170,600
5,320
200,000
180,000
800
4,200
3,000
11,700
8,400
180,000
115,000
717,000
271,000
415,000
197,000
188,000
477,600
284,000
14,400
165,000
3,000
616,000
557,000
2,540
15,750
9,060
24, 150
19,200
280.000
4,370,700
76,000
52 , 800
147,500
1,220,000
39,600
75,800
1,093,000
262,000
2,966,700
10,354,010
6,345,990
$16,700,000
84,500 500 7,800
23,400 4,670 48,480
11,300 3,020 34,240
12,000 2,800 6,OOO
17,200 8,000
6,000 800
365,400 5,960
25,300
1,240 5,300
140,000 2,740 6,000
770 320
32,000 530 7,000
27,000 7,000
20O 40 15
600 60
450 30
1,230 600
620 300
24. OOP 1,560 80 26,000
773,210 22,140 15,085 142,620
52,800
123,000 410,000
1,038,000 15,270
82,800 36,000
75,800
766,000
105,000
2,243,400
20 12
11 . . .
14 8
8
80 32
43 12
20 8
3
80 20
56 12
2
6
6
14
10
15
388 104
190
660
-------
APPENDIX B
Data Tables
Detailed tables of data obtained during the program described in the report
are contained in this section. They are:
B-l Composition of Exit Solids from Cycling Runs
B-2 H2S Content of Exit Solids from Cycling Runs
B-3 Attrition Data
B-4 H2S Content of Exit Gas - Batch Variable Studies
287.
-------
TABLE B-l
Composition of Exit Solids from Cycling Runs
Run
No.
A20A
A21
A22A
A23
A24
A2S
Cycle No.
0.4
1.5
4.2
6.9
12.4
14.4
19.1
22.5
26.4
30.0
31.8
3.3
6.1
13.4
16.0
19.1
22.2
25.4
28.7
32.1
35.8
36.7
2.3
5.2
6.2
7.7
10.6
13.5
17.6
19.5
22.3
3.2
7.0
8.0
10.8
14.5
22.1
1.1
3.0
4.0
5.0
7.0
9.0
11.2
14.4
15.2
15.3
1.1
2.3
5.6
8.0
10.4
12.3
16.1
17.2
19.7
22.6
25.8
28 .O
36.0
Gas
CaC03
74. 0
65.9
62.4
52.3
35 .5
36.4
23.2
18.9
17.3
16.5
6.4
66.5
53.7
31.0
27.8
24.6
21.5
18.6
.14.8
14.5
13.4
12.9
78.0
71.1
69.3
64.5
59.7
53.6
45.7
39.4
37.3
37.1
23.4
23.2
6.4
3.9
2.4
47.5
47.7
43.9
42.2
38.8
30.8
23.8
1.9
1.2
4.1
49.4
50.6
47 .7
47.0
40.9
37.0
1.6
4.0
7.5
6.8
5 .1
4 .3
4 .6
Desulfurizer
Mol 1°
CaS
19.8
20.0
30.6
42.4
60.1
55.3
68.4
77.1
80.0
83.6
93.6
33.6
46.3
66.1
69.5
71.7
74.2
77.0
79.0
79.5
79.8 ,
82.8 '
9.5
25.5
27.7
32.7
39.4
45.8
51.3
56.1
57.7
47.5
62.6
62.0
79.4
79.8
85.8
41.2
43.9
45.5
48.8
52.1
55.6
66.9
89.2
88.4
84.1
43.4
40.7
44.5
43.3
49.2
51.5
88.7
83.6
82.0
82.2
83.1
82.9
81.1
'CaO"
6.3
14.0
7.O
5.3
4 .3
8.3
8.4
4.0
2.7
0.0
0.0
0.0
0.0
3.0
2.8
3.7
4.4
4.4
6.3
6.0
6.8
4.3
2.6
3.4
3.0
2.8
1.0
0.6
3.0
4.4
5.0
15.4
14 .0
14.8
14 .2
16.5
11.8
9.2
7.6
9.5
8.3
9.1
13.6
9.4
8.9
1O.4
11.8
6.3
8.2
7.8
9.7
9.9
11.5
9.6
12.4
10.5
11 .0
11.8
12.9
14.3
Cyc le No ,
1.3
4.0
6.8
10.1
12.3
14.3
19.1
22.5
26.3
28.4
30.0
31.7
3.2
5.9
8 .7
13.4
15.9
19.0
22.0
25.3
28 6
32.1
' 35.8
36.7
2.2
6.0
7.6
10.5
13.4
17.5
19.6
22.0
1.7
3.1
6.8
8.0
10.8
14.5
22.1
1.0
2.9
3.9
4.9
6.9
8.9
11.2
14.4
15.2
15.3
O.9
2.4
5 .5
7.9
10.3
12.2
15.9
17.0
19.7
22.5
25.7
28.0
36.0
Regenerator
Mol 4
CaCO,
78.4
71.2
63.6
50 7
42.1
43.5
37.2
27.0
23.5
20.9
19.1
16.7
72.9
64.9
54.1
43.1
38.9
35.0
32.5
28.6
25.4
22.9
22.0
21.7
89.0
79.0
75.4
68.7
64.4
57.4
50.8
45.9
53.5
51.5
33.2
24.3
14.1
11.8
5.0
63.6
61.7
57.9
55.9
52.1
45.3
37.7
19.5
12.7
12.0
68 .9
72.2
61 .4
58.8
52.6
47.7
16.3
17.3
18.9
18.6
15.8
13.9
13.9
CaS
9.4
21.9
33.2
43.7
52.7
51.1
59.5
67.7
71.3
76.6
80.9
80.9
27. 0
34.6
43.3
55.9
60.4
64.6
64.9
70.0
70.2
75.0
74.0
72.2
8.2
18.0
21.4
28.0
35.6
38 .1
45.6
50.8
35.9
36.1
56.4
61.0
71 .5
71.8
81.5
28.3
30.9
33.3
34.6
39.4
47.4
50.7
69.7
77.2
78.0
26.5
25.4
32.6
35.1
40.9
45.5
73.4
72.3
71.6
72.5
72.6
76.8
75.5
"CaO"
12.2
6.9
3.2
5.6
5.2
5.5
3.2
5.3
5.2
2.6
O.O
2.5
0.0
0.5
2.6
1.0
0.7
O.4
2.6
1.5
4.5
2.2
4.0
6.1
2.8
3.0
3.2
3.3
0.0
4.5
3.7
3.2
10.6
12.5
10.4
14.8
14.4
16.3
13.5
5.8
6.8
8.8
9.5
8.6
7.3
11.6
10.9
10.1
10. 0
3.9
1.8
5.6
6.1
6.5
6.8
1O.3
1O.4
9.5
9.0
11 .6
9.3
11.5
288.
-------
TABLE B-l (Cont'd.)
Run
A26
A27
A28
A29
A30
A30A
A32
A33
A3S
A36
A36A
Cycle No.
3.3
5.7
8.O
11.0
12.0
14.3
16.9
19.7
21.5
2.5
4.7
6.4
8.8
11.3
14.0
16.6
20.4
2.2
4.4
6.8
9.7
2.2
4.4
7.0
2.3
4.9
7.7
1O.5
12.7
2.1
3.7
3.1
5.2
7.5
1O.8
1.7
4.1
6.6
9.2
11.5
2.1
4.3
6.5
8.0
9.9
12.1
13.4
1.6
2.7
3.6
4.6
5.7
6.6
1 .4
2.4
3.4
4.4
6.7
8.8
10.9
13.0
14.1
15.7
17.1
18.8
20.0
ComDosition of
Exit Solids
from Cycling Runs
Gas Desullurlzer
CaC03
49.5
43.7
36.6
27.1
24.4
15.1
7.1
0.5
O.2
45.3
41.9
37.6
28.3
19.8
12.1
4.5
0.4
53.9
47.3
40.7
34.7
60.8
55.7
49.5
58.4
50.1
42 .O
34.8
28.8
59.5
56.2
40.3
23.3
1O.1
1.5
60.9
51.7
48.0
32.4
26.4
44.4
29.5
19.3
13.1
2.3
2.1
0.7
72.8
69.2
64.1
61.2
54.5
51.2
76.5
71.5
66.2
62.4
46.4
44.0
37.6
27.8
23.8
21.7
17.3
13.5
10.8
Mol 4
CaS
42.7
50.2
54.1
67.3
66.9
73.8
83.0
88.7
93.1
49.0
51.5
54.3
64.8
68.8
78.8
83.9
90.6
34.8
40.1
48.2
52.6
28.9
32.8
37.9
29.1
37.2
41.1
54.2
60.3
29.3
31 .2
48.2
63.5
77.1
85.9
32.5
40.4
46.4
61.3
73.1
43.5
56.5
65.7
7O.7
84.1
84.7
84.4
22.6
24.6
29.7
34.4
35.9
41. 4
17.8
25.3
27.0
33.3
42.9
49.2
56.2
67.0
68.9
71.7
73.5
84 .O
83.6
"CaO"
7.1
5.5
9.3
5.7
8.7
11.1
9.8
10.8
6.7
4.8
6.0
8.2
6.9
11.5
9.2
11.6
9.0
9.9
11.8
11.1
12.7
8.6
10.7
12.6
12.5
12.7
-16.9
11.0
10.9
11.2
12.5
10.9
13.1
12.9
12.6
3.8
5.8
5.6
6.3
0.6
11.5
14.0
15.0
16.2
13.6
13.3
14.8
4.6
6.2
6.3
4.4
9.6
7.4
5.7
3.2
6.8
4.3
10.7
6.7
6.2
5.2
7.3
6.6
9.2
2.5
5.6
Cycle No.
3.2
5.5
7.9
11.0
11.9
14.1
16.9
19.7
21.5
2.4
4.6
6.3
8.7
11.2
14.0
16.4
20.6
2.1
4.3
6.6
9.7
2.2
4.3
6.9
2.2
4.7
7.5
1O.4
12.7
2.0
3.6
3.0
5.1
7.4
10.7
1 .6
3.9
6.5
9.1
11.5
1.9
4.2
6.4
7.9
9.8
11.9
13.3
1.5
2.5
3.5
4.5
5.5
6.6
1.2
2.3
3.2
4.2
6.6
8.7
10.8
12.9
14.0
15.6
17.0
18.7
20.0
2O. 0 + 3OO min
CaC03
66.3
59.2
51.8
41.8
37.0
30.8
21.4
14.6
9.7
64.1
55.9
52.0
43.2
34.8
27.0
19.3
12.1
71.4
66.3
56.2
49.9
79.8
71.9
67.0
76.3
66.0
59.3
50.5
42.8
77.8
72.4
49.7
33.9
19.7
10.9
69.8
66.1
55.9
48.0
38.4
53.2
40.2
28.1
23.1
13 .O
10.6
9.1
81.3
77.9
69.5
69.1
68.4
59.3
80.4
79.9
75.7
73.3
60.5
50.6
48.2
39.7
39.8
30.7
30.3
26.6
23.5
39.9
Hecenerator
Mol %
CaS
29.2
34.9
42.1
53.4
57.8
62.4
70.0
78.6
84.4
32.7
37.0
43.1
51.4
59.2
65.6
70.0
80.4
22.5
3O.1
34.8
41.0
13.9
19.1
27.0
13.5
23.2
31.3
39.2
45.6
14.4
22.9
40.0
53.2
67.2
78.2
24.6
28.2
35.0
42.1
55.1
34.7
45.3
58.1
62.7
73.5
76.6
77.2
11.8
16.5
19.9
22.8
25.5
33.8
8.0
14.9
20. 7
22.1
33.8
38.4
49.3
57.4
58.4
68.4
62.1
70.1
75.3
56.1
"CaO"
4.O
5.2
5.7
4.7
5.2
6.8
8.6
6.8
5.8
2.3
6.4
4.9
5.5
6.0
7.4
10.8
7.5
4.6
2.3
9.1
9.0
5.3
7.9
6.0
1O.2
1O.8
9.3
10.3
11.6
7.8
4.8
9.5
13.0
13.2
11.0
2.7 v
1.1
9.1
1O.O
6.6
11.7
14.5
13.7
14.3
13.6
12.8
13.7
7.0
5.6
1O.7
8.1
6.2
6.9
11.6
5.3
3.6
4.7
5.7
11.0
2.5
2.9
2.O
1.0
7.6
3.3
1.2
4.0
789.
-------
TABLE B-l (Cont'd.)
Run
No.
A37
A3B
A41
A42
A43
A44A
A45
A46
Cycle No.
2.2
4.6
6.7
8.1
9.2
1.0
2.1
3.1
5.9
8.3
10.8
12.0
0.9
1.8
3.0
3.4
1.8
3.1
5.5
8.5
9.7
1.0
2.1
3.2
5.3
7.6
10.1
12.9
16.2
1.3
2.1
3.1
6.3
8.1
10.1
11.6
1.1
1.8
2.8
5.0
7.5
10.0
12.8
1.4
3.1
5.3
8.3
11.1
14.1
17.2
19.8
23.4
26.8
30.5
33.5
37.7
42.1
Composition of
Exit Solids from Cvcllne Runs
Ga.s Desullurizer
CaCOj
45.3
26.1
12.4
5.8
1.3
37 .O
38.8
35.7
25.3
11.7
1.9
0.9
11.4
6.5
3.0
3.3
40.2
30.3
17.2
4.4
2.0
60.9
61.9
48.5
35.9
26.9
16.6
10.6
3.7
43.0
38.9
35. 0
22.8
12.8
4.5
1.3
49.5
45.0
38.4
26.1
13.1
4.3
3.2
51.8
44.3
34.1
19.4
10.1
2.8
3.8
3.4
3.0
2.9
2.0
2.0
1.8
2.0
Mol
CaS
42.2
59.1
75.8
84.9
89.7
53.3
50.7
53.1
62.8
74.5
80.0
81.2
73.7
84.4
84.0
85.7
•|
41.6
51.2
64.2
79.4
81.4
35.5
38.9
43.6
57.9
70.5
79.7
87.0
91.9
52.1
56.0
61.5
73.9
79.0
89.6
95.0
33.0
36.0
41.9
52.7
68.5
76.6
75.1
31.6
39.4
47.7
63.9
70.8
78.3
76.8
79.6
78.9
78.7
77.9
80.9
80.0
77.8
"CaO"
11.7
14.8
11.8
9.4
9.0
9.1
1O.O
11.2
11.9
13.8
18.1
17.9
14.9
9.1
13.0
11.1
18.3
18.6
18.6
16.2
16.6
3.0
0.0
7.9
6.3
2.7
3.7
2.4
4.5
4.4
4.6
2.7
3.4
8.2
6.0
3.7
16.9
19.0
19.8
21.2
18.5
19.1
21.8
15.7
16.3
18.2
16.8
19.1
18.9
19.4
17.0
18.1
18.5
20. 1
17.1
18.2
20.2
Cycle So.
2.0
4.5
6.8
8.0
9.2
9.2 + 100 min
0.9
1.9
3.0
5.8
8.2
10.7
11.9
12.0 + 180 min
0.8
1.7
2.8
3.3
3.3 + 180 min
0.7
1.7
2.9
"5.4
8.4
9.6
9.7 + 3OO min
0.9
1.9
3.0
5.2
7.5
10.1
12.9
16.2
1.2
2.0
3.0
6.2
8.0
10.0
11.5
11.6 + 240 min
0.9
1.7
2.7
4.9
7.4
9.9.
12.6
15.5
15.5 + 150 min
1.3
3.0
5.1
8.1
1O.9
13.9
17.1
19.7
23.3
26.7
3O.4
33.6
37.6
42. 0
CaC03
55.6
34.7
23.9
16.0
10.9
16.8
55.5
52.5
48.1
36.3
21.5
11.4
9.6
18.1
17.4
12.6
7.6
7.4
9.3
62.2
52.1
43.4
27.8
14.9
10.7
24.2
78.2
69.1
55.9
46.1
37.3
28.6
23.1
15.2
58.4
51.6
46.2
34.9
25.0
16.8
13.2
24.7
62.5
54.3
46.2
33.5
21 .9
13.9
11.9
11.5
18.0
59.4
49.4
4O.8
27.4
17.2
11.0
9.9
9.3
8.7
8.8
8.3
8.1
8.5
8.1
Renenerator
Mol i
CaS
33.2
51.8
65.5
75.7
81.6
77.7
35.1
39.2
43.4
55.9
66.7
72.5
73.6
66.5
73.4
78.8
83.5
83.4
79.7
25.4
33.7
4O.O
58.8
69.9
72.1
64.1
20.6
30.9
34.4
48.6
60.7
70.4
76.9
82.3
39.3
44.1
50.3
65.1
71.1
79.0
83.4
73 .0
23.7
28.4
35.6
46.5
57.3
65.1
68. S
70.2
65.4
25. 6
32.4
43.2
54.6
65.4
74.7
73.3
71.9
73.9
73.5
75.2
74.3
76.7
77.5
"CaO"
10.8
13.6
10.6
8.4
7.5
5.5
8.8
7.8
8.5
7.8
11.8
16.1
16.8
15.5
9.1
8.6
8.9
9.3
11.0
11.9
14.2
16.6
13.4
15.3
17.1
11.6
0.8
b.o
9.7
3.3
2.0
1.0
0.0
2.6
2.0
3.9
3.1
O.O
3.9
4.1
3.4
2.3
13.3
16.5
18.2
2O.O
20.9
21.0
19.6
18.3
16.7
14.5
18.2
16 .O
17.9
17.4
14.3
16.8
18.9
17.4
17.7
16.6
17.6
14.8
14.5
290.
-------
TABLE B-l (Cont'd.)
Run
No.
A47
A48
AS1
A52
A123
A124
A128
A129
Composition of Exit Solids from Cycling Runs
Gas Desulfurizer
Cycle No.
1.3
2.7
4.0
6.3
8.9
11.6
14.1
16.8
19.4
22.3
25.4
27.7
31.1
34.6
38.2
41.8
45.7
49.4
52.0
56.4
58.1
l.O
1.6
2.1
2.7
3.2
4.5
5.8
8.1
1O.5
12.8
14.9
17.4
19.7
22.7
2.4
4.7
7.1
11 .0
14.1
1.8
4.6
9.2
13.0
16.6
20.2
22.8
25.5
29.0
32.6
35.9
38.7
1 (feed)
2
3
4
1
1 (feed)
2
3
4
1
2
3
4
CaC03
47.5
48.1
43.8
36.4
31.3
22.9
16.6
11.3
6.7
3.9
2.2
3.8
1.1
3.6
5.O
5.3
6.3
6.O
6.1
6.1
5.5
22.4
8.1
12.3
12.8
11.0
1O.3
8.0
4.8
6.5
6.1
5.4
5.6
5.O
7.0
66.0
56.0
47.7
38.9
34.2
63.5
47.5
31.5
22.6
16.3
12.2
14.4
12.4
10.1
5.8
3.9
3.8
6.2
1.5
0.7
4.2
2.2
12.3
2.5
2.1
1.4
69.6
65.4
60.7
43.5
Mol ^
CaS
45.8
48.3
53.0
60.2
63.9
73.9
78.6
84.9
86.3
89.7
92.3
91.8
94.8
94.1
91.9
9O.2
89.7
9O.2
91.1
88.6
89.2
77.6
85.3
81.7
79.3
82.0
85.5
86.3
87.8
87. 0
87.3
88.2
87. 0
88.1
86.8
28.5
39.6
45.5
54.8
60.5
32.7
46.1
59.4
69.1
73.9
78 .O
80.6
80.8
85.9
86.1
89.2
90.2
87.9
94.2
91 .0
88.4
87.3
81.8
90.0
89.5
88.3
14.9
17.2
22.5
37.7
"CaO"
6.4
3.5
3.2
3.3
4.9
3.9
4.7.
3.9
7.0
6.4
5.5
4.4
4.1
2.3
3.1
4.5
4.1
3.8
2.9
5.3
5.3
O.O
6.6
6.1
7.9
7.1
4.2
5.7
7.4
6.6
6.6
6.5
7.4
6.9
6.2
5.5
4.4
6.9
6.3
5.4
3.9
6.4
9.1
8.3
9.8
9.8
5.0
6.8
4.1
8.1
6.9
6.O
5.9
4.3
8.4
7.4
10.5
5.9
7.5
8.4
1O.3
15.6
17.4
16.8
18.8
Cycle No.
1.2
2.5
3.9
6.1
8.8
11.4
14.0
16.7
19.3
22.2
25.2
27.5
31.0
34.5
38.1
41.8
45.6
49.3
52.0
56.3
58.1
58.1 + 135 ain
1 0
1.5
2.0
2.6
3.1
4.4
5.7
7.9
10.4
12.7
14.9
17.3
19.6
22.6
2.4
4.6
7.O
1O.9
14. 0
1.7
4.5
9.1
12.9
16.6
20.1
22.7
25.4
28.9
32.5
35.9
38.6
1
2
3
4
1
1
2
3
4
2
3
4
Regenerator
Mol I
CaCO;i
61.5
57.7
51.2
43.4
37.6
30.8
25.7
20.9
15.7
11.7
10.7
10.0
9.1
1O.7
10.3
10.6
10.7
11.3
11 .4
11.7
10.9
14.0
52.0
46.1
36.0
35.9
33.6
28.9
21.6
16.2
15.1
13.5
12.8
11.7
1O.9
11.7
73.9
67.2
58.1
51 .3
46. 0
77.2
61.7
48.4
39 .O
32.9
29.6
3O 7
26.7
24.5
20. 1
20. 0
18.6
24.9
17.2
13.5
9.0
11.6
48.5
26.5
18.2
20. 7
7O.7
65.6
54.6
CaS
36.8
41.0
46.5
53.1
61.1
66.7
71.0
73.3
8O.3
83.4
84.1
86.1
86.7
86.1
84.6
83.7
85.2
83.6
83.8
84.0
84.3
82. 0
43.4
48.9
64. 0
59.8
62.4
65.8
73.8
78.3
79. 0
81.4
81. 0
81.5
83.3
82.8
19.0
28.9
35.9
40.2
46.8
19.2
33 O
46.9
55.6
61.3
67.3
63.4
67.8
70.4
75.2
77.6
75.7
69.8
76.4
8C.3
84.1
8O.4
47 .?
67.5
75.1
70.9
13.8
16.8
27.8
"CaO"
1.4
1.3
2.3
3.5
1.4
2.5
3.3
5.8
4.O
5.0
5.2
3.9
4.3
3.2
5.1
5.7
4.1
5.2
4.7
4.4
4.8
4.1
4.6
5.O
0.0
4.4
4.1
5.3
4.6
5.5
5.9
5.1
6.2
6.8
5.9
5.5
7.0
3.9
6.0
9.0
7.2
3.6
5.3
4.7
5.4
5.8
3.1
6.0
5.5
5.1
4.1
2.4
5.6
5.3
6 .5
6.2
6.9
8 .0
4.2
6.0
6.6
8.5
15.5
18.0
17.6
29L
-------
TABLE B-l (Cont'd.)
Run
No.
A129A
A130
A131
A132
Cycle No.
1
2
3
4
5
7
9
11
1
2
3
5
7
9
11
1
2
3
5
7
9
11
1
2
3
5
7
9
11
Composition of Exit Solids from Cycling Runs
Gas
CaC03
72.8
69.5
57.8
51.9
41.9
24.0
17.9
7.8
73.7
73.2
71.5
68 .O
62.2
58.5
48.3
6O.6
44.3
27 .O
2.8
3.5
2.5
1.4
53.7
52.7
28.7
1O.7
1.8
16.3
9.O
Desulfurizer
Mol 4
CaS
1O.6
16.8
24.5
31.1
40.4
57.6
65.6
75.5
11.1
11 .0
11.9
16.6
17.9
25.4
33.9
24.4
39.5
57.5
85.7
76.9
77.6
79.7
26.5
34.4
54.1
71.9
83.6
66. 0
72.7
"CaO"
16.6
13.7
17.7
16.9
17.7
18.5
16.6
16.7
15.3
15.9
16.6
15.4
19.9
16.1
17.8
15. 0
16.2
15.5
11.5
19.5
2O. O
18.9
19.7
13.0
17.3
17.4
14.6
17.7
18.3
Cycle No
1
2
3
4
5
7
9
11
1
2
3
5
7
9
11
1
2
3
5
7
9
11
1
2
3
5
7
9
11
Regenerator
Mol %
CaC03
81.5
74.0
66.8
60.9
51.3
40. 0
31.8
25.7
89.6
87.2
86.8
82.0
75.3
7O.8
65.4
79.8
66.0
47.4
25.5
17.4
16.1
13.0
76.5
64 .O
53.1
30.7
22.6
25.0
22.2
CaS
3.7
11.3
17.3
23.4
30.5
35.4
47.5
58.5
2.1
2.3
2.4
5.1
8.9
13.5
19.9
7.6
21.3
39.1
64.0
,63.2
64.2
68.2
9.1
21.7
31.3
53.6
63.7
56.2
61.6
"CaO"
14.9
14.7
15.9
15.7
18.2
20. 5
20.8
15.8
8.3
10.5
10.9
13.0
15.8
15.6
14.7
12.7
12.7
13.6
10.5
19.4
19.7
18.8
14.4
14.4
15.7
15.7
13.7
18.8
16.3
-------
TABLE B-2
H.,S Content of Exit GaS from Cycling Runs
Dry Basis
Run
No.
A20A
A21
A22A
A23
A24
Cycle No.
1.5
4.2
6.9
9.9
12.4
14.4
19.1
22.5
26.4
28.4
30.0
31.8
3.2
5.9
6.9
8.8
13.4
15.9
19.0
19.0
22.2
25.4
28.7
32.1
33.5
35.8
36.7
2.2
5.2
6.2
7.7
1O.6
13.5
17.6
19.6
22.0
0.7
1.0
1.7
3.1
5.8
8.O
9.4
10.8
12.5
14.5
15.0
18.0
22.1
22.1
1. 1
3.0
4.O
5.0
7.O
9.0
11.2
11.2
14.4
15.2
Mol % HgS from Regenerator
.68
.88
.84
.96
.86
.76
.88
.82
.82
.80
.82
.80
2.40
2.48
3.07»
2.80
3.10
2.71
2.54
2.94*
2.44
2.42
2.41
2.24
2.62*
2.10
2.08
1.00
.98
1.00
.96
.94
.94
.98
.96
.93
2.4
1.82
1.3
1.08
3.04*
2.4
2.1
2.06
3.2
4.36
4.25*
1.12
.64
.76*
1.9
1.5
1.28
,26
24
,42
,43
49*
44
Uol $ HaS from Gas Desulfurlzer,
± .01% •
.02
.02
.02
.03
.02
.03
.05
.03
.02
.03
0.0*
.03
I
0.0*
.03
0.0*
.03
.03
.03
.02
.02
.03
1.20
.03
.03
.04
..03
0.0*
.04
.08
.12
.12
.08
0.0*
.06
.34
.25*
.02
.03
.02
O.O*
.20
.08
* Based on gas chromatograph analysis,
Other numbers, are via ultraviolet absorption.
293.
-------
TABLE 3-2 (Cont'd.l
Run
No.
A25
A26
Mil
A27
Cycle No.
0.9
2. -1
5.5
7.9
10.3
12.2
12.2
15.9
16.1
17.0
17.0
19.7
22.5
36.0
36.0
3.2
5.5
7.9
11.0
11.9
14.1
16.9
19.7
20.1
21.5
21.5
2.4
4.6
6.3
8.7
11.2
14.0
16.4
20.5
20.6
HgS Content of Exit Gas from Cycling Buns
Dry Basis
Mol % H,S from Regenerator
Mol % H2S from Gas Desulfurizer,
± .01%
15
85
38
06
16
14
88*
75
50*
91
66*
34
37
1.08
.90*
1.52
1.54
1.60
1.50
1.40
1.45
1.58
1.60
1.48
1.24
1.12*
,70
50
58
58
56
66
68
16*
1.34
A28
A29
A30
A30A
A3 2
A33
2.1
4.3
6.6
9.7
2.2
4.3
6.9
2.2
4.7
7.5
10.4
12.7
2.0
3.6
3.0
5.1
7. 4
10.7
10.7
1.6
3.9
6.5
9.1
11.5
11.5
1.68
1.54
1.60
1.52
70
63
58
1.64
1.56
1.4O
1.56
1.58
1.66
1.48
.28
.40
.52
.62
.53*
.51
.58
.54
.56
.58
.57*
.03
.02
.03
.03
.02
.03
.03
.03
.03
.04
.03
.03
.18
.11*
.03
.03
.03
.03
NA
< .02*
Based on gas chromatograph analysis. Other numbers are via ultraviolet absorption.
294,
-------
TABU B-2 (Cont'd.)
HaS Content of Exit Gas from Cycling Runs
Dry Basis
Run Hoi $ H2S from Gas Desulfurizer,
No. Cycle NO. Mol # HaS 'from Regenerator ± .Ol#
A35 1.9 .49 .03
4.2 .57 .04
6.4 .58 .04
7.9 .70 , .03
9.8 .72 .06
11.9 .58 .04
13.3 .68 .11
13.3 .72* .10»
A36 1.5 .46 .03
2.5 .56
3.5 .56
4.5 .62 -
5.5 .56 .03
6.6 .62 .03
A36A 1.2 .52 .03
2.3 .48 .03
3.2 .64 .03
4.2 .62 .04
6.6 .68 .03
8.7 .68 .03
15.6 .68 .04
17.0 .66 .03
18.7 .66
2O. O .68 .03
20.0 + 30 min. .32
20.0 + 60 min. . 18
20.0 + 120 rain. . 1O -
20.0 + 180 min. .08
20.0 + 300 min. .04
A37 2.O .72 .03
4.5 > .92
6.8 > 1.08
8.0 1.12 .02
9.2 1.18 .28
9.2 + 30 min. O.48 -
9.2 + 60 rain. 0.34
9.2 + 90 mln. O.32
A38 0.9 1.10 .05
1.9 .70 .03
3.0 .58 .03
5.8 .48 .03
8.2 .52 .04
10.7 .60 .10
11.9 .54 .22
11.9 + 30 min. .18 -
11.9 + 60 min. .12
11.9 + 90 min. . 1O . -
11.9 + 180 min. .08
A41 0.8 .70 .04
1.7 .56 .04
2.8 .60 .18
3.3 .74 .16
3.3 .68* .24*
A42 0.7 .88 , .03
1.7 .88 .03
2.9 .84 .04
5.4 1.O6 .06
8.4 1.24 .04
9.6 1.28 .10
9.6 ca. 0.9* .09
9.6 + 15 mln. .84
9.6 + 3O min. .60
9.6 + 60 min. .32
9. 6 -f 12O min. .26
9.6 -t- 180 min. .21
9.6 + 300 min. .12
* Based on gas chromatograph analysis. Other numbers are via ultraviolet absorption.
295.
-------
TABLE B-2 (Corxt'd.)
Run
No.
A43
A44A
A45
A46
11
11.5
Cycle No.
0.9
1.9
3.0
5.2
7.5
9.6
10.1
12.9
14.2
16.2
16.2
1.2
2.O
3.0
6.2
7.9
8.O
10.0
11.4
11.5
5 -f
+
11.5 +
11.5 +
11.5 +
11.5
15 min.
30 min.
60 rain.
9O min.
+ 12O min.
+ ISO min.
11.5 + 24O min.
0.9
1.7
2.7
4.9
7.4
9.9
12.6
15.5
15.5
15.5 + 15 min.
15.5 + 3O min.
15.5 + 6O min,
15.5 + 90 min.
15.5 + 120 min.
15.5 + 150 min.
3.O
5.1
8.1
10.9
13.9
17.1
19.7
23.3
24.8
26.7
30.4
33.6
36.4
37.6
42.0
HaS Content of Exit Gas from Cycling Runs
Dry Basis
Mo I % H,S fi-om Regenerator
l.OO
1.06
1.08
1.32
1.62
1.21*
1.5O
1.60
1.7O
1.44*
1.32
1.26
1.34
1.44
1.40*
1.66
1.58
1.68
.98
.62
.50
.44
.38
.30
.24
.76
.64
.54
.80
.92
1.06
1.24
1.02*
1.24
.94
.52
.36
.26
.20
.18
.62
.68
.80
.95
L.13
.76
.76
.81
.56*
.75
.80
.80
.59*
.80
.80
Mol % H8S from Gas Desulfurizer,
* .01%
.03
.03
ca. .04
ca. .04
.03
.08*
.12
.03
.03
.03
.03
.03*
.03
.06
.16*
.20
.04
.04
.03
.03
.03
.03
.10
.08*
.12
.02
.03
.03
.04
.10
.03
.03
.03
.02*
.03
.04
.03
.01*
.04
.04
* Based on gas chromatograph analysis.
Other numbers are via ultraviolet absorption.
296.
-------
TABLE B-2 (Cont'd.j
Run
N°jL
A47
Cycle No.
A48
A51
A52A
58.1
58.1
58.1
58.1
58.1
58.1
1.2
2.5
3.9
6. 1
8.3
6.8(0
11.4
14.O
16.7
19.3
22.2
25.2
27.5
29.0
31. 0
34.5
38.1
41.8
45.6
49.3
51.9
52.0
56.3
58.1
58.1
+ 15 mln.
+ 30 mln.
+ 60 min.
+ 90 min.
+ 120 mln.
+ 135 min.
1.0
1.6
2.0
2.1
2.7
3.2
4.5
5.3
5.8
8.1
8.8
1O.5
12.8
14.9
17.4
19.7
22.7
2.4
4.6
7.O
10.9
14.0
14.0
1.7
4.5
9.1
12.9
13.6
16.6
20.1
22.7
25.4
28.9
32.5
35.9
38.6
38.6
HaS Content of Exit Gas from Cycling Runs
Dry Has is
Mol % H.,5 from Regenerator
1.13
.97
1.02
.98
.81*
.76
.80
.86
.85
.86
.90
.92
.88
.88*
.80
.54
.60
.56
.50
.52
.60*
.56
.58
.64*
.56
.34
.20
.16
.13
.13
.11
1.9
1. 45
2.08*
1.85
,52
39
30
1.41*
1.46
1.21
1.52*
1.10
l.OO
l.OO
0.94
l.OO
0.56
1.78
2.2
1.67
2.10
2.53
2.53*
2.32
2.56
2.89
2.80
3.29*
3.53
3.35
3.36
2.96
3.29
3.26
3.21
3.50*
2.92
Kol % H2S from Gas Desulfurlzer,
± .01^
.04
.02
.02
.03
.04*
.02
.01
.01
.05*
.09
.01
.01*
.01
.01
.03*
.02
.02
.39*
.02
.01
.01
.02
.13*
.01
.03
.16*
.01
.01
.01
.01
.03
.01
.01
.01
.01
.02
.02
.04*
.01
.01
.03
.02
.03*
.02
.02
.02
.02
.02
.03
.05
.15*
.10
(>) Adjusted HaS analyzer.
* Based on gas chromatograph analysis. Other numbers are via ultraviolet absorption.
297.
-------
TABLE B-3
Run Cycle Number Attrition,
No. Start End *& of Feed Rate
A15 0 5.0 1.39
.64
.26
A20A 0 7.1 .73
.91
.74
.51
.54
A21 0 2.4 .74
.76
.88
.79
.93
1.05
.82
A22A 0 0.9 .57
.58
.56
.37
.31
A23 0 1.7 .72
.21
.20
.36
.34
A24 0 5.2 .75
.79
A25 0 3.1 .62
.60
.78
.76
.71
A26 0 8.5 .61
.52
.65
A27 0 6.4 .69
.74
.79
A28 0 9.8 .66
A29 O 2.2 1.14
.76
.66
A30 0 1.0 2.87
1.60
1.24
1.04
A30A 0.9 3.7 2.01
298.
Attrition Data
Cycle
Start
0
5.0
11.2
0
7.1
12.5
17.6
28.6
0
2.4
9.7
11.3
17.3
23.2
30.0
0
0.9
6.4
12.2
15.8
0
1.7
4.5
9.1
15.3
0
5.2
0
3.1
10.1
21.1
28.0
0
8.5
12.2
0
6.4
13.0
0
0
2.2
4.5
0
1.0
3.5
7.6
0.9
Number
End
5.0
11.2
17.7
7.1
10.9
17.6
28.6
31.8
2.4
9.7
11.3
17.3
23.2
30.0
36.9
0.9
6.4
12.2
15.8
22.3
1.7
4.5
9.1
15.3
22.2
5.2
15.3
3.1
10.1
21.1
28.0
36.1
8.5
12.2
21.7
6.4
13.0
20.7
9.8
2.2
4.5
7.3
1.0
3.5
7.6
12.8
3.7
-------
TABLE B-3 - Cont'd.
Run Cycle Number Attrition,
No. Start End # of Feed Rate
A32 0.8 5.0 1.79
1.65
A33 0.0 2.9 1.76
3.58
5.41
6.42
A35 0.0 3.6 .95
.76
.58
A36 O 3.7 1.98
1.42
A36A 0.3 5.5 1.52
1.53
1.24
1.18
A37 0 6.9 • .87
.53
A38 0 3.5 .86
1.13
.99
A41 0 3.4 .34
A42 0 3.3 2.16
2.25
1.26
A43 0 4.2 1.52
3.01
2.11
A44A 0 2.5 .67
.87
1.22
.67
A45 O 2.8 2.10
1.43
1.33
.99
A46 0 2.1 1.93
1.13
1.O4
.90
.70
.46
.44
.40
299.
Attrition Data
Cycle
Start
0.8
5.0
0.0
2.9
5.9
8.6
0.0
3.6
8.1
O
3.7
0.3
5.5
9.7
14.3
0
6.9
0
3.5
7.4
0
0
3.3
4.7
0
4.2
7.9
0
2.5
6.5
9.0
O
2.8
6.3
10.1
0
2.1
4.4
8.1
12.9
20.1
25.6
33.7
Number
End
5.0
11.0
2.9
5.9
8.6
11.6
3.6
8.1
13.6
3.7
6.6
5.5
9.7
14.3
20.0
6.9
9.3
3.5
7.4
12.0
3.4
3.3
4.7
9.9
4.2
7.9
11.4
2.5
6.5
9.0
11.7
2.8
6.3
10.1
15.6
2.1
4.4
8.1
12.9
20.1
25.6
33.7
42.5
-------
TABLE B-3 - Cont'd.
Attrition Data
Cycle
Start
0
4.3
8.3
13.0
18.3
22.6
27.7
33.2
40.3
48.8
1.0
3.8
15.6
19.9
0
4.6
9.8
1.2
7.0
15.3
20.2
28.7
Number
End
4.3
8.3
13.0
18.3
22.6
27.7
33.2
40.3
48.8
58.1
3.8
15.6
19.9
24.3
4.6
9.8
14.1
7.0
15.3
20.2
28.7
33.5
Run Cycle Number Attrition,
No. Start End 1o of Feed Rate
A47 0 4.3 .59
.98
1.03
1.02
.93
.41
.64
.52
.45
.48
A48 1.0 3.8 .47
.16
.26
.22
A51 0 4.6 .44
.64
.84
A52A 1.2 7.0 .72
.91
1.20
.87
.87
33.5 38.7 .82
300.
-------
TABLE B-4
H,S Content of Exit Gas - Batch Variable Studies
Run
No. Time, Min. Mol % H9S NO-*- Time. Min. Mol % H,,S
A68 8 1.77 A69 5 2.38
2.78
2.54
2.07
1.66
1.33
.88
.60
.40
.29
.18
.09
.07
.05
.03
A70 9 2.86 A71 8 .27
.36
.41
.41
.36
.27
.23
.17
.14
.13
.11
.08
.08
.06
.06
A72 8 .61 A73 7 .89
1.40
1.43
1.43
1.13
.72
.50
.35
.22
.14
.08
.05
.02
Time, Min.
8
14
24
41
71
101
129
159
189
249
309
369
429
9
19
28
45
58
75
90
105
120
150
180
240
300
344
8
15
30
45
60
75
90
120
150
180
210
240
285
345
405
465
Dry Basis
Mol # H9S
1.77
1.53
1.19
.92
.67
.56
.62
.40
.34
.28
.20
.17
.15
2.86
2.68
2.45
2.03
1.65
1.16
.84
.58
.42
.25
.14
.06
.04
.03
.61
.79
.77
.65
.51
.44
.37
.24
.20
.17
.13
.10
.08
.06
.04
.04
Run
No_.. Time, Min
A69 5
9
' 13
24
39
54
84
114
144
174
234
294
354
414
474
A71 8
23
35
44
53
68
83
113
143
173
233
293
353
413
473
A73 7
22
26
28
43
58
73
88
118
148
178
238
282
301.
-------
TABLE B-4 - Cont'd.
HoS Content of Exit Gas - Batch Variable Studies
Run
No.
A74
A76
Time, Min.
1
3
11
17
32
47
77
107
167
227
287
347
407
467
2
5
7
12
23
32
43
58
73
88
103
133
163
193
223
283
343
403
463
Dry Basis
Mol % H,S
.32
.38
.33
.25
.19
.15
.11
.10
.08
.06
.06
.05
.04
.05
.68
.87
.90
.82
.65
.60
.54
.48
.44
.41
.38
.33
.31
.27
.26
.22
.19
.16
.14
Run
No. Time, Min.
A75 5
8
16
25
40
55
70
85
115
145
205
265
325
385
445
505
A77 5
7
23
32
50
62
92
122
152
182
212
Mol % HgS
.54
.83
.53
.42
.34
.28
.26
.23
.21
.20
.16
.12
.10
.09
.09
.08
.063
.092
.080
.057
.048
.044
.032
.031
.029
.029
.023
302
-------
TABLE B-4 - Cont'd.
HgS Content of Exit Gas - Batch Variable Studies
Run
No.
A78
A80
Time, Min.
4
9
12
24
33
48
63
93
123
153
183
243
303
363
423
468
1
5
14
29
44
74
104
134
164
194
224
284
344
404
464
Dry Basis
Mol % H,S
.17
.195
.19
.14
.13
.10
.09
.07
.06
.05
.05
.04
.04
.03
.03
.03
.17
.25
.31
.29
.28
.23
.20
.20
.17
.14
.13
.10
.09
.07
.05
Run
No. Time. Min.
A79 3
7
1O
20
27
32
47
62
77
107
137
197
257
317
377
437
477
A81 7
9
12
15
28
37
55
82
112
172
232
292
352
412
472
Mol
.35
.39
.37
.28
.25
.22
.19
.17
.15
.12
.11
.10
.08
.07
.07
.06
.05
.33
.39
.44
.47
.46
.45
.42
.36
.31
.22
.14
.08
.05
.03
.02
303.
-------
APPENDIX C
Thermodynamic Data
Introduction
Therraodynamic data for equilibrium and heat balance calculations perti-
nent to the hot sulfur removal process were published previously.! 1 ) These
data are included here again for convenience to the reader.
Source of Data
A. Gaseous Phase
Nearly all of the standard free energy of formation values, heat
capacities at zero pressure, and the standard heats of formation were taken
from the JANAF Thermochemical Tables.(25) Exceptions in the sources of data
were;
1. The heat capacity and heat of formation for S6(g) were
calculated from the JANAF tables for S8(g) and the
experimental data in the equilibrium constants measured
experimentally by Preuner and Schupp.(26)
2. The vapor pressure and heat of vaporization of liquid
sulfur were taken from West.(27)
3. The heat content of steam above liquid water were values
given in the Keenan and Keyes steam tables.
B. Solid Phase
Sources of data involving calcium compounds are primarily experimental
equilibrium measurements which were supported privately by CCDC during the
1960's. CCDC(28) published equilibrium constant data in the reactions:
3/4 CaS04 + 1/4 CaS = CaO + S02 (l)
CaO + H2S = CaS + H20 (2)
and CaC03 = CaO + C02 (3)
The results were in agreement and extended the prior experimental measurements
of the above reactions.(29>30>3l> 32)
The reaction,
CaC03 + H2S = CaS + C02 + H20 (4)
by which gas desulfurizing and acceptor regeneration occurs is the sum of
reactions (2) and (3) above.
305'
-------
For reactions involving solids, we prefer to use experimentally determined
values of the equilibrium constants in order to avoid the errors involved in:
l) the use of the sometimes uncertain or discordant free energy values in the
literature, and 2) the extrapolation of free energy data to temperatures of in-
terest to the process.
Equilibria in many reactions for which we have no experimental data can be
calculated accurately by summing one or more of the reactions (l), (2), and (3)
above, with gaseous reactions for which the free energies are well-established.
For example, consider the reaction,
CaO + 3/4 S2 = CaS + 1/2 S02
This reaction can be expressed as the sum of the following reactions:
Equilibrium Constant
CaO * H2S = CaS + H2O K12
H20 + 1/4 S2 = H3 + 1/2 S02 (l/Ka)1/2
H2 + 1/2 S2 = H2S K2
CaO + 3/4 S2 = CaS + 1/2 S02 K13
Thus,
_ (K12)(K2l
~
The above equilibrium constants are those identified in Tables C-2 and C-7.
Presentation of Data
Distribution of sulfur vapor species are presented in Table C-l.
Equilibrium constants for gas reactions 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 6O°F, derived from
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.
Equilibria for the pertinent solids reactions involving calcium compounds
are given in Table C-7. Mean heat capacities above 60 °F are shown in Table C-8.
Heats of formation and heats of reaction are shown in Table C-9.
i
Plots of the equilibrium constants versus temperature for reactions in-
volving CaS are shown in Figures C-2 and C-3.
306
-------
TAIIIJ: c-i
Effect of Temperature and Sulfur Partial Preiiura
on Distribution of Sulfur Species
Mol Fraction
S,
.8575
.8472
.8080
.7567
.7022
.8249
.8OOO
.7473
.6916
.623O
.7966
.7196
.6501
.3841
.5072
.6903
.5903
.5204
.441O
.3682
. 6O79
.5651
.4937
.4133
.3401
.2623
.1958
.4614
.4146
.3404
.2612
.1934
.1279
.08O1
.3287
.2822
.2117
.1423
.0910
.0499
.0264
.2170
.1751
.1164
.O670
.O369
.O176
.OO84
.OO1 6
.0977
.0572
.0288
.0143
. OO63
.0029
.0005
.O055
.OO23
.0010
.OOO4
.0001
.OOOO
s.
. 1424
. 1527
.1917
.2426
.2963
.1749
.1997
.2520
.3068
.3735
.2428
.2831
.3474
.4108
.4819
.3080
.4045
.4693
.5374
.5906
.3873
.4274
.4913
.5558
.6O19
.6273
.6163
.5150
.5512
.5955
.6175
. 6O18
.5380
.4431
.59-19
.6O75
.5997
.5448
.4582
.3398
.2368
.5970
.5726
.5012
.3899
.2822
. 1817
. 1149
.0398
.4625
.3561
.2423
. libS
.09-16
.O571
.0189
.O87O
.0623
.O294
.0172
.OOiS
.0007
3^
.O001
.0001
.OOO3
.0007
.O015
.0002
.OO03
.OOO7
.O016
.0035
.0006
.0011
.0025
.0051
.0109
.0017
.0052
.0103
.0216
.0412
.0048
.0075
.0150
.0309
.O58O
.1104
.1879
.0227
.0342
.0641
.1213
.20-18
.3341
.4768
.0764
.1103
. 1886
.3129
.4508
.6103
.7368
•
. I860
.2523
.3824
.5431
.6809
.8007
.8767
.9586
.4398
.5867
.7289
.8269
.8991
.940O
.9806
.9O75
.9452
.9696
.982-4
.9944
.9993
Partial Pressure of
Total Sulfur, film Atci
250* F
4.H5 x 1O-"
3.54 x 10-1
1.24 x 10-"
.396 x 1O-*
.142 x 10-*
300° F
2.33 .1 10"
1.25 x lO"4
.401 x 10-*
.144 x lO'4
.O482 x ID"4
400"F
3. 11 x 1O"J
1.4O x 10"' ;
.461 x 10-'
.171 x 10-'
.O592 x 1O"'
SOO*F
2.4J2 x 10''
.508 x 10- J
. 192 x lO"1
.O680 x 10-'
.0272 x 10-'
6OO*F
9.81 x 10-'
5.31 x 10-'
2.02 x 10-'
.726 x 10-'
.294 x 10-'
.114 x 10-'
.O511 x 10-'
7OO°F
. 130O
.O724
.O294
.0115
.00517
.00234
.OO124
BOO'F
. 1825
. 1063
.O472
.0211
.0110
.00601
.00378
9OO'F
.2764
. 1714
.0859
.O448
.O27j
.0171
.011))
.00609
1000* K
.3O71
. 1749
.1042
.0698
.O475
.0345
.O186
1200" F
.54U
.398
..87
.216
.120
.0328
-------
TABLE C-2
Equilibrium Constants
for Gas Reactions
u>
o
GO
Kl CO + 1/2 S2 = COS
K2 H2 + 1/2 S2 = H2S
K3 2H2 + SO2 = 1/2 S2 + 2H2O
K4 2CO + S02 = 1/2 S2 + 2CO2
KG 3S8 = 4S6
K8 S8 = 4S2
K5 2H2S + S02 = 3/2 S3 + 2H20
K7 H20 + CO = H2 + CO2
K9 CO + 3H2 = CH4 + H20
Temperature in ° Kelvin
InK = A +
A
-15.5992
-10. 35O3
-3.3749
-4.9853
-4. 8O08
-1.4871
-15.7562
-14.3343
21.9713
-33.2782
1.99482
-6.21477
-33. 415
BT + CT2 +
B
.0122924
. OO06332
-.O027 18 4
-.0007315
.O03597O
-.OO2161O
.0052624
.0015907
-.003582
. 132324
.0089502
. O016890
.0030594
D/T
C x 1O6
-8.O1O42
-. 1OO12
1.01362
. 18766
-2.26137
. 43958
-2.65427
-.35856
3.5055
-63.9707
-4.6424
-.34401
-1.03191
D
12,017.2
11,327.0
10,039.6
10,480.5
15,353.9
14,786.0
25, 277. O
25,314.0
-15,135.4
-29,596.0
-4370.84
5,187.25
28,104
Temperature
Range. °K
300-900
9OO-1400
3OO-9OO
900-1400
3OO-900
900-1400
300-900
900-1400
300-80O
3OO-800
300- 90O
900-14OO
900-1300
-------
TABLE C-3
Numerical Values of Equilibrium Constants for Table C-2
Temperature in °F
°F
200
300
4OO
w 500
o
^ GOO
7OO
800
900
1OOO
1200
14OO
1600
1800
20OO
Kl
1.05 x 1O8
1.66 x 107
8.O4 x 105
73,90O
j
1O,60O
2,090
505
158
58 . 1
11.4
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 1O5
5.27 x 1O4
1. 42 x 1O4
4,62O
1,750
353
99.6
35.9
15.5
7.64
K3
1.80 x 1015
1. 45 x 1O14
2.26 x 1012
8.36 x 1010
5.72 x 109
6.15 x 108
9.29 x 107
1.82 x 107
4.39 x 10°
4.O1 x 1O5
6.35 x 104
1.38 x 104
3,900
1,340
5.
8.
9.
4.
5.
1.
7.
5.
6.
1.
9.
27
45
35
35
62
55
55
73
16
55
15
9
1
K4
x 1021
x 1019
x 1016
x 1014
x 1012
x 1011
x 109
x 108
x 107
x 106
x 104
,280
,420
296
K6
3. 17 x 1O~8
3.88 x 10~7
2. 44 x 1O~5
6.62 x 10~4
9.81 x 10~3
. .0937
.644
3.43
15. 1
188
—
—
— ' -
—
K8
2.O3 x
2.61 x
5.75 x
1.56 x
8. 48 x
1.26 x
6.30 x
1.25 x
10-29
10-26
10- 21
10- 1G
10~13
io-9
10"7
10~4
.0108
9.
54
K5 K7
.OOO749
. OO19O
.OO914
. O325
.O925
.221
.461
.856
1. 44
3.22 1.965
6.39 1.20
10.7 .819
16.3 .603
22.9 .471
K9
-
-
6.97 x 10~2
1.569 x 10~2
1.224 x KT3
1. 486 x 10~4
-
-------
TABLE C-4
Heat Capacities at Zero Pressure
Temperature in "Kelvin
Cp = A + BT + CT2 + DT3 + E/T2
B
C x 106 D x 109
Ss
s6
S2
COS
H2S
SO 2
N2
CO
CO 2
H2
H20, v
2 (g)
CH4
°2
Air
28.51
23.13
8.633
9.678
7.344
5.855
7.098
5.551
5. 1O1
7.219
7.757
1. 3609
6.055
5.8254
.03771
.01852
.OOO272
. 005978
.001851
.01534
-.001431
.003185
.015568
-.000674
.O000003
.021486
.0036021
. 002572
-25.25
-3. 137
-.1903
3.364
-11.002
3.490
-. 8609
-10.238'
.6840
3.219
-5.80317
-1.6265
-.52567
-1.623
2.842
-1.348
2.552
-1.125
.2825
-87,826
-123,658
48,814
-16,003
112,176
36,576
Temp. Range,
»K
300-800
300-1000
300-1400
300-1000
310.
-------
TABLE C-5
Mean Heat Capacities Above 60°F. Gases
Temperatures in °F
Cp = A + BT + CT2 + DT3
Btu/Lb Mol/°F
B
C x 106
D x 109
S8
S6
S2
COS
H2S
SO 2
N2
CO
C02
H2
H20, N
2 (g)
CH4
02
Air
36. 903
27.940
7.680
9.718
8.065
9.274
6.941
6.925
8.632
6.910
7.974
8.1915
6.9211
6.953
.006734
. 004678
.001251
.003033
.OO0916
. 002798
. 000037
. 000105
. 002958
. OOO135
. O00410
.00332O1
.OO077623
-.000003471
-2.598
-.321
-.640
-1.096
.214
-.898
.252
.266
-.844
-.044
.239
. 4754
-.1442
.3839
.126
.177
-.070
.121
-.058
-.070
.109
.022
-.048
-.2312
. 012092
-. 1108
Temp. Range,
°F
60-1000
60-13OO
60-2000
60-1340
Heat content of steam (l atm) above liquid water at 60°F.
Btu/Lb
Btu/Lb
1021.08 + .4854OT - 35.83 x 10~6 T2 + 34.46 x 1O 9
1O39.25 + .41916T + 46.6O x 10~G T2
220-100O
1000-160O*
* Can be extrapolated accurately to 2000°F.
311,
-------
TABLE C-6
Heats of Formation at 25°C. Gases
cal/g mol
AHf
S8, . +24,200
S6(g) +25,580
S2, . +30,840
COS -33,080
H2S -4,880
S02 -70,960
CO -26,416
C02 -94,052
H2°(g) -57,798
H2O, ^ -68,317
CH4 -17,889
Heats of Reaction at 25°C. Gas Reactions
cal./g mol
Reaction
1 CO + 1/2 S2 = COS -22,080
2 H2 + 1/2 S2 = H2S -2O,3OO
3 2H2 + S02 = 1/2 S2 + 2H20, ^ -29,220
4 2CO + S02 = 1/2 S2 + 2C02 -48,890
5 2H2S + S02 = 3/2 S2 + 2H20, . +11,380
6 3S8 » 4S6 +29,72O
8 S8 = 4S2 +99,160
6a 1/6 S6 = 1/2 S2 +11,160
8a 1/8 S8 = 1/2 S2 +12,4OO
5a 2H2S + S02 = 1/2 S6 + 2H20, . -22}09O
5b 2H2S + S02 = 3/8 S8 + 2H20, v -25,80O
Heat of Vaporization of Sulfur
Temperature in °F
Btu/Lb = 211.1 - .3302T + 434.6 x 10"6 T2 - 214.4 x 1O~9 T3
-------
TABLE C-7
u>
K1O
Kll
K12
K13
K14
K15
K1G
1/4 CaS04 + CO = 1/4 CaS + CO
CaO + H2S = CaS + H20
CaO + 3/4 S2 = CaS + 1/2 SO
CaSO4 + S2 = CaS + 2 .
CaC03 = CaO + CO2
CaCO3 + II2S = CaS + C02 + H20
Equilibrium Constants for Solids Reactions
CaO +
!aS +
/2 SO
02
!2 + H
Temperature in °
In K = A -f BT +
A
SO2 71.5O76
CO2 3.0O4O
17.5646
2 -7.22273
-6.55015
-.84O165
,0 16.6700
i
Farenheit
CT2 + D/T
B C
-.O3125O 5.
-.OOO751
-.OO92926 1.
.O02533O
.O118713 -1.
.O074O15 -1.
-.OO18535
x 106 D
79655 -67,104
01932 4, 232
7716 -1,07O
3688O 1 5,2OO
86265 -12,576
28243 -12,932
480912 -13.977.2
Temperature
Range, °F
1500-20OO
1500-2OOO
13OO-2OOO
15OO-2OOO
15OO-2OOO
all
11OO-2OOO
-------
TABLE C-8
CaS04
CaS
CaO
MgO
MgO-CaO
MgO-CaC03
Char Carbon
MgO-CaS
Ash
Mean Heat Capacities Above 60° F, Solids
.Temperature in "Farenheit
Cp = A + BT + CT2
Btu/lb Mol/T
A
23.57
10.14
10.98
9.77
20.75
30.64
1.447
19.91
.207
B
.00645
.QO23O
.OO071
. 00081
.00152
. 00506
.00291
.00311
Btu/lb/°F
. 000029
Heat Content Above
Btu/lb
C x 106
-.0383
-.502
-.0772
-.0387
-.116
-.193
-.643
-.541
-.0034
60° F
Temperature
Range. "F
60-2000
(33,34)
60-2OOo(34)
Coal
-12.2
.1879
.253
60-8OO
Also used for impurity content of acceptor.
i
Including products of pyrolysis.
314.
-------
TABLE C-9
Heats of Formation at 25°C, Solids
cal/g mol
AHf
CaS04 -344,090
CaS -113,550
CaC03 -288,280
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 + C02 -10^002
CaO + H2S = CaS + H20, . -14,57O
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. . +27,760
CaS + 3/2 02 = CaO + S02 -109,310
CaS + 2 02 = CaS04 -230,540
Heat of Combustion
Btu/lb mol
Char Carbon 178,540-4T* (35)
* T is maximum temperature to which the char has been exposed,
"Farenheit.
315,
-------
In-O
FIGURE C-l
O.OOO1
30O
500
316-
-------
FIGURE C-2
1000
11OO
12OO
13OO , 14OO
317.
15OO
160O
17OO
-------
FIGURE C-3
-------
APPENDIX D
Comparison of CCDC Process to Hot Carbonate Scrubbing
In 1973 there was a mutual agreement to expand the Scope of Work
under this contract to include an economic comparison between the
CCDC (then Consol) hot gas cleanup process and a conventional gas
desulfurization process. The study was completed in November, 1973,
issued as Consolidation Coal Research Memorandum RM-13451, submitted
to EPA on November 30, 1973, and is enclosed here.
The study compared the cost of the Consol (now CCDC) hot gas
cleanup process to the cost of wet scrubbing and gas desulfurization
by the hot carbonate process. The cost of desulfurization by the
Consol process was about 38$ less (see Table I, page 4 of the enclosed
memorandum).
319,
-------
CONSOLIDATION COAL COMPANY
Research Division
Library, Pennsylvania
November 28, 1973
MEMORANDUM- RM-13451
To : J. A. Phinney
From : G. D. Rutledge and.J. T. Clancey
Subject: Cost of Pressurized Producer Gas Desulfurization
I. Introduction
The scope of work under EPA Contract No. 68-02-1333 requires an
economic comparison of two methods of producer gas desulfurization. This
was done by developing the cost of conventional, low-temperature desulfuri-
zation technology for comparison with that of the lime-based hot gas de-
sulfurization system being developed under the above EPA contract.
Briefly, the hot process desulfurizes pressurized gasifier gas at
about 1650°F in a fluidized bed of a regenerative, lime-based sulfur acceptor.
The desulfurized producer gas is cleaned of particulates and alkali by high
pressure drop cyclones and then is cooled to 1300°F by heat exchange with the
water needed to generate the gasifier steam. The sulfided acceptor is con-
veyed to the regenerator in a stream of C02 and steam, where sulfur is
evolved as H2S at 1300°F. The regenerated acceptor is returned by gravity
to the desulfurizer. The H2S in the regenerator offgas is converted to
elemental sulfur with dilut aqueous H2S03 in the liquid-phase Glaus reaction.
Unreacted gas is recirculated to the regenerator. The spent aqueous H2S03 is
then reconcentrated in a packed column by absorbing S02. The S02 is generated
from burning one-third of the sulfur produced.
The Annual Report(l) under EPA Contract No. EHSD-71-15 details the
system under study and also gives an economic analysis of its cost. Two
methods of particulate and alkali removal were covered in that study; hot
removal only in high pressure drop cyclones and wet removal with water in
Venturi scrubbers. The description and the economic analysis for the hot
particulate removal method is used in this study as the basis for evaluation.
In the conventional desulfurization process, the hot gasifier gas
must be cooled by heat exchange and wet scrubbing. The gas then can be de-
sulfurized by contact with one of the commercially available chemical or
physical liquid absorbents. To recover heat from the hot gas, the cool
desulfurized gas is reheated to about 1290°F. Absorbed H2S is stripped
320.
-------
November 28, 1973 - 2 - RM- 13451
from the liquid absorbent and converted to elemental sulfur in the standard
Glaus reaction. The high sulfur content of the Claus tail gas also neces-
sitates the inclusion of a tail gas clean up step. Depending upon the
method selected, sulfur can be recovered directly as elemental sulfur or as
a concentrated stream of either S02 or H2S for recirculation to the Claus
converters.
The hot potassium carbonate process was selected as the conven-
tional method of sulfur removal. The selection v;as made primarily because
the technology for design is available in the literature.*2»3»4»55 Although
the basic hot potassium carbonate process is reported to be selective^)
for H2S, there are no known commercial applications. Lack of published
data on the selectivity prevents the evaluation of this design feature. The
potential problem with hot potassium carbonate may be the inability to
remove the non-acidic sulfur .gases (COS, CS2 and thiophene) formed in coal
gasification. If these gases are formed in sufficient amounts, then one of
the selective, physical solvents may be the only viable conventional desul-
furization process.
To complete the conventional system, the direct oxidation Claus
process was selected for sulfur recovery followed by the Wellman-Lord SOg
recovery process for tail gas cleanup. The low K2^ concentration (about
8 volume %) in the acid gas feed necessitates the use of direct oxidation.
The Wellman-Lord S02 recovery process was selected because low S02 effluent
levels (<100 ppmv) in the stack gas have been achieved commercially.(?)
These two conventional processes are comparable to the liquid-phase Claus
used in the hot gas desulfurization system.
II. Summary
An economic evaluation of desulfurizing pressurized producer gas
is required under this EPA contract. The purpose is to determine the
potential economic incentive for further development and pilot plant work,
The study is restricted to the cost of desulfurizing the producer gas. The
cost of gasification is not included, nor are estimates developed for the
cost of generating electric power from the desulfurized producer gas.
For the evaluation, the plant size was selected to supply a de-
sulfurized fuel gas to a conventional, gas-burning power station with an
approximate rating of 1400 MW. With respect to quantity of gas, this is
the same size plant that was considered in the previous studies.
In the study, a conventional, low-tercperature desulfurization
system is evaluated and compared to the hot gas desulfurization system. The
hot gas desulfurization system removes sulfur via regenerative-acceptor and
recovers sulfur via liquid-phase Claus. The conventional desulfurization
system removes sulfur via hot potassium carbonate and recovers sulfur via
direct oxidation Claus and Wellman-Lord S02 recovery.
The results of the study show a significant advantage is to be
gained over conventional technology by the development of hot gas desulfuri-
zation. The advantage is about $2.50 per ton of feed coal, or 13.1£/MM Btu
321,
-------
November 28, 1973 - 3 - RM- 13451
of product gas. About one-half of the cost saving (5.3£/MM Btu) would be
gained from the development of a hot gas clean up technology for alkali
and particulate removal to replace the conventional wet scrubbing and gas
reheat. (Note: the last Annual Report^) developed the cost of incor-
porating wet scrubbing and reheat into the hot gas desulfurization system.
These costs, $4.98 per ton of feed coal and 25.4? per MM Btu (HHV +
sensible heat) of product gas, can be compared with the desulfurization
costs in Table IV, Section V.) The remainder of the advantage (7.8£/MM
Btu) is gained from the lower direct operating cost of the hot system.
The higher operating cost of the conventional system is chiefly caused
by the following: 1) the reboiler import steam and cooling water require-
ment of the hot potassium carbonate process, and 2) the fuel gas necessary
to preheat the feed gas and incinerate the tail gas of the direct oxidation
Claus process.
The evaluation results, based on an arbitrary start-up date of
January, 1978, are summarized in Table I. In addition to the cost ad-
vantage shown in Table I for hot (Case I) over conventional desulfurization
(Case II), there is also a greater amount of excess power available from
expansion of the product gas (about 52 MIV equivalent). While it is beyond
the scope of this study to evaluate the dollar value of this excess powerD
it could amount to an additional credit of 0.3-0.6?/MM Btu of product gas.
The cost of the plant facilities per installed KW (1978 cost
basis) to desulfurize producer gas under pressure amounts to $32 for Case I
versus $46 for Case II. The annual operating cost corresponds to 1.5 and
2.4 mills/KWH of power generated for Cases I and II, respectively.
The make-up acceptor cost for this evaluation was assumed at a
relatively high value of $10 per ton of raw dolomite. The make-up acceptor
requirement for the regenerative acceptor case is, however, quite modest,
and the acceptor cost does not have a significant impact on the desulfuri-
zation cost. In this study, the rate of make-up acceptor required is
estimated as one per cent of the circulating acceptor. This is based on a
reasonable extrapolation of the laboratory data obtained to date. If the
make-up rate were doubled, the desulfurization cost would increase only by
4-5%.
Interestingly, the sensible heat content of the desulfurized
producer gas delivered to the power station is a significant factor. In a
conventional, gas-fired power station, the fuel gas is delivered to the
burners at ambient temperature. The combustion air is preheated to 250 to
350°F in an air heater by indirect sensible heat exchange with flue gas.
In this process the fuel gas is delivered from the expander to the power
station at 645-660°F. If the fuel gas is preheated, then either less air
heater surface is required, or more steam can be generated per unit of
fuel burned. An exact determination of the dollar value of sensible heat
in the fuel gas would require an extensive analysis of the design and cost
of the power station proper. This is beyond the scope of this study.
322.
-------
to
OJ
Table I
Summarized Economic Evaluation
70% Plant Operating Factor (6132 hrs/year)
Case
Method of Gas Desulf urization
Desulfurization Process
Coal Required:
MM Tons/Year (6% Moisture)
Higher Heating Value, Btu/Year
Producer Gas to Station (After Expansion):
Mo Is /Hour
Temperature
Pressure
Higher Heating Value, Btu/Year
HHV + Sensible Heat Content/1' Btu/Year
Desulfurization Plant Investment^)
/0\
Desulfurization Cost: v"'
$/Ton Feed Coal
C/MM Btu IUIV of Food Cool
Hot
Regenerative Acceptor
II
Conventional
Hot Potassium Carbonate
Btu HlfV of Product Gns
wor tiunurntod by Expander,
213,700
660°F
10 psig
65.19 x 1012
68.68 x 1012
$48.3 MM
4.03
16.9
21.2
20.1
160(/t)
—3,422,000
-81.71 x 10
,12
o
A
§•
(D
to
3C
CD
-»J
W
183,957
645°F
10 psi|
64.63 x 1012
67.41 x 1012
$68.8 MM
6.55
27.4
34.7
33.2
109
M)
(1) Sensible heot content above an assumed air heater outlet temperature of 300°F,
(2) January, 1978 operation - includes escalation and interest during construction.
Does not include gasification plant cost or the cost of coal,
(3) After deduction of 157 MW required for compression of process air for gasification plant.
(4) Assuming 91% of the isentropic efficiency for the expander, and 89% of the polytropic
efficiency for the air compressors.
GDR: mm
w
^
01
-------
November 28, 1973 - 5 - RM- 13451
However, a reasonable approximation of the value is to credit the system
for the sensible heat content of the fuel gas above the temperature of
the flue gas exiting the air heater. In this study this temperature is
assumed as 300°F. In Table I, the credit to the process for this excess
sensible heat is about 1.0-1.5$/MM Btu or about 5% of the system cost.
III. Recommendations and Conclusions
Based on the economic evaluation in this report, a significant
advantage exists for hot gas desulfurization over conventional desulfuri-
zation. The magnitude of the advantage is sufficient incentive for
further development and pilot plant work. Sufficient incentive exists
for the hot gas desulfurization process even if wet scrubbing and reheat
are ultimately necessary to produce adequate quality gas for an expander.
This is because of the large energy requirement (import steam and fuel
gas) of the conventional desulfurization process. The relative incentive
for (1) hot sulfur acceptance (including wet scrubbing and reheat + liquid-
phase Claus sulfur recovery) is $1.57 per ton of feed coal or 7.8C/MM Btu
(HHV + sen. heat) of product gas and (2) hot solids removal Is $0.95 per
ton of feed coal or 5.3£/MM Btu of product gas.
IV. Process Description
A. General
The simple block flow diagrams in Figure 1, hot gas desulfuri-
zation, and Figure 2, conventional gas desulfurization, define the steps
in the overall process. The design and evaluation of coal preparation
(crushing and sizing) and of coal gasification (Section 100 and 200,
respectively), are not covered in this study.
In the hot gas desulfurization process, Section 300, sulfur removal
by regenerative acceptor, and Section 400, sulfur recovery by liquid Claus,
have been completely evaluated. A relatively small quantity of make-up CC>2
is required in the process (Section 500). As shown in Figure 1, the CC-2
could be recovered either from a slip-stream of the desulfurized producer
gas before power generation, or from a portion of the power station stack
gas. Preliminary cost studies indicate approximately equal costs for the
two alternative locations. For the purposes of this study, it is assumed
that the make-up C02 is recovered from the pressurized producer gas before
delivery to the power station. A standard hot potassium carbonate system
is assumed to establish capital cost and utility requirements.
In the conventional gas desulfurization process, Section 300,
sulfur removal by hot potassium carbonate, Section 400, sulfur recovery by
direct oxidation Claus, and Section 500, S02 recovery by Wellman-Lord
process, have been evaluated. Section 600, wet scrubbing (particulate and
alkali removal), and reheat has been evaluated as an integral part of the
sulfur removal section.
-------
Figure 1
Block Flow Diagram - Hot Gas Desulfurization Overall 'System
Regenerative Acceptor Case
10
S3
Ul
Feed
Coal __
Coal
Preparation
Section 100
Gasification
Section 200
,
Desulfurized
Producer Gas
Producer
Gas
i
Make-up f
CO,
\ Removal !
1
Section 500 '
i
|
i
1
Make-up
co2
*
Sulfur
Removal I
"^
Section 300 i
KW
1
Power
"" Station
Sulfur
Recovery
Jcction 400
1
Stack z
Gas <
t i
o"
1 O
i i
1 Make-up ' »
•yl C0? '
Removal 3
1 ' -i
CO
I Section 500
i >
I
i
J
i
o
i
Ash
Note: Alternative locations of Section 500 shown by dotted lines.
Sulfur
I
l-i
CO
Ol
-------
Figure 2
Block Flow Diagram - Conventional Gas Desulfurization Overall System
Hot Potassium Carbonate Case
KW
o
ffi
a
cr
o
to
oo
OJ
N>
ced Coal
Coal
Preparation
Section 100
Producer
Gas
Gasification
Section 200
Wet Scrubbing
and
Reheat
Section 600
Power
Station
Desulfurized
Producer Gas
Sulfur
Removal
Section 300
Sulfur
Recovery
Section 400
Stack
Gas
S02
Recovery
Section 500
Ash
Sulfur
I
I-'
w
en
-------
November 28, 1973 - 8 - RM-13451.
B. Hot Gas Desulfurization
Sulfur Removal by Regenerative Acceptor
A schematic flow diagram of the proposed commercial embodiment
of the hot gas desulfurization process /is shown in Figure 3 (Consol Divg.
No. AF-3452).
Gasifier product gas (Stream No. 4) flows to the bottom of a
fluidized gas desulfurizer (D-301). The gas fluidizes a bed of sized
dolomite at a temperature of 1652°F and at 15 atrn. absolute pressure.
Most (97%) of the H2& in the gasifier gas reacts with the CaCOs component
of the dolomite as follows:
H2S + MgOCaCOa = MgO-CaS + C02 + H2O
The desulfurized gasifier gas then passes through two stages of cyclones
(G-301 and G-302, respectively), to remove substantially all of the en-
trained dolomite fines. A portion of the sensible heat content of the
desulfurized gas is used to generate and superheat the steam required for-
gasification of the feed coal in heat exchangers C-302 and C-301,
respectively. The desulfurized producer gas, at 1300°F and 210 psia
(Stream No. 6), is then delivered to the power station. Subsequent use
of this gas is dependent on the design of the power station, and is
discussed later in this report.
Regenerated acceptor (Stream No. 5) at 1300°F is continuously
charged to the top of the gas desulfurizer, and is continuously withdrawn
via a standleg (Stream No. 7). The sulfided acceptor is pneumatically
conveyed by a portion of the required recycle gas (Stream No. 9) to a
fluidized acceptor regenerator (D-302). Make-up acceptor (Stream No. 8)
is introduced into the system by a lockhopper system comprised of L-3Q1,
F-301, F-302, and L-302. This small stream is also pneumatically conveyed
to D-302 by means of recycle gas. Make-up C02 required in D-302 (Stream
No. 10) is pressurized by JC-301 and charged to D-302.
The fluidized acceptor regenerator (D-302) is maintained at 1300°F
and 15 atm. absolute pressure. The sulfided acceptor is recarbonated at
these conditions by the reverse reaction:
MgO-CaS + C02 + H20 = MgO-CaCOs + H2S .
The carbonated magnesium component of the make-up acceptor is also calcined
at these conditions:
MgC03-CaC03 = MgO-CaC03 + C02 .
The regenerated acceptor is returned to the gas desulfurizer by
gravity flow (Stream No. 5). Spent acceptor (Ifs of the circulating flow)
is withdrawn from D-302 via a lockhopper, F-303, and a rotary feeder, L-305
(Stream No. 11). This spent acceptor must be treated before disposal.
327.
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DIGITALLY
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November 28, 1973 - 10 - RM-13451.
About 75% of the calcium component of this stream is in the form of CaS.
If this were disposed of directly to the station ash pit, H2S gas would
slowly evolve as the CaS was hydrolyzed. To avoid this unacceptable
condition, the spent acceptor is directly contacted with C02 and water
in three stages of stirred reactors, D-303, to convert the CaS to CaC03-
At these conditions, the MgO component of the dolomite would also recar-
bonate. The overall reaction is,
MgO- CaS + 2 CC-2 + H20 = MgCOs-CaCOa + H2S.
The spent acceptor is conveyed hydraulically by a dilute re-
circulating slurry stream to a hydroclone, G-305. The underflow slurry
(controlled at 35 wt.% solids) flows to the three series converters,
D-303A, B and C, respectively. Each converter is agitated by a turbine-
type mixer, L-305, and some reaction heat is renoved In each stage by
cooler, C-305. The spent acceptor, stripped of sulfur and fully carbonated,
is pumped to the station ash pit by slurry pump, J-301. The overflow slurry
from hydroclone, G-305, is continuously recirculated via surge drum, F-304
and circulating pump, J-303. The sensible heat loss from the spent
acceptor is removed by cooler, F-304. Make-up water to this system Is
supplied by station pond overflow via J-302 (which also supplies the water-
to cool the gasifier ash and convey it to the station, ash pond).
Acid gas resulting from the acceptor stripping operation (Stream
No. 16), is compressed to system pressure by JC-302 and flows to the
liquid-phase Glaus reactor.
Sulfur Recovery by Liquid-Phase Glaus
The process gas exiting D-302 (Stream No. 12) Is passed through
two stages of cyclones (G-303 and G-304, respectively), to remove entrained
acceptor. The sensible heat content of this stream is then exchanged with
recycle gas in exchanger C-401, and with boiler feed water (required for
gasification steam) in C-402. An electrostatic precipitator, L-401, is
provided to remove all entrained dust. The gas then flows to the bottom
of the liquid- phase Glaus reactor, D-401.
The concentration of H2S in the gas from the acceptor regenerator
is limited by the equilibrium restriction at 1300°F and amounts to only
about 3.6 volume per cent. The liquid-phase Claus reaction developed in
this work is uniquely suited to processing this gas. The liquid-phase
Claus reactor, D-401, operates at 310°F and about 210 psia. Product gas
from the acceptor regenerator (Stream No. 12), gas from the reject acceptor
stripping section (Stream No. 16), and dilute, aqueous H2S03 (Stream No. 17)
are charged to the bottom of the liquid- phase Claus reactor. Flow is upward
through a sparged reactor containing only water, as was demonstrated during
the experimental work described in the annual report(i) to EPA. Liquid
sulfur is produced by the reaction,
2 H2S + H2S03 = 3 S + 3 H20.
330.
-------
November 28, 1973 - 11 - RM-13451.
Liquid sulfur and liquid water flow from the reactor to a decanter-
type separator, F-402. Unreactcd gas (Stream No. 9), saturated with water
vapor at 310°F, is compressed by JC-401, reheated in exchanger, C-401 to
975°F and returned to the acceptor regenerator reactor. The sensible heat
content of the liquid water from F-402 (Stream No. 19), is exchanged with
the feed acid (Stream No. 17) in C-404, further cooled to 90°F in C-405
and charged to the S02 absorption tower, D-402.
Approximately one-third of the sulfur from F-402 is burned with
stoichiometric air (Stream No. 20) in pressurized combustor, B-401, to
produce 803. Excess heat is removed to boiler feed water via cooling tubes
in the walls. The exit gas from the sulfur combustor (Stream No. 21) flows
to the base of the S02 absorption tower, D-402. tfater (Stream No. 19)
flowing down through the packed tower absorbs the S02 in the gas by,
S02 + H20 = H2S03 aqueous
Most of the exothermic heat of reaction is removed by side stream
cooler, C-406. The vent gas from the absorption tower (Stream No. 22) is
at 90°F and 205 psia. It probably would contain some residual S02 (assumed
in this case as 0.3 volume per cent). The most practical means of dis-
posing of this gas is to bleed it to an intermediate stage of the expansion
turbine required for the desulfurized producer gas in the power station.
This would decrease the sulfur removal efficiency of this process from
97.0% to 96.5%, but would increase the power output of the station slightly,
and would eliminate an entire stack gas scrubbing installation following the
SC>2 absorption step.
C. Conventional Gas Desulfurization
Sulfur Removal by Hot Potassium Carbonate
A schematic flow diagram of the proposed commercial embodiment
of the conventional gas desulfurization process is shown in Figure 4
(Consol Dwg. AF-3451).
Gasifier product gas (Stream No. 4) at 1750°F and 225 psia passes
through two stages of cyclones (G-301 and G-302, respectively), to remove
the bulk of the residual dust and is then cooled prior to entering the
absorber (D-301). A portion of the sensible heat content of the gas is used
to generate and superheat steam required for gasification of the feed coal
in heat exchangers C-301, C-302 and C-303, respectively. The remainder of
the sensible heat is used to reheat the desulfurized gas leaving the
absorber to 1288°F in heat exchangers C-601, C-602 and C-603, respectively.
The gasifier gas, now cooled to 425°F, nust be wet scrubbed for
particulate and alkali removal prior to contacting the hot carbonate solution.
The cooled gas flows to three parallel Venturi scrubbers (L-601) where it is
scrubbed by circulating water at the gas dewpoint of 228°F. The circulating
water flows to the Venturi scrubbers via the black water pumps, J-602. The
331.
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PAGE NOT
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DIGITALLY
-------
November 28, 1973 - 13 - RM-13451
excess sensible heat in the gas is removed via coolers, C-604, and make-up
water enters the system via water pumps, J-601. The make-up water rate is
based on an assumed solids loading in the process gas of 5 grains/cubic
foot, and a solids content in the circulating water of 2 wt.%. Net solids
removed from the gas flow result in a 2 wt.% slurry to the station pond.
Clean gasifier gas (Stream No. 5) at 228°F and 206 psia flows
to the bottom of the absorbers (D-301) where the gas is contacted counter-
currently in a packed bed with a 25 wt.% solution of hot potassium
carbonate. The flow of lean carbonate to the absorbers is split and
cooled in the lean carbonate coolers (C-304 and 305) to ensure adequate
absorption of H2S. Most (97.5%) of the H2& and the bulk (88.6%) of the
C02 in the gasifier gas is absorbed in the hot carbonate solution, leaving
a desulfurized gas with an H2S content of about 190 ppm by volume. The
desulfurized gas exiting the absorbers is then reheated to 1288°F by the
aforementioned exchangers (C-601, C-602, and C-603). The desulfurized
producer gas at 1288°F and 196 psia (Stream No. 6) is then delivered to
the power station. Subsequent use of this gas is dependent on the design
of the power station, and is discussed later in this report.
The rich carbonate solution leaving the absorbers is let down to
95 psia by power turbines, which drive in tandem with electric motors the
lean carbonate pumps (J-301). The rich carbonate then enters the strippers
(D-402) where the pressure is reduced to 27 psia and the solution Is
stripped with steam. Steam in the stripped acid gas is condensed in
water-cooled condensers (C-306) and collected in accumulators (F-301).
Some of the condensate is returned to the strippers via reflux pumps
(J-302) to maintain water balance and the excess is sent to condensate
recovery. The stripped acid gas (Stream No. 7) at 110°F and 21 psia
contains about 8% H2& and is suitable feed to a Claus sulfur recovery
unit of the direct oxidation type.
About one-sixth of the reboiler (C-307) low pressure steam re-
quirement is produced by the sulfur recovery and the 803 recovery sections.
The other five-sixths must be imported.
Make-up carbonate solution is stored in the carbonate storage tank
(F-302) and enters the lean carbonate line via the carbonate make-up pump
(J-303).
Sulfur Recovery by Direct Oxidation Claus
Acid gas (Stream No. 7) containing about 8% J^S flows to the
sulfur recovery section at 110°F and 21 psia where it is indirectly preheated
in a furnace prior to entering the first converter. Oxygen in an air stream
reacts directly with H2S over the converter catalyst to produce S02 and
sulfur. The hot effluent gas (except bypass for reheating) flows to the
first sulfur condenser where sulfur is condensed. The gas leaving the
condenser is directly reheated to reaction temperature with about a 15%
bypass of converter hot effluent gas. The gas then enters the second
converter to obtain equilibrium conversion of H2S and S02 to sulfur and
water. The sulfur produced here is condensed in a second sulfur condenser.
333.
-------
November 28, 1973 - 14 - RM-13451
»
Overall sulfur recovery is about 88%. The first plant*8'9' utilizing this
process obtained an 82% sulfur recovery from an acid gas feed containing
10% H2S (dry basis).
The sulfur condensers generate 90,000 Ib/hr of low pressure
steam (70 psia) while condensing 471 long tons per day of sulfur product.
The tail gas leaving the last condenser passes through a
coalescer to remove entrained liquid sulfur and is incinerated at 1000°F.
The incinerated tail gas then flows to the S02 recovery section. The re-
covered SO2 is returned to the converter where it reacts to produce
additional sulfur. This gives a net sulfur recovery of 99.8%.
SOa Recovery by Wellman-Lord Process
Hot gas from the incinerator is cooled in a waste heat boiler,
then quenched and fed to the S02 absorber, where 5O2 is absorbed in a lean
sodium sulfite solution. The clean gas (Stream No. 8) from the absorber
passes to the stack. The concentrated stream of SO2 recovered from the
sodium sulfite solution in the regeneration system is recycled to the
sulfur recovery section.
Ib/hr.
The process net production of low pressure steam is about 70,000
V. Economic Evaluation
A. General
This evaluation is restricted to the desulfurization of a producer
gas from a pressurized coal gasifier for delivery to a conventional gas-
burning station at a pressure of 10 psig. Power recovery from the desul-
furized producer gas from system pressure (about 200 psig) to 10 psig is
best accomplished in an expanding turbine-generator set. The power generated
by the turboexpander exceeds the power required for the coal gasifier air
compressor. The excess power recovered is a net credit to the process.
However, it is beyond the scope of this contract to define the actual dollar
value of the power credit.
In hot gas desulfurization, the technical question remains as to
the required quality of gas (with respect to content and size of particulate
matter and alkali) that may be charged to an expander. The maximum power
obviously will be recovered from the turboexpander if the producer gas can
be adequately "cleaned" by high pressure drop cyclones alone. Conversely,
the gas may be water scrubbed (in conventional venturi scrubbers) and then.
reheated by indirect heat exchange with hot gas to the scrubbers. This
assures the protection of the turboexpander from particle erosion and/or
alkali deposition. The penalty is a reduction in power recovery, because
gas reheat is only economical to about 1100°F, and an additional pressure
loss is incurred through the scrubber and reheat system. There is no
incentive to eliminate the venturi scrubber for conventional, low-temperature
334.
-------
November 28, 1973 - 15 - RM-13451
•
desulfurizntion because the gas must be cooled anyway. The Annual Report^ ^
to EPA covered this technical uncertainty in the hot gas desulfur.ization
process. The process Incorporates only the hot particulate removal method
for this study.
In both cases, the producer gas after expansion to 10 psig Is
delivered to the power station as preheated gas (645°F to 660°F). This
sensible heat content serves to improve the efficiency (or the heat rate)
of the station proper. In a conventional gas-burning station, the fuel
gas is delivered at ambient temperature, and sensible heat from the
products of combustion is recovered (via combustion air preheat) to an
air heater exit temperature of about 300°F. It seems most realistic to
credit the process with the excess sensible heat in the preheated producer
gas above 300°F as an approximate means of measuring the improvement in
the power station efficiency. .
B. Plant and Operating Cost Estimates
The capital cost estimate for hot gas desulfurization was de-
veloped by the Consol Engineering Department. The estimate for Sections
300 and 400 of Figure 3 is summarized in tabular form in the Appendix as
Tables A-l and A-2. A detailed estimate was not made for Section 500,
the make-up C02 system, because a conventional "hot pot" unit was assumed.
The costs and utility requirements for this type of system are adequately
defined in the literature.'*0) Similarly the Consol Engineering Department
developed a capital cost estimate for conventional desulfurization. The
estimate for Sections 300 and 600 of Figure 4 is summarized in the Appendix
as Tables B-l and B-2. For the remaining sections, Section 400, the sulfur
recovery system (a direct oxidation Claus unit), and Section 500, the S02^
recovery system (a Wellman-Lord unit), cost and utility requirements were
obtained from the literature.^7» -^
The capital cost estimates developed in the Appendix are defined
as inside battery limits (ISBL). To these costs must be added the off-site
facilities (OSBL) which include utility costs, electric substations, cooling
water towers, distribution piping, and boiler feed water treating. The
accuracy of the capital cost estimates in this study is estimated as - 20%.
It was required in this study to anticipate the escalation in
these costs to a presumed start-up date of January, 1978. The economic
climate at present defies a logical derivation of these factors. This work
is based on these arbitrary assumptions:
1. July, 1973 costs as a base point.
2. Begin design and construction, January, 1974.
3. Begin operation, January, 1978.
4. Escalation at 7^% per year average of construction labor
and material.
5. Interest during construction at 7^% simple interest on
cash flow.
6. Escalation of direct operating labor at 5^% per year.
7. Interest on working capital at 7^% per year.
335,
-------
November 28, 1973 - 16 - RM- 1345L
The total investment summary for the two cases described above
is shown in Table II. The installed plant cost (present day cost) for
the hot gas desulfurization scheme (Case I) is $35,800,000, whereas, that
for the conventional gas desulfurization scheme (Case II) is $49,600,000.
The greater plant cost for the conventional system is mostly associated
with wet scrubbing and reheating. Escalation of these costs to January,
1978 (as defined above) increases the total investment to $48,300,000
and $68,800,000, respectively.
Direct operating costs for these two cases were estimated as
shown in Table III. Note, that an average plant operating factor of 70%
was assumed reflecting power station practice. The unit values used to
develop operating costs are shown directly in Table III. The total direct
operating cost (January, 1978) for Case II ($10,834,000) is double that
of Case I ($5,321,000). This is a result of the large import steam re-
quirement for regeneration of the hot carbonate solution and of the fuel
gas requirement for preheat and incineration within the Glaus sulfur plant.
C. Evaluation of Desulfurization Cost
The Regenerative Acceptor process based on the results of this
economic evaluation has a potential for being substantially cheaper than
the conventional liquid absorption processes. Based on the cost estimates
presented in Tables II and III, the system costs are estimated as shown
in Table IV. The net annual operating costs for Case I (hot regenerative
acceptor) are $13,810,000 per year and for Case II (hot potassium carbonate)
the costs are $22,416,000 per year. Approximately from one-half to two-
thirds of these costs are represented by capital charges. The low value
for sulfur credit ($8.00/long ton) reflects a conservative view of the
sulfur market.
. The make-up acceptor cost has little impact on the hot regenerative
acceptor process cost. Since this study does not define a specific plant
site, a definitive cost for delivered acceptor cannot be derived. An
arbitrary cost of $10.00 per ton of raw stone (MgCOsCaCC^) was assumed.
This somewhat high cost is based on two facts: a) the laboratory studies
to date have shown that some dolomitic stones cannot be used in the
process (excessive attrition) thereby limiting the potential supply sources,
and b) transportation costs for raw materials are substantially increasing.
However, even with an acceptor make-up cost of $10.00 per ton, the annual
cost is only 4-5% of the total system cost. The imprecision of the acceptor
cost has only a minor influence on the evaluation of the total system.
The system costs presented in Table IV are more meaningful if -
related to either the coal required for the system, or to the quantity and
quality of the gas delivered to the power station. The net annual operating
costs for Case I are equivalent to $4.03 per ton of feed coal; for Case II,
$6.55 per ton. Expressed in terms of heating value of the coal, the cor-
responding costs are 16.9 and 27. 4C/;.c.! Btu, respectively.
336.
-------
Tibia II
Investment Summary
LO
Cose
Method of Gas Desulfurlcatlon
Desulfurlzatlon Proceso
Plant Section
.Utilities Required
Electricity, KW
Cooling Water, gpo
Low- (treasure Steaa, Ib/hr
Boiler Feed Water, gpn
Fuel Gas, MM Otu (HIIV)
Operating Labor Required
lien, Shtlt
Investment
Erected Cost (ISDL)
Off-altos fa Utilities (OSBL)
OU-aitoa
Electrical
Cooling Water
Doller Feed Water
Total OSUI,
Installed Plant Cost
(July, 1973)
Escalation to Jan. 1978
Sub-Total
Interest During Construction
Total Investment
(Jan., 1978)
I
Hot
Regenerative Acceptor
300
Sulfur
Removal
2,490
1,870
X
X
X
4
15,000
1,200
ZOO
X
_ *
1,400
16,400
400
Sulfur '
Recovery
10, 030
10,340
X
1,234
X
3
11,400
900
700
400
3OO
8,300
13,700
SOO
lloke-up
CO,
Syatea
344
16, 185
78,300
X
X
1
3,000
200
X
600
_JU.
800
4,700
Total
12,864
28,395
78,300
1,234
X
8
30,300
2,300
900
1,OOO
30O
4,500
94,800
8.500
41,300
^000
48,300
300
Sulfur
Reaoval
4,136
87, 450
746,000
1.234
X
1
(In $1000)
23.400
1,800
300
3,100
300
6,500
is.eoo
II
Conventional
Hot Potasalua Carbonate
400
Sulfur
Recovery
633
X
X
180 .
246
*
3,500
200
X
X
X
300
8,700
' SOO
802
; Recovery
970
9
85,500
271
X
*
4,200
300
100
X
.IPO
SOO
4,700
600
Bet Scrubbing
and Reheat
401
19.800
X
X
f
1
10,800
BOO
X
700
*
1,500
12,300
Total
6.142
107,259
811,500
1,685
X
3
41,900
3,100
400
3.800
•100
7,700
49,600
r9.300
58,900
9,900
68,800
GBRiBS
-------
U)
OJ
oo
Baals: 70% Plant Operating Factor
Case
Method of Cac Iteoulf urliatlon
Oesulfurlxatlon Process •
Plant Section
Direct Operating Labor
Direct Operating Cost
1. Operating Labor at
$63, 200/man/ohif t/yr
2.
Maintenance Labor at
1.6% Installed Plant Cost
3. Direct Supervision
107. of I * 2
4.
Indirect Overhead
SOT. of I * 2 » 3
Toblo III
Direct Operating Cost Sunaary
Occluding Acceptor Coat for Case I
3. Payroll Overhead
19". ol 1 « 3 « 3 * 4
6. Maintenance Uatorlal
3.4% Installed Plant Coot
7, MU<-ollaiioi«j« SupplleB
13% tUlnt, tutorial
8. Utilities
Electricity at 9 Bill/Km
Cooling Water at 3«/ti gal.
erv at aoo/u gal.
lov-Pressuro Steam at 73V/H Ib
Fuel Cas at SOe/tai Dtu (HJIV)
Cheaical* b Catalysta (ex Accepter for Ceoo I)
Sub-Total Utilities
Totala
Escalation to Jan., 1978
Total Plraet Operatins Ce*t (Jan. 1878)
CD Iteos 8-7 am investment eensltlvo and aro oeaalated ea tko eeao basla as plant
lavesteent, (teas 1 and 6 era at eoauaatf January 1978 values,
I
Hot
Regenerative Acceptor
300
Sulfur
•Removal
4
333
362
77
296
133
394
39
138
31 '
X
X
X
Jl
139
17633
400
Sulfur
Recove ry
3
190
220
61
233
106
338
49
933
114
136
X
X
g
~ "
803
T755?
300
Make-up 7
COz
System
1
63
79
21
80
30
113
17
19
179
U
360
x
13
T
371
""STe
Total
8
. 306
337
139
611
273
833
123
710
.314
138
380
X
13
1,333
«75o7
720
8,331
300
Sulfur
Removal
1
(in SlOOO/yr)
63
462
79
302
136
604
104
329
963
136
8,431
x
73
~~
4,336
3757?
II
Conventional
Hot Pot OB slum Carbonate
400
. Sulfur
Recovery
$
32
39
14
32
24
80
13
39
x
80
x
1,208
30
~"
1,293
TTsTf
300
5°2
Recovery
*
32
78
16
62
38
113
17
84
x
30
X
X
179
•
863
"~606
600
Vet Scrubbing
and Reheat
1
63
197
39
130
67
303
44
33
219
K
X
X
x
~
341
17096
Total
3
190
793
148
36G
333
1,101
178
340
1,184
180
3,01
1,208
284
6.633
97934
_»M
10,834
• |,
•,
8
M
U
1
M
CD
?
U
V
•
ODtihw
-------
November 28, 1973
- 19 -
RM-13451
Table IV
Bconomlc Evaluation
Hot vs Conventional Dosulfurizatlon Process
Basis: 70% Plant Operating Factor (6132 hr/yr)
Cose
Method of Gas Deaulf urizntion
Desulfurlzatlon Process
Coal Required:
Tons/yr (6% Moisture)
Higher Heating Value, Btu/yr
Deeulfurized Producer Gas to Station
Mols/Hr
Temperature
Pressure
Higher Heating Value, Btu/yr
HHV + Sensible Heat Content, Btu/yr(1)
Cost Analysis
Installed Plant Cost (1973)
Escalation to 1978
Interest during Construction
Total Investmant
Working Capital
Annual Operating Costs
Direct Operating Cost (1978 Basis)
Acceptor at SlO/ton
Interest on Working Capital at 7.3%
Capital Charges at 18% Investment
Sulfur Credit at $8/icetrlc ton
Ket Annual Operating Cost
Deaulfurization Cost Expressed;
In terms of feed coal
$/ton coal
C/MM Btu HHV
In terns of product gas to station:
C/ilM Btu HHV (u
C/MM Btu (HHV + Sens. Ht.)
Excess Power Generated by Expander
I
Hot
Regenerative Acceptor
II
Conventional
Hot Potassium Carbonate
• 3,422,000 —
81.71 x 1012-
213,701
660°F
10 psi|
r»"
,12
65.19 x 10
68.68 x 10
$34.8 MM
6.S MM
7.0 UM
$48.3 MM
S 1.33 MM
$ 5.321 MX/Year
0.648 rat/Year
0.100 tK/Year
8.694 NX/Year
(S 0.953 MM/Year)
313,810 lC.!/Year
4.03
16.9
21.2
20.1
160 IIW
183,957
643«F
10 psle
64.63 x 10"
67.41 x 1012
$49.6 MM .-
9.3 MM
9.9 UM
$68.8 MM
$ 2.IS UM
910.834 MM/Year
x «M/Year
0.161 KM/Yoar
12.384 UNI/Year
($0.963 MM/Year)
$22,416 UK/Yenr
6.53
28.4
34.7
33.2
109 UVT
(1) Sensible heat content above an assumed air heater outlet temp, of 300°F.
CDR:hw
339,
-------
November 28, 1973 - 20 - RM-13451
Considering only the higher heating value of the product gas
to the power station, the corresponding costs are 21.2 and 34.7
-------
Bibliography
1. Curran, C. P., et al, "Development of the C02 Acceptor Process
Directed Towards Low-Sulfur Boiler Fuel," Annual Report to
Control Systems Division Office of Research and Monitoring
Environmental Protection Agency under Contract No. EHSD 71-15,
March 15, 1972-June 30, 1973. (Not yet published.)
2. Benson, H. E. and J. H. Field, Petroleum Refiner. 39, 127,
April, 1960.
3. Bocard, J. P. and B. J. Mayland, Hydrocarbon Processing and
Petroleum Refiner, 41, 128, April, 1962.
4. Buck, B. 0. and Angus R. S. Leitch, Petroleum Refiner, 37, 241,
November, 1958.
5. Goolsbee, J. A. and H. K. McLaughlin, Petroleum Refiner, 39. 159,
January, 1960.
6. Garrett, E. Pack, the Oil and Gas Journal, p. 76, December 15, 1969.
7. Hydrocarbon Processing (NG/LNG/SNG Handbook) .52, 116, April 1973.
8. Grekel, H. et al, Chemical Engineering Progress, 61, 7O,
September, 1965.
9. Chemical Engineering, p. 38, April 1, 1964.
10. Katell, Sidney and John H. Faber, Petroleum Refiner, 39, 187,
March, 1960.
11. Grekel, H. et al., the Oil and Gas Journal, p. 88, October 28, 1968.
RM-13451
341.
-------
Appendix A
Detailed Investment Costs
for Hot Gas Desulfurization
342.
-------
Ul
10
-P-
u>
tnultaeot Buaber
Be Jar 8teaa Supar Haatora C-301
Vteto Heat Bailer C-JOJ
IntorcoclOM 0*303
Cuoncb »«ter Cooler C-304
Trio Cooler 0-101
QUOBCB »»tor Coaler C-104
Sulfur Beoctor D-301
Refea. Beactor D-303
Acceptor Converter 0-303
Acceptor feed Bappor P-30O*
Acceptor 3ur(0 BIB P-301
Acceptor Look Hopper f*301
Spent Acceptor Lack Boppar f-303
Queack Ifatar Surce Teak F-JO4
Dual Colloctara . O-301
Dual Colloctaro O-301
Ouat Colloctore O-301
Duet Cellectora O-104
Rydraoloaee O-103
Slurry Puap " J-101
Dlack eater Puap .1-101
0,ueoca Water Clra. Puap J-1O3
•eke-up CO] Coap'r JC-301
Ac Id Caa Coap'r JC- 301
CO, Bluior JC-301
Acceptor *ere* feeder L-30O*
Acceptor Iterator L-301
Acceptor feeder L-301
B Teep. latary feeder L-303
I Taap. totery feedar
B Taap. latary feeder
Controller Agitator L-1O8
Cent tattaat8;
•erlptloa
Section 3OUi Sullar Reaoval
Oaterlal
tabor
0 ta. ISJS^I-I fait, float fid., 169«*r,
110*. 318 S.B. Tube
» la. 315001-1 Plat, float Rd., I6OO*?,
ISO*. 1-1/4 Cr-lHB Tube
1 Sa. 830 a u-Tube "8"d f 1-. in o-JOS
73*. llO-f, 304 S.S. Tube
140OdT. 1-1 Paea. Pit. Hd., 113*
3O4 a. 8. Tuba
173 It), U-Tube, C.B., 10*8 Shell a
10* L. },-4"» Tube • 14 O.B.C.
1400$, 1-1 FBaa. ft*. Bd., 11M
304 a.a. Tuba
8 la., fl'-8"O.D. « 47'fl.l.«., ISO*,
1700V. Befr. Lined, C-SI
1 Ia. ll'-3"O.O. z 38'O.B.*., 1304,
1300'f. Befr. Lined, C-SI
t In. 3' I.D. « 7'-0"0.8.8., 1S»,
130*P. Cone Bat., faD, 104 3.8.
10'zlO>«49*Cono, 1/4 C.8. I(.*/Oratla<,
CORO. Baoe
* la. 7'O.p.xlO'O. 8.B. V/Cone BID, f IB Top,
1/4" C.S. t
4 la. a* I.D.llOOJ.S. D/Cona Ota,
Illlp. Top, 9/8* C.B., 190*
1 la. 7* I.D.»10'O.a.S, Illlp. Hda,
1/4" C.8. I.190ff, I330*f
I' O.D.zl3'0 8.8., 13* Doe, 3/18"-3O4 8.1.
a.,Caaa Bla.
IB la. 11.100 ACfU, C. fltl. t. 1700*r,
S" Oefr. Llnlttj
li la. 11,100 Acra. c-at 0/3- aafr. Lia.,
110*. i7oo*r
4 Be. 14.800 ACm. C.«. 8,. 1130*f,
9" Refr. Llalag
4 la. 14,800 ACrH, C-SI 0/3" Befr. Lla.,
190*, 1330V
171 OPK, 4 Ia. B" Die., Cubbar Lla.
Canted W/Ptpl«f
1 ta. u cm, 11 pal, Auat. 1. Ceatrlf.
at 1790 in
I Ia. 1114 GW, 98 pal. Auat. CBL-3"a
at liao am
1 Ia. 33* Cl«, 108 pal, Auat. Iron Ceatrlf.
et 1730 HIM
7810 ACm, 13-J40 pale, a lt|. w/Brecooler
end lotercool aO.OOO
MO Acrn. 13 »JU pale, I «l«. P'/lalercuuler,
Reolp. 34,ww
1430 ACm, > pel, Bodry «/*.C.
Dearlnta, 300O Rltl 3,>"U
t I*. 300 CTH, 33'IJ «*" 01*. al JO* Slope
I-M C.». 1,100
* ta. 300 Cn, Bl'C-C, Vetutar *C-I1M
11 B*C Caalfif 13,831)
4 ta. 110 cyn.8"8 lot. V«8e, 10O*f,l»0«,C.( 8,800
« ta. »io cm, ia"e lot. v.no, 3o"» ri|.
lafr. Lla. 190*. 170O*f 41,OOO
6 ta. 91 CTI.a'f Rat. Vane. U"B Tit.,
lefr. Lla. ISOt, 1700*f . 17,000
I la. lOlCre, l"B lot. Vane, lO'B Tig.,
lefr. LIB. 390*. 1340V 10,100
1 ta. 14*8 lapeller, 13*. lSO*f, 104 8.1. 1,300
Tatala Hajar leulpaent
Ponaaatioaa
.7
1,410
14.100
n.ooo
3,780
0,140
ia,ooo
11,000
8,940
*,140
1,800
1,800
10
130
80
1,300
10U
40
10
8
10
II
8
37"'
79.000
07,BOO
l.SOO
4,000
0,000
11.OOO
18,OOO
11,000
I1.1UO
I,OOO
11,740
1,400
7,100
4,100
1,100
.too
•lllej
30 T U.
10
1
I
1
138
111
*7
109
41
.1
7
1
11
II
1
1
113
mcier
PI 1108
foundatlona
Structural
Ileetrlcal
Inaulitloe
BulldlBia
•Iplni
InatpieentattOB
• » L at 300
I . 80. L • 80
• • .»». L e .11
•B e 80, L a SO
H . .80, L • 1.10
II e .80, L e .80 at 43,000 CT
«,. .3D ««J. B. L • .70 Up
Bj. .Oo Kaj. a. L - .40 Xj
Totala Minor (qulpnont
Total » » L
jarkup, lacl. by.., Bnpamtloa, Purch., field, nm Of f loo. Profit and Cost tnianay
Total loToataent
*.^ot ehovn on fl-ura 4.
147,000
88,000
403,000
101,000
m.ooo
17,000
831,000
190.000
9.113,DUO
8,730,010
19,000,000
18.000
. 117,000
171,000
408,000
17.000
838,000
78.OOP
I, 731. UUO
t«0,DUO
1,440
1,440,000
1.100
400
-------
Cost EatlmJtO! Section 4001 Sulfur Rjcovory
Equipment Sunber
Major Sulfur Coobustor > B-401
Claus Gas Roheater C-401A
Claus 3as Roheater C-401B
Heat Exchanger C-402
Tendering Gao Cooler C-403
Acid Cooler C-404
Recycle Cooler C-403
Sidostreaa Coolers C-406
Liquid Pfaaoo Clous Rx. D-401
S02 Absorber Coluono D-402
Dust Hoppers 7-401
Separators F-402
.p. Liquid Sulfur Storage Toni F-403
•
Recycle Gas Coap'r. JC-401
Centrifugal Air Coop'r. JC-402
Sulfur Punpa J-401
Acid Circ. Pumps J-402
Acid Sldeatream Pumps J-403
B.P. Water Puopa J-404«
Absorber R. C. Puops J-403
Elec.-Static Preclpltator L-401
Minor Piling
Foundations
Structural
Electrical
Insulation
Buildings
Piping
Instrumentation
Barfc-up. Including Eng'r,, Purchaalng,
Description
U"0 x 34' Lg., 250 pa Iff, Refr. Lined,
Vator Vail, 88 UK Btu/hr
4 El. 5400lf3, 1"0 X 18'BWG., 316 S.S.
Tuba, 22' tg,, 230*. 1300°?, Rafr. Lin.
4 Ea, 5400$, 1"0 x 16 BVO., lCr-£ Ho Tube,
22' Lg., 230*. esO°P, C-tto. Shell
4 Ea. 9000 fl, l"B x 14 BffG., C.S., 20* Lg.,
230*, 7SO»P, C.S. Sholl
1 Ea. 630{A, 3/4"0 x 14 OWO., C.S.,10' Lg.,
V Bond, 2500, 430°?, C.S. Sholl
4 EA. 2800H?, 316 S.S. Tubes, 3/4"0x 16 BtM.
. "0", 28~'0 C.S. Shall, 250*
4 Ea. 7S35I&, C.S. Tuboa, 3/4n0 x 14 BWC.
Fix. ltd., 1-2 Pasa, 230*
2 Ea. 3273$, 316 S.S. Tubes, 3/4"0 x 16 BVQ
"u", 24* Lg. 1-4 Pasa, 230*
2 Ea. 9'0 x 90'0.3.S o/oklrt, 316 S.S. Clad,
250*, 450° P, El Up. Hd.
2 Ea. 7'3"0 x 37'O.S.S.oXaklrt, 31U S.S. Clad,
230*, 450°F, 44' Peck.
3 Ba. 9*0 x 10'O.S.S., 250*. 400°F, C.S. fc
w/valvea
2 Ea. 14'6"0 x 16'6" O.S.S., Horlc, 316 S.S.
Clad, 2500, 4SO°F
3*0 x IS' O.S.S., Horli, v/Sta. Coll, 14*,
310°F. C.S. E
3 Ea. 13,200 ACTH, 310°7, 200—6-342 psia,
Centrlf., 8000 RPM
20,900 CFH, 14.7—*219.7 pslo. 2 Stg.
v/Intercool, 8000 8PM
2 Ea. 25 CPM. 213 psl Rot. Gear, 316 S.S.,
V.C. Bearings
3 Ea. 750 GPU, 210-*262 psia, Aust. I.
Centrlf. 1750 RFU
3 Ea. 750 CPU, 210-f216 polo, Aust. I,
Centrlf, 1150 RPM
2 Ea. 1232 GPU, 0—*2SO palg, C.S., 2 Stg.
Centrif, 1750 RPM
3 Ee. 730 CPU, 200—«-310 pslo, C.S., Centrlf.
1750 RPM
2 Ea. 14'. 400 ACFH, 380 °F, 210 psia, 1 Stg.
86-8"0 Tubas
Totals Major Equipment
II *• L • 300
« a 60, L a 60
H a .28, L *> .22
M o 60, L a 49
U a .80, L • 1.70
B - .60, L » .60 ot 100,000 C.7.
Hp= .29 Maj. X., L ° .60 Up
Mj« .OS IUJ. M., L a .40 HI
Totals Minor Equipment
Total M * L a
Superv,, field, Hose Office, Profit t Cont Ing.
V..l«ht
353,000
200,000
160,000
143,000
s.ooa
72.000
240,000
40,000
2G9, 000
308,780
13,600
170, 600
3,320
276,000
180,000
800
4,200
3,000
8,600
8,400
200,000
UaiorUl
83,800
368,800
192. 400
140,000
4,200
121,200
138,000
70,800
268,000
207,740
10, 300
120,200
2,190
630,000
411.000
2.540
12,630
9,060
13,400
14,070
• 236.000
- 3,056,530
63,000
42,000
113,000
1,033,000
24,000
60,000
764,000
92,000
2,193,000
7,086,280
4.313,720
Labor
$
61,800
13,000
8,000
8,000
300
4,000
12,000
2,000
22,200
18,440
* 900
102,400
360
30,000
'20,000
200
430
450
600
450
80. OOP
324,750
43,000
90,000
774,000
31,000
60,000
438,000
37.000
1,912,009
Insulation Electrical Structural pound's '
$ H. P. * C. Y..---S
300 7,800 20
3,270 32,360
2,040 34,240 17
1,360 2,000 IS
160 300 2
4, 030 8
7,900 13
600 8
5,800 75
43 :
3,300
2,740 6,000 20
320 3
9,000 106
7.000 56
40 IS 2
120 6
30 6
500 14
300 10
80 27.940 17
17,130 17,043 128,720 441 '
700
410,000
17,220
30,000
'Uini
.O'Sa,
12
14
8
30
12
8
24
12
120
210
Total Investment
$11.400.000
-------
Appendix B
Detailed Investment Costs for
Conventional Gas Desulfurization
345-
-------
TABLE I
Cost Eatlnatei Section 300; Sulfur Renewal
(Hot Carbonate HjS Ranoval)
Biulpaent
Major Clean Superheatero
Waste Heat Bolter
BTV Preheater
lean Carbonate Cooler
lean Carbonate Cooler
Reflux Condenser
-~
Reboller
Absorber
Stripper
-
Accumulator
Carbonate Storage Tank
Dust Collectors
w Duot Collector*
,p»
^ Lean Carbonate Punps
Ruflux Pumps
Carbonate Hnke-up Puapa
Minor I'lll MR
Kniniilnt lona
.Structural
liutlOlnKN
Electrical
Insulation
Piping
Inst nuauntotlon
llarfc-up Including Engineering,
Hunter
C-301
C-302
C-303
C-304
C-330
C-306
C-307
D-301
D-302
P-301
F-302
C-301
C-302
J-301
J-302
_
J-303
Purchasing
Daocrlptlaa
B eu. 1346 oq. ft., 316 S3 Tube,
l"x 14 BBC, 18*. 2750 1730eF
6 oa. 8020 oq. f t ..C-SL Tubo,
3/4"x 14 BOO, 24 ',2730. 800CF
3 ea. 3910 oq. ft., C.S. Tubo,
S/8"xl4 DUO, 24'. 275*. 6SO°F
3 ea. 5760 sq. ft., 304 SS Tubs.
3/4"xl6 BWO. 24',27W, 2SO°P
3 ea. 9170 sq. ft., 304 S3 Tuba
3/4nx 16 DBG, 24' ,2750. 2SO°F
9 ea. 11.800 sq.ft. ,304 SS Tubo
3/4"x 16 BBC. 24', 75*. 250°F
B oa. 14,150 oq. ft. .304 SS Tubo
3/4"x 16 BTC. 24* , 731.80tRefr. *3,SOO ot
Up " .24 MAj. M, L a ,70 lip
U. -.053 UaJ. tl, L o .40 U,
1 I
Total Minor Cqutpeont
Total H « L B
tfelght,
6
400,080
332.800
130,000
180,000
270,000
945,000
1.170,000
6.354,000
3,093.000
39. COO
54,000
.1,083,600
1,083,600
I 159,000
0,400
1.600
9,OOM*9.00L
, Supervisory, Field, Hona Off. Profit ft
Katerlol,
3
391,800
290,400
133.000
216,000
324,000
1.138.SOO
1.494,000
1,393,200
593,400
24,150
21,600
385,200
385,200
219, COO
13.280
4.-18O
7,031,810
35.000
!>U.5OO
ZGH.OOO
37,000
318,000
235.000
l.O'.'O.OOO
381,000
3,090,300
14.078.030
9.331,970
Labor, Inoulatlon Electrical,
S £6 H.P.
33,000 3,600
24,000 4,320
7, 500 2, 170
9,000
13,500
43.000
54,000 9,750
1,012,800
217,800 20,910
3.000 1,950
-
91,260 10,980
91,260 10.980
18,000 360 . 6,000
. 800 100
300 30
1,021,220 C3.020 6,130
.
93.5OO
835.000
30,000
284,000 4,320
332,000 118.000
1,1110,000
151.000
8,334,500
Bulldlnso, Structural, Foundatlono,
C.F. 0 C.Y.
13.700
12,100
6,660
3.960
3,160
26,500
31,500
25,200 468
23.200 244
12
42
»
30.000
30.000
36,000 96
10
4
30.0OO 211.980 876
1,700
1,070,000
61.000
Contlncency.
Total Invostnent o
23,400.000.
JW-134S1
-------
TABLE B-2
Cost gatlaate: Section 600; trot COB Scrubbing ami Reheat
Equipment
Major Gaa-Gas Heat Exchanger
Gas -Gas Heat Exchanger
Gas-Gas Heat Exchanger
Clrc. Water Coolers
Hake -Up Water Punp
Water Clrc. Pump
Ui
«»j Vanturl Dust Collector
»
Minor Pi !ln(;
Foundations
Structural
Buildings
Electrical
Insulation
Plplnc
Instrumentation
Herh-up Including Enjlnserlns,
(Hot Carbonat
Nuaber Dnacrio'lon Weight,
a
C-601 12 oa. 6600 sq. ft.. 1700°F. 230* 1.630.000
Refr. Lin, 316 SS Tubo
C-602 6 ea. 0480 sq. ft.. 13SO°F, 2300 741. OOO
Refr. Un, 2-1/4 Cr. - 1 Mo. Tuba
C-603 3 oa. 11,400 sq . ft., 350"F, 250
Labor.
3
108. 500
49, 100
12,000
13.000
750
800
201,000
387. 130
w
13.400
88,000
13,700
37,000
63,000
600,000
31.000
030, 100
•
Insulation. Electrical, Dulldlngo, Structural Foundation
tfl H.P. C.P. » C.Y.
7,440 23.800
4,400 14,500
1.830 7,200
8.600
430 13
200 13
36
13,670 6SO 36, 10O 64
280
400,000
22,800
740
23,000
Contingency
Total Investn«nt
*10.800.000
RM-13451
-------
APPENDIX E
Conversion Factors - English to Metric Units
Length
Area
Volume
Mass
Pressure
Temperature
Energy
English System
inch
foot
square foot
gallon
cubic foot
pound
pound per square inch
atmosphere
atmosphere
0 Fahrenheit
Btu
horsepower hour
Metric Equivalent
2.54 centimeter
0.305 meter
0.093 square meter
3.785 liter
28.32 liters
453.6 grams
51.70 millimeters Hg
760 millimeters Hg
101.3 kilopascals
1.8 (° Celsius) + 32
1055 joules
2.69 x 106 joules
349;
-------
TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
1. REPORT NO.
EPA-600/7-77-031
2.
3. RECIPIENT'S ACCESSION NO.
4. TITLE AND SUBTITLE
High-Temperature Desulfurization of Low-Btu Gas
5. REPORT DATE
April 1977
6. PERFORMING ORGANIZATION CODE
7. AUTHOR(S)
8. PERFORMING ORGANIZATION REPORT NO.
G..P. Curran, B.J. Koch, B. Pasek, M. Pell, and
E. Gorin
9. PERFORMING ORGANIZATION NAME AND ADDRESS
Consolidation Coal Company
Library, Pennsylvania 15129
10. PROGRAM ELEMENT NO.
EHB529
11. CONTRACT/GRANT NO.
68-02-1333
12. SPONSORING AGENCY NAME AND ADDRESS
EPA, Office of Research and Development
Industrial Environmental Research Laboratory
Research Triangle Park, NC 27711
13. TYPE OF REPORT AND PERIOD COVERED
Final; 7/73-1/76
14. SPONSORING AGENCY CODE
EPA/600/13
15.SUPPLEMENTARY NOTES EpA project officer for this report is S.L. Rakes, Mail Drop 61,
919/549-8411 Ext 2825.
16. ABSTRACT
repOrt describes and gives results of economic studies of a process for
desulfurizing low-Btu fuel gas. The gas is first desulfurized at high temperature in a
fluidized bed of half-calcined dolomite. It is then cooled to 700 C and passed through
high-pressure-drop cyclones to remove particulates and alkali. The gas is intended
for use as fuel to gas turbines in combined- cycle power generation. The sulfur accep-
tor is regenerated with steam and CO2. A liquid-phase Glaus reactor is used to pro-
cess H2S in the regenerator off gas into elemental sulfur. Experimental data are pre-
sented in several areas: desulfurization and regeneration activity of dolomites as a
function of cycles; batch studies to determine variable effects and rate data; particulate
and alkali removal at high temperature; Chance reaction studies; and process impro-
vement studies. Two economic studies were performed: one showed an incentive accru-
ing to the process versus a conventional wet desulfurization scheme; and the other, an
update of the process economics , showed that a plant designed in 1975 for 1980 oper-
ation would desulfurize gas from a high-sulfur coal for 38 jzf/MM Btu (HHV + sensible
heat) delivered to a power station.
17.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.lDENTIFIERS/OPEN ENDED TERMS
c. COSATI Field/Group
Air Pollution
Coal Gas
Manufactured Gas
Desulfurization
Regeneration
(Engineering)
Fluidizing
Dolomite
Fines
Alkalies
Air Pollution Control
Stationary Sources
Low-Btu Gas
Particulate
Chance Process
Claus Reactor
Fuel Gas
13 B
21D
07A,07D
13H
08G
18. DISTRIBUTION STATEMENT
Unlimited
19. SECURITY CLASS (ThisReport)
Unclassified
21. NO. OF PAGES
365
20. SECURITY CLASS (This page)
Unclassified
22. PRICE
EPA Form 2220-1 (9-73)
350.
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